Control of Sulfur Emissions from Oil Shale Retorts
by
R. J. Lovell, S. W. Dylewski and C. A. Peterson
IT Enviroscience
Knoxville, Tennessee 37923
Contract 68-03-2568
Project Officer
Robert C. Thurnau
Energy Pollution Control Division
Industrial Environmental Research Laboratory
Cincinnati, Ohio 45268
INDUSTRIAL ENVIRONMENTAL RESEARCH LABORATORY
OFFICE OF RESEARCH AND DEVELOPMENT
U. S. ENVIRONMENTAL PROTECTION AGENCY
CINCINNATI, OHIO 45268
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DISCLAIMER
This report has been reviewed by the Industrial Environmental Research
Laboratory, U. S. Environmental Protection Agency, and approved for publica-
tion. Approval does not signify that the contents necessarily reflect the
views and policies of the U. S. Environmental Protection Agency, nor does
mention of trade names or commerical products constitute endorsement or
recommendation for use.
ii
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FOREWORD
When energy and material resources are extracted, processed, converted,
and used, the related impacts on our environment and even on our health
often require that new and increasingly more efficient pollution control
methods; be used. The Industrial Environmental Research Laboratory -
Cincinnati (lERL-Ci) assists in developing and demonstrating new and improved
methodologies that will meet these needs.
Synthetic fuel processes under development must be characterized prior
to commercialization so that pollution control needs can be identified
and control methods can be integrated with process designs. Shale oil
recovery processes are expected to have unique air, water, and solid waste
pollution control requirements. This report describes an in-depth evaluation
of the control technology systems that are applicable to the removal of
hydrogen sulfide from retort off-gases. Further information on the environ-
mental aspects of oil shale processing and control technology can be obtained
from lERL-Ci, Oil Shale and Energy Mining Branch.
David G. Stephan
Director
Industrial Environmental Research Laboratory
Cincinnati
iii
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ABSTRACT
The objectives of this study were to determine the most applicable
control technology for control of sulfur emissions from oil shale processing
facilities and then to develop a design for a mobile slipstream pilot plant
that could be used to test and demonstrate that technology.
The work conducted included an in-depth evaluation of available gas
characterization data from all major oil shale development operations in the
United States. Data gaps and inconsistencies were identified and corrected
where possible through working with the developers and/or researchers in the
field. From the gas characterization data, duty requirements were defined
for the sulfur removal systems. Based on this information, Stretfbrd gas
sweeting technology was recommended, and the design of a 1000 CFM pilot
plant was completed.
iv
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CONTENTS
Paqe
I. PROJECT SUMMARY 1
A. Introduction 1
B. Objectives , 2
C. Approach 2
D. Recommended Available H2S Control Technology 3
E. Design of Pilot Plant 5
F. Conclusions 5
II. INTRODUCTION 8
A. Oil-Shale Resource 8
B. Environmental Constraints 8
C. Purpose of Study 9
D. Approaches and Limitations 11
E. References 13
III. OIL-SHALE GAS CHARACTERIZATION 14
A. Characterization of Oil Shales 14
B. Paraho Retort Gas 21
C. Occidental Vertical, Modified, In-Situ-Retort Gas 26
D. Geokinetics Horizontal In-Situ-Retort Gas 32
i
E. Union SGR-3 Retort Gas 37
F. TOSCO-II Retort Gas 40
G. References 46
IV. REVIEW OF SULFUR REMOVAL PROCESSES 48
A. Treatment Techniques 48
B. Process for Recovering Sulfur 54
C. References 64
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CONTENTS (Continued)
V. FACTORS INFLUENCING CHOICE OF PROCESS 65
A. Product-Gas Specifications 65
B. Acid-Gas Components 66
C. Influence of Impurities 67
D. Condition of Feed Gas 68
E. Capital and Operating Costs 68
VI. DUTY REQUIREMENTS FOR OIL-SHALE RETORT-GAS DESULFURIZATION 70
SYSTEMS
A. Classification of Oil-Shale Retorting Processes 70
B. Commercial Oil-Shale Operations 75
C. References 79
VII. SCREENING OF GAS TREATING PROCESSES 80
A. Direct-Conversion Processes 80
B. Indirect-Conversion Processes 84
C. References 96
VIII. EVALUATION OF CANDIDATE PROCESSES 97
A. Basis of Evaluation 97
B. Gas Pretreatment 99
C. Direct-Recovery Processes 103
D. Indirect-Recovery Processes 113
E. References 128
IX. COST COMPARISON OF CANDIDATE PROCESSES 130
X. RECOMMENDED AVAILABLE H2S CONTROL TECHNOLOGY 134
A. Adaptability 134
B. Operational Requirements 134
C. Sulfur Removal Effectiveness 135
D. Relative Costs 136
E. Wastes Generated 136
F. Reliability 139
vi
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CONTENTS (Continued)
Page
XI. PILOT PLANT DESIGN 140
A. Introduction 140
B. Sizing of Pilot Plant 141
C. Duty Specifications 142
D. Recommended System 142
E. Preliminary Cost Estimate 148
F. Advantages and Use of Pilot Plant ' 154
APPENDICES
A. STRETFORD DIRECT PROCESS
B. THREE-STAGE SELECTIVE ABSORPTION PLUS CLAUS SULFUR RECOVERY WITH
SCOT TAIL GAS TREATMENT
C. ONE-STAGE SELECTIVE ABSORPTION PLUS INDIRECT STRETFORD SULFUR
RECOVERY
D. DIAMOX PROCESS PLUS CLAUS SULFUR RECOVERY WITH BSRP TAIL GAS
TREATMENT
vii
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TABLES
Number
I Maximum Concentration of S02 to Meet Colorado and Federal 10
PSD Requirements
2 Properties of Green River Mahogany Ledge Oil Shale 15
3 Ash Composition of Oil Shale 15
4 Composition of Organic Matter in Green River Oil Shale 15
5 Gas Produced by Pyrolysis of Oil Shale by Modified 18
Fischer Method
6 Material Balance on Pyrolysis of Oil Shale by Modified 19
Fischer Method
7 Distribution of Elements in Products of Pyrolysis of 20
Oil Shale
8 Estimated Distribution of Sulfur in Oil Shale 22
9 Design Basis for Paraho Retort 25
10 Estimated Composition of Green River Oil-Shale for 25
Paraho PON Retort
11 Gas Produced by Paraho Direct-Fired Retort 27
12 Range of Critical Compositions in Gas from Paraho 28
Retort by Direct Mode
13 Design Basis for Occidental Vertical MIS Retorts 31
14 Gas Produced by Occidental Vertical MIS Retort 33
15 Design Basis for Geokinetics Horizontal In-Situ Retort 35
16 Gas Produced by Geokinetics Horizontal In-Situ Retort 35
17 Design Basis for Union SGR-3 Retort 39
18 Gas Produced by Union SGR-3 Indirect-Heated Retort 41
19 Design Basis for Tosco-II Retort 44
20 Net Gas Produced by Tosco-II Retort 45
21 Comparison of Claus Process Options 60
22 Comparison of Fuel Gases 7^
23 Range of Gas Compositions from Direct-Fired Retorts 72
24 Range of Gas Compositions from Indirect-Heated Retorts 73
2!> Selectivity Data 74
viii
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TABLES (Continued)
Number Page
26 Comparison of Gas Treatment Duty Requirements 76
27 Commercial Oil-Shale Operations 77
28 Selectivity of Absorbent Processes 82
29 Selectivity Absorption Using MDEA with Two Stages 87
of Adsorption
30 Selective Absorption by Benfield Process 89
31 Selective Absorption Using Aqueous Ammonia 91
32 Selective Absorption by Diamox Process 93
33 Selective Absorption by Selexol Process 95
34 Hypothetical Direct-Fired Retort Gas 98
35 Comparison of Selective Absorption Processes for 114
Treating Gas from a Direct-Fired Retort
36 Performance of Three-Stage Selective Absorption System 118
Using MDEA
37 Estimated Capital and Operating Costs for Fuel-Gas 131
Desulfurization Options
38 Variation of Process Effectiveness with Amount of COS 137
in Raw Gas
39 Relative Costs of Various Gas Desulfurization Options 138
40 Approximate Pilot-Plant Operation for Various Gases 149
41 Pilot-Plant Components
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FIGURES
Number -
1 Flow Diagram for Paraho Direct-Mode Retort 24
2 Flow Diagram of the Occidental Modified In-Situ Process 30
3 Flow Diagram of Geokinetics Retort No. 17 Off-Gas 34
Handling System
4 Flow Diagram for Retort System in Union Oil Retort B 38
Prototype Plant
5 Flow Diagram for Pyrolysis and Oil Recovery Unit 42
Tosco-II Process
6 Gas-Sweetening Processes
7 Typical Three-Stage Glaus Process 55
8 Glaus Tail-Gas Treatment Processes 62
9 Gas-Sweetening Processes 81
10 Gas Cooler/Ammonia Absorption System 101
11 Stretford Process for Fuel-Gas Sulfurization 105
12 EIC Sulfuric Acid—Copper Sulfate Process 111
13 Three-Stage MDEA Selective Absorption System 116
14 Sulfur-Burning Claus Sulfur Recovery System with SCOT 120
Tail-Gas Treatment Unit
15 One-Stage MDEA Selective Absorption System with 122
Stretford Sulfur Recovery Unit
16 Diamox Process with Claus Sulfur Recovery System 125
17 BSRP Tail-Gas Treatment Unit 127
18 Material Balance Flowsheet for Pilot Plant 144
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I. PROJfrT SUMMARY
A. INTRODUCTION
The future beneficial use of the nation's extensive oil-shale resources
depends not only on the development of suitable process economics but
also on the development of suitable environmental controls. Even
though the oil that would be produced is comparatively low in sulfur
content, the potential sulfur emissions from large-scale production of
shale oil could be enormous. Oil shale contains up to about 2% sulfur.
A typical shale in the Green River Formation in Colorado contains about
0.7% sulfur. When the shale is retorted, somewhere between 16 and 30%
of the sulfur is liberated to the gas stream, with the majority remain-
ing with the spent shale. The emissions from a 64,000-m^/day
oil-shale industry could be as high as 691-1273 tonnes per day if
emission controls were not applied. If conventional flue-gas scrubbing
systems were used to control the emissions for which the average reduc-
tion achieved is 90%, the controlled emissions would be 69 to 127 tonnes
day. However, if the sulfur could be removed before the gas is burned,
in which case the average reduction is 98%, the controlled emissions
would be in the order of 14 - 26 tonnes/day.
Gases produced by direct-fired retorts, either above ground or in-situ,
are significantly different from gases normally encountered in applica-
tion of desulfurization technology that the technology cannot just be
transferred. Gases from direct-fired retorts contain large amounts of
inert components and have a high ratio of carbon dioxide (CO2) to
hydrogen sulfide (H2S); they also contain large amounts of ammonia and
unsaturated hydrocarbons, such as acetylene, ethylene, propylene,
butylene, and butadiene. The gases are saturated with water and con-
tain some oxygen and trace amounts of sulfur species other than H2S.
The large amounts of CO2 in the gases and the high C02/H2S ratios make
it impractical to employ many of the desulfurization technologies.
Since the gases are produced in huge volumes at near-atmospheric pres-
sures, many other desulfurization processes cannot be economically
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applied. Those processes that can be applied may be only marginal in
performance because of the large amounts of C02 and the presence of
oxygen and/or unsaturated hydrocarbons in the gases or because the gas
may contain a large amount of organic sulfur.
Oil-shale developers are involved in a number of significant pilot-scale
activities for the development of retorting process technology, and
have indicated to the United States Environmental Protection Agency
(EPA) their willingness to cooperate on joint projects for sulfur con-
trol technology evaluation. To capitalize on this opportunity and to
explore the possibility that sulfur emission control will be more of a
problem than was orginally thought, EPA contracted with IT Envirosci-
ence, Inc., to investigate the various commercial sulfur-removal tech-
nologies and to propose a pilot-plant design based on the most cost-
effective process for the removal of gaseous sulfur compounds from oil-
shale retort gases.
B. OBJECTIVES
The objectives of this study were to determine the most applicable
control technology for control of sulfur emissions from oil
shale processing facilities and then to develop a design for a mobile
slip-stream pilot plant that could be used to test and demonstrate that
technology.
C. APPROACH
The work conducted included an in-depth evaluation of available gas
characterization data from all major oil-shale development operations
in the United States. Data gaps and inconsistencies were identified
and corrected where possible through working with the developers and/or
researchers in the field. From the gas characterization data, duty
requirements were defined for the sulfur removal systems. It was found
that oil-shale retorting processes fall into two broad categories:
idirect-fired-retort processes and indirect-heated-retort processes,
each category having distinctly different duty requirements.
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The overriding factor that separates the two categories of retorting
processes and that is dominant in the application of desulfurization
technology is the C02/H2S ratio of the gas produced from the retort.
Those from direct-fired retorts have C02/H2S ratios that range from 76
to more than 165 and thus would require that the sulfur-removal process
selectively remove H2S in the presence of large amounts of C02. Indi-
rect-heated retorts produce gases with C02/H2S ratios in the range of
4.3 to 5, which would allow a nonselective process to be usedl
During this study it was determined by the EPA that the greatest imme-
diate concern is control of sulfur emissions from direct-fired oil-
shale retorting processes and that the pilot-plant design should be
applicable to these retorting methods. Since application of desulfuri-
zation technology to gases from direct-fired-retorting processes is
more limiting, the screening of available process technologies was
based on the duty requirements for those gases.
RECOMMENDED AVAILABLE H?S CONTROL TECHNOLOGY
The class of processes that remove H2S and CO2 from fuel gases is
generically called acid-gas removal or gas-sweeting processes. Removal
of acid gases and/or other gaseous impurities from gas streams is
accomplished either by direct chemical conversion of the acid gas to
another compound that can be more easily separated from the gas, by
I
absorption into liquid, or by adsorption on a solid. The large volumes
of gas that must be processed in a typical oil-shale plant will limit
the application of desulfurization technology to high-capacity, liquid-
phase processes. Since C02 is absorbed to some extent by all liquid-
phase processes, the high C02/H2S ratio of the gas limits the selection
to those processes that can selectively absorb sulfur compounds in the
presence of large amounts of C02.
Of those processes that selectively remove H2S by direct conversion of
the H2S to elemental sulfur, the Stretford process is the most effec-
tive one. Of those indirect processes that selectively remove H2S by
separating the H2S as a concentrated acid-gas stream, the following
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processes were selected as the most effective in their separate process
classifications: the Selectamine and the Adip processes, which use
MDEA as the absorbent, the Benfield, the Selexol, and the Diamox proc-
esses. The Benfield and Selexol processes require the gas to be at
high pressure and thus were eliminated since compression of the gas for
the purpose of desulfurization could not be economically justified.
Except for the Diamox process, all the candidate processes are capable
of removing H2S down to about 10 ppmv. However, organic sulfur com-
pounds , principally COS, which exist in only trace amounts in -. the gas,
are not significantly removed or are only partly removed by the various
processes. However, the presence of those compounds may reduce the
overall effectiveness to 98%.
For desulfurization of gases from direct-fired oil-shale retorts the
Stretford direct process is the most cost-effective system. For the
model case used to evaluate the various processes the total estimated
cost of sulfur removed by the Stretford process would be about $0.50
per barrel of oil produced, which is less than half that projected for
the best of the other processes evaluated.
The Claus process is used to recover sulfur from the acid gas produced
by the indirect sulfur-removal processes. The large amount of CO2 in
the gas makes the best of the indirect processes only marginally capa-
ble of producing an acid gas rich enough in H2S for processing by the
Claus process. Thus to apply these processes multiple stages of selec-
tive absorption would be required to handle the gas produced by many of
the direct-fired retorts. The process, however, does not effectively
remove COS. Only trace quantities of COS (less than 50 ppmv) have been
found in gases produced by direct-fired retorts and thus an overall
effectiveness of 98% or more is projected for the Stretford process.
If the quantity of COS in the gas should be higher than indicated by
the currently available gas characterication data the removal efficien-
cy could be less than 98%.
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The Stretford direct process, on the other hand, is only minimally
affected by the quantity of C02 in the gas and therefore is adaptable
to the full range of gases produced by direct-fired retorts.
E. DESIGN OF PILOT PLANT
The pilot-plant design is based on the current state-of-the-art tech-
nology for commercial application of the Stretford process. The maxi-
mum design capacity of the unit is 28.3 sm^m of feed gas and 6.6 Kg of
sulfur per hour. The plant should be capable of reducing the H2S con-
tent of the gas to 10 ppmv or less, and C02/H2S ratios as high as 200
to 1 should be possible.
The pilot plant is sized primarily to remove H2S from oil-shale gas
produced by direct-fired retorts. However, use of an ejector-venturi
gas-scrubbing system affords wide gas turndown capability for the sys-
tem. The pilot plant thereby is capable of operating on a slip stream
from any of the currently proposed direct or indirect oil-shale retort-
ing processes in the United States.
it
To properly function, the feed gas to the pilot plant must be 120°F or
less, with most of the ammonia removed. A gas cooling column has been
incorporated into the pilot design for cooling and removing the ammonia
from the feed gas. The estimated cost of the pilot plant with all
equipment, instruments, and controls, assembled on skid mountings as a
complete and operable unit, is as follows:
With Cooler Without Cooler
$520,000 $338,000
400,000 260,000
308,000 200,000
F. CONCLUSIONS
The Stretford direct gas desulfurization process may be the only cur-
rently available commerical process capable of effectively removing H2S
from gases produced by direct-fired retorts. Application of the Stret-
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ford process or of any other process to the treatment of these gases
would extend the technology of the process into areas in which no
analogous experience is available. Many questions need to be answered
before the process can be applied with confidence to a full-scale com-
mercial shale-oil production facility.
The principal areas of concern are as follows:
1. absorption of CO2 versus gas characteristics,
2. capacity of the solution for absorbing sulfur versus gas charac-
teristics,
3. rate of by-product thiosulfate formation versus gas characteris-
tics,
4. disposition of COS and other organic sulfur compounds in the feed
gas,
5. effects of unsaturated hydrocarbons in the feed gas on process
operation, life of the Stretford chemicals, and quality of the
sulfur produced.
The viability of the oil-shale industry hinges on an environmentally
compatible sulfur-removal process. Although there are no federal
industry standards for emissions for the oil-shale industry at this
time, the State of Colorado has enacted legislation that limits indus-
try emissions to less than 0.3 Ib of sulfur dioxide per barrel of oil
produced and an equal amount per barrel of oil refined. To meet this
standard at least 96% of the sulfur in the gas expressed as SC>2
would have to be removed.
The area of air pollution compliance that is of the greatest concern to
industry and government is the Prevention of Significant Deterioration
(PSD) requirements of the Federal Clean Air Act. This concept was
enacted to prevent the addition of specified pollutants above a pre-
scribed baseline value in specified air regions. Colorado adopted a
more stringent plan, which limits the maximum level of sulfur dioxide
in the air to an annual average of 10 pg/m3. Thus the maximum quantity
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of shale oil that can be produced will be limited by the effectiveness
of the sulfur emission control system used.
Unless the Stretford process can be demonstrated as an effective and
reliable process for treatment of direct-fired oil-shale gases, indus-
try may have to resort to combusting the gas first and then using less
effective flue-gas disulfurization techniques. Because of the strin-
gent PSD requirements any increase in sulfur emissions could result in
reduction of the potential production capacity of the shale-oil indus-
try.
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II. INTRODUCTION
A. OIL-SHALE RESOURCE1
Recently, as the price of imported oil began to rise and gasoline ex-
ceeded $1.00/gal, interest was renewed in obtaining oil from oil shale
as part of the solution to the energy problem. The solutions to the
environmental problems that were once a cost deterrent to commerciali-
ssation have become more acceptable now that crude-oil prices are ex-
ceeding $20.00 per barrel.
The consumption of oil in the United States in 1978 was about 6 billion
r
barrels. It has been estimated that the known oil-shale resources that
could be tapped by use of existing technology are equivalent to
600 billion barrels, and the total resources in the Green River Forma-
tion in the west have been estimated at 2 trillion barrels of oil,
which would provide oil for 100 and 333 years, respectively, based on
the 1978 rate of consumption. There are very sizable deposits of shale
in the east and midwest, which are referred to as Devonian shale.
These shales are leaner in oil than western (Eocene) shales; but, if
they are incorporated into the resource estimate, the total amount
climbs to about 28 trillion barrels. With a resource of this magnitude
available, some people would ask why we import almost half of our oil
supply. Part of the answer lies in the fact that the organic fraction
of the shale is small and that, if the yield is assumed to be 24 gal/
ton of shale, about 1 cubic mile of rock would have to be processed to
yield the oil we consumed in 1978. Another part of the answer is that
sizable quantities of environmentally undesirable components could be
released to the air and to surface waters.
B. ENVIRONMENTAL CONSTRAINTS1
Sulfur, one of the undesirable components, is contained in oil shale up
to about 2%, and a typical shale in the Green River Formation may con-
tain 0.7%.2 Partitioning studies2'3 indicate that somewhere between 16
and 30% is liberated to the gas that is generated, with the majority of
the sulfur remaining with the spent shale. Under some circumstances
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oil produced from shale might be considered to be a low-sulfur re-
source; in this case, however, where immense volumes of rock are
processed, the potential for sulfur emissions is sizable. The uncon-
trolled sulfur emissions from an 8,000 m3/day oil shale retort
installation could be 86—.59 tonnes per day. The resulting impact on
the semiarid region in the west would be significant.
The National Ambient Air Quality Standard (NAAQS) for sulfur dioxide
limits the annual average to 80 |jg/m3 of air. If this were the only
regulation that industry had to comply with, all sulfur species would
probably be oxidized to sulfur dioxide and some form of flue gas desul-
furization would be used for control. However, the oil-shale industry
must also be concerned with regulations dealing with the amount of a
specific material that may be emitted by a specific source. Although
there are no federal industry standards for emissions for the oil-shale
industry at this time, the State of Colorado has enacted legislation
that limits industry emissions to less than 0.3 Ib of sulfur dioxide
per barrel of oil produced and an equal amount per barrel for oil
refined. The area of air pollution compliance that is of the greatest
concern to industry and government is the Prevention of Significant
Deterioration (PSD) requirements of the Federal Clean Air Act. This
concept was enacted to prevent the addition of specified pollutants
above a prescribed baseline value in specified air regions. Colorado
adopted a similar plan, but made the acceptable levels for PSD more
stringent. Table II-l summarizes the PSD requirements for the federal
government and the State of Colorado, which puts limits on sulfur
dioxide to an annual average of 10 pg/m3. The future potential of the
oil-shale industry therefore hinges on the effectiveness of the sulfur-
removal process.
C. PURPOSE OF STUDY
Oil-shale developers are involved in a number of significant pilot-
scale activities for the development of process technology, and have
indicated to the EPA their willingness to cooperate on joint projects
for sulfur control technology evaluation. To capitalize on this oppor-
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Table 1. Maximum Concentration of SO- to Meet
Colorado and Federal PSD Requirements
a
Class I
Annual average
24-hr maximum
3-hr maximum
Class IIb
Annual average
24-hr maximum
3-hr maximum
SO 2 Maximum Concentration
Colorado
2
5
25
10
50
300
(yg/m3)
Federal
2
8
25
20
,91
512
aClass-I areas are listed as national parks and as natural wilder-
ness and primitive areas.
Class-II areas are the rest of the country.
10
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tunity and to explore the possibility that sulfur emission control will
be more of a problem than was originally thought, EPA contracted with
IT Enviroscience, Inc., to propose a pilot-plant design based on the
most cost-effective process for the removal of gaseous sulfur com-
pounds, as well as one that is mobile. This type of design vjould
necessitate investigation of various commercial sulfur-removal tech-
nologies and their applicability to oil-shale gases.
APPROACHES AND LIMITATIONS ]
i •
Alternatives for Control of Sulfur Emissions ''
Two approaches can be taken to control sulfur emissions: to ;remove
sulfur compounds from the gas before it is combusted (a gas-sweetening
process), thereby producing a sulfur-free fuel; or to first combust the
gas and then remove the resulting S02 from the flue gases {flue-gas-
desulfurization (FGD) process].
i
The latter approach is generally less effective, because most of the
flue-gas scrubbers can remove only 90 to 95% of the sulfur in the flue
gas. The volume of gases handled by the FGD process is much larger
than that handled by the first process because the products of combus-
tion are more voluminous than the fuel gas burned. Also, th^ FGD sys-
tem generally consumes more water and chemicals and is more costly to
operate than the gas-sweetening processes.4 ;
i
Processes that remove sulfur from the fuel gas before it is fcombusted
are generally more effective, with most processes capable of removing
i
98% or more of the H2S in the gas. With these processes, however,
sulfur in other forms, such as COS, CS2, and mercaptans, may be only
partly removed or not removed at all. ;
To meet the Colorado air quality control regulations of 0.3 Ib/bbl of
SO2 equalivant emissions, at least 95% of the sulfur in the igas would
11
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have to be removed. For these reasons this study was directed; toward
the more effective fuel-gas desulfurization techniques.
Basis of Process Selection ;
During this study it was determined by the EPA that the greatest imme-
diate concern was control of sulfur emissions from direct-fired oil-
shale retorting processes and that the pilot-plant design should be
applicable to these retorting methods. Therefore the screening of
available process technologies was based on the duty requirements for
desulfurization of direct-fired retort gases.
12
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E. REFERENCES*
1. R. C. Thurnau, H^S/SOg Control Technology for Oil Shale Retort Efflu-
ents , Work Directive No. T-7012, EPA Contract No. 68-03-2568, June 18,
1979. !
2. K. E. Stannfield, I. C. Frost, W. S. McAlley, and H. N. Smith, Proper-
ties of Colorado Shale, U.S. Bureau of Mines, Report of Investiga-
tions 4825 (1951). :
3. 3. W. Ward, Analytical Methods for Study of Thermal Degradation of Oil
Shale, U.S. Bureau of Mines, Report of Investigations 5932 (1962).
4. T. Nevens et al., p 133 in Predicted Costs of Environmental Controls
for a Commercial Oil Shale Industry, Denver Research Institute, COO-
5107-1 (July 1979). [
*When a reference number is used at the end of a paragraph or,on a head-
ing, it usually refers to the entire paragraph or material under the
heading. When, however, an additional reference is required:for only a
certain portion of the paragraph of captioned material, the earlier
reference number may not apply to that particular portion. :
13
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III. OIL-SHALE GAS CHARACTERIZATION ;
I
A. CHARACTERIZATION OF OIL SHALES !
1. Introduction
In characterizing the gas produced by the various oil-shale retort
processes, it is necessary to have some understanding of the nature and
j
composition of the organic and mineral materials that comprise oil
shale and of how the materials become distributed among the resulting
products. The oil in oil shale can be neither squeezed nor drained out
at ambient or moderately elevated temperatures. When the oil; shale is
heated to temperatures in excess of 200°C and up to 500°C, the material
is destructively distilled, producing recoverable oil and gas and
leaving a spent mineral residue.
2. Composition of Oil Shale
Oil shales having similar assays of oil, in gallons per ton (gpt), were
found also to have a fairly narrow range of compositions, as can be
seen in Table 2. .1<2 However, even in the Mahogany ledge,; which is
about 21 m thick, the assay of oil varies from 37.5 to 312.3 Up tonne.
i
The mineral material, after being fired to an ash, has the composition
sshown in Table 3. 1 The minerals in their natural form in oil shale
are varied and complex, and on a point-to-point basis the pH 'can vary
from 8.5 to 10.3
3. Composition and Nature of Organic Matter in Oil Shale
The organic matter in Green River oil shale has a significant amount of
hydrogen (see Table 4). *»4 The ratio of about 1.5 moles of hydro-
gen per mole of carbon enabled about two-thirds of the organic matter
to be converted to oil during retorting. In contrast, high-volatility
bituminous coal has a hydrogen-to-carbon mole ratio of about :0.9 to I.3
The substantial amount of oxygen present in the organic material indi-
cates the likely presence of carboxyl groups. It is also likely that
I
14 ;
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Table 2. Properties of Green River Mahogany
Ledge Oil Shale
Component
Carbon, organic
Hydrogen
Nitrogen
Sulfur
CO_, mineral
Ash
Mineral content of spent shale
Amount (wt % of
Stanfield3
12.4
1.8
0.4
0.6
18.9
65.7
99.8
raw shale)
1 b
Smith
12.21
: 1.76
; 0.37
: 0.70
17.92
65.26
98.22
i
a
See ref 1; analysis based on sample that yielded 115 J,/tonne.
ref 2; analysis based on sample that yielded 105 l-Jtonne.
Table 3.
Ash Composition of Oil Shale
Component
SiO_
2
Fe-0,
2 3
Al_0_
2 3
C 0
a
M 0
g
so3
Na-0
2
K2°
Amount (wt % of raw shale)
27.8
|
3.0 :
8.6
15.1 !
i
6.5
1
1.2
2.0
•
1.5 i
aOil assay, 115 Jl/tonne (see ref 1) .
15
-------
Table 4. • Composition of Organic fatter in
Green River Oil Shale
Component
Carbon
Hydrogen
Nitrogen
Sulfur
Oxygen
, . ... _ — .._,_ _______
Amount
Stanfielda
76.5
10.5
2.5
1.2
9.3
(wt % of raw shale)
Smith
80.5
10.3
2.4
1.0
;5.8
a.
See ref 1; analysis based on sample that yielded 115 £/tonne i
See ref 2; analysis based on sample that yielded 105 H/tonne
16
-------
these carboxyl groups are saponified by the mineral material,; whose pH
after pyrolysis is about 11.s Therefore only through pyrolysis, or
destructive distillation, wherein organogenic CO2 is released, can oil
be recovered from oil shale. As further evidence of the change in the
nature of the organic material, as implied by the term "destructive
distillation," the oil produced from oil shale has a pour point of
about 29°C/ although oil will not drain from oil shale at this tempera-
ture . i
The presence of sulfur and nitrogen in oil shale and shale oil is dis-
cussed below. ;
4. Recovery of Products from Pyrolysis of Oil Shale by the Modified
Fischer Method i
A quantity of oil shale representing 94 fcpt assay was obtained from 10
locations in the Green River formation and was blended to provide a
quantity of uniform composition. Eight replicate pyrolysis runs were
\
made by the modified Fischer method to determine the reproducibility of
this method of assaying the yield of oil and gas from oil shale.2
Table 5 gives the quantity and composition of the gas produced
during pyrolysis; the data, in material balance form, that were ob-
tained are given in Table 6. i
The recovery of materials expressed as percent of element contained in
raw shale is shown in Table 7. These data show that the\ recovery
of nitrogen in the oil and gas streams combined is only 53%, of which
8.6% is in the gas stream. They also show that the combined, recovery
of sulfur is only 25%, of which 16.9% is in the gas stream, and that if
the shale oil were treated to remove nitrogen and sulfur the NH3 con-
tent of the gas would increase about sixfold and the H2S content would
increase by 50%.
