United States Industrial Environmental Research EPA-600/7-79-169
Environmental Protection Laboratory July 1979
Agency Research Triangle Park NC 27711
Hot Gas Cleanup Process
Interagency
Energy/Environment
R&D Program Report
-------
RESEARCH REPORTING SERIES
Research reports of the Office of Research and Development, U.S. Environmental
Protection Agency, have been grouped into nine series. These nine broad cate-
gories were established to facilitate further development and application of en-
vironmental technology. Elimination of traditional grouping was consciously
planned to foster technology transfer and a maximum interface in related fields.
The nine series are:
1. Environmental Health Effects Research
2. Environmental Protection Technology
3. Ecological Research
4. Environmental Monitoring
5. Socioeconomic Environmental Studies
6. Scientific and Technical Assessment Reports (STAR)
7. Interagency Energy-Environment Research and Development
8. "Special" Reports
9. Miscellaneous Reports
This report has been assigned to the INTERAGENCY ENERGY-ENVIRONMENT
RESEARCH AND DEVELOPMENT series. Reports in this series result from the
effort funded under the 17-agency Federal Energy/Environment Research and
Development Program. These studies relate to EPA's mission to protect the public
health and welfare from adverse effects of pollutants associated with energy sys-
tems. The goal of the Program is to assure the rapid development of domestic
energy supplies in an environmentally-compatible manner by providing the nec-
essary environmental data and control technology. Investigations include analy-
ses of the transport of energy-related pollutants and their health and ecological
effects; assessments of, and development of, control technologies for energy
systems; and integrated assessments of a wide range of energy-related environ-
mental issues.
EPA REVIEW NOTICE
This report has been reviewed by the participating Federal Agencies, and approved
for publication. Approval does not signify that the contents necessarily reflect
the views and policies of the Government, nor does mention of trade names or
commercial products constitute endorsement or recommendation for use.
This document is available to the public through the National Technical Informa-
tion Service, Springfield, Virginia 22161.
-------
EPA-600/7-79-169
July 1979
Hot Gas Cleanup Process
by
A. Bekir Onursal
Dynalectron Corporation/Applied Research Division
6410 Rockledge Drive
Bethesda, Maryland 20034
Contract No. 68-02-2601
Program Element No. EHE623A
EPA Project Officer: Robert A. McAllister
Industrial Environmental Research Laboratory
Office of Energy, Minerals, and Industry
Research Triangle Park, NC 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, DC 20460
-------
ABSTRACT
Twenty-two hot gas cleanup (HGC) processes for desulfurizing reducing
gases at temperatures above i»30°C were identified and generically classified
according to absorbent type into groups employing solid, molten salt, and molten
metal absorbents. Each process is described in terms of its status, chemistry,
operating characteristics, problems and uncertainties. The applicability of nine
selected HGC processes to a variety of coal gasification systems is assessed for
several end uses for the product gases. The advantages and disadvantages of HGC
are evaluated relative to conventional low temperature cleanup systems with respect
to thermal efficiency, the presence and/or emissions of tars, particulates and
NOX, and corrosion. Economic comparisons between HGC and low temperature desul-
furization are also presented.