5. Sulfur Chemistry Relating to Pyrolysis Gas
The oil shale in the Green River formations contains sulfur in two
major forms: organic (with -C-S- groups present) and pyrite, FeS2.
17 '<
-------
Table 5. Gas Produced by Pyrolysis of
Oil Shale by Modified Fischer Method
Component
C2
C2=
C_
3
V
C4
C4=
C4
C5=
C5=
C6=
co2
CO
H2
H2S
NH,
Total
Amount
(vol %) (DG)a
17.44
5.44
2.10
2.43
2.39
1.50
0.60
1.20
1.56
1.45
0.11
24.00
5.09
29.36
3.20
2.12
100.00
of Gas Produced ,
(sin?/ tonne of raw shale)
4.7
1.5 i
0.6
i
0.6 ;
f
0.6 ;
0.4 I
i
0.2 I
0.03
0.3 ;
0.4
0.4 ;
6.5
1.4
7.9 ;
0.9
0.6 ;
27.03 ;
See ref 2.
b
Calculated gas production.
18
-------
id
1
4J
0)
14
4J
(0
m
iO
o
1-1
.0-1
n
o
0)
r-l
rt
*
4-1
8
w
Q)
C
C
o
JJ
c -t.
H Q£
3 CD
5 rj
i 2
S x:
CO
4,
c
O
in in o dP
• • • *3*
1-1 •— I rH
XI
CO >— i en CM >— I O dP
— i in m
OO 1-1 •
o*
dP
i~- O oo T-( <)• O '"J
oo CM 1-1 m in co Jg
CM CM 0 00
OO 00
ON r~ oo oo oo o
r— ( f— CO VC £?> O O
CM i-l -* C) O
^i oo o ^-i
t— 1
u
o S
•H (B
§ M
g, o
° C C °to to
^ QJ Q) rH tH
O O O 3 M 4-1
i U W «4-l Ci) O
O S3 !Z W E
C
,Q
y
U
(0
0)
c
•a
E *
M O * O
M-I CM 0) O
G W I *S
^5 c *™^ 'O
Ll flJ flj
(tt «Q O
N c
•g 8 & s
J s i? s
• f* ^* c*
" * id °
c^ o^ ty* CD o^
c c c c
«W -H -H -H -iH
flj "o 7> " "o
10 M H __Q H
0) JQ O "d 0)
19
-------
Table "7. Distribution of Elements in Products
of Pyrolysis of Oil Shale3
Component
Carbon, organic
Hydrogen, total
Nitrogen, total
Sulfur, total
Spent Shale
23.6
11.4
46.6
75.1
Distribution
Oil
65.9
61.9
44.8
8.0
(wt %) !
Gas
10.5 :
18.0
8.6
16.9
Water
8.7
aSee ref 2.
20
-------
The -C-S- groups are pyrolyzed in the presence of hydrogen to!yield
s
H2S. Examination of the gas phase that is formed in this process indi-
cated that it went to completion; but since there are still sulfur-con-
taining organic compounds in the shale oil, the existence of small
quantities of low-boiling organic sulfur compounds in the gasjwould not
be surprising. ;
Pyrite can undergo hydrogenolysis to yield H2S as shown below:
FeS2 + H2 > FeS + H2S:
(pyrite) (hydrogen) (iron sulfide) (hydrogen sulfide)
Consequently one-half of the sulfur is available for recovery: as H2S
and can appear in the gas. Sulfate salts, which are found in spent
shales, are apparently formed in the combustion zone following pyroly-
sis.6 Table 8 shows the estimated quantity of sulfur in these two
forms and the sulfur available for recovery (3.89 kg/tonne of raw shale).
The sulfur recovered as H2S (1.15 kg/tonne of raw shale; see ;Table 6)
is estimated to be 30% of the sulfur available for recovery. '
B. PARAHO RETORT GAS ;
1. Introduction
Development Engineering, Inc. (DEI), a subsidiary of Paraho Development
Corp., have developed an above-ground retort to the point thajt they are
ready to design a commercial-size module. This retort can be designed
to operate in either the direct mode, wherein combustion air ;is brought
into direct contact with the shale in the retort, or in the indirect
mode, wherein only recycled reheated pyrolysis gas is brought into
contact with the shale. In the latter case the recycle gas is reheated
by some of the pyrolysis gas being burned in a separate unit.7 Recent-
ly, DEI concluded that the direct mode is more energy efficient, and
they are now emphasizing its development in preference to the indirect
mode.8 :
21
-------
Table 8. Estimated Distribution of
Sulfur in Oil Shale
Sulfur
Form
Organic
Pyrite
Unaccountable
Total
Distribution
Total Sulfur
1.5213
4.73C
.0.75d
7.00
( kg/ tonne of raw shale)
Available as HsS
1.52 j
2.37 i
3.89 i
aOil assay, 104 'i /tonne. i
bBased on 0.45 kg of S/36.6 kg of C; see ref 4. ;
c:Based on 1.41 kg. of S as pyrite/ 45 kg of S as organic; see ref 4.
Calculation based on 7.0 kg of total S per tonne of raw shale.
22
-------
RAW
SHALE
OIL MIST
SEPARATORS
MIST
FORMATION
AND
PREHEATING
RETORTING
ZONE
_ — — — — _ —
COMBUSTION
ZONE
RESIDUE
COOLING
AND
GAS
PREHEATING
^ 2
X /
_
Y
\
c
»->
)IL
N
A
r
^N
PRODUCT .,,,_
GAS i
(ELECTROSTATIC
PRECIPITATOR
:
C ^\ 1 RECYCLE GAS
V^_^7 BLOWER
f .• \ .,.._ /\in BLOWER
'GATE
SPEED
CONTROLLER
RESIDUE
OIL
Figure 1. Flow Diagram for Paraho Direct-Mode Retort
24
-------
2.
3.
BEI has sotted . proposal
(DOE) Program opportunity Notice (pON)
nd oerate an ^ove.ground retort
^
^ consti.uct .on
4.
^
^
to .develop technology to exploit oil shale as an energy resource. The
r9;;:;;:yhat such * ret°r
Paraho Retort Process Description
In the direct mode of operation of the Paraho retort (see Fig 1,
crushed „ shale is fed to the top of the retort and spent skali
yTherolthe "Tr- Tbe Shale P"S" d°™""d * *"^* ""cessive-
ly through . mist Nation and preheating 2,ne, a retorting .one. ,
co^usuon 2one. and finally, a residue-coding, g,s-prehe,ti,g zone.
The „ aro onaceous residue on the retorted shale is turned in the co^us-
taon zone to provide the principal fuel for the process."
The shale-oil vapors produced in the retorting zone are coole
-------
Table 9. Design Basis for Paraho Retort3 l
Shale-oil production _,ft •».„.
••960 nr?/day
Oil-shale throughput ., „ „„„ ' I
-10,000 "tonnes/day
Oil-shale assay
91.6-—158, £/tome
Method of heating '
Direct fired
Gas recycle ratio ,
498 sm-Vtonne
Gas production ratio - • -> • '
223 sm3/tonne
Gas mole ratio, CO_/H_S " i
Heat content, HHV ;
105,400 joules
Pressure <./, ..&.
Atmospheric
Temperature
60°C
See refs 9 and 11.
Table 10. Estimated Composition of Green River
Oil-Shale for Paraho PON Retort3
Component
Carbon, organic
12.2
Carbon, mineral CO
2 4.9
Hydrogen . i
1.76
Nitrogen
0.37
Sulfur, organic i
0.15
Sulfur, pyrite
0.47
Sulfur, sulfate '
0.08
Oxygen, except CO i
2 1.37
Balance, minerals
78.7,
loo. o:
Oil assay, 104 to 112 £/tonne(see refs 4 and 9)
"25
-------
TY?;:-:G G'Ji J2 SK'JET
llV.l- Oi
v. L;.. :.-'•.
5.
Retort-Gas Characterization • i :
: ; !
As the pyrolysis gas is produced, 498 sm^ of the gas per t:onne of
I -- •
shale is recycled with sufficient air to provide the necessary heat for
pyrolysis, and 223 sm-Vtorme is forwarded for uses external to the '
retort process. The gas produced is discharged_at_essentially atmos^
pheric pressure and at about 60°C (see Table 11)i . i
1 i ! i
1 ! i
Although the gas from the Paraho retort has been sampled, analyzed, and
reported over the years, only the most recent data are considered as j
adequately representative of the process because the techniques and
methods used now are considered to be better than those used'before
1977.8 Table . 11 shows the mean values of the composition of the
gas produced by the Paraho retort in the direct mode during 1977 and
1978. In 1977, 31"gas samples were taken and analyzed and in 1978
i :
there were 51. The high, low, and mean values for CO2 H2S, and NH3 are .
shown, in Table 12'.. -12
C. OCCIDENTAL VERTICAL, MODIFIED, IN-SITU-RETORT GAS
I
!
1. Introduction
For more than six years Occidental Oil Shale, Inc. (OOSI) (formerly 1
Garrett Research and Development Co.) has been engaged in modified, in- j
situ, process technology development. Occidental's philosophy in i
developing this process for producing oil from shale is to maximize the .
i • !
recovery of in-place oil while minimizing costs and the environmental :
impact.13 ,
I
OOSI prepared and processed its fifth and sixth retorts under a cooper-
i ' !
eitive program with the Department of Energy (formerly the Energy Re-
I ! i
search and Development Administration). The DDE/Occidental Cooperative
f *
Agreement is a two-phase agreement consisting of engineering develop-
ment of two specific retort designs for the Occidental modified in-situ
process as the first phase and a technical feasibility demonstration as
i BOTTOM OF
BEGIN
LAST LINF
OF TEXT £
the second phase.13
;
! i;3/8" i_ iCI'26'':"ll ;
PAGE NUMBER I
EPA-287 (Cin.) :
(4-76) '
OUTSIDE
DiMLNSION
FOR TABLE?
.-AND-ILLUS-
, TRATICNS
-------
Table 11. Gas Produced by Paraho Direct-Fired Retort'
Component
CH4
C2
C2
C3*S
C.'S . .
c5-s
co2
CO
H2
H2S
NH3
N2
°2
so2
NOX
CS2, RSH, COS
Thiophene
Thio-4- (methylthio) -
3-cyc:lo-pentene-2-one
Total, DG
H O
2
•
(sm-V tonne)
' 4.8
-1.2'
1.5
=1 ,6
-1..2
1.2b
' 50..6
- '5.5
. -10,5
0,7
..i.'e
: 141.5
0,2
' 222.2
Amount !
(dry vol %)
[2.16
!0.56
0.66
;o.7i :
:0.56
0.57b
22.81
2.50
;4.74
0.30
0.70
63.80
0.09
(17 ppm)
(168 ppm)
None detected
(50 ppm)
(200 ppm)
i
100
20d
Oil assay, 100 to. 108 H /"tonne.; .see ref 12.
Total for C ' and higher hydrocarbons.
Identified sulfur compounds.
20% H,.O on wet-gas basis; see ref 9.
j*
27
-------
Table T2. Range of Critical Compositions
in Gas from Paraho Retort by Direct Mode3
Component
co2
H2S
NH
High
27.7
0.55
1.23
Amount (dry vol %)
Low
18.0
0.19
0.50
Mean
22.8
0.30
; 0.70
"Represents 82 sample analyses taken in 1977 and 1978; see ref 8.
28
-------
TP
BEGIN
LAST LiNE
OF TEXT
Description of Process j i
The modified in-situ process consists of retorting a rubblized column '
of broken shale, formed by expansion of the oil shale into a previously
Unined-out void volume. In retort 5 the mined void volume is removed
^.frorn the^retort zone in_ the form of_a vertical _slice_along_'die_center '_
of the room, extending from the top of the rubble pile to be formed, to :
the production level. This system is known as a vertical fr|ee-face i
retort system (VFFR). In retort 6 three horizontal levels are mined !
out similarly to room and pillar mining and then the oil shale is \
blasted above and below the mined-out sections to the horizontal rooms.
This system is the horizontal free-face retort system (HFFR).13 !
After the column of shale has been rubblized, connections are made to
both the top and bottom and retorting is carried out (Fig. 2).
Retorting is initiated by heating the top of the rubblized shale column
I
with the flame formed from compressed air and an external heat source,
such as propane or natural gas. After several hours the external heat
source is removed and the compressed air flow is maintained, with the
f ;
carbonaceous residue in the retorted shale used as fuel to sustain air
combustion. In this vertical retorting process the hot gases from the
combustion zone move downward to pyrolyze the organic matter! in the
shale below that zone, producing gases, water vapor, and shajle-oil
mist, which condense in the trenches at the bottom of the rubblized
column. The crude shale oil and by-product water are collected in a
sump and pumped to storage. Part of the off-gas is recirculjated to
control the oxyg;en level in the incoming air and the retorting tempera-
ture. The^ remainder of the off-gas passes on to the desulfurization
train.10 ;
1
Hasis of Retort Design and Operation '•
i ; '
The projected bases for design and operation of Oxidental retorts 7 and
B are given in Table 13. "'is The design basis for retort 6 is
also given'.
BOTTOM OF
IMAGE AREA.
OUTSIDE
EPA-2S7 (Cin.
.(4-75)
29.
l*l*l/vJLi
PAGE NUf.-BER
FOR TABLES
•AND'ILLUS-
TRATIONS
-------
-OIL RECOVERY
RAW
SHALE
OIL
RECYCLE GAS
COMPRESSOR
FUTURE RETORT
CENTER SHAFT
AIR MAKE-UP-A
COMPRESSOR
OIL SHALE RUBBLE
OIL SUMP AND PUMP
figure 2. Flow Diagram of the Occidental Modified"in-Situ"Process"
"V." 30 .~~~ "" ~— — — - -
TYP'JG
-------
Table 13. Design Basis for Occidental Vertical MIS Retorts
Parameter
Retort 6
Retorts 7 and 8
Retort size (tonnes)
Oil-shale assay (fi,/tonne)
Method of retorting
Gas injected
Gas production rate (snv
Gas productions ratio (sm-V tonne)
Gas ratio, CO2/H2S (mole/mole)
Heating value, HHV (joules'/sm^)
Pressure (psig)
Temperature '(°C)
329,058
62.5
Direct in-situ
Airc
408
588"
330 to 130
i;6x!06 to 2.2xl06
Atmospheric
57 to .66 .
618,120
62.5
Direct in-situ
50% air/50% steam
1100
f
296 !
165
4.1xl06
Atmospheric
57 to ,77
See ref 13.
^Estimated from Oxy No. 6 and refs 14 and 15.
°Steam injected at end of run.
31
-------
4.
"-. i
Retort-Gas Characterization ! ;. —
The composition of the gas produced by retort 6 at midpoint of the run
is given in Table 14. The composition of the gas varied throughout
the run. In the initial stage of operation the C02/H2S molar' ratio
averaged about 330 to 1. After steady-state operation was established,
the CO2/H2S molar ratio averaged about 165 to I.13 Retorts 7 and 8
will be operated using a 50% oil/50% steam mixture injected into the
retort. The use of -steam will affect the composition of the; gas pro-
duced. The quantity of COS in the gas is expected to increase.
! i
GEOKINETICS HORIZONTAL IN-SITU-RETORT GAS !
-•* f -, -'
• I -.
- 'Introductic.--.-
of recovering shale: oil from
shallow formations of oil shale. The method consists of detonating
implanted explosives to rubblize the shale. This raises the overbur-
.. t
den, thereby developing void spaces among the broken shale. The bed of
broken shale (retort) is then ignited with a fuel-air mixture injected
at one end, and the retort products are expelled through a production
i
hole at the other end (Fig. 3).
LAST LINE
OF TEXT I
The retort burn advances horizon--
tally. When the retort is well ignited, the fuel is shut off and only
air is fed to the injection hole. i
Retort 18 was fired from November 16, 1979, to February 5, 1980.
1 i ;
Retort firings will continue in an effort to develop this technology.
The aim of this program is the production of 320 m^ of oil ;per
day over about a 10-year period from small shallow retorts, i
.-.-Basis of Retort Design and Operation ;
-\ies~ign basis for geokinetics retort 18 is given in Table 15.16,17
3. Retort Gas Characterization
i ! :
The average composition of the gases produced by retort 18 is given in
Table 16.
A T./R"
EPA-287 (Cin.)
X":'"''
BOTTOM OF
IMAGE ARE.-*
OUTSIDE
DI.V.EMSIOfJ
FOR TABLE;.
"yAND ILLUS
TF; AT IONS
PAGE NUf/.SER
-------
Table 14.
Gas Produced by Occidental Vertical MIS Retort1
Component
H2
CO
co2
CH4
C2H4 ..
C2H6
C3H6
C3H8
C4
C5 +
H2S
NH
N2
°2
so2
NOX
cs2
RSH
COS
Total, DG
H20
Composition
• (sm-Vtonriek)
45.0.
-5.2-
.'188.8
8\4
0:4-
1,3
0.4
.0.6
0.-6
' 0.1
:0.7-
c
'330.5
0.5 .
6.9
0.2
c
. c
0.02
• 585.9
: C
(•vol. %) (DG)
7.65
j 0.89
I 32.26
1.44
; 0.07
\ 0.224
0.0
.0
; o
0.01
0.1
c
56.41
0.08
0.15
0.03
c
' c
: (1 — 40 pprn)
! 99.611
! 38.3d
Gas produced by Oxy No. 6 at midpoint of run (Jan. 15, 1979); see ref 13.
Average calculated values. ,
C '
No data.
vJet-gas basis. :
33
-------
4J
(Q
>i
(0
S1
a
a:
10
o
(4-1
U-l
O
-• w
o
-<-l
JJ.
0>
o
Ol
en
(0
Q
0)
VI
D
cr>
FAGL r-.L-Vi,;.;;
TV-'*^1'
1 > r i
-------
15. Design Basis for Geokinetics
Horizontal In-Situ Retort
'"t size
On *hale
""''"/a of
Injected
-*-
<4^.~Z,
V^C.
10,025
83.3— 91.6
Direct in-situ
Air
45.3
j
c i
181
3,35xl06
Atmospheric
54
-------
Table 16.
Gas Produced by Geokinetics
Horizontal In-Situ Retort
Component
H2
CO
2
CH4
°2H4
C2H6
C3H6
C3H8 ^' '' ':'''^'
C4
c5+b
H2S
1!IH3
2
°2
so2
!NOX
cs2
:RSH
'COS
Total
H20
aGas produced by Geokinetics retort
C- to C,n hydrocarbons present but
o 1*J
Composition (dry yol %)
7.47
8.03 I
23.48 ;
1.61
0.085
1
0.297
0.09
0.13 !
0.15 !
0.071
0.13 ;
0.060
57.4
1.13
c :
c !
c '
c ;
(40 ppm) ;
100.13 '
15. Od .
No. 18; see ref 16.
not measured. i
CNo data.
wet-gas basis.
36
-------
'•'•- FA
fur (,
, ;.VAC::
'Ar;HA;
.E. UNION SGR-3 RETORT GAS
\-:-.\r*- \
i 2:
BEGIN
LAST LINE
OF TEXT if;
3.
Introduction j !
The Union Oil Company of California (Union) have been involved in oil !
shale activities for more than 50 years. The_ development_of_their j
retorting technology was initiated in the early 1940s, and several i
variations of a vertical-kiln retorting process, with upward flow of .
1 ! !
shale and countercurrent downward flow of gases and liquids, have been. [
developed. Two variations are known as retort A and retort B. Union
' i
Oil now proposes to construct a 9090 tonnes/day demonstration; plant, i
using the retort-B process, together with all the necessary auxiliary ;
facilities.10
i
Description of Process | i
In the retort-B process (shown in Fig. 4 . ' ) as the crushed oil shale
I . !
flows upward through the retort it is met by a stream of hot (510 to !
538°C) recycle gas from the recycle gas heater flowing downward. The j
rising oil-shale bed is heated to retorting remperature by countercur- •
rent contact with the hot recycle gas, resulting in the evolution of j
the shale-oil vapor and product gas. This mixture of shale-oil vapor !
and product gas is forced downward by the recycle gas and is cooled by j
contact with the cold incoming shale in the lower section of the retort ,
i ; ; j
cone. The liquid level in the lower section is controlled by with-
'• i
drawal of the oil product. The recycle and product gases are removed
from the space above the liquid level. The product gas is first sent
to a venturi scrubber for cooling and removal of heavy hydrocarbons by
oil scrubbing. That portion of the product gas not recycled is then
• i
sent to the desulfurization train. Before the gas reaches the desul-
l i |
furization train, it may be compressed and oil-scrubbed to recover
additional hydrocarbons.10
- I
Basis of Retort Design and Operation
The basis for design for the Union SGR-3 retort is given in;Table
17.18'19 '
i
_1_,
EPA-287 (Cin.)
(4-76)
!^.l?i?.:.:.. *&
PAGE NUMBER
BOTTOM OF
IMAGE AP.3-
OUTSIDE
DI.MBMSIOU
FOR TABLE;:
VAND ILLUS-
TRATIONS
-------
ro
rH
fU
0)
o
LI
OQ
JJ
LI
O
3
O
1-1
D
V
4->
W
>i
(O
Ll
O
1-
LI
0)
LI
fa
PiI-JG GUIDd
-------
Table 17.
Design Basis for Union SGR-3 Retort
Shale-oil production
Oil-shale throughput
Oil-shale assay
Method-of retorting
Method of heating
Gas production ratio
Gas production rate
Gas mole ratio, CO_/H S
£ 2
Heat content, HHV
Pressure
Temper at ure ( °C )
1440 m3/dayb ' '
9090 tonnes/day
157.4 8,/tonne
Indirect mode i
Indirect heating; of recycled
retort gas !
29.9 sm3/tonne [
188.7 sm3/min j
4.35 : i
42.38xl06 joules/sin3
5—10 psig ',
65.5 71.1°C
See refs 18 and 19.
Stripped shale-oil production is 1419 m-Vrday.
39
-------
GUiL-J SHL^i
4.
Retort Gas Characterization ' j I
The estimated composition of the gas that will be produced by the
Union SGR-3 retort is given in Table 18.
i
TOSCO-II RETORT GAS
i.
BEGIN
LAST LINE
OF TEXT
2. .
Introduction
Tosco II is a process developed by the Oil Shale Corporation (Tosco).
Initial development work began in 1955; in 1964 a 909 tonne/day semi-
works plant was constructed and tested. A full-scale commercial plant
is planned by the Colony Development Operation, a joint venture of
several companies who formed or own Tosco.
Description of the Process ; | I
The heart of the processing sequence is the Tosco-II pyrolysis (retort-
ing) unit and associated oil recovery equipment. The flowsheet for a ;
single unit (or train) is shown in Fig. 5. The raw shale from the
final crusher is fed to a fluidized bed, where it is preheated to about '
1 i i
SOO°F with flue gases from a ball heater. The residual hydrocarbons in
i i '
the flue gases are burned in the ball heater.10
The preheated shale is fed to a horizontal rotating retort (pyrolysis
i I ;
drum), together with approximately 1.5 times its weight in hot ceramic:
balls from a ball heater in order to raise the shale to pyrolysis tem-
perature (482°C) and convert its contained organic matter to shale-oil
j i I
vapor. The oil vapors are withdrawn and fed to a fractionatjor for
hydrocarbon recovery. The mixture of balls and spent shale is dis-
charged through a trommel, in order to separate the emerging warm balls j
from the processed shale.10
\
The warm balls are purged of dust with flue gases from a steam pre-
i I
heater, and the dust is removed from the flue gases by wet scrubbing.
The dust-free warm balls are returned to the ball heater via the ball
elevator. 'In the ball heater they are reheated to about 704°C and
_spent_shaie_is_cppled
then recirculated to the
/8"
EPA-237 (Cin.)
(4-7S)
W
BOTTOM OF
'.'.'AGE Ar.E-
OUTSICI
Q1N!ENS':G\
•'"<-. TABLE;
-AND ILLUS-
TRATIONS
PAGE NUM3ER
-------
Table .18. Gas Produced by Union SGR-3
Indirect-Heated Retorta
Component
H2
CO
co2
CH4
C.H..
2 2
C H
2 4
°2H6
C H
3 6
r H
^38
C.H-
4 6
i-C..HQ
4 8
n-C H
4 8
f-C4H10
n-C_H_
5 8
i-CcH-
5 8
C5H10
C5H12
C '
6+
H2S
NH3
N2
°2
so2
NOx
cs2
RSH
COS
. Total, DG
H 0
2
Oil assay, 157£/tonne;
No delta.
(sm-V tonne)
6.95
1.44
.. 4.95
- 6.63
0.06
0.53
2.28
1.13
b
1.02
-0.09
T -
0-.'79
0.34
0.09
'T
0.12
0.06
0,28
1.84
- 1.14
b
0.03
c
. • _
b
• . •-.
-
0.-02
29.77
5.95
see ref 19.
Amount i
; (dry vol '%)
23.34
: 4.85
16.62
| 22.28
1 0.19
1.72
; 7.67
i 3.81
'
b
i
i 3.43
1 0.29
2.66
!
1.14
0.29
0.39
0.19
i 0.94
6.19
i 3.82
b
0.11
c
\
; (125 ppm)
b
j (20 ppm)
; (164 ppm]i
I ^542 ppm)
i 100.00
20d
\
No data; likely nil.
Estimated value, wet-gas basis.
. 41
-------
0]
(0
V
u
o
Ll
cu
8
a
u
0)
o
u
r-J
-M
O
•o
IS
03
-•H
01
O
Ll
CU
10
u
D>
10
in
LI
3
t,
-------
to about 149°C in a rotating drum cooler and moisturized with water
recovered from the plant's foul-water stripper unit.10 \
Unlike some U.S. oil-shale facilities, the Tosco-II/Colony commercial
plant will be designed not only to produce shale oil but also to up- :
grade it on-site to produce synthetic crude oil and liquid petroleum
gas (LPG), with ammonia, sulfur, and coke produced as by-products. The
shale-oil hydrocarbon vapors from the pyrolysis drum are separated into
water, gas, naphtha, gas oil, and bottom oil in a fractionatpr. The
water is sent to a foul-water stripper, the gas and naphtha are sent to
si gasoline recovery and treating unit, the gas oil is sent to a hydrog-
enation unit, and the bottoms oil is sent to the delayed coking unit.10
3;
4.
ScG i N
LAST Li,\E
OF TEXT I
Basis of Retort Design and Operation
The basis for design and operation of the Tosco-II retort is jgiven in
Table .19. ,20/21
r
Retort Gas Characterization
The average composition of the gas produced by the Tosco-II retort is
given in Table 20. .22
A O '0 • •
43
EPA-2S7 (Cin.)
(-5-76)
PAGE NUMBER
BOTTOM OF
IVAGE ARE-
OUTS'DE
DIMENSSG:.
FOR TASLEG.
AND ILLUS-
TRATIONS
-------
Table .19. . Design Basis for Tosco-II Retort
Shale-oil production -8,872 m3/dayb
Oil-shale throughput (6 retorts) • 60,000 tonnes/day
i
Oil-shale assay 158 Si /tonne
Method of -retorting Indirect mode;
Method of heating Pyrolysis by heated balls
Gas production ratio 38.5 sm3/tonnec
Gas production rate (single retort) 26.7. f» sm3
Gas mole ratio, CO /H_S 4.95 !
rt
Heat content, HHV - 35.26xl06 joules/sin3
Pressure : Slightly positive
Temperature 60°C- ,
asee refs 20 and 21. !
v» ' -
Based on pilot study; see ref 21. Production as low sulfur distillate is
7,360 nrVday; see ref 20. j
°Calculated from data in ref 21. . i
Calculated valve based on C$ and higher MW hydrocarbons removed. |
44
-------
Table
Component
H-
2
CO
CO-
2
CH
C-H-
2 2
C' H
24
C' K
26
C3H6
f * TT
38
C.
4
c
H2S
NH
NO
2
°o
2
so_
-2
NO
cs_
2
RSH
COS
Total
H20
. • • . •
20. . Net Gas Produced by Tosco-II
Amount
•(sm^/ tonne)
7.54
1.31
t 7.83
• 7.75
b
3.21
2.91
.2.73
-1.28
2.00
c
,1.58
d
b
b
b
b
Nil6
.- - .-
"
38.38
9.59
s r •
a
Retort ;
i :
|
(dry yol %)
20.2
3.40
20.38
'
20.2
b
8.39
7.01
. 7.14
3.35
5.22
.
c
4.12
!d
b
b
\T°
t
!b
Nile ;
(35 ppm)
(135 ppm)
100 ' ;
20b!
*See ref 21. i
None reported.
cReported in oil recxjvered.
Absorbed in water phase and does not appear in gas.
e
See ref 22.
wet-gas basis.
45
-------
i OP o;
"G.
REFERENCES*
1. V. Kalcevic and J. Lankford, Pilot-Plant Operation of Gas-Flow Oil- :
Shale Retort, Report of Investigations 5507, U.S. Department^of Interi-,
or. Bureau of Mines (April 1959). |
. 2. J. W. Smith, Analytical Method for Study of Thermal Degradation of Oil
Shale, Report of Investigations 5932, U.S. Department of Interior,
Bureau of Mines (1962).
i h
3. J. W. Smith,, Geochemistry of Oil-Shale Genesis in Colorado's, Piceance
Creek Basin, Rocky Mountain Association of Geologists—1974 Guidebook.
4. J. W. Smith, Ultimate Composition of Organic Material in Green River
j Oil Shale, Report of Investigations 5725, U.S. Department of! Interior,
| Bureau of Mines (September 1960).
i 5. E. R. Bates and T. L. Thoem, editors. Pollution Control Guidance for
• Oil Shale Development, compiled by Jacobs Environmental Division, July
U.:_ —1979. — _ 7 — "j -
EEG'.N
LAST L!\<
OF TEXT
6. J. W. Smith, "High Temperature Reactions of Oil Shale Minerals and
Their Benefit to Oil Shale Processing in Place," pp 100—112' in llth
Oil Shale Symposium Proceedings, November, 1978.
7. J. B. Jones, "The Paraho Oil-Shale Project," 81st National Meeting of
AIChE, presented at Kansas City, MO, April 11—14, 1976. ;
i i
1 i i
8. Private conversation on Sept. 27, 1979, between S. W. Dylewski, IT
Enviroscience, Inc., and R. N. Heistand, Development Engineering, Inc.
9. ' Telephone conversation on Dec. 19, 1979, between S. W. Dylewski, IT
Enviroscience, Inc., and R. N. Heistand, Development Engineering, Inc.
: l '
10. C. Shih et al., Technological Overview Reports for Eight Sha'le Oil
Recovery Processes, EPA-600/7-79-075 (March 1979).
11. Development Engineering, Inc., Air Emission Source Construction and
Operation Permit Application, submitted to USEPA Region VIII, July 5,
1978.
-I
12. K. N. Heistand and R. A. Atwood, Development Engineering, Inc., "Paraho
Environmental Data," prepared as Part II — Air Quality, under USDOE con-
tract EP-78-C-02-4708.AOOO (February 1979).
i i
13. R. A. Loucks, Occidental Vertical Modified In Situ Process for the
{Recovery of Oil from Shale Phase I, Occidental Oil Shale, Inc., Grand
Junction, CO, DE-FC20-78LC10036 (November 1979).
I i
14. Private conversation on Sept. 26, 1979, between R. Lovell, S. Dylewski,
_ _ and V. Kalcevic of IT Enviroscience, Inc., and C. Bray of Occidental
-- ' Oil Shale, 'Inc. i _ •• _ ; __
!
f. -3 'O"
i *.• O O
EFA-2S7 (Cin.)
i.
46
PAGE NUMBER
BOTTOM OF
IVAGE ARE.-
FOR TABLED
•AND ILLUS-
TRATIONS
-------
15.
16.
R. E. Thomason, Occidental Oil Shale, Inc., letter request for a Pre
vention of Significant Deterioration Permit, addressed to T.iL. Thoem,
USEPA, Mar. 24, 1980.
Telephone converstion on Feb. 5, 'l979, between S. W. Dylewski, IT Envi-
roscience. Inc., and L. Morriss, Geokinetics, Inc.
"177
18.
19.
20.
22.
~r.~Morriss, Geokinetics,"inc., "letter" dated Feb. 13, 1980, to
S. Dylewski, IT Enviroscience, Inc. i
R. Lathrop and T. Thoem, Environmental News, news release by U.S. Envi-
ronmental Protection Agency, Region VIII, Aug. 8, 1979.
t i
J. Pownall, Union Oil Company, Long Range Experimental ShalejOil Plant,
PSD Permit Application, Apr. 7, 1978, and addendum Apr. 13, 1979.
! i
A. Merson, USEPA, letter dated July 11, 1979, to M. Legalskiy Colony
Development Operation. . ;
— J-.- Whitcombe, The Tosco II -;ril "Shale Process ^presented at the~79t'h~~;
National Meeting of the American Institute of Chemical Engineers
March 16—20,. 1975. !.
M. Legatski, Colony Development Operation, letter dated Jan. 26, 1979,
to T. Thoem, USEPA.
EEGiN
LAST LINE
OF TEXT I
*When a reference number is used at the end of a paragraph or Ion a head-
ing, it usually refers to the entire paragraph or material under the
heading. When, however, an additional reference is required for only a
certain portion of the paragraph or captioned material, the earlier j C.-.TTOM ,-
reference number may not apply to that particular portion. j i'^GE A^
' CUTSI-OE
A 3 'R"
EPA-237 (Cin.)
(fl-76)
:•:$•:•...•. 47.....'i::
FOR TABLE
>AN'D ILLU5
TRATIONS
PAGE'NUV.BER
,n
-------
IV. REVIEW!OF SULFUR REMOVAL PROCESSES
TREATMENT TECHNIQUES ! .....
!•••.. ... . |
The class of processes that remove sulfurous compounds and Carbon diox-
| . »,i. . . , d.,..,, , _ ,, . . [
ide from gases is generically called acid-gas remoyal_or_gas/-sweeting __
_. - |. . - - • |
processes. Removal of acid gases and/or other gaseous impurities from
gas streams is accomplished by chemical conversion to another compound,
i . •
by absorption into a liquid, or by adsorption on a solid. ;
BEG!N:
LAST L!M
OF TEXT J
In the first method the gas is passed through a fixed bed or is con-
tacted with a liquid, whereby the impurities are converted to another
compound that can be morei easily removed. Conversion is either by
direct chemical reaction with the bed or liquid or by catalytic reac-
tion'
i direct chemica
T
I
In the second method the gas stream is contacted with a liquid, and the
I.. , : . \
gaseous impurities are either chemically or physically dissolved in the
' I ' '
liquid absorbent. The absorbent is subsequently regenerated to strip
i ' •<
the absorbed gas and is then recycled.
When the method of adsorption is used on a solid, the gas is passed
through a fixed bed of granulated solid material. The impurities are
adsorbed and held by the solid adsorbent. When the bed becomes satu-
rated, it is replaced or regenerated. ;
Numerous processes for removal of sulfur compounds have been developed
I ' i I i
from these methods and more than 30 of them have achieved cpnunerical :
I ' ' i •
importance. Each commercial process has specific advantages, disadvan-
! '- ' • i
tages, or limitations and] was developed to satisfy specific needs. i
1 i ; ' !
Some processes are designed to remove both H2S and CO2, whereas other ;
i ! : ' i
processes are more selective toward H2S or C02 removal. All liquid- i
I . | .: , i | '
phase processes will remove C02 to some degree, along with the H2S. :
Some processes are economical for only bulk removal of acid gas and ; ;y.\GE ARE-
will not achieve a high degree of acid-gas removal. Other processes ; ^"^ . ,"c.
will achieve a high degrep of acid-gas remoyal, to the ppm level, but_____; FC-.-: TABLF:
I
if-!L _ I.
EPA-2S7 (Cin.)
(4-73)
48
•AND ILLUS-
TRATIONS
PAG
-------
TOP Or
BEGIN
LAST LIME
OF TEXT
a.
are not economical for gases containing large amounts of acid gases. —
Still other processes are effective only if the C02/H2S ratio is small,
whereas others are more effective if the C02/H2S ratio is high. Some
processes cannot be.used at a low pressure, but some are equally effec-
tive at all pressures. Most processes'must be_carriedi_out jat_near^____
ambient temperatures; others require cooling or chilling of the feed
gas; some can be used at elevated temperatures, whereas others must be
tased at elevated temperatures.
&11 H2S removal processes can be divided into the general categories of.
direct- and indirect-conversion processes. In the direct-conversion
processes the sulfur compounds are directly oxidized to elemental sul-
fur or another compound that can be separated and recovered.]^ In_the__
indirect conversion processes the acid-gas components are removed from
I
the feed gas and recovered as a separate stream, which is subsequently
processed for recovery of the sulfur. Sulfur is recovered from the
concentrated acid-gas stream by the Glaus sulfur recovery process or by
one of the direct-conversion processes. Both these categories can be
further broken down into either gas-phase dry-bed processes pr liquid-
phase (wet) processes (see .Fig. 6).
Direct-Conversion Processes
Dry Bed Several dry-bed processes have been described for 'direct
removal of H2S from hydrocarbon gases based on the .classic Claus reac-
tion of H2S and S02 to form elemental sulfur.1 The application of
these processes is limited because of possible plugging of the beds by
condensable hydrocarbons and/or other impurities in the gas.. None of
these processes are known to have achieved commercial importance.
a
The Claus process, while a dry-bed process, is not in the true sense
gas-treating or sulfur-removal process and therefore is not included in ;
! i i j bOT lUi.s v_'~
this discussion. The Claus process is used to recover sulfur from j ;MAGE ARC
hydrocarbon-free acid-gas streams containing large amounts of H2S and j ^.r^fn-
is discussed in Sect. IV-B. i i 1 FOR TABLE
i
PAGE NUMBER
EPA-287 (Cin.)
(4-70.)
-------
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1
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-------
G GUif'L L/l-iu::)
t'LO I
b" Liquid-Phase Direct-Conversion Processes — Liquid-phase direct-conver-
sion processes usually are best suited to gases containing low concen-
trations of H2S, especially in the presence of substantial concentra-
tions of C02. Very little CO2 is absorbed by the solution, and thus
t;.hej .Processes s?.lectively remove the H2S._ The principal disadvantage
is the relatively low capacity of the solutions for H2S, which results
in large liquid circulation rates and requires large equipment for
separating and processing precipitated sulfur. Generally plants of
this type are not economical when the sulfur production rate 'exceeds
9 tonnes per day. Another disadvantage is that the considerable reac-
tion heat generated by oxidation of H2S has to be dissipated I at low
temperature and cannot be recovered.2 The quality of sulfur -produced
_1S *owfr than that of the sulfur obtained with the indirect processes,
and the wastewater from purge streams is a major problem.
i '• i
Direct-conversion liquid processes have substantial economic ladvantages;
over direct-conversion dry-bed processes in that they require less
space, eliminate the high cost of bed replacement, and produce higher
quality sulfur. Some liquid processes are capable of producing a
treated gas of high purity equal to that obtained with dry-be!d proc-
esses.
2. Indirect-Conversion Processes
a.
[JEGIN
LAST LINE}.
OF TEXT •••••>-
I
Dry Bed The dry-bed indirect-conversion processes are generally
applied on a batch-loaded basis, where the bed is removed from service
when it is_loaded and the sorbent is replaced or regenerated. The dry-
bed processes are selective for H2S and generally achieve a high degree
of removal. These processes are limited to the treatment of gas
streams of small volume or to gas streams containing relatively low
i • i
amounts of sulfur compounds. Regeneration of the bed is only partly
achieved, and the bed eventually becomes plugged with sulfur and must
be replaced. The economics generally limit application of thtese proc-
esses to plants recovering less than 10 tons of sulfur per day.3
t? 3/s"
J
EPA-287 (Cin.)
(4-7G)
I
51
PAGE MLVBER
-------
TV .•••'• ("iMit!" Cl. ,-''-'|
I i . .•_•_: '. !•_- > tJi- -'. : . - !
;IN
. i .'s
.c
jb.
The advantages of dry-bed processes are simplicity, low cost, and ease
[
of operation. Essentially complete removal of H2 can be achieved, and
some processes will selectively remove H2S without removing 'any CO2.
The disadvantages are the large space requirements for the equipment,
high maintenance costs for bed replacement, and problems with_sulfur
recovery. ;
Liquid-Phase Process Involving Chemical Solvents Chemical solvents
involve absorption by reversible chemical reaction of the H2S with a
water solution of the absorbent. The following three types of chemical
solvents are used: ••
3£G!N
LAST LINI
OF TEXT :
Alkanolamines Many types of alkanolamines are used as absorbents
(Eig. 6 ), and numerous processes and variations of them liave been
applied, depending on the physical conditions of the gas to be treated,
its composition, and its purity requirements. Generally alkanolamine
processes exhibit high reactivity with H2S and achieve a high degree of
removal. The processes are not sensitive to pressure, and operate at
near-ambient temperatures. Organic sulfur compounds are removed to
some extent, but the lighter molecular weight amines may forjn nonregen-
erative compounds when absorbing organic sulfur or cyanide compounds.
The lighter amines, though less expensive and capable of higher absorp-
( J
tion capacity, are not selective in absorbing H2S, present corrosion '
problems in stripper and heat-exchanger surfaces, and are plagued by ]
higher solvent vaporization losses than are the heavier amines. i
Alkanolamine processes are generally preferred for treatment! of gases
containing moderate amounts of acid gas at near-ambient temperature and
pressure. ;
Alkaline salts Alkaline salt processes employ an aqueous solution of
a potassium or sodium salt that forms a buffered solution with a pH of
about 9 to 11. The weak alkaline solution chemically absorbs the acid-
gas components. These processes are of two types: those carried out
at low (ambient) temperature and pressure and those carried but at
elevated temperature and pressure. The low-pressure processes have
\1 Z's2''?m :
i p.-
AND -iLL'
—r- • T '~ • '
i h — i i,_..
EPA-287 (Gin.)
(4-7C)
PAGE NUMBER
-------
; 1.1."!
£ I.1--:': Of
TVPt.'.'G '-:
BEGIN
LAST LINE
OF TEXT ::>-
c.
given way to more efficient processes and are not currently used. The
hot processes have an advantage when the feed gas is hot (up, to 149°C )
and under pressure (<100 psig). Stripping is accomplished partly by .
flashing, and thus the energy required for regeneration is Ipwer than
that for amine systems. The alkaline salt solution does not degrade _
significantly, and only minimum purge and makeup are required. Organic
sulfur compounds are removed by the process, and it can be m|ade partly
selective toward H2S removal. The alkaline salt processes are usually
the best ones to use when bulk removal of C02 is required, j
' I :
i i
Aqueous ammonia Aqueous ammonia processes are generally applied when
the gas to be treated contains ammonia. The ammonia contained in the
gas is absorbed simultaneously with the H2S and thus serves as the^ ac-
~tive agent in the absorber solution. No chemical additives are re-
quired. Ammonia can be recovered in addition to the acid-gas compo-
nents. Under certain operating conditions H2S can be selectively
removed/ Removal of up to about 97% of the H2S can be achieved; how-
ever, organic sulfur compounds are only partly absorbed.
Liquid-Phase Processes Involving Physical Solvents Physical-solvent
processes are based on H2S, CO2, and minor gas impurities being more
soluble in certain anhydrous organic solvents than are the fuel-gas
components', i.e., hydrocarbons and hydrogen. Since their solubilities
increase with pressure, the processes generally require high pressure
to be economical. The solvent is regenerated by pressure reduction,
gas stripping, or heat applied to produce a concentrated stream of
absorbed gas.
i
Most physical solvents have higher solubility for H2S than for CO2, and
a high degree of selectivity can be obtained with some processes. The
solubility of hydrocarbons, however, increases with molecular weight.
Consequently hydrocarbons above C4 are also largely removed with the
acid gas.
3/3"
!_
53'::::?::::>
ECT7O?.' 0=
OL'TSICE
.1 FC", TAE.E:'
•;-AND ILLUS-
| TRATiGNS
PAGE
EPA-227 (Cin.)
(4-7G)
-------
-—;•<: p |
h£AD. i
B.
BEGIN
LAST LINE
OF TEXT i..
The advantages of physical solvents over chemical solvents are lower —
heat and power consumption, greater removal of organic sulfur compo-
nents, higher selectivity for H2S, no purge or wastewater streams, and
little corrosion. The disadvantages are that a high-feed gas pressure
i
:Ls required, heavier hydrocarbons are absorbed, the solvent loss is
high, and the solvent cost is high.
i • !
il :
Physical-solvent processes are most economical when the feed gas is
available at high pressure. These processes are usually usep for bulk
removal of acid-gas constituents when the acid-gas impuritie's make up
an appreciable fraction of the total gas stream, making the cost of
removing them by heat-regenerable chemical solvents less attractive.
When high-purity treated gas is required, additional solvent; regenera-
tion steps beyond simple flashing are required.
PROCESS FOR RECOVERING SULFUR
The Claus sulfur process is the principal process used for recovery of
sulfur from acid gases produced by the indirect-conversion processes.
i 1
The Claus process is discussed in some detail since it is important
that the operation, capabilities, and limitations of the process be
understood before fuel gas desulfurization processes are described.
A flow diagram of a conventional Claus system is shown in Fig. 7.
The basic Claus reaction is
2H2S + S02
2H2O
The conventional process is carried out in stages: the firslt stage is
' i '
at high temperature in a thermal oxidation furnace, where one-third of
i ; . ;
the H2S is combusted to SO2 in the presence of air; the reaction in the
i ! !
succeeding stages occurs at lower temperatures in the vapor phase by
I ! i
catalytic oxidation. Usually two or more catalytic stages are em-
! ! '-.
ployed. Sulfur is recovered after each stage by the gas streams being
!• i !
cooled and condensed. The gas is then reheated to the operating tem-
! i "—
perature of the succeeding stage. When the acid-gas feed is mostly
ri o 'Q"
i y3'8 •
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H2S, sulfur conversions of up to 70%4 can be obtained by th6 thermal .
oxidation step alone, and overall conversions of up to 99%5 have been
obtained by use of four catalytic stages. The average sulfur recovery '
yield obtained with properly designed and operated systems usually
ranges from 93—97%.6 The remaining 3 to 7% of the_sulfur_is present
in the Claus unit off-gas. ; • • > j
Recent developments have improved the process so that sulfur recovery
levels of above 99% are routinely achieved by employing tail-gas treat-
ment;5 acid-gas streams containing as little as 5% H2S have,been proc-
essed with modified Claus systems.7 ;
A necessary condition for successful operation of the thermal oxidation
furnace is that the acid-gas mixture be rich enough (high in H2S con-
tent) to ensure stable combustion at the required reaction temperature.
The gas stream entering the first-stage catalytic converter should be
at its stoichiomatic ratio of 2 parts of H2S to 1 part of S02 and be
free of oxygenated impurities (S03), unburned heavy hydrocarbons, soot,
and excessive ammonia. An imbalance of either H2S or S02 will cause
the excess component to pass through the system unreacted. Oxygen
: * , ,
causes sulfonation of the catalyst, which results in loss of catalyst
activity. Soot and carbonaceous material and ammonia, through forma-
tion of ammonium sulfate, cause the catalyst beds to become!plugged.
Unburned hydrocarbons that condense and stick to the catalyst will also
cause loss of catalyst activity. \
The operating temperature required for the thermal oxidation furnace is
largely governed by the impurities in the acid gas. If the acid gas
consists of H2S, COS, CS2, and C02, a furnace temperature of 700 to
800°C is sufficient. If the gas contains hydrocarbon compounds,
I ! '
ammonia, and other combustible impurities, temperatures of 1000 to
1200°C are required to assure complete combustion of the impurities and
elimination of soot formation. The presence of combustible1 impurities
i . i
in the acid gas increases the amount of air that must be introduced to
_JL_
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oi- ,"•-/::"
the combustion chamber, and their products of combustion greatly add to
the quantity of process gas entering the catalytic stages. ;
i i
I ' ! o
Methods for Processing Acid Gases Containing 16 to 100-6 H2S !
If the acid gas contains 50 to 100% H2S, the total yolume of|acid gas
is normally fed to the combustion chamber (the classic, or the
straight-through, process). The amount of air fed to the reaction
furnace is controlled to combust one third of the H2S and all the
ammonia and hydrocarbon impurities in the feed. If the temperature is
maintained at the proper level (1000 to 1100°C), organic imparities are
oxidized and the catalytic stages operate without significant problems.
As the H2S content of the acid gas decreases, the lower the temperature
obtained in the reaction furnace becomes. Longer residence^time_in_the
furnace is required to compensate for this lowered temperature, and
thus larger chamber volumes are needed to maintain the dyanmic equilib-
rium. Below 35% H2S, the combustion required to produce the;SO2 for
the Claus reaction is no longer possible without special modifications
being employed.
t
I
Below 50% H2S, the Claus process is usually modified. If the H2S con-
tent is between 16 and 50%, the split-flow process is often applied.
j ' ;
With this process only one-third of the acid gas is fed to the thermal
oxidation furnace. Stable combustion is obtained by the H2S; being com-
pletely combusted to S02. The other two-thirds of the acid gas by-
passes the furnace and is mixed with the gases from the therknal oxidiz-
er before they enter the first catalytic converter. However!, impuri-
ties in the acid-gas stream are detrimental to the downstream catalytic
section of the Claus plant and can affect the service life of the
catalyst and the purity of the sulfur produced. If impurities are
present in the acid gas, the complete stream should be fed to the com-
bustion chamber as in the straight-through process. Preheating the
combustion air and/or acid-gas stream or burning supplemental fuel or
sulfur will enable stable combustion to be obtained.
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b.
c.
Methods for Processing Acid Gases Containing Less Than 16% H2S ,
•JLf the acid gases contain less than 16% H2S (called lean acid gases),
the methods for obtaining stable combustion involve the following proc-
esses:
Preheating Process The acid gas and/or combustion air is preheated
before it enters the thermal oxidizer. Preheating is accomplished by
heat exchange with the hot gases exiting the combustion chaniber, by
steam or by externally fired heaters. Only a limited amount of heat
can be obtained by preheating. For lean acid gases it may be nessary
to use supplemental fuel in direct-fired burners. However,Jburning the
fuel adds to the cost. x>.f. operation and to the volume of gases that must
be handled by the downstream components. j
~T ^-£.'&*! :
Sulfur Burning—By-Pass Process The Claus thermal oxidizer as such .is
eliminated. The S02 required for the reaction is obtained by burning
liquid sulfur in a separate combustion chamber;. This has the advantage
that, stable combustion is obtained in the sulfur burner and' that the
gas produced is mostly SO2, depending on the purity of sulfur burned.
Off-grade sulfur can be burned since the carbonaceous material is
oxidized to carbon dioxide and water. The acid-gas stream is preheated
in a gas-fired heater. The hot acid gas and the S02 from the sulfur
burner are then mixed and fed to the catalytic converter. Application
; i
of this process is limited to relatively pure acid gases since even
small quantities of hydrocarbon can cause carbon and tar to deposit on
the catalyst and/or the sulfur quality to be degraded. ;
Oxygen Process—Oxygen in place of air is used in the thermal oxi-
dizer. This eliminates the ballasting effect of the nitrogen; thus j
less heat is required to maintain equilibrium and the volume of gases :
t * i • '
passing to the catalytic stages is reduced. Depending on the quantity '-.
of combustible material in the feed, oxygen is mixed to the acid gas or I
to a portion of the acid gas, which is then fed to one or more burners, j
If insufficient heat is generated because of lean-acid-gas Imixtures, j
the supplemental fuel must be_burned._ —; _ ,-J
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d. Sulfur Burning—Straight-Through Process A portion of the acid gas is
mixed with excess air, depending on the quantity of combustible impuri-
ties in the acid gas, and is then fed to a special burner. The remain-
ing acid gas is fed into the combustion chamber. Liquid suljfur is
burned in a special burner to supply most of the S02 required for the _
Claus reaction and the heat required to stablize the combustion proc-
i
ess. The quantity of sulfur burned is limited since the H2S:SO2 ratio
i . ' 1
of 2:1 must be maintained. Supplemental fuel may have to be, used if
the heat input is insufficient for thermal equilibrium to be: maintained
in the furnace. •
3.
BEGIN
LAST LINE
OF TEXT J-
Cost Data
Table IV-1 gives the relative capital and operating costs of the vari-
ous process options, based on cost data reported by Fischer.8 The
sulfur-burning by-pass process results in the lowest capital and oper-
ating costs of all processes available for processing lean acid gases.
However, its application is limited to relatively hydrocarbon-free acid
gases. When the H2S and hydrocarbon contents are low, the oxygen proc-
ess is preferable provided that the oxygen can be obtained at low cost.
The preheating process is used for those applications in which the acid
gas contains hydrocarbon impurities and the H2S content is moderately
low. '
I
I t . ;
The sulfur-burning straight-through process is the most expensive proc-
ess. Since the entire gas stream is introduced to the combustion cham-
i . i
ber, its temperature must be raised to the furnace operating tempera-
ture. The higher the furnace operating temperature, the higher the
operating cost and the larger the equipment required.
4. Claus Tail-Gas Treatment Processes
i
The tail gas from a properly designed and operated Claus sulfur recov-
ery plant may contain from 8,000 to 28,000 ppmv of sulfurous, compounds '
(H2S, SO2,'COS, CS2, S vapor).6 In order to comply with the; more j ^ ;j.-".-.--
i i ; ...','.!"%-- '
stringent current regulations, many processes have been developed to < '-'-'i- i c
: i ; 1 Di'.'ENiiO'.
clean up Claus tail gases. ! | j PQR TAL.LE".
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Figure 8 shows the principal Claus tail-gas treatment processes.
Three basic approaches to tail-gas cleanup are employed. With the
first approach all sulfurous compounds are reduced to H2S in; a high-
temperature catalytic converter. The gas is then cooled and the con- '.
densed water removed. After the H2S is cooled,_i^is_eitherabsorbed~3.
and subsequently regenerated as a concentrated acid-gas stream for
recycle back to the Claus unit or is reacted and recovered as elemental
I
sulfur. With the second approach all sulfurous compounds ark oxidized
\
to S02 in a thermal oxidation furnace. The SO2 is then either absorbed
eind subsequently regenerated as a concentrated S02 gas stream for re-
cycle back to the Claus unit or is reacted to form another compound,
which is then recovered as by-product. The third approach is to extend
the Claus reaction at lower temperatures. When a catalytic system is
operated at a lower temperature than normal (near or below the sulfur
condensation temperature), the thermodynamic equilibrium for the forma-
tion of sulfur from H2S and S02 is further enhanced. Processes based
J
on this approach are carried out in either the liquid phase with a
catalyst present or a gas phase, in which a solid catalyst bed is used.
i i i
' : :
The choice of which tail-gas treatment process to use depends on many
factors, principally those given below:
1. the composition and volume of the tail gas, ;
2!. whether the system can be integrated as part of a new installation
or whether it is an add-on to an existing Calus plant, i
; j
3. the initial investment cost and/or the operating costs,,
4. the utilities required, the sulfur product produced, and/or the
wastes produced,
5. the performance, reliability, and operating range of the system.
|
The Shell Claus off-gas treating (SCOT) process, the Beavon sulfury
recovery process (BSRP), and the Wellman-Lord (W-L) process are the
most widely applied processes. If the tail gas is low in C02, the SCOT
process usually produces the best economics of the three. If the gas
1
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^ .^MAGI
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contains large amounts of C02, the W-L or the BSRP may be thje best
choice. !
1m excellent discussion of the various Claus tail-gas treating proc-
Cisses is given in a paper presented by Goar at the Fifty-Seventh Annual
Convention of the Gas Producers Association.6 I
BEGIN
LAST LINE
OF TEXT -i-
J. _ 6-1/2" }-
I
9-1/8"
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;c.
REFERENCES*
SECTIONS '.?!
' ;:3T
A. L. Kohl and F. C. Riesenfeld, Gas Purification, 3d ed., pp 421—422,
Gulf Publishing Co., Houston, 1979.
1
Ibid., p 441.
" I
"Ibid.', p 382. ~ '
! *• Sulfur Recovery Catalysts for the Claus Catalytic Process. :Pro-
Catalyse, Technical Documentation, Rhone-Poulenc, Paris (September
1978). : . i
5.
! 6.
7.
8.
BEGIN
LAST LINE
OF TEXT t
A. E. Chute, "Tailor Sulfur Plants to Unsual Conditions," Hydrocarbon
Processing 56(4), 119—124 (April 1977). ! >
.— < ; !
G..Goar, "Current Claus Tail Gas Clean-Up Processes," pp 152—163 in
Proceedings of the 57th Annual Convention of the Gas Producers Associa-
tion, March 20—22, 1978, New Orleans, LA. | j
P. Grancher, "Advances in Claus Technology," Hydrocarbon Processing ;
57(7), 155—160 (July 1978). .
' i 1 i
H. Fischer, "Sulfur Costs Vary with Process Selection," Hydrocarbon
Processing 58(3), 125—129 (March 1979>.
*When a reference number is used at the end of a paragraph or on a
heading, it usually refers to the entire paragraph or material under
the heading. When, however, an additional reference is required for
only a certain portion of the paragraph or captioned material, the
earlier reference number may not apply to that particular portion.
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BEGIN
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OF TEXT l'.-r
,4 V. FACTORS INFLUENCING CHOICE OF PROCESS
i • i . :•
i .. !. • ' , .
In selecting a gas-treating process to remove sulfur compounds many
factors have to be considered. The most important are the following:
1. the product-gas specifications,
2. the quantity of acid-gas components (CO2 and H2S) contained in the
gas,
i
3. the influence of impurities such as COS, CS2, mercaptans (RSH),
NH3, and HCN, !
4. the quantity of heavier hydrocarbons in the gas,
i
5. the condition of the feed gas (temperature and pressure),
I ,__•_
6. the capital and operating costs of the process.
A. PRODUCT-GAS SPECIFICATIONS
i
1
1. Sulfur Compounds
i
Sulfur compounds are removed to prevent air pollution by SO2i when the
i i i
gas is combusted. They are also 'removed for safety and to prevent
1 !
corrosion and odor problems. For^ natural-gas distributed byi pipeline
• ' • - • • 3
the total sulfur content is reduced 4.6 grams, or less, per 100 m
1 ; - . • I
of gas. If the gas is used as chemical feed stock, sulfur compounds
i i
often have' to be reduced to less .than 1 ppm to protect sensitive cata-
lyst systems.
Generally the cost of removing the sulfur increases as the degree of
removal required increases. Achieving a high degree of removal often
i [
requires sacrificing other desirable process features, such ^s H2S
selectivity, process simplicity, 'or favorable operating economics. The ;
i i '. I
higher the degree of sulfur removal required the more limiting the
process selection becomes.
65
....o ••
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K-2AD;
,' >""7 ! *
co, .. j :
Carbon dioxide and inert gases have no harmful effect other than that
of reducing the heating value of the gas. Often C02 can be jtolerated
and does not have to be removed. i
BEGIN
LAST LINE
OF TEXT r.
B.
TOP Or
IMA&c
"A!-FA
j 3. COS, CS2/ and RSH j i
Carbonyl sulfide, carbon disulfide, and mercaptans contribute to the
total sulfur content of the gas and generally must be removed to the
same degree as H2S.
i
4. Dew Point
In a saturated gas a drop in .temperature or an increase in pressure
_.._ will_cause water to be condensed in, the..lines.. ..Therefore-the .gas J
should be precooled to the lowest temperature to be encountered in the ;
system. Similarly, condensable hydrocarbons could condense out if the I
gas conditions dropped below their condensation point.
b-i/o • •
I
5. INH3 j
If not removed, ammonia cbuld be a problem in that NO compounds could
be formed when the gas is combusted. Also, ammonia could be detri-
mental to some desulfurization processes. :
ACID-GAS COMPONENTS
The most important factor in choosing a gas-treating process is the
quantity of acid-gas components, CO2 and H2S, contained in the gas and
I ! !
the volume ratio between these two components. If the total quantity
! i .
of sulfur contained in the gas is small, less than 9 tonnes/day, a dry-
' ' ]
bed or direct-oxidation process may be more economical. If•the quan-
• i j
tity of sulfur is greater than 9 tonnes/day, an indirect process may be
more economical.
Although the Claus process is preferable for recovering sulfur from
;
concentrated acid-gas streams produced by indirect processes, it re-
EOTTOM Oc
iVAGE ARt"-'.
, OUTSIDE
quires an acid-gas feed containing greater than 15% H2S for ]effective—J DIMENSION'
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OF TEXT i'r
__ operation. The higher "the concentration of H2S in the acid gas the
more effective the Claus system will be. !
! .r "
Since all liquid-phase indirect processes absorb some C02 alpng with
~~ the H2S, an indirect process ^elective J^owar^d H2S should_be_used ifjthe
CO2/H2S ratio in the feed gas is greater than 1 and is essential if the
ratio is greater than 6. If it is necessary to remove the Cp2 along
with the H2S, when the C02/H2S ratio is high, a process selective
toward H2S removal should be used followed by a process thati can remove
i l
bulk quantities of C02. If the acid-gas stream contains less than
about 5% H2S, it is usually more economical to process the acid-gas
• . . j • '
stream by a liquid-phase direct-oxidation process than by thfe Claus
process. r -
C. INFLUENCE OF IMPURITIES
Host direct-recovery processes will not remove COS or CS2 and will only
partly remove mercaptans. The alkanolamine processes will remove j
organic sulfur compounds, but the lighter alkanolamines (MEA|) and to j
'• ' ; !
some extent DEA form nonregenerable thiosulfates, which degrade the |
1 * | !
solvent when COS or CS2 is present in the feed gas. Most alkaline salt;
processes remove COS and CS2 by hydrolyzing them to H2S, but mercaptans .
are only slightly absorbed. The physical-solvent processes [effectively .
i
absorb organic sulfur compounds.