ii
-------
Contents
Abstract ii
Figures v
Tables vii
Acknowledgement ix
1. INTRODUCTION AND SUMMARY 1
1.1 Introduction 1
1.2 Summary 1
2. CONCLUSIONS AND RECOMMENDATIONS 3
2.1 Conclusions 3
2.2 Recommendations 6
3. STATUS OF HGC PROCESSES 7
3.1 HGC Processes Using Solid Absorbents 7
3.1.1 HGC Processes Using Calcium Based Absorbents 8
3.1.1.1 Air Products Process 10
3.1.1.2 Atlantic Refining Process -I 18
3.1.1.3 CCNY Process 20
3.1.1.4 CONOCO Process 25
3.1.1.5 U.S. Steel Process 31
3.1.2 HGC Processes Using Iron Based Absorbents 32
3.1.2.1 Appleby-Frodingham Process 34
3.1.2.2 Sabcock and Wilcox Process 33
3.1.2.3 BatteIte Columbus Process 46
3.1.2.4 IMMR Process 47
3.1.2.5 MERC Process 51
3.1-3 HGC Processes Using Copper Based Absorbents 57
3.1.3.1 Atlantic Refining Process -I 59
3.1.3-2 Esso Process 60
3.1.3.3 Johns Hopkins Process 64
3.1.3.4 Kennecott Process 67
3.1.4 HGC Processes Using Zinc Based Absorbents 68
3.1.4.1 Catalysts and Chemicals Process 70
3.1.4.2 IFP Process 73
3-1.5 HGC Processes Using Other Solid Absorbents 74
3.1.5.1 Exxon Process 7&
3.1.5.2 Foster Wheeler Process 78
ill
-------
Contents (cont'd)
3.2 HGC Processes Using Molten Salt Absorbents 30
3.2.1 Battelle Northwest Process 80
3.2.2 HRI Process 90
3-2.3 Pullman Process 93
3.3 HGC Processes Using Molten Metal Absorbents 95
3.3.1 IGT-Meissner Process 95
4. APPLICABILITY OF HGC TO COAL GASIFICATION SYSTEMS 100
4.1 Gasifiers 100
4.1.1 Fixed Bed Gasifiers 102
4.1.2 Fluidized Bed Gasifiers 103
4.1.3 Entrained Bed Gasifiers 104
4.2 Potential End Uses For Low and Intermediate Btu Gases 105
4.2.1 Industrial Applications 105
4.2.1.1 direct Process Heat 105
4.2.1.2 Industrial Boilers 106
4.2.1.3 Gas Turbines 107
4.2.1.4 Other Industrial Applications 107
4.2.2 Electric Utilities 108
4.2.2.1 Base and Intermediate Load Power Plants 108
4.2.2.2 Peak Load Power Plants 109
4.2.2.3 Combined Cycle Power Plants 109
4.3 Gasifier-End Use Compatibility 110
4.4 Applicability of HGC Processes for Selected
Gasifier End-Use Pairs 112
5. ADVANTAGES AND DISADVANTAGES OF HGC 122
5-1 Thermal Efficiency 122
5.2 Tars 124
5.3 Participates 124
5.4 MOX 126
5-5 Corrosion 127
5-6 Economic Analysis 128
References 133
Appendix ll»l
A. Personal Communications 141
iv
-------
FIGURES
Number Page
1 Equilibrium Relationship for the CaO-CaCo^-CO^ System 12
2 Equilibrium Constants for H£S Reacting with CaO and CaCCK 13
3 Multicycle Absorption Data 15
k Flow Diagram of the Air Products Process '6
5 Cross Section of Panel Bed Filters 22
6a Gas Cleaning 23
6b Puffback Cleaning of Panel Bed Filter 23
7 CCNY Process 24
8 Equilibrium Constant for the Reaction of H2S with CaCO^ 27
9 CONOCO Process 2S
10 Appleby-Frodingham Process 36
II Equilibrium Constant for h^S Absorption by Iron Oxide 39
12 Large Scale Test Configuration 42
13 H2$ Removal Efficiency for Large Scale Experiments *»3
]k Sulfur Concentration Versus Temperature ^5
15 Sulfur Removal ^5
16 B & W Regenerative Desulfurizer 46
17 Equilibrium Constants for Reduction and Desulfurization
Reactions ^9
-------
FIGURES (cont'd)
Number Page
18 Absorotion Capacity of Sintered Pellets of 75%
Fly Ash - 25% Fe203 54
19 Regeneration of Sulfided Sorbents 56
20 Equilibrium Constant for the Reaction:
2Cu + H2S - Cu2S + H2 61
21 Esso Process 63
22 Equilibrium Constants for Desulfurization with Copper-
Containing Absorbents 69
23 Equilibrium Data for the Reaction:
ZnO + H2S - ZnS + H20 72
24 Exxon Process 77
25 Equilibrium Constants for Ni-Containing Sorbents 79
26 Temperature Dependence of Equilibrium Constants for
Reactions Between H2S and Molten Carbonates 84
27 Batch-Mode Test Apparatus 86
28 Continuous-Mode Test Apparatus 87
29 Schematic Flow Diagram of Salt Bleed Recovery 89
30 HRI Process 92
31 Pullman Process 94
32 Equilibrium Constants for Desulfurization with Lead
Absorbents 96
33 IGT-Meissner Process 98
34 Temperature/Compos it ion Diagram for Lead-Lead Sulfide
System 99
35 Combined Cycle with Unfired Waste Heat Recovery System '09
vi
-------
TABLES
Number Page
1 HGC Processes A
2 HGC Processes Using Calcium Based Solid Absorbents 9
3 HGC Processes Usina Iron Based Solid Absorbents 33
k Typical Gas Analysis Across an Absorber 37
5 Equilibrium Constants for Reactions (15) and (16) *»8
6 HGC Processes Using Copper Based Solid Absorbents 58
7 HGC Processes Using Zinc Based Solid Absorbents 71
8 Fluidized Bed Absorption-Regeneration Experiment 7^
9 HGC Processes Using Other Solid Absorbents 75
10 HGC Processes Using Molten Salt Absorbents 81
11 Promising Low and Intermediate Btu Gasifiers 101
12 Gasifier - End Use Compatab i 1 i ty Ill
13 Characteristics of Selected Gasification Options 113
1^ Minimum Sulfur Removal for Gasifier/End Use Pairs 115
15 Assessment of HGC Process Applicability 116
16 ^2$ Product-Gas Concentrations Below Which No Clean-up
Is Required 118
17 HGC Processes Applicable for Selected Gasifier-End Use
Combinations 120
vii
-------
TABLES (cont'd)
Number Page
18 Summary of Estimated Thermal Efficiencfes 123
19 Summary of Estimated Capital Requirements for 1,000 WM
Gasification - Combined Cycle Power Plants . 130
20 Summary of Estimated Costs of Electricity for Gasification-
Combined Cycle Power Plants of 1,000 MW Capacity 131
21 Summary of Estimated Capital Requirements for the Puri-
fication System Components of Each Configuration
Investigated 132
vlii
-------
ACKNOWLEDGEMENT
The author is grateful to Mssrs. Chester A. Voqel and William J. Rhodes
of the Industrial Environmental Laboratory, EPA, for their continuing interest
and support.
ix
-------
SECTION I
INTRODUCTION AND SUMMARY
1.1 INTRODUCTION
Raw product gases from coal converters generally contain a variety of
impurities which must be removed either because of process constraints or environ-
mental regulations. Hydrogen sulfide, the major sulfur impurity in the raw
product gas, can be removed by many commercially available technologies at temper-
atures below 150° C. However, for applications such as combined cycle electric
power generation, cooling of the gas stream prior to combustion results in a
reduced overall thermal efficiency. This energy inefficient step can be eliminated
by removing hydrogen sulfide at elevated temperatures (e.g., above ^30° C), namely
by hot gas cleanup (HGC).
The general objectives of this study of hot gas cleaning processes,
carried out under EPA Contract 68-02-2601, are to identify and determine the
status of HGC processes in general, determine the applicability of these HGC
processes to coal gasification systems, and compare the advantages and disadvan-
tages of HGC generically with established low temperature cleanup processes.
1.2 SUMMARY
Section 2 of this report presents the conclusions derived from this
study and the recommendations for further study. The status of HGC processes is
described in detail in Section 3- The applicability of HGC processes to coal
gasification systems is assessed in Section k. Section 5 compares the relative
merits of HGC with respect to low temperature desulfurization.
Literature and patent searches were conducted to identify HGC processes
at all stages of development. The twenty-two identified HGC processes were then
classified in Section 3 according to the following type of absorbents: solids,
molten salts, and molten metals. The advantages and disadvantages of each generic
type of absorbent were discussed. The solid absorbents were then classified
according to the following types of active material employed in the sorbents:
calcium, iron, copper, zinc and other ions. Information gathered on each HGC
process was reviewed and employed to develop process descriptions and to describe
process chemistry, research and development, and process status. Uncertain!ties
and problems associated with each HGC process were also identified.
In Section 4, the applicability of HGC processes to various coal gasifi-
cation systems is assessed. First, the type of gasification processes and the
-------
possible end uses were studied and matched. Then, the applicability of representa-
tive HGC processes to selected gasifier-end use pairs was evaluated. These evalua-
tions were based on the thermodynamic equilibrium relationships for the hydrogen
sulfide removal reactions, using the environmental regulations and process require-
ments as restrictions on the sulfur contents of the purified gases.