If ammonia is not removed before the acid-gas absorption step, it will
be absorbed with most processes and end up in the acid-gas stream.
Most ammonia in the gas can be removed with the condensates ±n the pre-
i i
cooling step. Small amounts of ammonia can be tolerated by ;the Claus
unit if the quantity of ammonia is small compared to the quantity of
1H2S in the, acid-gas stream.
I
Hydrogen cyanide is largely co-absorbed with the other acid gases and, . ^
if the absorbent solution does not contain oxidants, will end up with j :;;AG=AR£-
I i I rii 'TS'DE
the acid-gas stream. In the direct-recovery processes HCN is convert*^ D:"-NSIGiv
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to stable thiocyanates, which will build up in the solution 'and require
that a portion of the solution be purged.
! .••••! I '
If condensible hydrocarbons are absorbed, poor-quality black sulfur
will be produced by direct-recovery processes.. Hydrocarbons in the
D.
E.
5-ZG1N
LAST LINE
OF TEXT :-,
acid-gas feed to the Claus unit will increase the combustion air re-
quired and necessitate that the Claus thermal oxidizer be operated at a
higher temperature to assure complete combustion of hydrocarbons. If
the hydrocarbons are only partly oxidized, the catalyst beds can become
plugged or a poor-quality sulfur can result. ,
* ' i
Generally the alkaline salt and alkanolamine processes will not appre-
ciably absorb hydrocarbons. The physical-solvent processes,; however,
will readily absorb hydrocarbons heavier than butane.
!
CONDITION OF FEED GAS
i !
Most processes are carried out at near-ambient temperatures 1(16 to
54°C). The physical-solvent processes are most effective at lower
temperatures, and some require that the gas be chilled. In most of the
currently applied alkaline salt processes elevated temperatures of up
i
to 149°C can be used.
!
j
The physical-solvent processes arid most alkaline salt processes require
a high pressure, above 150 psig, to function efficiently. The alkanol-
1 | !
amine and direct-recovery processes are generally independent of pres-
* •
sure and can be used equally well over a wide range of pressures.
CAPITAL AND OPERATING COSTS
The bottom line in choosing a process is the capital and operating
costs. Generally the more sulfur contained in the gas the less the
\ ; - ;
cost on a unit basis ($/tonne) to remove it. If sufficient sulfur com-
pounds are present in the gas, an indirect-recovery process 'followed by
i I
a Claus sulfur recovery unit will provide the best economics. The
physical-solvent processes are more economical when the acid-gas con-
tent is high or when the gas exists at high pressure. The alkaline
I
d 3/8"
.1.
J
EHA-287 (Cin.)
(4-76)
PAGE NU'V.BER
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: L;1-' t I.V_«.«. V_
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salt processes have an advantage over other processes when the gas is-
at high pressure and high temperature and contains moderate amounts of
acid gas. The alkanolamine processes are most economical whbn the gas
is at low pressure and contains little CO2. For gases that contain
little H2S or that have very high CO2/H2S ratios the direct-conversion.
processes may be more economical than the indirect-conversion proc-
esses. If complete removal of H2S is required and the gas contains
little H2S, the solid-bed processes may be the most economical ones.
9-1 "8"
BEGIN
LAST LINE
OF TEXT i
« •: /o-
5 - °
I_
EPA-2^7 (Cin.)
(4-76)
PAGE NL1.V.CER
BOTTOM OF
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OUTSsDE
^_l FOR TABLE'!
•ILLUS-
J TRATiONS
-------
EEGIN
LAST UN:
OF TEXT J
i. •'. i •.;•:
.!'.„ VI. DUTY REQUIREMENTS FOR OIL-SHALE RETORT-GAS-DESULFURIZATION SYSTEMS —
A. CLASSIFICATION OF OIL-SHALE RETORTING PROCESSES j
From the standpoint of removing sulfur compounds from oil-shale retort
\ gases all retort processes can be divided into two_broad_and_ widely
separated categories: those in which direct-fired retorts are used
either above or below ground (in-situ), where combustion occurs within
'the retort; and those in which indirect-heated retorts are used and the
shale is retorted in the absence of air. [
: ' ' •
The application of desulfurization technology for gases produced from
(direct-fired oil-shale retorting processes is distinctly different from
that required for natural gas, coke-oven gas, refinery gas, tor gas ;
produced by coal gasification or, for that matter, for gas produced :
from indirect-heated oil-shale retorting process. Typical compositions
of the gases are shown in Table 22. -1—3 The gas from direct-fired .
retorts contains large amounts of inert components, over 63%'N2 and 22%
CO2, but the H2S content is less than 0.3%; it also contains large •
amounts of ammonia and unsaturated hydrocarbons, such as acetylene, j
ethylene, propylene, butylene, and butadiene. More than 40 hydrocarbon
• compounds have been identified in the gas; it is saturated with water j
and contains some oxygen and trace amounts of sulfur species other than j
H2S.
Table 23 ' gives the range of gas compositions from direct-fired re-
torts, and Table 24 gives the range of those produced by indirect-
heated retorts. The overiding factor that separates the two groups of
processes and that will be dominant in selection of sulfur removal
processes is the C02/H2S ratio. For direct-fired retorts the C02/H2S
ratio ranges from 76:1 to more than 165:1, thus requiring a!sulfur
removal process that will selectively remove H2S. Indirect-heated
retorts produce a CO2/H2S ratio in the range of 4.3:1 to 5:1 which
would allow a nonselective process to be used. As defined in Table 25
H2S selectivity is the molar ratio of H2S absorbed to CO2 absorbed.
If an indirect recovery process were to be used in conjunction with_a.
, j...-...•.•.-.•.•.
\T &w'"'76'-x$:
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EPA-287 (Cin.)
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-------
Table 23... Range of Gas Compositions from Direct-Fired Retorts'
Oil assay
Gas ratio
Composition (dry basis) ,
Carbon dioxide (CO_)
Hydrogen sulfide (H_S)
Carbonyl sulfide (COS)
Ammonia (NH )
Water (HO) (wet basis)
Gas mole ratio
C02/H2S
NH3/H2S
Temperature
Higher heating value (dry gas)
62.5 to 125 5,/torine
224 to 405 smVtojnne
23 to 39 vol %
0.07 to 0.24 vol %
0 to 40 ppm i
0.6 to 0.7 vol % i
13.5 to 15 vol % '
76:1 to 165:1° i
2.3:1 to 2.4:ld
57 to 77°C j
1.52xl06 to 4.13xl06'joules/sm3
Based on Paraho PON, Occidental Oil Shale (Gxy) retorts 7 and 8,
and Geokinetics retort 18. i
Ratio for Geokinetics is 0.07:1. i
C '
High ratio occurs at start of Oxy burn and averages about 330:1
during the startup phase. At steady-state conditions an average
ratio of about 165:1 is obtained.
!
Ratio reduces to 2.4:1 at steady state for the Oxy retort.
72
-------
Table 24. Range of Gas Compositions from
Indirect-Heated Retorts
Oil assay ' 158 S, /tonne i
Gas ratio 29,9 to 38.5 sm3/tonne
Composition (dry basis) ;
Carbon dioxide (CO ) 16.6 to 20.4 vol %
Hydrogen sulfide (H2S) 3.8 to 4.1 vol %
Carbonyl sulfide (COS) 135 to 550 ppra
Carbon disulfide (CS_) 0 to 20 ppm ;
2 i
Alkyl mercaptans "(RSH) 35 to 165 ppm
Sulfur dioxide (SO ) 0 to 125 ppm !
b c ;
Ammonia (NH ) 8 vol % ;
3 i
Water (HO) (wet basis) 20 to 25 vol $
Gas mole ratio
CO2/H2S 4.3:1 to 5.0:1
NH3/H2S 2*1
Temperature 60 to 71.1 °C;
Higher heating value (dry gas) 35.20xl06 to ^2.46 joules/sm3
'Based on Tosco-II and Union SGR-3 retorts.
Absorbed in sour water condensed with the oil.
« '
'Calculated value.
73
-------
Table 25 - Selectivity Data
Selectivity
(H S feed - H S treated gas) fR^> feed H2S Claus gas/H2S feed
(CO feed - CO treated gas) /CO., feed ~ CO Claus gas/CO_ feeii
22 22 2 ,
CO_ feed H S Claus gas mole % H S absorbed
H S feed CO_ Claus gas ~ snole % CO, absorbed
22 ^
For direct-fired retorts
CO_ feed
feed 11
For indirect-fired retorts
CO2 feed
H2S feed = T
For Claus feed gas
>0.08 to 0.25
Selectivity required
Direct- fired retorts
.76
X (0.08 to 0.25) = 6
.165 .
1
X (0.08 to 0.25) = 13 to 41
Indirect- fired retorts
• X (0.08 to 0.25) = 0.4 to 1.25
•74
-------
Tti'{{•;(.] GUIDE SH£i:T
V-r<.£ or
Ti-xr __
i
,.».^ |
BEGIN
LAST LINE
OF TEXT E
Claus sulfur recovery process, the acid-gas feed to the Clau|s unit
should be as rich in H2S as possible and consist of at least 8 to 25%
H2S. Gases from direct-fired retorts with CO2/H2S ratios in the range
of 100:1 to 200:1 would require H2S selectivity in the range' of 8 to 50
~T_ to produce an_ acceptable Claus gas (see Table 2_5_);_ however;, gas from
indirect-fired retorts would produce an acceptable Claus gas' only if
all the CO2 were removed along with the H2S. I
1 • :
The quantity of H2S and NH3 in the gas per ton of shale processed is,
as shown in Table 26 about equal for both direct- and indirect-
retorting processes. The quantity of sulfur contained in an equal
volume of gas, on the other hand, is as much as 60 times greater for
indirect-fired retorts than for direct-fired retorts. Thus a gas
treatment system for a direct-fired retort would have to process up to
SO times more gas per ton of sulfur recovered than would a system proc-
essing gas from indirect-heated retorts.
9-1/3"
I
B. COMMERCIAL OIL-SHALE OPERATIONS
i i
The projected sizes of several commercial oil-shale processing facili-
i i !
ties are given in Table 27. All full-scale processing facilities
i i ;
will consist of multiple retort systems with several trains jof gas
treating equipment.
i
' I
The data given in Tables 22 and 23 are based on the gas discharged
from the retort after the oil is separated. The actual composition and
i i !
condition of the gas as it reaches the sulfur removal train may be
i ;
somewhat different for a full-scale commercial plant if additional
j :
product recovery and gas treating steps are incorporated. For most
direct-fired-retort processes water scrubbing will be used to remove
1 i
ammonia from the gas and to cool the gas before it enters the sulfur
removal equipment. Some developers propose to compress and ithen cool
i '
the gas to recover additional hydrocarbons. Some may elect to upgrade
the raw crude by coking and hydrotreating and/or direct hydrjogenation.
Residual gases from these operations would likely be added to the feed
•- to the sulfur removal units.
3/8"
SIS-' 75 /xg:
D:,V:=,\S;O-
FOR TAE_=
'AND'iLLUS
TRAT!G.\S
EFA-287 (Cm.)
(4-7G)
PAGE NUMBER
-------
Table 26. Comparison of Gas Treatment Duty Requirements
Duty Range Per 45,450
Gas rate (Msm3D)
Hydrogen sulfide (tonnes)
Other sulfur compounds (tonnes)
Carbon dioxide (tonnes)
Ammonia ' ( tonnes )
Total sulfur (tonnes)
Duty Range Per Million
Direct-Fired
Retorts
Tonnes of Shale
10. 2-- 18. 4
33.6—64.5
o;Q4— 2,.o
;4, 545— 32,087
38—95.
32—62
Indirect-Heated
Retorts
Processed
1.4—1.8
78.. 2— 85.4
0.6—3.3.
443-r-r545
83—106
74—83-
Cubic Feet of Gas Treated
Shale processed (tonnes) 70—126 " 734f-947
Hydrogen sulfide (kg) -30—103 .1,636—1,773
Other sulfur compounds (kg) ".5—3 12-r64
Carbon dioxide .(kg) •12i727—21,818 9,250—11,363
Ammonia (kg) M30--151 l,^?--1,818
Total sulfur (kg) 29—99.- 1-547—i; 710
76
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BEGIN
LAST L!N=
CF TEXT ::
Indirect-retort gases that have high heating values may be combined
with residual gases from other process steps and may undergo' several
treatment steps before they reach the sulfur removal equipmept. For "
instance the shale-oil hydrocarbon vapors from the pyrolysisj drum of
.J*®_Pr°Posed. Colonv.irosco":!:.I.._facility wi.ll be j?.ePar2ted_in_ a| fraction^
ator into sour water, gas, naphtha, gas oil, and bottom oil. The gas
and naphtha are piped to a gas recovery and treating unit. The other
streams are further treated to upgrade the product. The residual gases
from product upgrading units are combined with the gas strewn from the
fractionator and then compressed and fed to the gas recovery! and treat-
ing unit along with raw naphtha recovered from various treatment units.
Stabilized naphtha, LPG, butanes, butadiene/butylenes, and ammonia are
separated and recovered. The remaining gas then enters the sulfur
removal equipment. Acid gas with residual ammonia from the sour-water
r.
stripper and from the ammonia separation unit is sent directly to the
Claus sulfur recovery unit. [ <
9-1/3" " :
i ]
For the purpose of this report it is assumed that the composition of
the gas is as discharged from the retort after the oil has bien sepa-
rated, l
T
_1
bOTTCV. CF
IMAGE AF.l
OUTSIDE
DI.V-iNSiOX
FOR TA3L.U.
.78
EPA-^37 (Cin.
(4-76)
PAGE NUMBER
TRATiOX'S
-------
i.
"27'
3.
REFERENCES*
M. Gmassemi et al., TRW, Inc., Applicability of Petroleum Refinery
Control Technologies to Coal Conversion. EPA-600/7-78-190 (October
31978). ! !
~H7 Hiraoka,'Mitsubishi"Chemical" InduYtYiesTLtd., ami E. Tana"kaTnd~H.~
Sudo, Mitsubishi Kakoki Kaisha, Ltd., "DIAMOX Process for the Removal
of H2S in Coke Oven Gas," Proceedings of The Symposium on Treatment of
Coke-Oven Gas. May 1977. McMaster Univeristy Press, Hamilton; Ontario,
Canada. ; ;
I
W. L. Scheirman, "Sour Gas Treating at the World's Largest Natural Gas
Processing Plant," p. 0-8 in Proceedings of the Gas Conditioning Con-
ference , March 7—9, 1977, sponsored by Continuing Engineering Educa-
tion, University of Oklahoma, Norman. i
BUG IN
LAST LINE
OF TEXT -V
*When a reference number is used at the end of a paragraph or on a head-
ing, it usually refers to the entire paragraph or material under the
heading. When, however, an additional reference is required ifor only a
certain portion of the paragraph or captioned material, the earlier :QVTOM OF
reference number may not apply to that particular portion. '••
_ J
ECA-237 (Cin.)
(4-70)
PAGE NUMBER
•'.'AGE ARE
OUTSIDE
-C^ TAELL.
•.-AND ILLUS
' -;-.-.Tio:.s
-------
M::-\D.
A VII. SCREENING OF GAS TREATING PROCESSES —
i ! :. -
i i ;
Gases from direct-fired retorts are more limiting in the application of
sulfur-removal technology than are other gases (see Sect. VIrA). In
the course of this study it was determined by the EPA_Jthat_the greatest
immediate concern was the control of sulfur emissions from direct-fired
i
processes and that the pilot-plant design should be applicable to these
retorting methods. Therefore process selection was directed toward the
i m ;
requirements for desulfurization of gas from direct-fired retorts. j
BEGIN
LAST LINE
OF TEXT :••£_
I
As is indicated in Table .23, treating gas from direct-fired retorts
I
requires a process with high selectivity toward H2S. The initial
screening of principal commercial gas-conditioning processes! was there-
fore based on the ability of each process to selectively remove H2S in
the presence of large amounts of CO2. Those processes most capable of
selectively removing H2S were selected as candidates for morje detailed
evaluations. The processes referred to are shown on Fig. 9, a
duplicate of Fig. 6. For the purpose of the initial screening,
iselectivities near the midpoint of the selectivity range reported in
the literature were used (see Table 28 ). :
_ «
A. DIRECT-CONVERSION PROCESSES ;
I ! !
In the direct-conversion processes the sulfur compounds are directly
oxidized to elemental sulfur or are converted to another compound,
which is then separated or recovered.
i
I
1. Dry Bed j
Several dry-bed direct-conversion processes have been tested but none
are known to have been commercially applied. These processes are
I ; i
similar to the Claus process except that the S02.necessary for the
reaction is obtained by burning sulfur in an external burner, and the
t ;
gas to be treated is preheated to the operating temperature >of the
catalyst beds. To maintain high conversion efficiency the catalyst
beds are regenerated more frequently than is required by the Claus
process. One version is designed so that the catalyst is continually_.
Sx 8b"3:S:|: ;
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-------
Table 28. Selectivity of Absorbent Processes
Absorbent
MEA
DEA
MDEA
DIPA
NH_
3
Diamox
Benfield
Selexol
See ref 5.
See ref 6.
°See ref 10.
d
See ref 13.
See ref 9.
See ref 14.
Selectivity Range
0.89
1.2 to 2.27
1.2 to 7.0
3.33 to 5
1.00 to 28.9
5.7 to 9.4
2.25 to 9.8
5 to 22
Reported By
Pearce ;
Pearce
Pearce ;
Naber et al.
Kohl and Riesenfeld '
- ja
Hiraoka et al.
Parrish and Neilson '
! f
Kohl and Riesenfeld
!
i
'
i
|
82
-------
r^i.-iG GUIDE wrn
BEGIN
LAST LINE
OF TEXT I:*
.__ withdrawn from the bottom of the bed and the regenerated catalyst is --
added at the top.1 j \ .
i I . • ' !
'. i
The Haines process uses a bed of synthetic zeolite to adsot-b the H2S.
rl The. b.fd.,is. theri_.^?e.nerated._wit.h high-temperature gas that • contains^'
sulfur dioxide. The sulfur formed is condensed and recovered.
i
I !
The application of these processes is limited due to the possibility
that the beds will be plugged by condensible hydrocarbons and/or other
impurities in the gas. Since these processes have not been commercial-
ly proven, they are not considered further. i
' I :
!
2. Liquid Phase j
Except for the EIC process the principal liquTd^haIe~di"rie"ct^c7nvIrIion'
processes shown on Fig. 9 selectively remove H2S by converting it
directly to elemental sulfur. Carbon dioxide is only slightly absorbed
and largely'remains in the treated gas. With the EIC process, sulfur
is removed as ammonium sulfate and C02 is not removed. ; |
1 I l \
The Stretford, Giammarco-Vetracoke (G-V), and Takahax processes are *
based on similar oxidation-reduction chemistry but use different acti- '
vator chemicals. The G-V process uses an arsenic compound/ but because '
of arsenic's toxicity the process is not applied in this country. The
Takahax process, developed in Japan, has not been widely applied.
The Ferrox process, which is based on the reaction between iron oxide
and H2S, is outdated and has been replaced by the more efficient Stret-
ford Process. In the Stretford process, the direct-conversion process
most often employed, H2S is selectively removed by direct oxidation to
elemental sulfur but CO2 is only 'slightly absorbed. The process is
usually the preferred choice when the H2S concentration in the raw gas
is very low or when a high degree of selectivity is required. Complete
removal of H2S to about 1 ppm can be obtained in the treated gas.
Since the sulfur absorption capacity of the Stretford solution is low,
the process is not competitive with indirect sulfur removal!processes'"
3/8" v ssrsi "m ' ' "
_ J ¥::::r:"-x-2:::-:-:::':-':'::: • !
n ..
I
BOTTOM Or
, ~CR TABLE;
.-AND ILLUS-
| TRATIONS
EHA-237 (Cin.)
(4-76)
PAGE NUMBER
-------
BEGIN.
LAST LINE
OF TEXT i
B.
1.
for volumes of H2S in the feed .gas that exceed about 16% of the total
acid-gas content of the gas.2 : ;
I • ! i
! ! i
Organic sulfur compounds are not removed by the Stretford process but
remain in the treated gas. Stable thiosulfates form in theJoxidizer
' " ——>— -• • - • • • .*. .... . ,,_... . - - — .»- — -- — ^ • —«
equivalent to about 1% of the sulfur in the feed gas. A portion of the
circulating solution must be purged to control the buildup of these
compounds.' The quality of sulfur produced is low compared to that
produced by the indirect-recovery processes. In spite of its: shortcom-
\
lings the Stretford process has been widely applied, and when, the CO2/
H2S ratio is very high, it may be the only process available! that is
1
capable of economically purifying the gas.
i !
i l
I i
The EIC process is an absorption-oxidation process using copper sulfate
as the absorbent that produces by-product ammonium sulfate from the
hydrogen sulfide and ammonia in the treated gas. The process reported-
ly removes H2S and COS and ammonia but does not remove C02.3: A high
removal efficiency of 99% is obtained in one absorption step and the
i . I
system can operate at elevated temperature. The process is in the
development stage and has not been commercially proven. However, since
it appears,to match the requirements needed for desulfurization of
i . i
direct-fired oil-shale gases, it will be considered as a canidate for
further evaluation.
i
!
INDIRECT-CONVERSION PROCESSES
' ' " ' ;
In the indirect conversion processes the acid-gas components! are re-
moved from the fired gas and recovered as a separated stream; that is
subsequently processed for recovery of sulfur. i
Dry Bed ,
The dry-bed indirect-conversion processes provide excellent selectivi-
ty, with most of them capable of removing sulfur compounds with little
or no C02 removal. These processes use a fixed bed of solid, material.
Processes that use molecular sieves or carbon adsorb the sulfur com-
pounds, which are subsequently released^unchanged during regeneration
5C7TCM 0
'.'. -, JE A.-.~
_ _____ i
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1
of the bed. With the iron sponge and zinc oxide processes- H2S reacts
with the bed to form iron or zinc sulfide. The Katasulf process is
based on the catalytic oxidation of H2S to SO2, followed by removal of
S02 with an ammonium sulfite—bisulfite solution.
Dry-bed processes, however, are limited in their application and become
economically impractical for those cases where large amounts of gas
have to be treated or the total quantity of sulfur to be rejmoved is
high. Since full-scale oil-shale facilities will be processing huge
amounts of gas beyond the practical limit for application of dry-bed
processes, these processes are not considered further
1 ! I ; •
Liquid Phase -
,. _ _ _ ;• ". ••-;" J_ . '
With liquid-phase indirect-conversion processes the acid-gas components
are removed from the feed gas by absorption into liquid. T|he liquid is
subsequently regenerated to produce a concentrated stream o'f acid gas,
which is'then processed by the Glaus process for recovery of the sul-
fur. The principal liquid-phase indirect-conversion processes are
shown on Fig. VII-1 and are classified by type. •
I J ' '
Alkanolamines The alkanolamine processes are based on the1 reaction of
a weak base (alkanolamine) and a weak acid (H2S, organic aciLds, and/or
C02) to form a water-soluble salt:
RNH2 + HS
•*• RNH3 • HS
RNH2 '+ C02 + H20
RNH3 • HC03
These reactions are reversible, and the equilibrium may be shifted by
adjustment of the solution temperature. The absorption of CO2 involves
the formation of carbamates as intermediate compounds. With primary
and secondary amines the carbamates are formed nearly instantly and the
reaction rate of CO2 with the amine is nearly equal to that;of H2S with
amine. With tertiary amines carbanates form at a much slower rate and
therefore ^f^^gb^or^ed_mgre_readly_than_is_CO2. This difference ii~
i "• ~] ~" ' " i " ""
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CF TEXT
reaction rates between H2S and C02 accounts for the higher selectivi-
ties that are obtained with tertiary amines. Methyldiethan^lamine
(MDEA) reportedly produces the highest selectivity of the tertiary
amines4—6 (see Table 28 ).
The selectivity at which H2S can be absorbed in the presence of C02
depends on absorption kinetics rather than on equilibrium effects. The
actual selectivity obtained by a process is dependent on several vari-
ables, the most readily controlled being the absorber contact time. In
practice high selectivity is obtained by limiting the number of contact
stages in the absorber and the residence time per stage. Absorption
kinetics are also affected by a variety of other process variables,
including competition between H2S and C02 when they are absorbed simul-
taneously. The degree of H2S removal decreases as the contact time is
reduced in an attempt to increase the selectivity. The extent of
selectivity that can be obtained will be restricted by the parity re-
quired of the treated gas. To reasonably predict these effects for a
specific gas composition and at a specified degree of desulfurization,
the process developers have devised proprietary computer programs for
modeling the absorption process.7'8 These models were developed and
verified through extensive laboratory and field testing programs.
The average selectivity of MDEA as reported by Pearce5 is 3.85. If a
two-stage selective absorption system were used and a selectivity of
3.85 per stage were obtained, the H2S in the acid-gas stream would be
about 12.9 vol % for a feed gas with a CO2/H2S ratio of 100:1 (see
, . I
Table 29 ). If the ratio were 200:1, the acid gas would contain
about 6.9% H2S, which is marginally acceptable. A possible process
option would be a system consisting of two or more stages of selective
absorption, with MDEA used as the solvent. !
Alkaline Salts The principal alkaline salt processes, shown on
.Fig. VII-1, are most successfully applied for bulk C02 removal but are
not generally considered when selective absorption of H2S is required.
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! TRATiONS
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Table 29. Selective Absorption Using MDEA
with Two Stages of Absorption
Feed
©
Treated Gas
t
C02 Vent
R
«-
(3)
A
t
Sulfur
R
fe
(^)
C
RED
i !
t
CO2 Vent
A e ab'sorption; R = regeneration; C = Claus.; RED = reduction.
{ i L i
Streeun
C02/H2S
Ratio3
(vol
Case 1
Q) Feed 1st stage
^y Feed 2nd stage
(3) Claus gas
Case 2
(l) Feed 1st stage
(2^ Feed 2nd stage
@l Claus gas
Case 3
(l) Feed 1st stage
(£• Feed 2nd stage
(5} Claus gas
100
26
6.7
200
51.9
13.5
4
1.04
0.27
3.7
12.9
0.5
1.9
6.9
20.0
49.0
78.7
aBased on obtaining selectivity ratio of 3.85 per stage.
As vol % of total CO2 and
-------
r , .. > {
i.!;'.'- Or
CE.;.;iLn
r,r iv1.-..-.'•
not
If.'AGu
buik
ess using DIK solvent
seect ,
y
The Alkacid proc-!
wlutlon, the
can b. „
"
carbonate
B,.,1 but, as
iVreaTesT ~
-
the M
produced selectiviti
rate of HzS. has
To apply the Benfield process.
the selectivity that i
P-Posed oil-sh,le retorts would be produc'ed^aTZ; ™"^
sure, the pas would have to be compressed near-atmosphere pres-
costs of compressing the oas ' P a and operating
process. kiso/ since „ ^^ * Justif"d by use of the jBenfield
"' Si«ce H2s is less soluble in th*>
at higher temperatures, the advJi™ ^ . .
be
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Table 30. Selective Absorption by Benfield Process
Compressor
Feed
<*
1
t Treated Gas A Sulfur A CO- Vent
1 T
A
R
©
4 i • 4
1 ,
ppn
11
A*
IP
t'7* |
_i__ _ _j
A « absorption; R = regeneration; C = Claus, RED = reduction.
Stream
Case 1
^tti^
ij) Feed
Q) Clau:; gas
Case 2
(l) Feed
^^
^T) Claus gas
Case 3
• ^) Feed
(5) Clau:; gas
CO2/H2S
Ratioa
• *
;* f
100 , ~
17
200
33
4
0.7
i
H?S (vol %)
!•'
; i.o
5.7
i
! 0.5
1
2.9
20.0
60.0
*Based on selectivity ratio of 6.
As vol % of total CO2 and H^S.
89
-------
T.--
BEGIN
LAST LINE
OF TEXT 2
Aqueous Ammonia Processes—Aqueous ammonia processes have Ibeen used -
since the end of the nineteenth century for removal of hydrogen sulfide
and nitrogen compounds, primarily ammonia from coal gas. Many ammonia-!
based processes have been developed over the years; the principal ones
*.-a5e._?_.°wil_0.n.,f^9: j. s _ Ammonia-based processes are par-
ticularly applicable to gases containing both H2S and ammonia. The ~"!
ammonia contained in the gas can;serve as the active agent for removal \
of H2S and can be recovered as a by-product. ': \
i j .'
If the physical parameters of the absorption process are controlled, !
aqueous ammonia solution can selectively remove H2S from gases contain-
ing C02. Selectivities of up to 28.9 have been reported." Hydrogen
. __ _._ 1- _. „ : - I •* " * **c 0 w^U (-^ VII CXXlU t 6 CiC wS
rapidly with the hydroxyl ions~, wher^aTdrbon dioxide'mustTfirst" react'
with water, forming carbonic acid, before it can react ioniklly with '
ammonia. The rate of the C02 hydration reaction is quite low compared '
to the reaction rate of H2S.11
i
The absorption of ammonia into water is quite rapid and is governed
principally by the gas-film resistance. The rate of absorption of H2S
. into aqueous ammonia solutions is also rapid although it is 'dependent
on the ammonia concentration and 'is probably also governed by the gas--
film resistance. Carbon dioxide absorption into water or weak alkaline
solutions, on the other hand, is|governed by its liquid-film resist-
ance, which is very much greaterjthan that for H2S. As a result when
gases containing H2S, ammonia, and CO* are contacted with wJter or
ammonia solution, the ammonia and H2S are absorbed much more! rapidly
than is C02. This difference can be accentuated by operating under
conditions that reduce the gas-film resistance or increase the liquid-
film resistance. Thus to achieve maximum selectivity, spray! columns in
combination with relatively short contact times are used." ! Table 31 |
shows the^results that could be obtained if a selectivity of 28.9 '
were achieved. However, the degree of H2S removal decreases1as the !
selectivity is increased. Removal of about 90% of the H2S is the maxi- '
mum efficiency that can be attained
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Table -31.' Selective Absorption Using Aqueous Ammonia
B2°i
Feed^ A
(l) (H2S)
" i!
!!
| |
+ 1
R —
1
-fr A
(NH3)
1
1
D
1
l_ -
**
J *
„
1 H2
JNH
'
Waste
^ Water
-
fr Treated Gas
B Final Absorber
Batch Loaded
.
A 4 C02 Vent
r * ' r } BEP fr A » R — i
C^)
^k ^ 11
T ' J
D = distiilation, .
n, C » C1aus, BED - „***!«.
2 Claus gas
C3.aus gas
Claus gas
-._-----. __
BBased on selectivity ratio of 28.9.