The advantages and disadvantages of HGC are assessed generically relative
to low temperature cleanup systems in terms of thermal efficiency, the presence
and/or emissions of tars, particulates, and NOX, and corrosion. Economic compari-
sons between HGC and low temperature cleanup have previously been made by several
investigators, and are summarized in Section 5.
-------
SECTION 2
CONCLUSIONS AND RECOMMENDATIONS
2.1 CONCLUSIONS
As a result of literature and patent searches, twenty-two HGC processes
were identified. Table 1 summarizes these processes. Among them, only five are
presently undergoing further development, with the others being terminated for
various reasons. The IFF Process, one of the ongoing five processes, is being
developed in France and is proprietary. Among these twenty-two HGC processes, only
the Appleby-Frodingham Process has been employed on a commercial scale.
The following conclusions are drawn from this study:
• HGC processes are best suited for combined-cycle and on-site combus-
tion applications coupled with low or intermediate Btu-gasifiers. The
applicability of HGC processes for treating gases to be used as synthesis or
reducing agents depends on the specific industrial process under considera-
tion. The applicability of HGC processes for desulfurizing gases to be trans-
ported over long distances does not appear to be economically feasible.
• The applicability of nine HGC processes, each having a different
active material in the absorbent, to nine gasification options and two end
uses (namely, combined-cycle and on-site combustion) is presented in Table 17
of Section k. Typical gasification characteristics, environmental regulations,
process limitations, and theoretical thermodynamic relationships were consid-
ered in making the assessments that are summarized in Table 17- It can be
seen from Table 17 that the Kennecott and MERC processes (or HGC processes
using absorbents having Fe203 of Cu as active materials) are applicable
for desulfurizing gases at high temperatures for any pair of the selected
gasifier-end use pairs considered in this study.
• HGC processes provide greater overall thermal efficiencies than low
temperature desulfurization. The difference in thermal efficiencies increases
with increasing tar and steam contents in the raw product gas. In this
respect, the Lurgi gasifier, coupled with combined cycle application seems
especially attractive. For combined cycle applications which employ a gasifi-
cation system producing tars and steam, HGC also provides a distinct advantage
over low temperature desulfurization in terms of total capital requirements
and cost of power. The difference in capital requirements and costs decreases
for gasification systems producing gases which contain no tars and small
amounts of steam.
-------
TABLE i- use PKOCESSES
HCC PROCESSES
EMPLOYING
SOLiO ABSORBENTS
Calcium Boscd Absorbents
Air Products
Atlantic Refininj-XC
CCNY
CONOCO
U.S. Steel
1 ron Based Absorbents
Appl etty-f rod 1 ngham
Sabcock I Wilcox
Sattelle Columbus
!MMR
"ERC
Copper Based Absorbents
At'ancic Refining*;!
ESSO
John Hook I is
Kenrecott
STATUS
Terminated
Terminated
Terminated
Terminated
On-Going
Terminated
Terminated
Terminated
On-Going
On-Going
Terminated
Terminated
Terminated
Terminated
ABSORBENT
HgO.CaO
Ca(OH)2
Dolomite/
Fe203
MgO.CaC03
HgO.CaO
FejOj
FeO
Supported
Fe20j
Gasifier
Ash(Fez03)
Supported
Ft-}0*
•^
Cu/Pb/Zn
al umi no
si lieate
Supported
Cu
Supported
Cu/Cr/V
Cu/CuO
ABSORPTION
TEMP. (°C)
649-871
316-538
900
Above!) 6
360-i<27
1127-61.9
538-816
371-816
538-8)6
93-538
299-982
Up to 816
482-495
PRESSURE (Pa)
Ixlo5-2xlo5
7.0x105 -
I.SxIO6
1x105
1x105
IxtO5
lixlO5 -
9x105
1x105-
2x106
2xl06 -
6AS FLOW RATE
(normal m*/h)
79.3
17.7
85.0
t
O.llxlO6
170
0.06
1.22
212
METHOD OF
REGENERATION
HjO-COj.Calclnat Ion
NaOH solutlon-02
K20-C02/AI r
H20-C02
Proprietary
Air
Air
Air
Air
Air
steam/air
CujO
A!r
Air
REGENERATION
TEMPERATURE IOC)
"i27-593 ;61i9-927
below solution b.p.t.
Panel Bed
593-760
593-816
538-6<«9
593
It27-6ii9
538-816
371-538
3S9-738
816
COMMENTS
No experi-
mental work
Commercial
Experim'n-
tal work
s us pendcd
Large P'ant
Tests
Cu(OH}2
I.SxIO6
(continued)
-------
TA3LE I (eont'd)
HOC PROCESSES
EMPLOYING
2irc Based Absorbents
Catalysts t, Chemicals
IFJ>
Other Solid Absorbents
Exxon
Foster Wheeler
MOLTEN SALT ABSORBENTS
BatteLle Nonn«csc
HRI
Pu* Iman
WJLTEV ,»;TAL ABSOR3ENTS
ICT-Meissner
STATUS
Terminated
On-Going
Terminated
Terminated
Terminated
Terminated
Terminated
On- Go Ing
ABSCR3ENT
ZnO
Supported
ZnO
Supported
La20j
Supported
HI
Li, Na. K,
Ca Carbon-
ates
Supported
fli2tOj
Li, Ka, K
carbonates
Pb
ABSORPTION
TEMP. (°c)
150-650
J.00-600
149-927
538-738
600-910
307-1093
816-1093
327-1200
CAS FLOW RATE
PRESSURE (Pa) (normal mJ/h!
6x10* -
2.6xl06
Ixlfl5 -
6xl05
0.10
IxloS 127
l.ZxlO5 - 0.01
lixlO6
IxlO5 -
5x!05
rtETHOO OF REGENERATiON
REGENERATION TEMPERATURE (°C)
0} Containing Gas 600-500
Steam; Alr/Oj
Air 538-788
H20-C02 500-600
H20-C02 907-I09J
Hetal Blcarbonate-
C02
electrolysis Above 327
COMKENTS
Regenera-
tion not
studied
Proprie-
tary
Process
PO'J
Operated
Proprie-
tary
Process
-------
• HGC processes add complications and uncertain!ties concerning the
removal of participates at high temperatures, whereas the removal of particu-
lates at low temperatures can be considered as commercial technology.
• Processing of gases at high temperatures results in increased NO
emissions from the downstream end use systems. In the case of low temperature
desulfurization processes, most of the ammonia in the raw product gas is
removed by water quenching.
• Another disadvantage introduced by HGC is the potential corrosion of
metals by hydrogen sulfide at high temperatures.
2.2 RECOMMENDATIONS
Further research and development is recommended in the following areas:
• HGC Process R & D: None of the identified HGC processes presently
appear to be ready for commercialization. The problems and uncertainties
associated with HGC processes, as outlined in Section 3, need to be further
studied. HGC processes employing molten salt or molten metal absorbents are
complicated by serious corrosion and material handling problems. For this
reason, future research and development should be focused on those processes
employing solid absorbents. Among these HGC processes, those using iron and
copper based absorbents are most likely to be applicable for a broad range
of gasification systems.
• Particulates Removal/. There is presently no commercial particulate
removal equipment which could be operated at gasifier temperatures to reduce
the participate loadings in the fuel gas to levels required by environmental
regulations and/or process restrictions for the end uses under consideration.
Further research is needed in this area.
• NOX Control: There is no available technology which can remove ammonia
at high temperatures. However, the decomposition of ammonia into inert nitro-
gen and hydrogen might be a solution to the NOX control problem. Research in
this area should be emphasized.
• Corrosion: Research and development should be carried out to find
suitable materials of construction to withstand corrosion by hydrogen sulfides
and other chemical compounds such as alkali metal sulfides and chlorides.
-------
SECTION 3
STATUS OF HGC PROCESSES
A comprehensive literature and patent search was carried out to identify
HGC technologies at all stages of development. This search has been directed not
only at HGC technologies developed specifically for coal conversion, but also for
other related industrial applications such as coke-oven and refinery gases.
A computerized literature search was carried out at the Technical Informa-
tion Center of Oak Ridge National Laboratory. The data base for this search includes
Chemical Abstracts, Engineering Indices, U.S. and foreign patents and government
publications. A thorough U.S. patent search was also conducted separately. HGC
process developers and evaluators, HGC project officers at the U.S. Department of
Energy, and other EPA contractors were contacted for additional information. Further
references were obtained through the available publications.
Twenty two HGC processes were identified as a result of the information
gathering efforts discussed above. These processes can be classified according to
the type of absorbents, as follows:
• sol id absorbents,
• molten salt absorbents, and
• molten metal absorbents.