As vol % of total C0 and
-------
sorption. A final purification step must be added to remove the resid-
ual H2S if it is desired to achieve a high degree of removal with a
high degree of selective absorption. >
BEG'N
LAST LiNE
d.
The selectivity_of _aqueous_ ammonia solutions_decreases_mar_kedly_ at
temperatures above 27°C.12 Also, the selectivity decreases as the
ammonia concentration of the solution is increased. However; the
degree of H2S removal increases as excess available ammonia is in-
creased. The Diamox process, recently developed in Japan, was designed
to take maximum advantage of these factors. Hydrogen sulfide removal
in the range of 97 to 99%, with good selectivity still maintained, can
be achieved by the Diamox process. Selectivities in the order of 5.7
to 9.4 have been obtained based on data reported_by_Hiraoka,_Tanaka,_
andTsudo.13 If an average selectivity of 7.6 were obtained with the
Diamox process (Table 32 ), an acid gas containing 7.1% would be ob-
tained for a feed gas with a CO2/H2S ratio of 100:1, or 3.7% H2S would
be obtained if the ratio were 200:1. The Diamox process would be
marginally acceptable for gases with C02/H2S ratios of up to about
100:1 but would not produce sufficient selectivity for direct-fired
i j '
ire tort gases with higher CO2/H2S ratios.
Physical-Solvent Processes - Selectivity of physical-solvent processes
depends on the relative solubilities of CO2 and H2S in the solvent.
The capacity of the solution for absorbing H2S increases with increas-
ing pressure and decreasing temperature; generally the higher the par-
tial pressure of H2S the higher the selectivity attainable will be.
Selectivity can also be enhanced by favorable absorber kinetics.
The Selexol process appears to produce the highest reported selectivity .
of the physical-solvent processes; H2S is 9 times more soluble in the j
Selexol solvent than C02 is. Selectivity ratios of up to 22 have been
' ' 1*1
obtained based on data reported by Kohl and Riesenfeld.1* '•-
To apply the Selexol process the retort gas would have to be compressed
1_^AO I 1-1 l\ — \ ; t ' _ . . I
OF TrvT t W^ and cooled before it entered the absorber. If^an_Jave_r.a^e_sele.ct_iyity._r.;
\J t I fc_-T I ^^| ,1111 i " • ~ '" "", ~~~ -' " " ».
I |^''^ _J |C.'.:92.'.'''|j;
BOTTOM OF
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DIMENSION
"OR TABLE::?
•AND ILLUS-
TRATIONS
PAGE NUMBER
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(4-76)
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Table 32. Selective Absorption by Diamox Process
Feed
A Treated 1 Water A
I GaS i I
Sulfur
1
CO2 Vent
A
T """ 9
t_T~
1
1
I,
^f
R
_-J
D
C ^ RED
(2)
T
JNH
A -fr R U
: I
1 : 1
i
Waste Water
A « adsorption; D - distillation; R = regeneration; C - Claus; R = reduction.
Stream
Case 1
(l) Feed
(2) Claus ga.s
Case 2
(T) Feed
(?) Claus gas
Case 3
(l) Feed
^ Claus gas
CO2/H2S
Ratio
100
13.2
200
26.3
0.53
,H2S (vol %)
, 1.0
7.1
0.5
3.7
20.0
65.5
aBased on selectivity ratio of 7.6.
bAs vol % of total C02 and H2S.
93-
-------
of 13.5 were obtained as shown in Table 33 the acid-gas stream
produced would range between 11.9 and 6.3% H2S for a feed gas with
CO2/H2S ratios in the range of 100:1 to 200:1. i
9-1/3"
I
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y J/b
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Table
..33.- Selective Absorption by Selexol Process
•Treated Gas
Compressor
t
Sulfur
1
C02
Vent
R
-••
©
i
C
t
k
RED
"" *
*
', 4
— H
' R
d
A -
, F - flash, R - «,—»tl», C - Cl». «ED « ruction.
200
14.8
4
0.30
Sased on selectivity ratio of 13.5.
0.5
6.3
20.0
77.1
bAs vol % of total C02 and
95
-------
or p/••<;?
10!' OF
.;;.-' A cr
C.
1.
REFERENCES* j :
|
A. L. Kohl and F. C, Riesenfeld, Gas Purification, 3d ed., pp 421 and
422, Gulf Publishing Co., Houston, 1979.
S. Vasan, "Holmes-Stretford Process Offers Economic H2S Removal," The
Oil and Gas Journal 76(1). 78—80 (Jan. 2,-1978). r~
4.
S.
BEGIN
LAST LINE
OF TEXT r-
9..
10.
11.
12.
13.
14.
Letter dated Apr. 1, 1980, from W. Dyer, EIC Corporation, to R. LovelJ '
IT Enviroscience.
i
A. L. Kohl and F. C. Riesenfeld, op. cit., p 38.
i .. •
R. L. Pearce, "Hydrogen Sulfide Removal with Miethyl Diethanolamine,"
pp 139—144, in Proceedings of the 57th Annual Convention of the Gas
Processors Association. March 20—22. 1978, New Orleans. i i
J. Naber, J.-Wesselingh, and .W. Groenendaal, "New Shell Process Treats
Claus Off-Gas," Chemical Engineering Progress 69(12), 29—34 (December
; ~"~;" 1 ;~~ ~~~"~ : 1
C. Ouwerkerk, "Design for Selective H2S Absorption," Hydrocarbon Proc-
essing 57(4), 89—94 (April 1978). ; |
' i '•
Telephone conversation Jan. 28, 1980, between C. A. Peterson, IT Envi-
roscience, and R. L. Pearce, Dow Chemical USA. l
I. ! i '
R. W. Parrish and H. B. Neilson, Synthesis Gas Purification Including
Removal of Trace Contaminants by the Benfield Process, presented at the
167th National Meeting of the American Chemical Society, Division of
Industrial and Engineering Chemistry, Los Angeles, California, March 31 ';
to April 5, 1974.
A. L. Kohl and F. C. Riesenfeld, op. cit., p 148.
Ibid., p 134.
Ibid.. p 151.
I
H. Hiraoka, Mitsubishi Chemical industries, Ltd., and E. Tan
-------
VIII. EVALUATION OF CANDIDATE PROCESSES
I A.
BEG:--:
LAST LINE,
OF TEXT P-
I
As determined from the data given in Sect. VII the processes that
produce the highest H2S selectivity or appear to be most applicable to
_lreatrile™ of_ 9ase.s_.frorn .dHect-fired oil-shale ..r.e.to.rts. are _the follow::.
ing: ; : '
Direct-Conversion Processes
Stretford |
EIC [
Indirect-Conversion Processes
Alkanolamine processes MDEA
Alkanline salt processes Benfield
Aqueous ammonia processes Diamox
Physical solvent processes Selexol
i
BASIS OF EVALUATION
I :
To furthur evaluate these processes their capabilities were jcompared,
with a hypothetical feed gas composition used as the basis for compari-
son. The gas composition chosen (Table 34) is similar to the gas
produced by the Paraho direct-fired retort. Trace organic sulfur com-
pounds, not included in Paraho's ^data, were added so that their effects
could be evaluated. The quantity of trace sulfur compounds jshown was
extrapolated from data compiled from sources reporting on trace com-
pounds .1 "i
The indirect sulfur removal processes are only marginally capable of '
producing an acceptable Claus gas when the C02/H2S ratio of the feed ':
gas exceeds about 100:1 (see Tables 29 through 33 ). The hypothe-
tical gas, with a CO2/H2S ratio of 73:1, was chosen so that it would be
I } j ;
within the capabilities of the more marginal processes. To treat this
gas, using an indirect removal process, an H2S selectivity ratio of at :
'. ' '
least 6 would be required to produce a minimally acceptable acid gas
for the Claus unit.
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.1
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Table "34.
Hypothetical Direct-Fired Retort Gas
Physical Properties
Product Gas
Composition
(wet basis)
Amount
(wet vojl %)
Retort off-gas temperature:
66°C
Retort off-gas pressure:
20 psia
224 sm3/tonne
Gas density: 1.2 kg/in3
Gas rate:
.•of shale
Molecular weight: 28.43
II
2
°2
N^ -3- Ar
2
CO
co2
CH4
CJH,
2 4
C-H,,
26
CH,
3 6
C,H_
3 8
C S (MW = 57.0)
C^ (MW = 71.5)
Cf + (MW = 96.2)
6
H S
2
NH,
3
H2°
3.539
0.743
51.342
1.600
17.572
1.858
0.743
0.768
0.364
0.380
0.323
0.129
0.622
0.242
i
0.566
,
19.210
lOO.OIDO3
Includes the following trace compounds: COS, 36 ppm; CS_!,
7 ppm; RSH, 10 ppm; and SO_, nil. |-
98
-------
Uf l£ Or
"J • p\ l"
i- = .:•' ' '-
BEGIN
LAST LINE
OF TEXT £-?
B.
Projected plans for full-scale commercial oil-shale plants based on —-•
direct-fired retort technology range from Geokinetics1 facility pro-
ducing 2.83 million m3 of gas per day from 16 retorts processing about
5454 tonnes of shale per day2 to Occidental's MIS plant producing
48 million m3 of gas per day from a series_qf_4p_retorts_prqcessing
about 149,078 tonnes of shale per day.3 : ;
The model plant selected for evaluation of a sulfur removal process has
a capacity of 45,450 tonnes of shale per day and produces 10.2- million m3
of gas a day based on the composition of the hypothetical gas. A de-
sulfurization plant in this size range will require multiple . trains to
process the volume of gas produced, as would be the case for most
projected commercial oil-shale facilities. The individual components
required for the model-plant sulfur removal train should therefore be
in the size range of that required for a full-scale commercial oil-
shale facility and should be representative of such a facility.
9-1/8"
i
GAS PRETREATMENT
1 !
Host sulfur removal processes will require some form of pretreatment of
the gas before it.enters the sulfur removal equipment. For most proc-
esses the gas must be cooled and ammonia, condensible hydrocarbons, and
excess water be removed. Other processes may require compression of
the gas. ^articulate matter is largely removed in the oil separation
! ; ' '
and gas cooling steps. Additional particulate removal is not required
I
before the gas enters desulfurization equipment. ;
1.
Gas Cooling and Ammonia Removal
The temperature of raw gases produced by direct-fired oil-shale retorts !
range from 5/7° to 77°C. For most sulfur removal processes the gas
I !
must be cooled before it enters the absorber. Since the gases from
direct-fired retorts are saturated with water, large amounts; of water
will be condensed out as the gas is cooled. Calculations indicate that
sufficient water can be condensed to simultaneously absorb and remove
the ammonia contained in the gas if the gas is cooled to about 32°C.
• I —• -|
I
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The gas cooling unit for the model plant is shown in Fig. VIII-1. The
raw fuel gas (stream 1), which is at 60PC, is directly contacted with
the cooling water in a packed tower to bring the temperature down to - .
32°C. AS the gas cools, a large portion of the water vapor, 849 kg/ ^
min, and a small amount of hydrocarbons are condensed into the__circu-—..
lating cooling water. Essentially all the ammonia, a stoichiometric
amount of CO2, and a small amount of H2S are absorbed into the cooling
water. The cooled gas (stream 3) is sent to the desulfurization system
1 {
absorber. j I '
BEGIN
LAST LIKE
OF TEXT Z.-"
2.
A portion of the cooling water solution equilivent to the quantity of
gases and vapors absorbed is purged (stream 2) from the system. The
remaining cooling water solution, saturated with^COg,_is_^assed_ through
a heat exchanger to remove the heat absorbed and then recirculated. .
The operating conditions and the composition of the recirculating cool-
ing water are controlled so that minimum H2S is absorbed into the cool-
ing water. :
The purge "stream will be treated (not shown) by being passed, through an
oil separator, where condensed hydrocarbons are recovered, and then
through a sour-water stripper and ammonia recovery unit. The H2S re-
covered in the sour-water stripper is returned to the process gas
stream. The ammonia is stripped and recovered as a by-product. The
remaining water, which contains trace amounts of absorbed organics,
H2S, and ammonia, is sent to a conventional wastewater treatment
!
system. >
I
Condensible Hydrocarbons
The heavier hydrocarbons with boiling points above 90°F will condense
and be largely removed in the gas cooling unit. However, components in
the gas include unsaturated hydrocarbons such as acetylene, >thylene,
propylene, butadiene, and butylene. These components will not condense
as the gas is cooled. They can, however, polymerize and form tarry
substances that can foul the acid-gas absorber, discolor the sulfur, or_
ej^er^unjj^:y[jgQ^
clog the cataly£t_beds of a Glaus su!
1 i " I................
ft 3/8"
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PAGE MJV.CER
Ei'A-2K7 !
(4-76)
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90° F
GAS
COOLER
CW.
COOLED GAS TO
DESULFURIZATION
SYSTEM
"SOUR" WATER
TO AMMONIA
RECOVERY
Figure 10. Gas Cooler/Anunonia Absorption System
101
-------
BEGIN
LAST LINE
OF TEXT :•
I
I
tion takes place in the acid-gas absorber. There is no economical way
' i
i of removing these compounds before the gas enters the acid-gas ab-
| sorber, and their effects on the desulfurization train can be deter-
1 mined only by testing of the process. ,
| 3. Gas Compression j
j Both the Benfield and the Selexol processes require elevated;pressure
to operate and thus would require gas compression. For the model plant
producing 10.2 million sm3d of gas (Table 34 ) a compressor rated at
! 84,500 hp would be required to boast this pressure to 150 psig. A gas-
turbine-driven compressor system of this capacity with intercoolers and
I condensate separation is estimated to cost $6 to $8 million and would
consume about $34,500 of fuel per day, or about 32% of the treated gas
produced by the facility. ] I
If the treated gas were combusted in a gas turbine used for power
generation, a portion of the available energy in the compressed gas
would be recovered as the gas expanded through the turbine. iThus if
the gas is compressed, it would be advantageous to remove as|little C02
as possible in the desulfurization process so that more gas will be
sivailable to expand through the turbine. However, after compression
the gas would have to be cooled to the operating temperature i of the
sulfur removal process, and so a large portion of the available energy
; '
in the form of heat would be removed before the compressed gas reached
the power generation turbine. From an energy standpoint it would be
more efficient to compress the gas as it enters the power generation
turbine rather than before sulfur removal. j
Since the Benfield process can operate at up to 121°C, less heat would
have to be removed from the compressed gas than with the Selexol or
cither processes that operate below 32°C. Approximately 4673.6x109
i : '
joules of heat per day would have to be extracted to cool the compressed
t ;
gas to "12P.C, at a cost of about $1750/day for cooling water! To cool
| i
the compressed gas to 32°C approximately 10,llSxl69 joules/day must be
i ! ' —
extracted at a cost of $3850/day. _^ ;
I
3/8"
_ 1
EPA-287 (Cin.)
Sx;.;-:.: 102 : :-:'A
PAGE KUV.liER
3CTIO.V. OF
".'<-'• GE ARE
7LTCICE
Ci?.'tNSiO\
FOR TABLE;'.
•A;-*D ILLUS-
TRATIONS
-------
Compressing the gas to 150 psig would reduce its volume by a\factor of
6 to 10, depending on the exit temperature. Thus the number or size of
absorbers required to handle the gas would be reduced porportionately
for those processes in which the absorber size is gas limiting i
(Selexol, Benfield, and to some extent the alkanolamine processes)-_.For
those processes in which the absorber size is liquid limiting (Stret- j
ford, Diamox, and to some extent MDEA) the size of the absorbers may :
' ! i
not be significantly different if the gas were compressed. | !
I ; j
Due to the high capital and operating costs of gas compression it is :
not economical to compress the gas for the purpose of sulfuri removal
alone. Some full-scale commercial oil-shale plants employing indirect-
heated retorts may compress the gas before sulfur^ removal_for _the_pur- .;
pose of recovering LPG and other condensible hydrocarbons or for other
process reasons. In that case the application of the Benfield or
Selexol process may be attractive. i
C. DIRECT-RECOVERY PROCESSES j
i I i
1. The Stretford Process (See Appendix A) j
. The Stretford process was developed in England jointly by the North
Western Gas Board and the Clayton Aniline Company, Ltd.4 The process
i '• '•
has been proven commercially over a period of many years in ;more than
80 plants built around the world.5 Several modifications arid improve-
ments to the basic process have been made by various companies that now
'. '• I
license their proprietary versions of the Stretford process. The
principal licensors in this country are Peabody Process Systems of
Stamford, Connecticut,, the Pritchard Corporation of Kansas City,
Missouri, and the Ralph M. Parsons Company of Pasadena, California.
The chemistry of the Stretford process is relatively complex. The
idealized reactions illustrate the absorption, oxidation/reduction
cycle as follows:6
LAST LINE
OF TEXT r
BOTTCV. Or
CUTSlDE
Di.VEN'SlG'
FOR TABLE.
-AND iLLli
TF-.ATiONS
PAGE fvUY.L.ER
EHA-ib? (Gin.)
{4-76)
-------
— Reaction 1
Na2CO3 + H2S
NaHS + NaHC03
Reaction 2
4NaV03 + 2NaHS + H2O -
i
Reaction 3
Na2V4O9 + 2NaOH + H2O
i
Reaction 4
— — 2ADA (reduced) + O2 —
Na2V4O9 + 2S + 4NaOH
2 ADA
4NaVO3 + 2ADA (reduced)
2ADA + H2O
A flow diagram of the Stretford process for fuel gas desulftirization :Ls
shown in Fig. 11. The sour fuel gas (stream 1) enters the gas
cooler, where it is cooled and a large portion of the water land most of
the ammonia are removed (stream 2) and sent to foul water treatment. ;
The cooled gas (stream 3) enters the Stretford absorber at about 32°C.
. . i
The feed gas (stream 3) is countercurrently washed with an aqueous ;
solution containing sodium carbonate, sodium metavanadate, and anthro-
quinone disulfonic acid (ADA). The H2S rapidly dissolves and ionizes
in the alkaline solution, forming a small amount of sodium hydrosulfide
:
(reaction 1). Hydrosulfide loadings in the solution range from 500 to
1000 ppm; thus relative large amounts of solution must be circulated.
The sodium metavanadate in the solution readily reacts with the hydro-
sulfide to produce elemental sulfur (reaction 2). i •
Sodium carbonate in the solution provides a pH buffer to prevent rapid j
pH changes as the acid gases are absorbed. The alkalinity of the i
Stretford solution causes some of the C02 to be absorbed albng with the
i
H2S. The treated gas (stream 4) with essentially all the H2S removed
exits from the top of the absorber. The aqueous solution enters the ;
delay tank, with sufficient residence timeprovidedto assure that any
r."-V2'// (Cm.)
-------
§
-f«
.4-)
.<0
N
•rt
hi
m
• to
0)
3
fa
n
co
': 0>
O
o
V4
.cu
•o
ij
t 0)
4J
"
L-
i-i
• o>
;fe
ii/. 105 :---
TYPI;
/*"* I ' * J"1"- '-" *- --* ^
vz- \_;. I-' i- o i * v
-------
remaining hydrosulfide is reacted so as to prevent hydrosulfide from '
entering the oxidizer and becoming oxidized to stable sodium thiosul-
fate, an undesirable by-product. Thiosulfate can be formed!by oxygen
dissolved in the Stretford solution, oxygen in the feed gas; or hydro-
sulfide carryover to the oxidizer.. An excess..of vanadate.is maintained
to avoid overloading the solution with hydrosulfide. Formation of
thiosulfate can be controlled somewhat by controlling the pH and the
i i
temperature of the solution. Thiosulfate can be allowed to concentrate
in the solution to about 20% by weight. A small portion of;the solu-
tion (stream 14) is then continually purged and replaced to jprevent
further accumulation and crystallization of salts. ;
! :
i ' !
s
The reduced solution (stream 6) is regenerated in the oxidizer tower_ by
being sparged with an excess of air. The reduced vanadium is reoxi-
dized to its original state through oxygen transfer via the;ADA (reac-
tion 3). The reduced ADA is reoxidized by contact with air'(reac- ;
tion 4). Reoxidation of the ADA can be appreciably accelerated by the
i
presence of small amounts of iron salts kept in solution by a chelating
agent, ethylenediamine tetracetic acid (EDTA).7 ! i
The nitrogen and the excess oxygen from the air bubble upward through
the solution, floating sulfur particles to the top as a froth and
stripping the dissolved gases. The froth, which contains about 10%
solids, overflows the oxidizer to a slurry tank. A damp cake contain-
ing 50 to 60% solids is produced by filtration or centrifugation of the
slurry (stream 10). The sulfur cake can be further processed by wash-
ing to remove the Stretford chemicals and by drying and melting to
produce liquid or solid sulfur (stream 17). \
The regenerated solution (stream 5) is relatively free of sulfur and is
pumped back to the absorber. In route the solution must normally be
cooled to maintain the desired operating temperature, which'is usually
accomplished in an evaporative cooling tower. To maintain water
balance in the system due to the addition of sulfur and/or filter wash
..water and water produced by the reaction, water is evaporated with, the .
cri'A 2:17 iC.-.r. 1
-------
treated gas or in the cooling tower. A solution heater is sometimes
required for winter operation or to assure sufficient evaporation.
i •! .
Carbon dioxide is partly absorbed by the alkaline solution, resulting
in the formation of sodium bicarbonate and consequent lowering of pH.
The solution eventually reaches equilibrium with respect to jthe concen-
tration of C02 in the gas stream, after which only relatively small
i
amounts of CO2 are absorbed. Since the vanadate-ADA system functions
at a lower pH, decarbonation of the solution is not required. When the
gas contains high concentrations of CO2, the absorption efficiency of
the solution may be lowered to the extent that an appreciable increase
in absorber height will probably be required. Peabody has developed a
proprietary venturi-absorber that is insensitive to the H2S/CO2 ratio
. T . .. . • • .... ,
and is capable of handling fuel gas containing 85% C02 with!minimum
absorption of CO2.8
., Ac,"
The principal side reactions that are detrimental to the process are
the formation of thiosulfates and thiocyanates. If hydrosulfate is •
contacted with oxygen before it is converted to sulfur, thiosulfate :
will form, the amount depending on the pH of the solution and on the .
operating temperature. The rate of thiosulfate formation under favor- :
able conditions can be held to less than 1 wt % of the sulfur in the '
feed gas by controlling the pH and temperature of the solution and by ,
assuring that sufficient delay time has occurred for the conversion of
H2S to sulfur before the solution reaches the oxidizer tower. However,
gases from direct-fired retorts contain oxygen and a large amount of :
COo. Calculations show that for those gases about 3 to 4% pf sulfur in
- i i '
the gas will go to thiosulfate as the result of oxygen in the feed gas
and the buffering action of the CO2. Fortunately, oil-shale gases con-
tain only trace amounts of hydrogen cyanide (HCN). Hydrogen cyanide
will react with elemental sulfur to form stable thiocyanide, which also
accumulates in the solution.
BOTTO"! C
;"AGE AF:
For small'plants producing less than about 9 tonnes of sulfur per day, j ,;" l""'..-^-
_the_wet sulfur cake is. often disposed of by approved methods in a land-
i ,
£'.?'.
{•!-
(Cin.)
PAGE iM'Mi^H
-------
jfill. Since the wet cake contains sufficient thiosulfate, additional ~
purge of the solution is not required to control thiosulfate buildup.
For plants producing more than 9 tonnes per day it may be more economi-
cal to recover the sulfur, in which case a small stream of solution
must be continually purged. • ;. !
As thiosulfate is purged from the system it carries with it !a propor-
tionate amount of Stretford chemicals, which must be replaced
(stream 18). It is important therefore to permit thiosulfatje to con-
centrate in the solution to as high a level as the process will allow
(20 to 22%) to minimize the costs of replacement chemicals and waste
1 . I
disposal.
The purge stream containing toxic vanadium salts with an average COD
(chemical oxygen demand) of 20,000 mg/liter can be difficult to treat
and dispose of. Treatment methods used include evaporation lor spray
,'.-,' • i
drying, biological degradation, oxidative combustion, and reductive
incineration. Peabody Holms has successfully developed a zero-dis-
charge reductive incineration process that cracks the bleed liquor into
a gas stream containing H2S and CO2, along with a liquid stream con-
taining reduced vanadium salts that can be returned to the system.8
The problem of disposing of wastewater containing vanadium is elimi-
nated and no makeup of vanadium or sodium salts is required^ However,
makeup water to the system must be demineralized and the feed gas to
the Stretford absorber must be free of soluble minerals to prevent in-
organic solids from building up in the system.
The Stretford process is insensitive to pressure and can operate in the
range of atmospheric to maximum pipeline pressure. Operating tempera-
tures range from ambient to 49°C . Any type of gas-liquid contacting
device may be used as an absorber. Problems of plugging have been en-
countered in packed towers, especially with gases containing high con-
centrations of H2S (above 1%).9 .Therefore large rings or saddles are
; • l
recommended when packed towers are used. Venturi scrubbers have been
ft 3/tT
£T'i,-?o7 (Cin.)
(4-7G)
PAG;:
-------
!
BEGIN
LAST LINE
OF TEXT Vs-.
2.
successfully used for gases containing larger amounts of H2S, up to
15%. 10 |
i
Stretford plants are remarkably free of corrosion tendencies ; and can be
constructed entirely of .carbon steel, witK inert linings, e.g., cold- ._
cured epoxy resins, used for oxidizers and sulfur froth tanks. Stain-
less steel linings are recommended for solution and sulfur slurry
pumps . * * -
' j1 •
The Stretford process is extremely flexible in operation and can ;
tolerate wide variations in both gas feed rate and H2S concentration in
the feed gas. Startup and shutdown are relatively simple and can be •
accomplished in short periods_of time.__Good process control is_ main- .'
tained by simple analytical testing with little technical supervision. ;
t * ; I
Since all process streams are handled as liquids,, the process is easily
automated and requires little operator attention.5
9-1/8" ' |
i • :
Hydrogen sulfide removal to less than 10 ppmv can be obtained with
normal operation of a Stretford plant. Assuming that the gas contains
COS in the range of 10 to 50 ppmv the overall sulfur removable achiev-
able by the Stretford process would be in the range of 98.0 ;to 99.3%.
I
The EIC Process
The EIC process was developed by EIC Corporation of Newton, 'Massachu-
setts, for removal of H2S from geothermal steam. The process has been
t | i
field tested in a pilot plant processing 45,454 kg of steam per hour
! : i
at The Geysers in California.12 Although the process has not yet been
pilot tested on desulfurization of fuel gases, EIC is promoting the
process for that purpose.
I
Geothermal steam contains noncondensable gases such as carbon dioxide,
hydrogen sulfide, ammonia, methane, nitrogen, hydrogen, traces of
t » •• - - —
radon, argon, and mercury vapor, as well as rock dust, boron compounds, •••:.;; AF
and other particulates.12 In The Geysers test desulfurization to about '...-. '„-"-
•i. I ' """ ' -'•••••—- -
ll,ppm H2S in.the exhaust steam was achieved.
_
13
i L :
A
Sft:...:. 109.......v
PAGE NUV.BER
PlATiCr-IS
F.='A-237 (Gin.)
-------
HaS Removal:
^2 Removal:
H2S
H
CuS + 20Z
\
cuso4
^JJesult
«2S + 2NHS + 20
-------
m
to
o>
u
O
Wl
m
3
CO
Wl
I
•O
•rH
.a
! O
'•H
U
• 3
•M
w
0)
3
tn
^- X -• *\ » <• j*"' /"* *
t f i-.v'o t_-
-------
Oh '- M •'.-•
-.1
j ___ _ 'with
primarily ammonium sulfate, can be recovered by removing the water in a
vacuum crystallizer.13 ' •
Extensive use of titanium is required in the construction of the system
because of the corrosiveness of the sulfur'ic acid—copper,sulfate solu-
tion.14 The size of the components, however, is much smaller than that
of components for competing processes as the result of irreversible
reactions occurring in the absorber and of the reactiveness land concen-
.
tration of the solution. The smaller size of the equipment joffsets a
large part of the cost involved in the use of special materials of con-
struction. EIC believes that the capital and operating costs of the
system, including ammonium sulfate recovery equipment, are competitive
with those of the Stretford process.12 ; _.
The process appears to be well suited to desulfurization of oil-shale
gas since it removes both H2S and COS, as well as ammonia, and would
not require that the feed gas be cooled. However, there are many un-
answered questions, and more development work needs to be done before
the process is applied to oil-shale gas. A potential major ;problem is
that oil-shale gas contains various unsaturated hydrocarbons, which
could react with the strong sulfuric acid—copper sulfate solution and
1 ; :
form undesirable by-products. Of particular concern is the:large
amount of acetylene in the gas. :If the highly unstable compound
I • :
cuprous acetylide were to form and accumulate in the system, it could
1 1 !
become very hazardous. Cuprous acetylide is explosive and can be
i i
detonated by percussion or can explode when heated above 100°C. If
! i !
warmed in ;air or oxygen for several hours, it will explode when brought
in contact with acetylene.15 The potential for formation of cuprous
s i .. - i
acetylide in the EIC copper sulfate solution has not, to our knowledge,
been defined.
Another potential problem could be entrainment of the solution in .the
i i ; > BOTTOM
treated. gas. Carryover could cause corrosion of the downstream piping ' v.j-A
i i i ! .- .T-O r-.p
and gas turbine blades. The process has not been sufficiently proved :,..Vr-~
.to.recommend its use. at this time. The process could become a viable
' i
* Si?'' 112'."; i
3/8"
T::AT,C\S
(Cin.)
PAGE NU.V.3ER
-------
candidate for desulfurization of oil-shale gases if through testing and
development the adverse effects of corrosion and undesirable side
reactions are proved not to create a problem.
i
i 1
INDIRECT-RECOVERY PROCESSES '.
Table 35 gives the approximate volume and H2S concentration of the
acid-gas stream that the model plant would produce for each of the
indirect-recovery-process options. The selectivity data used for this
table are either the average selectivity data reported in the litera-
ture16—18 for the process or, as in the case for the MDEA19 and the
Diamox20 processes, are the actual selectivity calculated for the proc-
ess based on the gas composition assumed for the model plant. For the
purpose of this comparison it was assumed that all H2S and a portion of
the CO2 relative to the average selectivity obtained by the process
would be removed and end up in the acid-gas stream to the Claus unit.
' i ! . '
Based on the requirements of a minimum acceptable volume of .about 8% of
H2S in the acid-gas feed to the Claus unit and of the need for at least
25% H2S for effective operation, the table indicates that a[one-stage
selective absorption process using DEA or MDEA would not produce a suf-
ficient level of H2S in the acid-gas stream for operation of the Claus
sulfur recovery process. Two stages of selective absorption using MDEA
as the absorbent would be marginal, as would one stage of selective ab-
sorption using either the Diamox, the Benfield, or the Selexol process.
Three stages of selective absorption using MDEA, however, would produce .
an acid gas rich enough in H2S for economic operation of the Claus >
i . i • i
unit. !