Each absorbent reacts with the hydrogen sulfide present in the raw product gas to
form sulfides, and thus purify the gas stream. The sulfided sorbents are usually
regenerated by various means such as air, oxygen, steam, carbon dioxide, or a
mixture of steam and carbon dioxide.
The status description presented herein contain comparisons of the advan-
tages and disadvantages of the processes using the different generic types of
absorbents, summarized process descriptions and related data, and detailed process
descriptions, wherever possible.
3.1 HGC PROCESSES USING SOLID ABSORBENTS
The HGC processes using solid absorbents can be further classified accord-
ing to the type of the active cation in the sorbent, as follows:
• calcium based absorbents,
-------
• iron based absorbents,
• copper based absorbents,
• zinc based absorbents, and
• other solid absorbents.
Both summarized and detailed information on these processes are given
in the following sections.
The processes using solid absorbents generally have the advantage of
incorporating process equipment and technologies (e.g., solids handling and dis-
posal) employed on a commercial basis in a variety of industries. On the other
hand, processes using molten absorbents can be considered to be more novel, and
the operational difficulties associated with these processes seem to be more compli-
cated than those using solid absorbents. For example, the corrosive nature of the
melt often requires the use of special materials of construction whereas in the
solid absorbent processes commercially available materials of construction are
usually adequate. Other process difficulties encountered with molten absorbents,
such as the solidification of the melt at cold spots in the system and the subse-
quent plugging of lines and equipment, do not occur in processes employing solid
absorbents. Also, certain of the HGC processes using solid absorbents have been
developed to a more advanced stage that those using molten salt absorbents. For
example, the Appleby-Frodingham Process was used commercially for a number of
years, and the Johns Hopkins Process was tested at commercial scale.
One of the most common problems associated with processes using solid
absorbents is the loss of the physical strength of the absorbent. This was a
major technical difficulty during the operation of the Appleby-Frodingham Process,
which employed iron oxide absorbents. To minimize this problem, supported sorbents
were developed. For example, in the MERC Process, which is a more recent process
using iron oxide as active material, flyash or silica is incorporated as a support
material to prevent the crumb 1 ing, carry over, and loss of sorbent material. Al-
though supported sorbents reduce the amount of sorbent make-up and, therefore, the
raw material costs, the preparation of such sorbents adds another process step
and cost item.
A loss of activity was also observed for some processes at increasing
numbers of absorption-regeneration cycles. Poor regenerabi1ity of the sulfided
sorbent in the Air Products Process was one of the reasons for the termination of
that process. For processes using oxygen or an oxygen containing gas for regenera-
tion, temperature control in the regenerator bed can become a significant technical
difficulty.
3.1.1 HGC Processes Using Calcium Based Absorbents
There are five HGC processes using calcium based solid absorbents. These
processes are the Air Products, Atlantic Refining-U , CCNY, CONOCO and U.S. Steel
Processes. All these processes have been terminated, with the exception of the U.S.
Steel Process whose regeneration step is proprietary. A summary of information for
these processes is presented below and in Table 2. Sections 3.1.1.1 through
-------
TABLE 2 - HGC PROCESSES USING CALCIUM BASED
SOLID ABSORBENTS
HGC
PROCESSES
Air Products
Atlant ic
Refining -U
CCNY
CONOCO
U.S. Steel
STATUS ABSORBENT
Terminated MgO.CaO
Terminated Ca(OH)
Terminated Dolomite/
Terminated MgO.CaCO
On-Going MgO.CaO
ABSORP-
TION
TEMP.
871
3f6-
538
900
Above
816
PRES-
SURE
(Pa)
IxlO5-
2x1 05
7.0x10;"-
4.3x10°
l.SxlO6
IxlO5
GAS
FLOW
RATE
(normal)
m3/h
79.3
17.7
85.0
METHOD OF
REGENERATION
H20-C02,
Calcination
NaOHsoln.
°2
H 0-CO,/
Air i
H20-C02
Proprietary
REGENERATION COMMENTS
TEMPERATURE
^27-593
6^9-927
below solution
boi 1 ing point
Panel No experi-
Bed mental work
593-
760
-------
3.1.1.5 contain detailed information on these processes.