It can be'noted that the volume of acid gas produced decreases as the
selectivity of the process increases. Thus when high selectivity is ;
achieved, the volume of gas that must be processed by the succeeding
absorption stage or by the Claus unit is considerably reduced, thereby _
reducing the equipment size and its installation and operating costs.
£!'A-2G7 (Cm.)
(4-76)
PAGiL NV-.'.BES
-------
Table 35. Comparison of Selective Absorption Jrocesses for Treating
Gas from a Direct-Fired Retort ......
Ac*jri— f?as Stream to Claus Unit
A V* e= n r'V* ^ n t
DBA
MDEA
MDEA
MDEA
Diamox
Benfield
Selexol
Number
of
Stages Selectivity
• . O^
1 2.4C
2 8.2°
3 27. 4C
1 4. 7*^
1 6e
1 9f
CO2
895,129
.745,812
219,466
65,316
379,531
299,782
198,921
H2S
(sm3d)
24,649
24 ,-649
24,649
24,649
24,649
24,649
24,649
. Total
919,778
770,467
244,172
89,994
404,181
323,016
280,170
Total
; 640
535
; 170
i 62
: 280
i 224
: -156
H2S (dry
vol %)
2.7
3.2
10.1
27.4
6.1
7.6
11.0
C°Gas rate"=45 450 tonnes/day X 224 sm3/tonne
1,782,900 sm^d; H2S = 0.247 vol % = 24,649
bSee ref 16.
CSee ref 19.
dSee ref 20.
eSee ref 171
fSee ref 18.
= 10,188,000 sm3d;C02 = 17^.572 vol % =
114
-------
The Benfield and Selexol processes, while only marginally capable of
producing an acid gas of acceptable concentration, require high pres-
I
sure to operate. The high cost of gas compression caused these proc-
esses to be categorically eliminated from further considerations (see
discussion in Sect. VIII-B3). : :
The indirect-process candidates therefore were narrowed to a three-
stage alkanolamine (MEDA) process and the Diamox process. A third
candidate that was evaluated is a one-stage selective system coupled
with a Stretford unit for recovery of sulfur from the acid gas. A
discussion of these systems follows.
1. Three-Stage Selective Absorption Using_MDEA as the_Absprbent (See .
Appendix B). j ; /
Two principal commercial processes are available based on the use of
amine solutions for selective absorption of H2S in the presence of C02:
the Selectamine process developed by Dow Chemical USA at Freeport,
Texas, and the Adip process developed by Shell International Petroleum
Company, The Hague, The Netherlands. The Adip process uses'the second-
ary amine diisopropanolaraine (DIPA) or the tertiary amine methyldietha-
nolamine (MDEA). The Selectamine process uses MDEA as the absorbent.
! : I
The basic flow steps for all alkanolamine acid-gas absorptipn systems
' i
are similar. As shown in Fig. 'tl3, the gas to be purified
(stream 3) enters the bottom of the absorber and passes upward counter-
current to the aqueous MDEA solution. The lean solution reacts with
and chemically absorbs the H2S and a portion of the CO2 as it contacts
the gas in the absorber. The purified gas (stream 4) exits1 at the top
of the absorber and the rich solution containing the absorbed acid gas
.(stream 6) is drawn off at the bottom. To achieve the highest practi-
i t i
cal selectivity, absorption kinetics are carefully controlled through
absorber design and operation (see discussion in Sect. VII-B2a). The
selectivities achieved and the approximate volumes of gas handled by
j p
each stage are given in Table 35. :
SOTTO.V. Oi-
•MAGE ASE.
OUTSIDE
DSf.'.nJSIO::
FOR TABLED
AND ILLUs
1 RATIONS
EPA-237 (Gin.)
•,-•,-76)
PAGE
-------
CO
s»>
CO
o
•1-1
o
01
0)
r-l
0)
CO
W
O
0)
CO
<0
4-1
CO
o
:X
t
1
0)
•1-1
PAuE r:U?.'^E:i
TYP'.MG GU'Dc xS
-------
YVr l:«:o y.!.'JC C-i -...: i
The sulfur removal effectiveness usually decreases as measures are
taken to increase the selectivity. For the process describe^ the H2S
content of the treated gas can be reduced to about 10 ppmv and the COS
will be reduced by about 60%. The total sulfur reduction of the
treated gas, assumed to contain. 10 to 50 ppmv COS, .would be .in the —
range of 99.2 to 99.5%. !
0- T-:XT i?-
The rich solution (stream 6) is heated by heat exchange with the hot
lean solution from the bottom (stream 5) of the stripping column and
then enters the stripping column near the top. The lean solution par-
tially cooled in the heat exchanger is further cooled by exchange with
water and then fed into the top of the absorber.
__ ___ !' e .j .••)*• __
In the stripper the absorption reactions are reversed as the tempera-
i . (
ture of the solution is increased. The desorbed acid gas exiting at
the top of the stripping column is cooled to condense out the water
'-"».- •;-•• " '
vapor as the acid gas (stream 11) continues on to the next absorption
I . !
stage. The condensed water is fed back at the top of the stripper
i !
above the point of the rich-solution entry. This serves to.absorb the
amine vapors carried out by the acid-gas stream.
The reaction between MDEA and organic sulfur compounds is not readily
i \
reversed. A side stream (7) of lean solution is sent to a reclaimer,
where a portion of the degraded solution is recovered by high-tempera- .
i i |
ture distillation. The bottoms (stream 10) accumulated from the i
! i 1
reclaimer are sent to waste disposal. i
The acid gas produced by the first-stage system ( 535 sm^/m ; becomes
the feed for the next absorption stage. Compared to the gas entering
i • •
the first-stage absorber (see Table -36 ) the quantity of gas that
1 i
must be treated by the second-stage absorber has been reduced by 92%
! !
and the CO2/H2S ratio reduced to 30.8.
The second-stage absorber operates nearly identically to the first-
.stage absorber The treated gas'(stream.12), however,~is high-purity
3/8" tf :•:<•*•••••••••••••••••+: 1
Q 3/s" M
BOTTOM O.r
If.' AGE ARE
OUTSIDE
"- TABU:
•••:.\D ILLUS
£PA-237 (Cin.)
(4-73)
PAGE
-------
-'-' — -*'• .
Table 36. Performance of Three-Stage Selective Absorption
System Using MDEA i
Stage
1
2
3
Feed Gas
Volume
(sm^m)
7,075
' 535
170
Acid-Gas Stream 1
Volume Volume Reduction H?S
r "* \ *•
(snpnu (%) (%)
535
. 170
. 62.
92.4
97.6
99.1
3.2
10.1
27.4
C02
96.8
89.9
72.6
CQ2/H2S
1 b
Ratio
30.8
8.9
2.6
% Reduction
57.6
87.7
96.4
\folume reduction related to volume of feed gas to first-stage .absorber..
CO2/H2S ratio of feed gas to first-stage absorber is 72.6.
c
CX)2/H2S ratio reduction related to feed gas entering first-stage absorber.
'118
-------
G'J'DE 3i I[!
— CO2, which contains about 10 ppmv H2S and is often used as a by-product
when a local market exists for it.
' ' ' ' '
The acid gas (170 snHm ) (stream 17) produced by the second-stage sys-
tem contains about 89.9% CQ^ and 10.1% H2 and has a CO2:H2S ;ratio.of
.about 8.9.
The third-stage absorber also produces high-purity CO2 (stream 18).
The acid gas (62 sm^m ) (stream 23), containing about 73% CO2 and 27%
H2S with a CO2:H2S ratio of about 2.6, is rich enough in H2S content to
be processed by a modified Claus system for sulfur recovery.
1 ;
I ! i
In the operation of. the Claus unit (describedjLn Sect_._IV-B) the_sulfur
burning— by-pass process is used since it is the least costly option
i •
and since the acid gas produced by the three-stage selective absorption
process is free of hydrocarbon contamination (see Fig. 14). '
9-1/3"
The SCOT tail-gas treatment was chosen because it can be integrated
with the selective absorption system to save total installation costs-
Since both systems use MDEA as the absorbent, the regeneration of the
rich solution coming from the SCOT unit can be combined with regenera-
tion of the absorbent from the primary absorber.
The SCOT process consists essentially of two parts: a reduction stage
in which all sulfur compounds and elemental sulfur in the off-gas are
reduced to H2S and an absorption stage in which, after water is removed
by condensation, H2S is selectively removed by MDEA absorption/regener-
ation and is recycled to the Claus unit. The total sulfur recovery of
i '
the Claus 'unit can be increased to above 99.8% of the sulfur in the
acid-gas feed. The SCOT unit is flexible, having a wide operating
range, and no secondary waste streams are produced.21
The estimated overall sulfur reduction calculated for the total desul-
I ! : ! ' 7--1-:
furization train is about 98.8 to 99.3%, depending on the quantity of . ~, /. ~~
organic sulfur in the gas. j : ! i r C". T--2
D 3/8" »*
J j
EPA-2S37 iCin.)
PAGE r:i//.s=R
r
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r.L.'.v.r:
I AST LI WE!
One-Stage Selective Absorption Followed by Treatment of Acid Gas by
Stretford Process (see Appendix C) •
As can be seen in Table 35 one stage of selective absorption using.
MDEA would produce an acid-gas stream containing approximately 3.2%
H2S._This gas.is.too lean to be effectively processed with a Claus
system. However, sulfur can be recovered from gases containing 3.2%
H2S by the Stretford process. Thus a possible option would be to use
one stage of selective absorption to remove the sulfur compounds from
the gas followed by a Stretford unit to recover the sulfur from the ;
acid gas. j
I 1 ^ '
Such a system (see Fig. -, 15 ) would have several advantages. The
acid gas (stream 11) produced by the^ selective absorption step would be
relatively clean, that is, free of hydrocarbons. Thus a much higher
quality sulfur would be produced since many problems inherent with the
Stretford process relative to impurities in the feed gas would be
.-..-•• !
alleviated. The MDEA absorber would remove most of the organic sulfur
compounds, COS, CS2 and mercaptans, but they would pass through the
Stretford absorber and end up in the C02 stream (16) discharged to the
atmosphere.
The volume of acid gas (stream 11) fed to the Stretford unit would be
i
reduced to about 535 sm^m or about 7.6% of the original! volume of
the raw gas (stream 3). Howeverj since stream 11 contains essentially
the same amount of H2S as stream 3, the amount of Stretford solution
circulated; therefore except for the absorber the size of the Stretford
t ',
equipment required would be the same as that for the Stretford direct-
absorption case. Thus there is no cost advantage in this system over a
•
Stretford direct system.
I
The estimated overall sulfur reduction calculated for the total desul-
furization train is in the range of 98.0 to 99.3%, depending on the
quantity of organic sulfur in the gas.
3 V
.121
EPA-237 iCm.)
-------
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: AST LIME:
OF TEXT ir«-
The Diamox Process (See Appendix D) [, ->
The Diamox process is a commercially proven process developed jointly
by Mitsubishi Chemical Industries (MCI) and Mitsubishi Kakoko Kaisha
(MKK), both of Japan, for desul'furizing coke-oven gas. Five facilities
_employing the process have been built, and their performance has been
excellent. MCI reports20 that the Diamox plants have achieved coke- j
oven gas desulfurization to levels as low as 13.7 to 18.3 gratis of H2S per
100 sm3 (about 95 to 126 ppmv). The Ralph M. Parsons Company of '
Pasadena, California, is the licensee for the process in the United
States. • s : •
i '
i- i .
The Diamox process is particularly applicable for removal of H2S from
•gases Containing ammonia since ammonia is used as the absorbent in the
Diamox solution..- AMP oia ^.S'sent in the raw gas is absorbed to gener-
ate the Diamox solution, and consequently no chemicals are required for
the process. Once the Diamox solution is saturated with ammonia, addi-
tional ammonia will not be absorbed but instead will pass through with
the treated gas. Additional ammonia scrubbers must be added if ammonia
has to be removed from the treated gas.
i i i
A large part of the proprietary design of the Diamox process is em-
bodied in the design and operation of the absorber and stripping towers
and the temperature and solution concentrations maintained throughout
the system.22 The process was developed to selectively absorb H2S from
coke-oven gas with greatly improved removal effectiveness. Up to 98%
H2S removal has been obtained in commercial applications of the process
compared to removals of only 90% achieved by conventional processes
using ammonical solutions. Selectivity ratios of up to 9.4 have been
obtained by the process based on data reported by Hiraoka, Tanaka, and \
Sudo.23 Organic sulfur compounds, however, are not appreciably
removed. ]
I
High-quality, 99.9% pure, bright-yellow sulfur can be recovered from
the acid gas "produced by the Diamox process since the concentration of ! '"'"'y?.'".
1 i i —s L ..'.;^,',-;:(
BOTTOM
/-
w
(4-76!
-------
ammonia, hydrocarbons, and other contaminants in the acid gas is ex- —
tremely low. | ;
A simplified flow diagram of the process is shown in Fig. 16. The
Diamox process operates as follows. The. incoming raw gas (stream .1) is
cooled to about .52°C by direct contact with circulating water in a
i
wash tower. Water, with a small amount of absorbed ammonia; and a
small quantity of hydrocarbons condensed from the gas are purged from
the system (stream 2) and sent to foul-water treatment. Cooling to
52°C permits the bulk of the ammonia to pass through with the-cooled
gas. The cooled gas (stream 3) then enters the H2S absorber, where it
countercurrently contacts the freshly stripped, lean ammoniacal solu-
tion. Nearly all the H2S and a portion of_ the CO2 are absorbed before
the treated gas (stream 4) exits from the top of the absorber. The
rich solution (stream 5) enters the acid-gas stripper, wher^ the ab-
sorbent solution is heated to expel the absorbed acid gases. The
stripped or lean absorbent (stream 6) is cooled and then returned to
the H2S absorber; the stripped acid gas (stream 9) is sent to the Claus
unit for recovery of sulfur.
A small amount of solution (stream 7) is purged from the recirculating
absorbent solution to control buildup of impurities and then is sent to
foul-water treatment.
The goal of obtaining a high degree of sulfur removal conflicts with j
that of simultaneously obtaining high H2S selectivity; one goal must be[
compromised to obtain the other. To obtain a high degree of desulfuri-
1 ! 'j
zation the partial pressure of H2S in the absorbent must be low and the :
acid-gas concentration in the regenerated lean solution musjt be at the ;
lowest possible level. Thus large amounts of solution must be circu- ;
lated through the absorber and greater amounts of utilities, consumed in
regenerating and recirculating the solution.
1 i :
BOTTCM O
The system proposed for the model oil-shale gas disulfurization plant
I : " i
3__based,on a. spray .tower, absorber.with six absorption stages, (see _.-;
f i , ^ _-
{: 3/3"
•AND ILLuc-
•RATIONS
tr>A-?i,7 (Cin.)
PAGE
-------
;i
-i .p
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I
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01
cn
s
(8
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TYFiKG Gv.!;l" S- ifHT
LAST LIME
•)r TEXT j.>~
Appendix D).20 The high C02 content results in increased Cb2 absorp- -
tion. Calculations indicate that the H2S content of the treated gas
will be about 63 ppmv and that-the acid gas stripped from the rich
Diamox solution in the acid-gas stripper contains about 6 mole % H2S.20
JThis equates to an H2S removal of..97.9% while achieving a selectivity ..
ratio of 4.7. With the gas assumed to contain COS in the range of 10
to 50 ppmv the overall sulfur removal achievable would be 96.3 to
97.6%. The acid gas is fed to a Claus sulfur recovery unit!for conver-
sion of the H2S to elemental sulfur (Fig. 16 ). The Claus unit
, !
tail-gas feeds to a Beavon sulfur removal process (BSPR) for final sul-
1 • !
fur recovery (Fig. 17 ). The BSPR unit was chosen due to the large
> i
amount of CO2 in the acid gas. ' ',
The Beavon sulfur removal process for treatment of Claus plant tail
gases consists of two main steps: catalytic hydrogenation and hydroly-
sis of all sulfur species to H2S and conversion of the H2S to elemental
;*';•**•• t !
sulfur by the Stretford process.: ;
The overall sulfur recovery in the Claus and BSRP units is 99.8% per-
cent.20 The estimated overall sulfur reduction calculated for the
total desulfurization train is about 96.1 to 97.4%, depending on the
quantity of organic sulfur in the gas.
The Diamox process, while suitable for coke-oven gas containing large
i • !
amounts of ammonia, carbon dioxide, and hydrogen cynide, is;only
marginally capable of processing oil-shale gas because the ratio of C02
; t '
to H2S in the gas is very high. The process would not be capable of
: i ' .
producing an acid gas of sufficient H2S content for effective operation
of the Claus system for most direct-fired oil-shale retort gases with-
i
out considerable sacrifice of H2S removal effectiveness. ]
IMAGE A-
_!,_! FOrr-TAEi.'.'
1
EPAO37 (Cin.)
14-76)'
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-E.
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! 4.
OF TE
LiME
XT
6.
7.
8.
9.
10.
11.
12.
13.
14.
15.
REFERENCES* • j
R A Loucks, Occidental Vertical Modified in Situ Process for the
Recovery of Oil from Shale. Phase I, Occidental Oil Shale, Inc., Grand
Junction, CO, prepared for EPA under Contract No. DE-FC20-78LC10036
(November 1979). j , ....
"Telephone conversation Feb.~8Tl98o7""between S.~~W. Dylewski, !IT Enviro-
science, and M. Lekas, Geokinetics.
T. Nevens 'et al., Predicted Costs of Environmental Controls for a Com-
mercial Oil Shale Industry, Denver Research Institute, COO-5107-1 (July
1979). | j
i i
A. L. Kohl and F. C. Riesenfeld, Gas Purification, 3d ed., p^476, Gulf
Publishing Co., Houston, TX, 1979.
Sulfur Recovery Qualifications and Experience, Technical Documentation,
The Pritchard Corporation (November_1978)._
A. L. Kohl and F. C. Riesenfeld, op_. cit., p 477.
I
Ibid., p 479.
S. VasanY "Holmes-Stretford Process Offers Economic H2S Removal," The
Oil and Gas Journal 76(1), 78—80 (Jan. 2, 1978).
A. L. Kohl and F. C. Riesenfeld, op_. cit., p 484.
S. Vasan, *The Holmes-Stretford Process for Desulfurization of Tail-
Gases from Acid-Gas Systems," presented at the Ammonia-from-Coal
Symposium, held at TVA National Fertilizer Development Center, Muscle
Shoals, AL, May 9, 1979. ,; i
A. L. Kohl and F. C. Riesenfeld, p£. cit., p 485. :
R. Dagani, "Cleaning of Geothermal Steam Simplified," Chemical and
Engineering News 75(12), 29—30 (Dec. 3, 1979).
Letter from W. Dyer, EIC Corporation, dated April 1, 1980, to R. Lovell.
IT Enviroscience.
i
G. Allen,'pacific Gas and Electric Co., and F. Brown, EIC Cprporation,
Highlights of the Test Results From the Operation of a 5 HW, Pilot
Plant Demonstration of the EIC Process at The Geysers, a report pro-
vided by the EIC Corporation, Newton, MA. i
i I
S. Miller^ Acetylene. Its Properties. Manufacture and Uses,, Vol. 1,
Academic Press, New York, 1965.
! V-.GE f-~~-
TA2L-
EPA-237 (Cin.)
-------
-16. R. L. Pearce, "Hydrogen Sulfide Removal with Methyl Diethanolamine,"
pp 139 144 in Proceedings of the 57th Annual Convention of the Gas
Producers Association, March. 20—22, 1978, New Orleans. }
17. R. W. Parrish and H. B. Neilson, Synthesis Gas PurificationiIncluding
Removal of Trace Contaminants, presented at the 167 National Meeting of
- the American Chemical Society, Division of Industrial and Engineering
Chemistry, Los Angeles, California, March 31 to April 5, 1974.
i , i
' * i
18. A. L. Kohl and F. C. Riesenfeld, op. cit., p 784.
i t """"~~~ i
19. Telephone conversation Jan. 28, 1980, between C. A. Peterson, IT Envi-
roscience, and R. L. Pearce, Dow Chemical USA. !
' :
20. R. E. Meissner, III, Diamox Desulfurization Process for Treating Oil
Shale Retort Gases, RMP File No. 180G-1714; unpublished report provided
by the Ralph M. Parsons Company, Pasadena, CA (Mar. 25, 198Q).
I ; '
I " .
21._ J. Naber, J. Wesselingh, and W. Groenendaal, "New Shell Process Treats
Claus Off-Gas," Chemical Engineering Progress 69(12). 29—34 (Dec."
1973). i __ ;
! i
22. T. Shiboya et al., Process for Preparing Purified Coke Oven Gas, United
States Patent 3,880,617, Apr. 29, 1975. 1
23. H. Hiraoka, Mitsubiski Chemical Industries, Ltd., and E. Tanaka and H.
Sudo, Mitsubishi Kakoi Kaisha, Ltd., "DIAMOX Process for the Removal of
H2S in Coke Oven Gas," Proceedings of the Symposium on Treatment of
Coke-Oven Gas, May 1977. McMaster University Press, Hamilton, Ontario,
Canada.
*When a reference number is used at the end of a paragraph or on a head-
ing, it usually refers to the entire paragraph or material under the
heading. When, however, an additional reference is required for only a
certain portion of the paragraph or captioned material, the 'earlier
reference number may not apply to that particular portion. '
I
TT
r"
EPA-2S7 (Gin.)
(4-7S)
PAGL
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BEGIN
LAST LIN
O:T TEXT
TIX COST COMPARISON OF CANDIDATE PROCESSES [
I ! :
Estimated capital and operating costs and cost-effectiveness data for
the candidate processes described in Sect. VIII are presented *nth»_
section The estimates were based on a hypothetical directed oxl-._.
"shale plant processing 45,450 tonnes of shale per day .and producing
10.2 million s^d of gas having the composition shown in Table ;34
Material balance flow sheets were developed for each of the selected
candidate processes. The major equipment components were defined and
sized and the capital and operating costs estimated. :
I I - ; ;
The estimlted capital and operating costs and net annualised tost of
the various options evaluated.are_given.in_Table_ 37. _ Thejnaterxal.
"balance flow sheets and a summary of the equipment requirements and ;
cost estimate details for each process are given in Appendices A, B, C, ^
and D. ! !
9-1/3" . |
The estimated costs are based on a new-plant installation and represent
the total investment, including all indirect costs such as engineering
,nd contractors' fees and overheads, required for the purchase and
installation of all equipment and material to provide , facility as
described: These are battery-limit costs and do not include^provisions
for bringing utilities, services', or roads to the site; backup facili-
ties; land; required research and development; or process piping inter-
con»ectio,is that may be required^within the process that generates the
gas fed to the desulfurization systems.
capital cost estimates were developed by suction of installed capital
costs for'.the individual components of each system. These installed
capital costs are based on n Enviroscience experience adjusted to the
January 1980 basis. « addition'to the sum of the itemized capital
costs a contingency allowance of 20% is included in overall, capital
cost estimates.
Er'A-257 (Cin.
-------
Table 37. Estimated Capital and Operating Costs "for
Fuel-Gas Desulfurization Options21
Item
Capital cost9
Operating cost (per year)
Chemical makevip
Steam
Power
Fuel gas
. Cooling water
Waste treatment
Operating labor
Maintenance , capital
recovery and misc.
Total operating cost
Sulfur recovery credit
(per year)
Net annualized cost
Per year
Per ton of shale
Per barrel of oil
Stretford
Direct
Process with
Purge Stream
Disposal'3
$13,322,000
$1,088,000
70,000
303,000
526,000
No cost
600,000
3,864,000
$6,451,000
$1.083,000
$5,368,000
0.29
• O.50
Stretford
Direct
Process with
Purge Stream
Recovery0
$14,840,000
$ 911,000
107,000
345,000
328,000
578,000
Neg.
720,000
4,304,000
$7,293,000
$1,116,000
$6,177,000
0.34
. 0.57
Costs
1-Stage MDEA
Selective
Absorption
with
Stretford
Sulfur
Recovery"
$17,200,00
$ 1,190,000
3,812,000
234,000
1,261,000
No cost
630,000
4,988,000
$12,115,000
$1,130,000
$10,985,000
O.60
1.01
1
3-Stage MDEA.
Selective
Absorption
. with Claus
Sulfur
Recovery3
$15,551,000
$ 24,000
.7,183,000
158,000
1,545,000
Neg.
. 480,000
4,510,000
$13,900,000 •
$1,171,000
i
$12,729,000
0.70
1.17
.
Diamox
Process
with Claus
Sulfur,
Recovery""
$32,741,000
$ 84,000
11,530,000
1,618,000
1.950,000
No cost
480,000
9,495,000
$25,157,000
$1,150 ,000
$24,007,000
1.32
2.21
"l-or 360-million-scfd plant based on Paxaho gas (see Table VZII-1) . bstretford purge stream disposal at no cost.
and gases, water, and decomposed sodium salts recycled. Sulfur recovered with
S^^euZOXu PUiQi; StiecJa iJiUJ.*iei.€»uc«j. OJ*M. y&^«&^ / «»*»,^— , -^. -
Stretford system. ' eSulfur recovered with Claus system with SCOT tail-gas treatment. Sulfur recovered with Claus
system with BSHP tail-gas treatment. Includes process royalty fee.
131
-------
The actual cost of applying any of these options at any specific loca-
tion could vary considerably from that given here. The purpose of
these estimates was to determine the relative cost differences between
the various options. Since the estimates for each of the options eval-
uated were calculated on the same basest .the relative costsjof applying
the systems should be in the order shown. ''••
\ '•
! I ' j
Special cautions must be used if an attempt is made to extrapolate from
the cost-effectiveness data given since these costs will be;largely
dependent on the composition- and volume of gases to be treated, on the
oil yields, and on the cost factors that are pertinent to installation
" l , .
site. I ; i
For desulfurization of gases from direct-fired oil-shale retorts the :
Stretford process is the most cost-effective system. For the model ',
case shown the cost of sulfur removal would be about $0.50 per barrel
of oil produced, which is less than half that projected for -the better '
of the indirect sulfur-removal processes. The model case is based on '
Paraho gas, which has the lowest C02/H2S ratio (76:1) of any gas |
considered for the direct-fired retorts. For direct-fired gases having
higher CO2/H2S ratios the gap between the cost of applying the Stret- •
j ford process and the indirect desulfurization processes would be ex-
pected to increase, making the other options even less competitive.
• f
For gases with lower CO2/H2S ratios the indirect sulfur-removal proc- :
esses would become more competitive. For gases from indirect-heated
retorts, such as Tosco with a C02/H2S ratio of 4.1:1, the indirect
processes could very well become more cost effective than the Stretford
process. !
I
The operating costs given are based on the assumption that there is no
1 i
cost for wastewater treatment. The sour water stream in each case i
! ; i
would be stripped of H2S and processed for recovery of hydrocarbon and
•ammonia by-products. The cost of this treatment, which is partly off- '. *-;:.j£ A"
set by the valve of the by-products recovered, was assumed £o be equal
I ; •• , U:,\U.: .biU
-r_in..each case, and therefore is not included The. Stretford purge I "'.'• 7*-2Lr
1 ."i -s .-,.. .... ' •»/• f ' "> 11 ! I • •
i •; j.-c *•* •:•:•:-. •:•:•:•:•:•: • i ;-.i>.^.-i._u.
i V '••y'-y.:-. .132 :::'.:::::: • iPAliO?,1'^
PAGE KI-V3ER
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cr,\'rrn
O:-~ PAi.:i.
streams, containing essentially dissolved salts, were assumed to be •— :
sicceptable for disposal through moisturizing of the spent shale at no
cost. The cost of evaporating -the purge streams to recover dissolved
salts would be about $557,000/yr ($0.05/bbl) for the Stretford direct ''•
systems and_about^$37,000/yr ($0.003/bbl) .for...the BSRP..system used.for...;
: {
treatment of the tail gases from the Diamox process Claus unit.
The cost of a Stretford system incorporating a purge-stream ;reductive-
incineration system for recovery and recycle of salts is also given in
Table IX-1. This option would result in a sulfur removal cost of about
i. i
$0.57 per barrel of oil produced, which is about $0.07 more per barrel
than the case based on disposal of the purge stream at no cost.
i
I
9-1/8"
EEG'N
LAST LINE
OF TEXT
_ i!'!l _ L
BOTTOM C:
IMAGE AFi:
OUTSIDE
DiMENSIC
I'OR TABL-:
j-AND ILLU2
! TRAT!G;-J§
EPA-2S7 (Cin.)
(4-70)
PAGE NUMBER
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BEGIN
LAST LINE
OF TEXT
X. RECOMMENDED AVAILABLE H2S CONTROL TECHNOLOG-Y I —.';
i . i ; ; ,..!
The Stretford direct fuel-gas desulfurization process was recommended as
i . ,
the most applicable technology for removal of sulfur from direct-fired
_oil^_shale_ retort_gases. The basis_for__this, selection, is as ;follows.. „
A.
ADAPTABILITY
Gases from direct-fired retorts have a low heating value (1-52 to 4.13x10"
->' ' ' • ''
joules/snP) because of the large amount of nitrogen, carbon dioxide and other
i • i .
inert gases in the gas. It is not economically feasible to upgrade the
quality of the gas so that it will meet pipe-line quality standards and
! i . • ' !
can therefore be sold as a substitute for natural gas. Practical uses
i ' ', I
of the gas are limited to combustion for generating processjheat _or_
electrical power, for which purpose only the ammonia and sulfur need to .
be removed.
I
0.1 •'-"
The large volumes of gas that must be processed in a typical oil-shale
plant will limit the application of desulfurization technology to high-
capacity, continuous-liquid-phase processes. Since C02 is absorbed to
some extent by all liquid-phase processes, the high C02/H2Siratio of :
1 : I
the gas limits the selection to those processes that can selectively
i i i
absorb sulfur compounds in the presence of large amounts of•CO2• Of
those direct processes that selectively remove H2S by direct conversion
of the H2S to elemental sulfur, the Stretford process is the most
effective. Of those indirect processes that selectively remove H2S by
! i !
separating the H2S as a concentrated acid-gas stream, the Selectamine
I i I
or the Adip processes using MDEA as the absorbent, the Benfield, the
i j <
Selexol, and the Diamox processes were selected as the most effective
in their separate process classifications. These processes\were than
considered as candidates for further study.
B.
OPERATIONAL REQUIREMENTS
I
Oil-shale .retort gases are produced at near-atmospheric pressures; thus
those processes requiring the gas to be at a high pressure (Benfield
BOTTOM Oc
FOR
VAND ILLUS-
! TRATIC.'vlS
EPA-?37 (Cin.)
(4-76J
PAGE NUMBER
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EEC IN
LAST 'Uf.E
OF TEXT r
and Selexol) were eliminated since compression of the gas for the pur- -
pose of desulfurization could not be economically justified.; j
The Claus process is used to recover sulfur from the acid gas produced
Jby the indirect sulfur-removal processes... The ..large amount of C02 in _
'the gas makes the best of the indirect process only marginally capable
of producing an acid gas rich enough in H2S for processing by the Claus
process. Thus to apply these processes, multiple stages of iselective
•absorption would be required to handle the gas produced by many of the
direct-fired retorts.