• Air Products Process: The hydrogen sulfide from coal gases is removed
with fully-calcined dolomite (MgO.CaO) at temperatures between 649 and 871°C (1200
and 1600°F). The sulfided sorbent is regenerated by first treating with steam and
carbon dioxide at 427-593°C (900-1 100°F) and then calcining at 649-927°C (1200-
1700°F). Bench scale experiments were terminated in 1975 due to the poor regenera-
bllity of the sorbent, coke and tar deposition, and unfavorable process economics.
• Atlantic Refining Process -H : This process was developed to the
bench scale level to remove hydrogen sulfide from the effluent gaseous streams of a
hydrodesul furizat ion process or the first stage of a catalytic reforming process.
In this process, an alkaline earth metal hydroxide, such as calcium hydroxide, is
used as sorbent at temperatures from 316 to 53S°C (600 to 1000°F) and pressures
Between 7.0 x 10$ and 4.8 x 10° Pa (100 and 700 psi). The spent sorbent is regenera-
ted by reacting it with an alkali metal hydroxide, such as NaOH, or sulfide in the
presence of an oxygen containing gas. Research work was terminated in the mid
1950's.
• CCNY Process: A mathematical model was developed at the City College
of New York to evaluate hydrogen sulfide removal using an absorbent, such as dolomite
or iron oxide, in a panel bed filter, which was originally designed and tested for
part icul ate removal from gases. Bench scale experiments were also carried out to
study the rate of absorption of sulfur species with half and fully-calcined dolomite.
However, the panel bed was never tested for sulfur removal, and there are presently
no plans for further research in this area.
• CONOCO Process: Raw producer gas is desulfurized with hal f -calcined
dolomite (MgO. CaCOi) at about 900°C (1650°F) and elevated pressures (about 1.5 x
10" Pa; 15 atm) . The spent dolomite is regenerated to calcium carbonate with carbon
dioxide and steam at 700OC (1300°F). The regenerator outlet gas containing hydrogen
sulfide is contacted with dilute hydrosul furous acid to produce elemental sulfur
at 154°C (310°F) . Laboratory and continuous bench scale experiments were performed
to test the process feasibility. Research work was terminated in 1977 due to lack
' funding.
• U.S. Steel Process: Fully-calcined dolomite (CaO.MgO) is used to de-
furlze coal gas at temperatures above 8l6°C (1500°F) and at atmospheric pressure.
i regeneration method is proprietary. Work is ongoing at U.S. Steel at the
ot plant level.
3.1.1.1 Air Products Process: Air Products and Chemicals, Inc. developed and
tested a process to desulfurize coal gases with fully-calcined limestone (1,2).
The research program.which started in October 1972 and ended in January 1975, was
funded by the Energy Research and Development Administration.
Air Products Process contains three major process steps:
1. Calcination, at temperatures between 649 and 927°C (1200 and 1700°F)
CaCO- - CaO + CO.- (endothermic) .
10
-------
2. Absorption, at temperatures between 6*»9 and 871°C (1200 and 1600°F)
CaO + H S = CaS + HO (exothermic).
3. Regeneration, at temperatures between k2J and 593°C (800 and 1100 F)
CaS + CO + HO = CaCO_ + H S (exothermic).
The calcination temperature depends solely on the carbon dioxide partial
pressure. The equilibrium relationship for the CaO - CaC03 - C02 system is shown
in Figure 1. A high calcination temperature is required when the carbon dioxide
partial pressure is high.
The amount of h^S removed from the producer gas depends on the feed gas
C02 and h^O contents and the absorption temperature and pressure. The absorp-
tion temperature should be kept above the CaCOo calcination temperature ( a function
of the C0£ partial pressure) to prevent the formation of CaCOo. Figure 2 shows the
dependency of the equilibrium constant on temperature for the reaction between CaO
and H2S. Lower absorption temperatures enhance increased H2$ removal efficiencies.
The H2S removal is adversely affected by the presence of steam in the inlet gas
stream, although it is not influenced by a change in the pressure.
Regeneration of the sulfided sorbent is achieved with steam and carbon
dioxide. The regeneration product is CaC03, which is then calcined to obtain
CaO for use in another absorption cycle. The variation of the equilibrium constant
with temperature for the reaction between CaS, C02 and H20 is included in Figure 2.