The Stretford direct process, on the other hand, is only minimally
affected by the quantity of C02 in the gas and therefore is adaptable
to the full range of gases produced by direct-fired retorts. .
A Stretford plant incorporating a venturi absorber can operate with a
i ;
wide-volume turndown range. The system capacity is limited"by the
maximum design-basis throughput capacity of the sulfur. The quantity
of sulfur in the gas can drop to nearly zero without affecting the
i ;
sulfur removal efficiency.
The indirect selective sulfur-removal processes have only limited turn-;
i • ii
down capability. The selectivity of the process (preferencial absorp- [
'. i : j
tion of H2S over CO2) is largely achieved through control of the ab- '•
sorption kinetics. Therefore large changes in absorber gas or liquid
throughput can adversely affect the selectivity achieved and conse- •
quently the composition of the acid gas produced. The Claus process !
i • i
has a limited range of operation and can be upset by large variations i
! \ I
in the composition of the acid-gas feed. The wide range and constant :
: ! i
variation of the CO2/H2O ratio of gases from in-situ retorts would be
very difficult to handle with an indirect selective absorption system.
BOTTOM Or
IMAGE ARE-'
C, SULFUR-REMOVAL EFFECTIVENESS
A very high degree of H2S removal can be achieved by all the candidate ^ '*'~''.^'~~.
processes, discussed..._ Generally the higher the amount of .sulfur, that
--.. ' •--•.•• '-,..•,.-... i
JL
\<-AND ILLUS-
_j Th AT IONS
I4-7G)
(Cin.:
PAGE .NUf.'Si-R
-------
imust be removed the more it will cost to install and operate the desul-
furization train. The cost of applying these systems then becomes the
overriding factor rather than the ultimate sulfur-removal effective-
\
ness. i . •
KEG1N
LAST LINE
r*- —
D.
Except for the Diamox process all the candidate processes are capable
of removing H2S down to about 10 ppmv. However, organic sulfur com-
pounds , principally COS, which exists in only trace amounts iin the gas,
are not significantly removed or are only partly removed by the various
process. Table 38 gives the variation in process effectiveness with
the amount of COS in the raw gas. Current Colorado emission standards
limit sulfur emissions to an equivalent of 6.3 Ib of SO2/bbl of crude
oil produced. The estimated S02 equivalent emissions_for_each of the
candidate process are shown in Table 38. ;
RELATIVE COSTS
The relative costs of the candidate processes studied are given in
Table 39. i
I
For desulfurization of gases from direct-fired oil-shale retorts the
Stretford process is the most cost-effective system. The cost ratios
shown, based on Paraho gas, would be expected to increase for other
direct-fired-retort gases having higher CO2/H2S ratios. ;
The three-stage MDEA process produces the highest degree of37 (Cm.)
PAGE NUMBER
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Stretford direct j
|
3-stage MDEA with Claus/SCOT
1-stage MDEA with Stretford
Diamox with Claus/BSRP
7570 Jl/day
37.85.A/day
7570 a/day
246 H/day
F.
PEGS::
LAST UNE
These estimates do not include sour water, which would be processed for
hydrocarbon and ammonia recovery prior to disposal. The quantity of
sour water in each case would be about the same (1.13 x 10° to 1.51 x
106 it/day).
RELIABILITY
All the candidate processes were commercially developed and\are used in
1 ; J
various applications around_ the world._ The Stretford process Jhas been '.
widely applied, with more than 80 commercial plants currently in opera-
tion. ;
As with any process application operating problems can develop, and if
the system is misapplied or poorly designed, the problems can be insur-
mountable. The principal problems reported by Stretford users are
plugging of packed absorbers by precipitated sulfur; loss of scrubbing
performance, when processing gases with large percentages of CO2, be-
cause CO2 is absorbed and the solution pH is consequently lowered; and
i i
poor-quality sulfur as the result of hydrocarbon and other contaminants
in the sulfur.
i
The process licensors feel confident that with proper design and
through use of proprietary Stretford chemicals these problems can be
managed. ;
EPA-2S7 (Gin.)
J4-7SJ
:*:.:.:.:.:•:•*•,:•:•: •:•:•:•:•
&;&, .139' -_; xj
PAGE KUV.bER
BOTTCV 0!
ivAGE AR:
OUTSIDE
__J -CP. TABLL
*>AND !LLU5
f-vriON?
-------
G'-'i . •-'. S''.- • i
XI. PILOT-PLANT DESIGN
LA of LIN
Or 1GXT
A. INTRODUCTION
Gases produced from direct-fired oil-shale retorting processes are
~ ______ sufficiently different from gases previously encountered in commercial _.
t
application of gas-desulfurization systems that the technology cannot
be simply transferred. Application of the Stretford proces;s or any .
other process to the treatment of these gases would extend the technol- ,
ogy of the process into areas in which no analogous experience is
available. Many questions need to be answered before the process can
be applied with confidence to a full-scale commercial shale-oil produc-
tion facility. . I
i ! ;
_ _ _; __ I __ r,.-> ">'•' ______ J ___ _ _ __ ___ _ • ______ -j.
The primary function of the pilot plant would be to test the technical
feasibility of the Stretford process for treatment of gases from oil-
shale retorting processes, to prove the engineering assumptions and
calculated forecasts made for the process, and to generate scaleup
i
design data for commercial applications of the process.
The pilot plant should be portable and capable of operating on a slip
stream from any of the currently operating or proposed direct-fired
oil-shale retorting facilities in the United States. The unit should
be flexible enough to accommodate inputs from surface and in-situ
retorts with flow rates ranging from 2.83 to 28.3 sm^m while retaining
adequate design efficiencies; if possible it should have sufficient
: i ;
turndown capability to also operate on gases from indirect-heated
i
retorts. '•
1
The pilot plant would be a Stretford system based on the current state-!
of the-art technology for commercial applications of the process. The '
unit should be the smallest size that will operate on actual shale-oil |
permit study
retort-gas feeds under typical plant constraints and still
of the process dynamics.
X T /Q ' •
,., O. O
j._
'A-?37 (Cin.)
-7bj
PAGE KUV.SLR
t>0 F i <-/'•- Or
IMAGE />Rr
O'jTS'nz
-------
B.
SIZING OF PILOT PLANT •' i :
i !
Based on the gas characterization data compiled the CO2/H2S mole ratios
are widely separated in the gases from indirect- and direct-fired
retorting processes, and it is unlikely that a fixed-size pilot plant
could handle the full range of gases from both types of retorts. With
direct-fired retorts, either above or below ground where combustion
occurs within the retort, gas rates vary from 68.5 . to 405 sm3/tonne
of shale processed and the H2S content varies from 0.07 to 0.30 vol %
(DG). The CO2/H2S molar ratio varies from 76 to 165 or more,
•c!—
For indirect-heated retorts, where heat to the retort is generated
externally, the gas rates vary from 27.2to 35.0 sm3/ tonne of; shale proc-
essed and the H2S content varies from 3.8 to 4.1 vol % (DG) .^ The CO2/
H2S molar ratio varies from 4.3 to 5.0.
• 1 " ' • i . :
Thus .for a volume turndown ratio of 10 (28.3 to 2.83 smty the pilot
plant would have to have a sulfur turndown capability of 586 to 1 if
the H2S content of the gas ranged from 0.07 to 4.1 vol %. If the plant
were designed for only gases from direct-fired retorts, a sulfur turn-
down capability of 34 to 1 would be required. Similarly if ithe plant
were designed for only gases from indirect-heated retorts, a| sulfur
turndown capability of 11 to 1 would be required. However, 'if the gas
volume were decreased as the sulfur content was increased, then a sul-
fur turndown capability of only 6 would be required to handle the full
range of gases from both direct and indirect oil-shale retorts (i.e.,
2.83 snAnof gas from an indirect-heated retort containing 4.il% H2S to
28.3 sm3m of, gas from a direct-fired retort containing 0.07% H2S).
In reality the Stretford system is limited in both gas throughput
i ; ;
capacity and sulfur loading capacity. Therefore to provide maximum ;
! i
turndown capability, the pilot plant should be sized to handle the i
maximum case for direct-fired retorts (28.3 sm3m of Paraho gas contain-
ing 0.30 vol % H2S). Gases containing less sulfur can be handled up to
i !
28.3 sm3m since the system is capable of operating with reduced sulfur
ii 3/8"
w
_ 1
'.'J \LL-J'
triA-2:.17 (Cm.)
-------
s;-c i •»'••
EEGirj
IA SI L'.ME
OF TEXT i-J
loadings. For gases containing more sulfur, the gas volume can be —
turned down to within the limits of the absorber operation while main-
taining maximum sulfur loading.
! I ' .
' j —
DUTY SPECIFICATIONS
The pilot plant should be designed primarily to remove H2S from oil-
shale gas produced by direct-fired retorts. The maximum capacity of
the pilot plant should be 28.3 sm^m of feed gas. The maximum sulfur
capacity should be 6.6 Kg of sulfur per hour based on a loading of
800 ppmw of hydrosulfide ion (HS ) in the solution. The plant should
be capable of reducing the H2S content of the gas to 10 ppmv or less,
and CO2/H2S ratios as high as 200 to 1.
The pilot plant should include equipment for cooling the fee|d gas from
60° to 32°C and for removing the ammonia before the feed gas enters the
Stretford absorber. The pilot-plant design should include only those
equipment components, instruments, and controls necessary for safe
operation of the equipment and as required to adequately test and
demonstrate the Stretford process and to carry out the intended re-
search program. Most of the data required for determining the perfor-
mance and operation of the system would be obtained by manual sampling
and laboratory analysis of the process streams. -
D. RECOMMENDED SYSTEM
1. Basis of Design
i
The bases for designing and sizing the pilot-plant components are as
follows: I
28.3 sm3m
Atra
60°C
Gas characteristics
i
Maximum volume (dry gas)
Pressure
i
Temperature
i!H_l
BOTTC'.' ••'•
"•••Aro ILL ,
F"A-2S7 (Cin.)
(4-76;
-------
Critical components
H2S
C02
NH3
I
Cooling water temperature
Sour-water treatment
Sulfur recovery
Stretford solution sulfide loading
0.3 vol % ;
22.8 vol %
0.7 vol % !
Saturated ;
2i°c ;
Not included
Disposal as wet cake
800 ppmw !
Description of Process
A material balance flow sheet for the Stretford pilot plant is shown
Fig. XI-1- -The idealized process reactions are as follows: - —
on
Reaction 1
i-"= Na2CO3 + H2S
NaHS + NaHC03
Reaction 2
-4NaV03 + 2HaHS + H2O
.1
Reaction 3
i
!Na2V409 + 2NaOH + 2ADA*
i
Reaction 4
i
!2ADA (reduced) + O2 -
Na2V4O9 + 4NaOH + 2S
4NaVO3 + 2ADA (reduced)
2ADA* + H2O
*Anthraquinone disulfonic acid.
The sour fuel-gas feed (stream 1) enters the gas cooler, where it is
cooled by direct contact with circulating water solution. The flow
rate of feed gas is controlled to the rate set for the test and is
continually recorded. As the gas is cooled, a large portion of the ;
.water.and most.of.the.ammonia are removed (stream.2) Heat
3 8
ICin.)
$$$.,l\3 ';:i;:i;::
F'AGC NL-'.'Cn
is .ex- ..
-------
CAS
COCKER
ABSORBER
DELAY TANK
SOLUTION
PUMP TANK
SLURRY TANK
OXIDIZER AND CHEMICAL SLURRY
TANK MAKEUP TANK PUMP
SULFUR
SULFUR TRANSFER
FILTER BOX
N, *n*r>
O.«- \. "0-
0.
*,
-?i- -
f.u.
<•».
£,".
<.-.
i'"> 8
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,-y l
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fftM
*-« *V,
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*•-•
•«•
15.10
3.29
78J5.53
11068
ISS667
54 72
2928"
2661
48.33
S055
6854
_
16 17
1883
712.03
<*€S38
10OOOO
140' f
<5>
48.67
Trace
of H-C'i
0.12
1883
581-87
649.49
90° F
15.10
2.29
2825.53
11068
1538.00
54.72
29.28
26.61
48.38
50.55
68.54
16.05
Trace
130.16
4915.89
9S6.00
90= F
<£>
15.10
2.29
2825.53
110.68
1538.00
54.72
29.28
26.61
48.38
50.55
68.54
00.54
_
130.16
4899.89
Oxygen
Nitrogen
NaHS
Water
Sulfur
Vanadate
ADA
EDTA
Iron (Soluble)
NaHCO,
Ma. CO,
Na,S,0,
Hydrocarbons
Total. Ib/hr
Total. SCFM
Total, gpm
^^
20000.00
114.90
214.90
54.00
2.00
1016.00
59.90
382250
25285.00
42.20
^^
25.57*
20006.00
I
114.90
214.90
54.00
2.00
1055.60
10.00
3824.00
Trice
25301 .00
<>>
36.70
120.90
2.80
160.40
34.50
29.40
120.90
4.80
Trace
155.10
19832. OO
—
113.90
213.10
53.50
138
1046.70
9.92
3791 SO
25063.00
^>
16S.OO
14.60
0.96
1.80
0.45
0.02
8.86
0.08
32.10
226.87
<1>
i
92,13
—
O.E5
1.02
0.2S
0,01
5.53
0:23
18.42
.
118.25
.
.
1
<2>
8.79
8.79
X7
24.60
0.03
0.06
0.02
Trace
0.31
Trace
1.11
40.73
^5*
0.03
0.05
0.02
Trace
0.29
6.21
*•*,(«•* MWOT bttxMfl [HS] and Sulfur
18. Material Balance Flowsheet for Pilot Plant
>v-: 144 -.-...<
TYPING GU!D2i:
-------
tracted by circulating the water solution through a water-cooled heat
t ' ' I
exchanger. i i I
The temperatures of the incoming sour gas, the cooled gas, and the
cooling water discharge streams are continually monitored..,. The .sour .__
!
water (stream 2) is discharged. Treatment of the sour water is outside
the scope of the pilot plant. ',
'
LAC!
The cooled gas (stream 3) then enters the ejector-venturi gas absorber
at about 90°F. The ejector-venturi absorber (see Appendix E) operates
on the jet principle, in which the Stretford solution is under pressure
to create a draft and simultaneously provide intimate contact between
the liquid and gas as it passes through the yenturi_. On contact with
the alkaline Stretford solution, which is an aqueous solution contain-
ing sodium carbonate, sodium metavanadate, and anthraquinone difulfonic
acid (ADA) , the H2S dissolves rapidly and forms a small concentration
of hydrosulfide ion (reaction l)j Sodium carbonate in the Solution
enters into the reaction and provides a buffer to prevent rapid pH
i
changes as the acid gases are absorbed. i
i I :
The solution will absorb some CO2, resulting in formation of sodium
bicarbonate and consequent lowering of the pH until equilibrium is
reached with respect to the level of C02 in the gas. The pH of the
Stretford solution will therefore be a function of the level of CO2 in
the gas. !At the point when equilibrium is reached, there must be suf-
ficient Na2CO3 remaining to convert all the H2S in the solution to
hydrosulfide ions .
I '
For gases "with large C02/H2S ratios the pH of the solution ^ay be suf-
f I
ficiently lowered to affect the absorption rate of H2S. In that case
1 »
addition of proprietary additives to the solution may be required for
the hydrosulfide loadings to be as high as 800 ppm.
COTTCf1 C
IMAGE AF,"
The venturi discharges into a mist eliminator, where gases are sepa-
_rated.from the liquids.- The treated.gas (stream 4),-in which essen-
(Cin.)
(4-76)
-------
*
LAST LINE
OF TEXT
tially all the H2S has been removed, will be returned to the oil-shale
process. The separated liquid will flow into the delay tank.
The sodium metavanadate in the solution readily reacts with'the hydro-
sulfide to produce elemental sulfur (reaction 2) and a reduced form of.
vanadium. Sufficient residence time is provided by the delay tank to
assure that any remaining hydrosulfide in the solution is reacted so as
i
to prevent hydrosulfide ions from entering the oxidizer and; becoming
i • !
oxidized to sodium thiosulfate, an undersirable, stable, by-product.
Flow (stream 6) from the delay tank is by gravity to minimize sticking
of the sulfur when the solution is in a transitory Redox stage. A
liquid-level controller is provided_ to control the _f low of solution
from the delay tank to the oxidizer. The reacted solution containing
reduced vanadium and reduced ADA is regenerated in the oxidizer tower
according to reactions 3 and 4 by being sparged with an excess of air
(stream 7) supplied by a blower. The flow of air can be adjusted so
that sulfur particles will properly rise, forming a froth at the sur-
face but allowing the Stretford solution to settle clear of sulfur.
The accumulated sulfur froth overflows (stream 10) to a slurry tank.
Trace amounts of volatile hydrocarbons that may be absorbed in the
Stretford solution will be stripped by the air rising in the oxidizer.
A hood provided over the oxidizer tank will lift and direct the dis-
! '
charge gases (stream 8) away from the operators.
t
The regenerated Stretford solution (stream 9) is essentially free of
! i
sulfur and is returned to the solution pump tank. The solution pump
tank is sized to hold all the required solution when the pilot plant is
shut down.
The regenerated solution (stream'4) is then pumped back to the venturi
' » ' f
absorber at a preset and controlled rate. The flow rate, pH, and tern- . ..,_.,.
I 1 , j ..:..• i I
perature of the solution are continually monitored and recorded.
r>-:M--N0-!G'
•:QS TA=L"
'-•-••JD iLLUl
EPA-7S7 (Cin.)
14-76)
PAGE
-------
rvr;i:
HEAL:.
Equipment for handling the sulfur slurry has been reduced to; the mini- ,
mum in an attempt to keep the cost of the pilot plant to a minimum and •
yet provide sufficient capability to test and demonstrate the Stretford
i • i i
process. , i I
BEGIN
LAST LiNE
OF TEXT r
1 ' i >
The sulfur produced will be disposed of as a wet cake. A sulfur filter .
(see Appendix E) operated on a batch basis will filter the sulfur from •
i " J
the slurry, The filtrate solution (stream 11) is returned to the solu-
tion pump tank. The wet sulfur cake accumulated in the filter will be !
discharged to a transfer box, which is sized for 4 days of operation at
: i
maximum utilization. The amount of Stretford solution lost with dis- ;
i
: ' ; t
charge of the wet sulfur cake (stream 13) should be sufficient to keep :
__ the buildup of sodium thiosulfate in balance. The cost of Stretford .
chemicals lost in the sulfur cake at the pilot level of operation is !
i
not a significant economic factor. The filter system is arranged so
t i
that the filter cake can be rinsed (stream 12) before is is discharged.
9-1/8"
I
3. Turndown Capability
The described capacity of the pilot plant ( 6.6 Kg of sulfur/hr) is
based on an average solution loading of 800 ppm. Sulfur loadings of
500 to 1000 ppm are commonly achieved. The maximum solution loading
i
obtained will depend on many factors but will largely be a function of
the solution pH, which is influenced by the amount of CO2 in the gas.
Solution loadings can be anywhere from zero up to the maximum capacity
of the solution, which therefore allows infinite turndown in the quan-
tity of sulfur contained in the feed gas. However, the cost of a
I ;
Stretford plant is largely a function of the amount of solution that
must be circulated, and therefore a goal in piloting the process is to
1 i :
determine .the maximum solution loading possible that will result in
adequate operation of the system.while obtaining a satisfactory level
of H2S removal effectiveness.
BOTTOM C~
' 'AGE AP>
! A 3/8"
i J— —
The use of the ejector-venturi gas scrubbing system (see Appendix E)
affords wide gas turndown capability for the system. Since :the flow of .:. J.'Jv
gas..is .motivated by. the. liquid spray, .the scrubbing system functions _.. J ' 'H T-
i ' i •* "-, \iP* i: I ! :
H X'XvX'XvX'X'X'X' f '•'•l^:'>-l-U,
V i£x£xi47.. x-x-': _ i TRATIG'JS
PAGE NUMBER
EPA-287 (Gin.)
(4-7C)
-------
largely independently of the gas flow rate. The scrubber, which has a
maximum gas capacity of 28.3 sm3m, can be turned down to nearly zero.
For gases containing larger amounts of sulfur the solution loading can
be adjusted by reducing the gas throughput. ;
By changing the ejector liquid-spray nozzle the rate of liquid circula-
tion can be reduced. Therefore gases with less sulfur can be accom-
modated while maintaining high solution loadings.
!- I :
The operation of the pilot plant would approximate that described in
Table 40 for various gas feeds. '•
i 4. Equipment Description_ j_ _ , . .
! "* '
The pilot plant would be a portable skid-mounted unit, completely
assembled with all equipment, controls, and instrumentation;required
for a complete and operable unit! Utilities and services needed to
operate the pilot plant must be supplied at the test site and be field
connected.
I
The maximum demand for services is as follows:
28.3 sm3m (30.5cia-diam duct)
28.3 sm3nr <30.5cm-diam duct)
30 hp (460 V, 3-phase, H2)
0.11 fcpm
1100 fcpm
4.9
Raw retort gas feed
Treated gas return
Electrical power
Process water
Cooling water
Sour-water disposal
A list of the major equipment required for construction of the pilot
i
plant is given in Table 41.
E. PRELIMINARY COST ESTIMATE
The total estimated cost of the pilot plant with all equipment, instru- I '.: '.CI A~:
i I ' , i Q! 'T^QC
ments, and controls as described, assembled on skid mountings as a com- .^,',~,"r~,
J~ , . I ---...- -.-.-.
; _plete and operable unit,, is as .follows:-
"PV3E N'JVFER
.4-76;
-------
Table . 40. Approximate Pilot-Plant Operation for Various Gases
Gas
Tosco
Union
Paraho
Geokinetics
Oxy
Type of
Retort
Indirect
Indirect
Direct
Direct
.Direct
H2S
(vol %)
4.12
3.82
0.30
0.13
0.119
Gas Flow
2.37
2.55
32.4
31.1
31.1
Sulfur
(KB')
6.6
.6.6
6.6 .
4.3
4.0
Liquid Flow
159
159
159
102
102
Solution
Loading
(ppm wt)
800
800
800
518
476
149
-------
cn
B
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I
8
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.
•
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?
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a
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^1 Each)
S 6 ^"
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£ rH *O
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Complete skid-mounted assembly as
5*1 cm, suction and discharge; 31
at 32°C; 1.27 to 2. 54cm water gagi
CQ
CO
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transmitter, controller, and re<
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*
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Flow control
n
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high-level alarm
pH monitor
Level control
o>
3
"8
i»
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%
4^
Temperature sensor, connected to.
recorder for Item No. 100 .
n
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As required
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151
-------
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lifliHO"._*rH<8.^r| jj «*|j O (j, . ~ ,ri o M 1 1 3
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^^
13
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c
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iH
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c
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•H
fV
•H
U
to
£
c
.3^.
•n rn
jj id
Construe
Materi
c —
8-S
am
u
SvH
^
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IH n ft
4J *
Sw
in
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^t o *
tH | O
*9 ""*"
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153
-------
Item-100 gas cooler module
Item-200 absorber module ;
i
Item 300 oxidizer module :
Item 400 sulfur handling module
Total cost (December 1980 capital)
$200,000
74,000
68,000
58,000
__$400,000_
If the gas were provided at 22°5,the gas cooler module could be
eliminated, in which case the total cost of the system would be about
$260,000. The cost of the gas cooler module (Item 100) includes the
instrumentation and controls required for proper functioning of the
pilot unit. These instruments ($60,000) would be required even if the
cooler module were eliminated. : !
__ . : ..--•" ... .... ..'_., j __--.
The probable cost range is estimated as follows:
jLow
With Cooler
$520,000
400,000
j
308,OOO
Without Cooler
$338,000
260,000
2OO,OdO
F.
ADVANTAGES AND USE OF PILOT PLANT :
The primary purposes of a pilot plant are to prove the technical feasi-
i ' '
bility of the process application and to obtain data needed!for scaleup !
of the process to production-size equipment. If the performance of the'.
pilot plant both at optimum and extreme conditions matched the I
predicted results and if there were no unreconcilable problems, then !
commerical application of the process could be carried out with confi- j
BEGIN
LAST LINE
1.
2.
3.
4.
5.
Or itXT i-'f-
i ' :
determine the characteristics of the raw retort gas.
determine the characteristics of the treated gas.
determine the characteristics of the waste streams.
predict the characteristics of controlled emissions.
determine the process efficiency and reliability.
-... ,....'
BOTTOM OF
IV. AGE ARE
OUTSIDE
DIMENSIOK
, _
OJA-7iJ7'(Cin.)
(4-70)
-------
TYPING GUlL-if SHLEI
BEGiN
LA£T LIN
or TEXT i:£-
I
1_ 6. evaluate transient conditions of startup, normal and emergency
i i
shutdown, process upsets, etc.,
• 7 evaluate the economics of the Stretford process. :
; I • i
Piloting of the process can aid in commercial application of the proc-
ess by providing demonstration of continuous operation of the system,
i '
as well as the following data:
i
1. Level of H2S and other sulfur compounds in the treated gas and
percentage of overall removal of sulfur compounds by the Stretford
process
2. Disposition of COS and other organic sulfur compounds in the feed
"™ 3^ ~ Effects of unsaturated hydrocarbons in the feed gas on process
operation and life of the Stretford chemicals
4. Amount of hydrocarbons or other contaminants in the oxidizer vent
5. Quality of sulfur produced
6. Solution sulfur loading versus gas characteristics ,
7. Rate of thiosulfate formation versus gas characteristics
8. Absorption of C02 versus gas characteristics
9. Design parameters for the absorber ,
a. HS" loading in the solution
b. 'Thiosulfate concentration
c. .Na2CO3 concentration ;
10. Design parameters for the oxidizer
a. 'Ratio of air to sulfur loading !
I i '
b. (Ratio of air to tank diameter
c. JHeight of oxidizer j ;
d. ICharacteristics of sulfur froth
i i I
11. Design parameters for filter
a. [Filter selection
b. iPumping rates and pump type ,
c.
•Filtration rates and wash cycles required
I.
A ' -5 .
y J
T
POTTO'.' C'
IMAGE AF/.
OUTS! 17 =
r'!MEr';'~'C
J^
.155
PAGr. NUMBER
TRATIO: Jt
£PA-2S7 (Cin.)
(4-76)
-------
12.
13.
d. Sulfur quality ;
e. Washwater evaporation rate required : •_-
Design parameters for gas cooler
a. NH3 and H2S absorption versus gas characteristics |
b. Characteristics of sour water _T
Evaluation of corrosion problems and materials of construction
OH TEXT
i —
37 iCir..) '
PAGE Nbr.'.SER
-------
9-1/5-
;. , U
j
EPA-:;:? !Cm.)
APPENDIX A
STRETFORD DIRECT PROCESS
PAC-E N
-------
[SHEET
PRELIMINARY CAPITAL - DETAIL SHEET I 1 OF 6
PROJECT
r^TT'-n'n'nTT'rYDT^ TYTP'PPT PPOPKS.S
' JOB NUMBtK
-. ; 9212
PHASE jCAbt BY
SECTION NUMBER ft NAME *
FACTORS
— i —
ITEM
101
102
103
104
201
202
203
'
204
205
206
207
208
209
210
211
212
ECTOR NUMBER & NAME |
BASE; x ESC. x CAPACITY • x QUANT, x
NAME OF FACILITY
Gas Coolers
Cooler Cir. Pumps
Exchangers
Trace S Cover
For Freeze Protection
Absorbers
Fwd. Pumps
Oxidizer Tower
Air Diffusion Sys.
Sulfur Slurry Tk.
Tank Agitator
Slurry Pump
Rot. Vac. Filter Sys.
Sulfur Melter
Sulfur Fwd. Pump
Sulfur Store Tk.
Agita.tor
QUANT.
6
6
6
16
16
1
1
1
1
1
1
1
1
[MATERIAL +
CONSTRUCTION
MATERIAL
304 St.Stl
304 St.Stl
304 St.Stl
Steel
FRP
Stl. &
Epoxy
Steel
Steel &
Epoxy
Steel
FRP.
Stl.
Steel
Steel
Steel
Steel
CONDITION + COMPLEXlTYj — Mi (19 )
DESCRIPTION j
i
12'$x42' Bl to Bl.; 10 PSI
18' Packed section .
2000 ft - 3 1/2"$ ;Pall Rings
1000 GPM, 100' HD, 40 HP
5000 ft2, U Tube, 14' LG Tube
All Items
Subtotal (100 Series)
12'$x42' Bl to Bl |
18 * Packed section .
2000 ft - 3 1/2" * Pall Rings
500 GPM, 40' HD, 10 HP
30'$x24' HI w/foam
Trough & Air Ring ;
7500 SCFM, Blowers',
Elec., Air Piping, Bldg.
15000 Gal, 25 PSI
Medium Ag., 15 HP.|
100 GPM, 40 HD, 3 HP
Incls; Filter, Vac. Pump,
Filtrate Receiver '& Pump,
Conveyor, Wash Tk, & Pump,
Building
1200 Gal., 25 PSI,: JKT. &
Internal Coil, Insulate
& Baffle i
10 GPM, 1 HP
1
30,000 Gal. 1 WK
API. S.G. = 2 !
Heavy, 5 HP !
$4,028
FORM 39650 PRINTED IN U.'i-A. Rl-69
158
-------
PRELIMINARY CAPITAL - DETAIL SHEET
PROJECT
STRETFORD DIRECT PROCESS
SECTION NUMBER t NAME
Case A
213
214
215
216
217
218
219
220
SECTOR NUMBER & NAME
BASE X ESC. X CAPACITY
NAME G
Drier Sys.
Make-up Mix Tk.'
Fwd. Pump
Lean Sol. Fd. Tk.
Tower Cool Pumps
Purge Pump
Mix Tk. Agitator
Trace & Cover for Freeze
Protection
XQUANT. X [MATERIAL +
CONDITION + COMPLEXITY]
Y
QUANT.
1
1
CONSTRUCTION
MATERIAL
Steel
Steel
20
Va
Co
27
1
1
8
1
1
St.Stl.
Steel
Epoxy
FKP
FRP
Steel
200 ft W/Heating, Air or
Vac. Sys., Feed & Disch.
Conveyor, Bldg. j
2700 Gal, 8 Hr. Store
100 GPM, 100' HD, 7 1/2 HP
36,000 Gal, 5 min. ;SG - 1.2
1000 GPM, 8Q'_ HD
100 GPM, 40' HD., 3 HP
Medium, 25 HP ',
All Items ;
As Required ;
Subtotal (200 Series)
Total (All Equipment)
= MS(19
$5943
$9971
FORM 39650 PRINTED IN U.S.A. R W9
159
-------
PRELIMINARY CAPITAL - DETAIL SHEET
f 6
PROJECT
STRETFORD DIRECT PROCESS
9212
PRE-1
Case A
2-1-80
SECTION NUMBER & NAME
SECTOR NUMBER 1 NAME
FACTORS
BASE X ESC. X CAPACITY
XQUANT. X [MATERIAL
CONDITION + COMPLEXITY]
= MS(19
ITEM
KAME OF FACILITY
QUANT.
CONSTRUCTION
MATERIAL
DESCRIPTIO'N
Case Al - Disposal of Purge Solution
Equipment as above (Sheets 1 and 2)
Allowance (20%)
Total
Royalty (10% of capital)
Initial chemical charge
1980 Installed Capital Cost
$ 9,971
1,994
11,965
1,197
160
$13,322
Case A2 - Recovery and Recycle of Purge Solution
Equipment as above (Sheets 1 and 2)
Purge stream reductive incineration system
(300 lb/hr salts)
Sub-total
Allowance (20%)
Total
Royalty (10% of capital)
f
Initial chemical charge
1980 Installed Capital Cost
$ 9,971
1,150
11,121
2,224
$13,345
1,335
160
$14,840
FORM 39650 PRINTED IN U.S.A. RI-6I
160
-------
CAPITAL - DETAIL SHEET
PROJECT
STRETFORD DIRECT PROCESS
V PHASE.
PRE-1
ITEM
; & NAME
CASE
Case Al
BY
SECTOR NUMBER & NAME
BASE X ESC. X CAPACITY
X QUANT. X [MATERIAL
NAME OF FACILITY
QUANT.
CONSTRUCTION
MATERIAL
Disposal of Purge Solution
CONDITION + COMPLEXITY]
DESCRIPTION
==MS(19 )
1. Raw Materials:
Vanadate
ADA
EDTA
Sodium Carbonate
Iron (Soluble)
$ 174,000/yr
887,000/yr
24,000/yr
3,000/yr
Neg/yr
$l,088,000/yr
2. By Products
Sulfur
12028 tons/yr x $90/ton = $1,083,000 (carbon)
3. Water Treatment or Disposal
"Sour" Water
"Purge" Water
"Filter Wash"
132x10 gal/yr
9.28x10 Ibs/yr
5x10 gal/yr
No charge
No charge
No charge
Utilities
10.1x10 KwH/yr
Electrical Power
150 Psig Steam 28x10" Ib/yr
Process & Cooling Water 10,000 GPM
Manpower
14 men x 2000 hr/yr x $.15/MnHr
1 supervisor 8760 hr x $20/MnHr
$303,000
70,000
526,000
$899,000
$420,000
180,000
$600,000
Maintenance,,(5%), Capital Rec. (20%), Misc. (4%) •
29% x capital $13,322,000 $3,864,000/yr
FORM 39650 PRINTED IN U.S.A. RIO
161
-------
CAPITAL - DETAIL SHEET
5 OF 6
PROJECT
n
, PHASE
PRE-1
Case Al
JOB NUMBER
9212
(UMBER
SECTION NUMBER * NAME
SECTOR NUMBER <• NAME
BASE X ESC. X CAPACITY
xQUANT, x [MATERIAL
CONDITION + COMPLEXITY]
MS(19 )
ITEM
NAME OF FACILITY
QUANT.
CONSTRUCTION
MATERIAL
DESCRIPTION
Disposal of Purge Solution (Cont'd)
7. Net Annualized Cost
8. Total Installed Capital
High, Dec. 1980
Probable, Dec. 1980
Low, Dec.. 1980
$5,368,000/yr ]
$0.294/ton of shale
$0.494/bbl of oil
$17,320,000
13,322,000
10,260,000
FORM 39650 PRINTED IN U.S.A. Rl-69
162
-------
CAPITAL - DETAIL SHEET
SHEET
G OF 6
STRETFORD DIRECT PROCESS
PRE-1
Case A2
BY
JOB NUMBER
9212
EF NUMBER
SECTION NUMBER * NAME
FACTORS
ITEM
2.
3.
SECTOR NUMBER
BASE X ESC. X CAPACITY
XQUANT. X [MATERIAL
CONDITION -f- COMPLEXITY]
= MS{19
NAME OF FACILITY
QUANT.
CONSTRUCTION
MATERIAL
Recovery and Recycle of Purge Solution
1. Raw Materials:
KDA
EDTA
By Products:
Sulfur, 12400 TPY
Water Treatment or Disposal
"Sour" Water 132x1.0 gal/yr
"Filter Wash" 5x10 gal/yr
Utilities:
Elec.
Stin.,
Water
Gas
11.5x10 KwH/yr
42.9x10 Ib/yr
11000 Gpm
15 M Btu/hr
5. Manpower
18 men x 2000 hr/yr x $15/MnHr
1 super. 8760 hr/yr x $20/MnHr
6. Maint. (5%), Capital Rec. (20%), Misc. (4%)
7. Net Annualized Cost
8. Total Installed Capital
High
Probable
Low
Dec. 1980
Dec. 1980
Dec. 1980
FORM 39650 PRINTED M U.S.A. Rl-69
DESCRIPTION
$887,000/yr
24,000/yf
$911,000/yr
$l,116,000/yr (Credit
No charge
No charge
$ 345,000/yr
107,000/yr
578,000/yr
328,000/yr
$l,358,000/yr
$540,000/yir
180,000/yr
$720,000/yr
$4,304,000/yr
$6,177,000/yr
$0.338/ton of shale
$0.569/bbl of oil
$19,292,000
14,840,000
11,426,000
163
-------
I . ! . i i li
O!" f-At;;?
LAST LINE
OF TEXT t-
I APPENDIX B :
THREE-STAGE SELECTIVE ABSORPTION PLUS GLAUS SULFUR RECOVERY
WITH SCOT TAIL GAS TREATMENT
9-1/8"
I
3/8'
J
EPA-287 (Cin.)
(4-7G)
PAGE NUMBER
BOTTOM OF
IMAGE ARE,:
OUTSIDE
D!f,'.Er;siot.
FOR TABLES
ILLUS-
-------
PRELIMINARY CAPITAL - DETAIL SHEET
1
SHEET
1 OF 4
PROJECT 3 STAGE SELECTIVE ABSORPTION PLUS GLAUS SULFUR RECOVERY WITH SCOT JOB u BER
QO1 *?
TAIL GAS TREATMENT , • „ :!„ „„
i ' ^
PRE-1 • Case B
SECTION NUMBER & NAME
FACTORS
ITEM
101
102
103
104
201
202
203
204
205
206
207
208
209
210
211
212
213
214
301
302
303
304
BASE X ESC. X CAPACITY X
MAME OF. FACILITY
Gas Cooler
Cooler Pump
Exchangers
Trace & Cover
For Freeze Protection
1st. Stage Absorber
Exchanger
Pump
Interchanger .. ....
1st Stage Desorber
Condenser
Pump
Pump
Reboileir
Compressor
Amine Recovery Reboil
Sludge Pump
Sludge Accumulator
Trace & Cover
For Freeze Protection
2nd Stg. Absorber
Pump
Interchanger
Cooler
SECTOR NUMBER & NAME
QUANT. X
OUANT.
6
6
6
16
16
16
16 ..
3
3
3
3
3
3
1
1
-1
1
1
1
1
[MATERIAL •+•
CONSTRUCTION
MATERIAL
304 St.Stl.
304 St.Stl.
304 St.Stl.
Steel
Steel
D.I.
Steel
Steel
Steel
D.I.
D.I.
Steel
Std.
Steel
D.I.
Steel
Steel
D.I.
Steel
Steel
CONDITION + COMPLEXITY] (
DESCRIPTION
12'$x42' B1-B1, 10 PSI, 18'
of Packing, 2000 ft
1000 GPM, 100' HD; 40 HP
5000 ft2 U Tube, 14' Lg.
.All Items ...
Subtotal (100 Series)
12'$x42' B1-B1, 10 Trays
829 ft2 ; '"'
150 GPM, 100' HD, 7 1/2 HP
2 :
287 ft ; , .
12 '$42' B1-B1, 10 Trays
3000 ft2, Float, 14' Lg
716 GPM, 150' HD, 50 HP
639 GPM, 60' HD, 15 HP
820 ft2
6200 ACFM, 2 PSI,: 25 HP
2300 Gal., 154 ft2
1/2 HP
10 Day, 150 Gal
All Items \
As Required
Subtotal (200 Series)
9'$x36' Bl-Bl, 10 Trays
506 GPM, 150' HD/ 25 HP
2
700 ft ;
2200 ft2
$4,028
-
$6,205
•
FORM 396311 PRINTED IN U.S. A. KI-69
165
-------
PRELIMINARY CAPITAL - DETAIL SHEET
MEET
2 OF 4
PROJECT 3 STAGE SELECTIVE ABSORPTION PLUS GLAUS SULFUR RECOVERY WITH SCOT : ^ goi?
'PHASE' ' CASE BY OATE , • EF NUMBER
Case B '
SECTION NUMBER & NAME
FACTORS
ITEM
305
306
307
308
309
310
311
401
402
403
404
405
406
407
408
409
410
411
500
BASE X ESC. X CAPACITY X
NAME OF FACILITY
2nd Stg. Desorber
Pump
Pump
Reboiler
Condenser
Compressor
Trace & Cover
For Freeze Protection
3rd Stg. Absorber
Pump
Interchanger
Cooler
3rd Stg. Desorber
Pump
Pump
Reboiler
Condenser
Compressor
Trace & Cover
For Freeze Protection
3 Stage Glaus Unit
Trace & Cover
As Required
SECTOR NUMBER & NAME
QUANT. X
QUANT.
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
Sys.
[MATERIAL •+•
CONSTRUCTION
MATERIAL
Steel
D.I.
D.I.
Steel
Steel
Std.
Steel
D.I.
Steel
Steel
Steel
D.I.
D.I.
Steel
Steel
Steel
DESCRIPTION
9'$x36' Bl-Bl, 10 Trays
432 GPM, 150' HD, 40 HP
432 GPM, 150' HD, 40 HP
900 ft2 ;
2000 ft2 !
6200 ACFM, 2 PSI, 25 HP
All Items
As Required
Subtotal (300 Series)
6'$x30' Bl-Bl, 10 Trays
200 GPM, 100', 20 HP
187 ft2
823 ft2
6'$x30' Bl-Bl, 10 Trays
160 GPM, 150' HD, 15 HP
160 GPM, 150' HD, 15 HP
300 ft2 "
731 ft2
2123 CFM, 7 1/2 HP;
i
All Items 1
As Required
Subtotal (Series 400)
2123 scfm, 27.4 moie % H2S
35.8 TPD of Sulfur'
$957
$462
$566
foot 396SO PRINTED IM U.S.A, Rl-69
166
-------
PRELIMINARY CAPITAL - DETAIL SHEET
PROJECT 3 STAGE SELECTIVE ABSORPTION PLUS GLAUS SULFUR RECOVERY WITH SCOT
TAIL GAS TREATMENT ' — 1
PHASE
PRE— 1
SECTION NUMBER i NAME
FACTORS
ITEM
500
BASE
CASE BY
Case B
DATE
SHEtT
3 OF 4
9212
SECTOR NUMBER i NAME
X ESC. X CAPACITY X QUANT. X [MATERIAL + CONDITION + COMPLEXITY!
NAME OF FACILITY
SCOT tail Gas
Treatment System
Trace & Cover
As Required
Allowance (20%) -.
Royalty (Per Dow Chemical U£
Initial Chemical Charge
QUANT.
Sys.
A)
198C
CONSTRUCTION
MATERIAL
Installed C
DESCRIPTIOJN
3225 scfm, 0.035 mole % H2S
164 bb/day of sulfur
Total (All Equipment)
1
apital Cost
i
$ 358
$12,576
$ 2,515
300
160
$15,551
FORM 39650 PRINTED IN U.S.*, RIO
167
-------
PRELIMINARY CAPITAL - DETAIL SHEET
SHEET
4 OF 4
•PROJECT 3 STAGE SELECTIVE ABSORPTION PLUS GLAUS SULFUR RECOVERY WITH SCOT
TATI. GAR
SECTION NUMBER * NAME
CASE
Case B
SECTOR NUMBER * NAME
9212
BASE X ESC. X CAPACITY
XQUANT. X [MATERIAL
CONDITION + COMPLEXITY3
ITEM
NAME OF FACILITY
OUANT.
CONSTRUCTION
MATERIAL
DESCRIPTION
1. Raw Materials:
MDEA 13.14 tons/yr
2. By Products:
Sulfur 13,016 Tons/yr
3. To Water Treatment or Disposal:
"Sour" Water 133.4x10 gal/yr
"Sludge" Water 3x10 Ib/yr
5.3x10 KwH/yr
2873x100 Ib/yr
1.5x10 gal/yr
4. Utilities:
Elec. Power
Stesim, Net
Cool Water
5. Manpower: •
10 Operators
1 Sxipervisor
6. Maintenance (5%), Capital Rec. (20%), Misc. (4%)
20% x Capital $15,551,000
7. Net Annualized Cost . .
8. Total Installed Capital
High, Dec. 1980
Prolb., Dec. 1980
Low, Dec. 1980
$ 24,000/yr
$l,171,000/yr (Credit
No charge
No charge
$ 158,000/yr
7,183,000/yr
1,545,000/yr
$8,886,000/yr
$300,000/yr
180,000/yr
$480,000/yr
$4,510,000/yr
$12,729,000/yr
$0.069/Ton of Shale
$1.172/bbl of Oil
$20,216,000
15,551,000
11,974,000
FORM 39690 PRINTED IN U.S.A. RI-69
168
-------
EEGIN
LAST LINE
OF TEXT '
, APPENDIX C . ;
ONE-STAGE SELECTIVE ABSORPTION PLUS INDIRECT STRETFORD SULFUR RECOVERY
!
9-1/8"
{5 3/8"
1
_s
BOTTOM.1 C:-
OL~S:C =
EPA-.2S7 (Cin.)
(4-70)
PAGE ivi'^VBER
-------
CAPITAL - DETAIL SHEET
SHEET
1 OF 3
I STAGE SELECTIVE ABSORPTION PLUS INDIRECT STRF.TFORD SULFUR ^RECOVERY
1 ft ec RY 1
9212
EF NuwSES
FACTORS
ITEM
101
102
103
104
201
202
203
204
205
206
207
208
209
210
211
212
213
214
301
BASE X ESC. X CAPACITY
XQUANT. X [MATERIAL
Gas Coolers
Cooler Pumps
Exchangers
Trace & Cover
For Freeze Protection
Selective Absorber
Exchanger
Pump
Interchanger
Strip Column
Condenser
Column Pump
Reboil Pump
Strip Col. Reboiler
Compressor
Amine Recv. Reboil
Sludge Pump
Sludge Accumulator
Trace & Cover
For Freeze Protection
H»S Absorber
ACILITY
QUANT,
6
6
6
CONSTRUCTS
MATERIAL
304 St.St
304 St.St
304 St.St
CONDITION + COMPLEXITY!)
16
16
16
16
Steel
Steel
D.I.
Steel
Steel
Steel
D.I.
D.I.
Steel
Strd.
Steel
D.I.
Steel
Steel
302
Oxidizer Tower
1
1
Steel
DESCRIPTION
12'$x42' Bl-Bl, 10 PSJ,
18' Pack Sec. 2000. ft
3 1/2" Pall Rings '
1000, GPM, 100' HD, 40 HP
500O ft2, U Tube, 114' LG
All Items i
Subtotal (100 Series)
12'$x42' Bl. to.Bl ;10 Valve
Trays
57/ft2
150 GPM, 100' HD, 7 1/2 HP
287 ft2
I
12 • 3>x42' Bl-Bl 10 Plate
3000 ft2. Float, 14' LG
716 GPM, ISO1 HD, 50 HP
639 GPM, 50' HD, 15 HP
722 ft2
6200 ACFM, 2 PSI ;
2300 Gal., 154 ft
1/2 HP
10 Day, 150 gal. |
All Items -;
As Required ;
Subtotal (200 Series)
10*x28' Bl-Bl, 13' Pack, Sec
2200 ft -3 1/2" Pall Rings
30'$x24' W/Foam Trough
Air Ring
= MS(19
$4,028
$5,957
FORM 13650 PRINTED IN U.S.A. Rl-69
170
-------
PRELIMINARY CAPITAL . - DETAIL SHEET
SHEET
2 OF 3
PROJECT ; ~~~~ ; !JOB NUMBER
I STAGE SELECTIVE ABSORPTION PLUS INDIRECT STRETFORD SULFUR RECOVERY 9212
"PHASE CASE BY DATE EF NUMBER
PRE-1 C 2-4-80
SECTION NUMBER * NAME
FACTORS
ITEM
303
304
305
306
307
308
309
310
311
312
313
314
315
316
317
318
• «" .
BASE X ESC. X CAPACITY X
NAME OF FACILITY
Air Diffusion System
Sulfur Slurry Tk.
Tank Agitator
Slurry Pump
Rot. Vac. Filter
. . - -
Sulfur Melter
Sulfur Fwd. Pump
Sulfur Store Tk.
Agitator
Drier Sys
Make-up Mix. Tk.
Fwd. Pump
Lean Sol. Fd. Tk.
Purge Pump
Mix Tic. Agitator
Trace & Cover
For Freeze Protection
•
SECTOR NUMBER 1 NAME
QUANT. X
QUANT.
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
[MATERIAL +
CONSTRUCTION
MATERIAL
Steel
Steel Epoxy
Steel
FRP
Std.
Steel
Steel
Steel
Steel
Steel
Steel
St.Stl.
Steel Epoxy
FRP
Steel
CONDITION + COMPLEXITY] f
DESCRIPTION
7500 SCFM, Blowers, Elec.
Air Pipe, Bldg. ;
15000 Gal. 25 psi ;
Med., 15 HP ;
100 GPM, 40' HD, 3 HP
Filter, Vac Pump, Filtrate
Receiver & Pump, Conveyor,
Wash Tk. & Pump, Bldg.
i
1200 Gal., 25 Psi,-JKT. &
Internal Coil, Insulate
& Baffle '•
10 GPM, 1 HP, T&C ;
30,000 Gal. T&C, 1 wk
API s.g. = 2
Heavy, 5 HP
200 ft w/Heating, Air or
Vac. Sys., Feed & Disch.
Conveyor , Bldg .
27,000 Gal., 8 Hr Store
100 GPM, 100' HD, 7 1/2 HP
36,000 Gal. 5 min. , s.g. =1.2
100 GPM, 40; HD., 3 HP
Medium, 25 HP
All Items
As Required i
Subtotal (300 Series)
Total (All Equipment)
— MSi'19 >
,
$ 2,924
$12,909
FORM 39650 PRINTED M U.S-A. Rl-69
171
-------
PRELIMINARY CAPITAL - DETAIL SHEET
SHEET
3 OF 3
PHASE -
f
PRE-1
SECTION NUMBER * NAME
CASE
•c
BY
.
SECTOR NUMBER & NAfc
;TPOT?n ^TTT/FTTR p-pri'~nrc"DV !
DATl:
2-4-80 •
JOB NUMBER
9212
EF NUMBER
FACTORS
ITEM
BASE X ESC. X CAPACITY
X QUANT. X [MATERIAL
CONDITION + COMPLEXITY]
NAME OF FACILITY
QUANT.
CONSTRUCTION
MATERIAL
DESCRIPTION
Case Bl - Disposal of Stretford Purge Solution
Equipment as above (sheets 1-3)
Allowance (20%)
Royalty (10% of capital)
Initial Chemical Charge
1980 Installed Capital Cost
Total
= MS(19
$12,909
2,582
$15,491
1,549
160
$17,200
Case B2 - Recovery and Recycle of Stretford Purge Solution
Equipment as above (sheets 1-3)
Stretford Purge Stream Reductive Incineration System
Subtotal
Allowance (20%)
Total
Royalty (10% of capital)
Initial Chemical Charge
1980 Installed Capital Cost
$12,909.
1,150
$14,059
2,812
$16,871
1,687
160
$18,718
FOKM 3XSO PRINTED IM U.S.A. *!•«»
172
-------
PRELIMINARY CAPITAL - DETAIL SHEET
PROJECT
I STAGE SELECTIVE ABSORPTION PLUS INDIRECT STRETFORD SULFUR RECOVERY
SECTION NUMBER * NAME
Cl
BY
SECTOR NUMBER * NAME
Disposal of Purge Solution
1. Raw Materials:
Methyl Diethanolamine
ADA
EDTA
Vanadate
Sodium Carbonate
Iron
By Products
Sulfur 12,516 ton/yr
To Water Treatment or Disposal:
"Sour" Water
"Purge" Water
"Sludge" Water
"Filter Wash"
Utilities:
Electrical Power
150 Psig Steam
Water
132x10 gm/yr
9.65x10 Ib/yr
3x10 Ib/yr
5x10 gal/yr
7.8x10 KwH/yr
1525x10 kb/yr
24M GPM
-, — +.
ITEM
BASE X ESC. X
NAME OF
CAPACITY ^(JUANT. X [MATERIAL •*• CONDITION + CC
FACILITY
QUANT.
CONSTRUCT ON
MATERIAL
DESCRIPT
1MPLEXITY]
ION
==MS(19 )
$ 28,000/yr
954,000/yr
24,000/yr
181,000/yr
Neg/yr
3,000/yr
§l,190,000/yr
$l/130,000/yr (Credit
No charge
No charge
No charge
No charge
$ 234,000/yr
3,812,000/yr
1,261,000/yr
$5,307,000/yr
FORM 39CJO PRINTED IN U.S.*. RI-69
173
-------
fiBT.TiCTIVE ABSORPTION PLUS
'
CONDITION + COMPLEXITY]
BASE X ESC. X CAPACITY
Manpower
15 men x 2000 hr/yr x $15/yr
1 supervisor
$450,000/yr
180,000/yr
$630,000/yr
6. Maintenance (5%), Capital Rec. (20%), Misc. (4%)
29% x $17,200,000
$4,988,000/yr
Net Annualized Cost
$10,985,000/yr
$0.602/ton of shale
$1.011/bbl;of oil
8. Total Installed Capital
High, Dec. 1980
Probable, Dec. 1980
Low, Dec. 1980
$22,360,000
17,200,000
13,244,OOJO
FOFBl 39t50 PRINTED IN U.S.*. Rl-69
174
-------
PRELIMINARY CAPITAL - DETAIL SHEET
TAGE SELECTIVE ABSORPTION PLUS INDIRECT STRETFORD DESULFURI2ATION
SHEET
3 OF 4
JOB NUMBER
9212
CASE
C2
SECTION NUMBER 4 NAME
SECTOR NUMBER «• NAME
BASE X ESC. X CAPACITY
X QUANT. X [MATERIAL +
CONDITION + COMPLEXITY]
ITEM
NAME OF FACILITY
OUANT.
CONSTRUCTION
MATERIAL
DESCRIPTION
Raw Materials:
Methyl Diethanolamine
ADA
EDTA.
2. By Products
Sulfur 12,903 TPY x $90/T
3. Watertreat or Disposal
"Sour" Water
"Sludge" Water
"Filter Wash"
Utilities:
Elec:. Power
Steam
Water
Gas
133.4x10 gal/yr
3x10^ Ib/yr
5x10 , gal/yr
9.2x10 KwH/yr
1540x10 Ib/yr
25M GPM
15M Btu/hr
Manpower
19 men x 2000 hr/yr x $15/hr
1 supervisor
Maintenance (5%), Capital Rec. (20%), Misc. (4%)
29% x $18,718,000
$ 28,000/yr
954,000/yr
24,000/yr
$l,006,000/yr
$l,160,000/yr (Credit
No charge '
No charge '
No charge
$ 276,000/yr
3,850,000/yr
1,314,000/yr
328,000>yr
$5,768,000/yr
$570,000/yr
180,000/yr
$750,000/yr
$5,428,000/yr
FORM 396SO PRINTED IN U.S.JL Rl-69
175
-------
CAPITAL - DETAIL SHEET
SHEET
4 OF 4
C2
JOB NUMBER
9212
EF NUMBER
ECTION NUMBER «. NAME
BASE X1SC. X CAPACITY
I SECTOR NUMBER* NAME
XQUANT. X [MATERIAL
CONDITION + COMPLEXITY]
DESCRIPTION
7. Net Annualized Cost
8. Total installed Capital:
High, Dec. 1980
Probable, Dec. 1980
Low, Dec. 1980
$ll,792,000/yr
$0.646/ton oif shale
$1.086/bbl of oil
$24,333,000
$18,718,000
$14,413,000
= MS(19
CONSTRUCTION
QUANT. u MATERIAL
FORM 3M50 PRINTED IN U.S.A. Rl-69
-------
G'JiL-[:
BEGIN
LAST LINE
Or TEXT
APPENDIX D
I
-1-
JDIAMOX PROCESS PLUS CLAUS SULFUR RECOVERY WITH
| BSRP TAIL GAS TREATMENT
I
9-1/8"
I_
BOTTOM C
ir.'AGE A:-;.
OUTS'DE
':-.\r--D !LLi
EPA-2S7 (Gin.)
(4-76)
PAGE NUW3LR
-------
PRELIMINARY CAPITAL - DETAIL SHEET
SHEET
1 OF 3
PROJECT JOB NUMBER
DIAMOX PROCESS PLUS GLAUS SULFUR RECOVERY WITH -BSRP TAIL GAS TREATMENT i 9212
PHASE CASE BY DATE i EF NUMBER
Case D
SECTION NUMBER 1 NAME
FACTORS
ITEM
101
102
103
104
204
202
203
-. .
204
205
206
207
208
. .
209
210
BASE X ESC. X CAPACITY !
NAME OF FACILITY
Gas Coolers
Cooler Circ. Pumps
Exchangers
Trace & Cover
For Freeze Protection
Absorbers
Absorber Pumps
• •• - » •
Stripper Feed Pumps
Stripper Bottoms Pumps
Acid Gas Strippers
Heat Interchanger
Cooling Water Exchanger
Refrig. Exchanger
Absorber Liquor
Purge Pump
Absorber Liquor Storage
SECTOR NUMBER & NAME 1
(QUANT. )
QUANT.
6
.....
6
6
6.....
30
-•
6
4
4
4
4
4
4
1
: [MATERIAL
CONSTRUCTION
MATERIAL
304 St.Stl.
304 St.Stl.
304 St.Stl.
Steel
D.I.
D.I.
• - -
D.I.
Steel
Steel
Steel
Steel
CI
Steel
CONDITION + COMPLEXITY]
DESCRIPTION
12'*x42' Bl to Bl, 10 psi
18' packed section, i
2000 ft -3 1/2"$ Pall Rings
1000 gpm, 100" Hd, 40 HP
5000 ft2, U Tube, 14' LG Tubes
All Items
12'$x72' High w/101: skirt,
6 sets spray nozzles deliver
3,000 gpm; 6 sets bulkheads
and overflow wiers for liquor
collection and visors for
gas flow. :
3000 gpm pumps to deliver
liquor to spray nozzles at
40 psig pressure
3000 gpm pumps to deliver
liquor to strippers at
30 psig pressure
4500 gpm pumps @ 30 jpsig
pressure ;
12'*x7 trays @ 24" spacing
18,000 sq.ft. multipass,
cross flow ;
18,000 sq.ft. multipass,
cross flow
2200 sq. ft., 40°F chilled
water — -
180 gpm @ 30 ft. head
!
540,000 gal., 48'4>x40'h
= MS(i9 )
-
$4,028
'
- . -. ...
FORM 39650 PRINTED IN U.S.A. Rl-69
178
-------
PBELIMfNARY CAPITAL - DETAIL SHEET
.iHEET
2 OF 3
DIAMOX PROCESS PLUS GLAUS SULFUR RECOVERY V7ITH BSRP TAIL GAS TREATMENT 9212
PHASE" " CASE BY °ATE : EF NUMBER
Case D i
SECTION NUMBER & NAME
FACTORS
ITEM
211
212
213
300
400
BASE X ESC. X CAPACITY X
NAME OF FACILITY
Absorber Feed Pump
Chilled Water Refrig.
System
Trace & Cover for
Freeze Protection " :
3 Stage Claus Unit
Trace & Cover as Required
BSPR Tail Gas
Treatment Unit
Trace s Cover as Required
• Allowance
Royalty
Initial Chemical Charge (BS
I
SECTOR NUMBER i NAME
QUANT. X
QUANT.
6
1
Sys.
Sys.
(Pe
PR Uni
198
[MATERIAL •+
CONSTRUCTION
MATERIAL
CI
Std.
,.
(20%)
: R. M. Pars
:)
) Installed
CONDITION + COMPLEXITY]
DESCRIPTION
3000 gpm pumps @ 60 psig
head
1250 Tons, 0°C
!
All Items :
As Required :
Subtotal (200 Series)
10,045 scfm, 6 mole % H-S
35.2 TPD of- Sulfur : •-•- - -
11,137 scfm, 0.031 mole % H-S
427 Ib/day
Total (Equipment)
Total i
>ns Co.)
Capital Cost
!
i
1
$19,769
$ 1,647
$ 1,021
$26,465 .
5,293
$31,758
900
83
$32,741
FORM 39CSO PRINTED IN U.S.A. RI-69
179
-------
PRELIMINARY CAPITAL - DETAIL SHEET
OF
PROJECT •
DIAMOX ABSORPTION PLUS CLAUS SULFUR RECOVERY WITH BSRP TAIL GAS TREATMENT
JOB NUMBER
9212
Case D
EF NUMBER
SECTION NUMBER fNAME
SECTOR NUMBER & NAME
FACTORS
BASE X ESC. X CAPACITY
XQUANT. X (MATERIAL
COND1TION + COMPLEXITY}
— MS(19 )
ITEM
NAME OF FACILITY
QUANT.
CONSTRUCTION
MATERIAL
DESCRIPTION
1. Raw Materials: (Parsons' Info)
Catalysts, Chemicals, Initial Charge
Consumption $230/day x 365
2. By Products
Sulfur 12830 TPY
3. Water Treatment or Disposal:
148 x 106 gal/yr
4. Utilities:
Electric Power
Steam, Net
Cool HO
53.9x10" KwH/yr
4480x101' Ib/yr
1.95x10 gal/yr
5. Manpower:
10 Operators
1 Supervisor
6. Maintenance (5%), Capital Rec. (20%), Misc. (4%)
29% x capital $32,741,000
7. Net Annualized Cost
8. Total Installed Capital
High, Dec. 1980
Prob., Dec. 1980
Low, Dec. 1980
$82,680
$84,000/yr
$l,150,000/yr (Credit)
No charge
$ 1,618,000/yr
$ll,530,000/yr
$ 1,950,000/yr
$15,098,000/yr
$300,000/yr
180,000/yr
$480,000/yr
$9,945,000
$24,007,000/yr
$1.315/ton of shale
$2.210/bbl of oil
$42,563,000
$32,741,000
$25,211,000
FORM 39650 PRINTED l« U.S.A. B1-C9
180
------- |