The regeneration reaction is favored by low temperature, high pressure, and high
C02 and ^0 concentrations in the regeneration gas. These absorption and regeneration
reactions were previously studied by Squires, et al . (3,*0-
Initial bench-scale experiments were performed in a 100 cc fixed-bed
quartz reactor at 1 atm. total pressure. No difficulty was observed during cal-
cination of limestones. Absorption tests were performed by passing a tar-free and
part i culate-f ree hot synthetic producer gas (containing 0.6 percent H2S) through
the bed. High calcium (98 percent CaCO^) , magnesian (83 percent CaC03) , siliceous
(6^* percent CaC03) , and dolomitic (5*» percent CaC03) limestones were found to be
active for absorption of H2S; however, no limestone appeared to be clearly superior
to the others. Up to 75 percent calcium utilization was observed at 10 percent
breakthrough (the point at which the exit H2S concentration is 10 percent of the
inlet gas H2$ concentration). Typical operating conditions were 8l6°C (1500°F),
1x105 Pa (1 atm) pressure, and a gas flow rate of 6,000 normal m3/h per cubic meter
(6000 SCFH per cubic foot) of limestone. The 10 percent breathrough point was
observed after three or four hours of continuous on-stream absorption.
Regeneration was first attempted at low temperatures (66-93°C or 150-
200°F) , in aqueous phase, with water and carbon dioxide. However, these experiments
were abandoned due to the pressure drop build-up resulting from bed agglomeration.
Attention was then directed to vapor phase regeneration, varying the temperatures
between 1**9 and 621°C (300 and 1150°F), since higher temperatures favor high reaction
11
-------
(lo')io6
(10°).O5
01
UJ
cc
CL.
h-
CoO
SOLID PHASE
(10~2)103
-------
10.000
10000
CoO SOLID PHASE
CoS SOLID PHASE
100.0
S
io
<:
o
S
w.o
CaC03 SOLID PHASE
CaS SOLID PHASE
[^PARTIAL PRESSURE. ATM.
(1 atm i.., I05 Pa)
1.0
538 593 6**9 704 760 8)6 871 927
(1000) (1100) (1200) (1300) (HOG) (1500) (1600) (1700)
TEMPERATURE, °C (°F)
FIGURE 2. EQUILIBRIUM CONSTANTS FOR H2S REACTING WITH CaO and CaCO?
-------
rates and lower temperatures favor high regeneration efficiencies. The optimum
temperature range was found to be between 1»82 and 566°C (900 and 1050°F). At
518°C (96*»°F) , the regeneration rate was 18 percent in three hours and 23 percent
In five hours. In this run, the regeneration inlet gas contained 58 percent C02
and 18 percent reducing gas on a dry basis, with a steam to carbon dioxide mole
ratio of 1:9. The corresponding instantaneous \\2$ concentrations, in the dry
effluent gas, were 5-5 percent and 1.2 percent. The effects of other process varia-
bles on the regeneration level were also studied. It was observed that higher cal-
cination temperatures yielded higher regeneration levels. The presence of reducing
gases (CO and H2) during regeneration prevented calcium sulfate formation; however,
increasing the reducing gas concentration beyond 18 percent (dry gas basis) did
not have an incremental effect.
Multicycle absorption/regeneration experiments were carried out up to
three cycles to determine cycle-to-cycle absorption performance and regenerabi 1 i ty .
About one- third reduction in the regeneration level was observed from cycle to
cycle. This poor regenerab i 1 i ty of the sorbent, exhibited by both calcium and
dolomitic limestones, reduced the subsequent absorption capacity corresponding to
a given h^S breakthrough level.
Further bench-scale tests (2) were conducted above atmospheric pressure
with a tar and part iculate-free synthetic producer gas in an Inconel 601 reactor.
The results of the absorption tests were similar at both 105 and 2xl05Pa pressures.
Figure 3 shows the multicycle absorption data for a typical experiment. The regen-
eration experiments were carried out at 2x10? and lxTOOpa (2 and 10 atm) . The
total amount of regeneration accomplished was higher for the first cycle regeneration
than the second cycle regeneration. It was also higher for lower space velocities,
higher steam rates and higher pressures. Coke formation, from the tar-free syn-
thetic producer gas, presented a serious problem as it caused plugging in the
transfer lines and limestone bed which, in turn, forced frequent shutdowns.
Experiments were also performed with a real producer gas at the coal
gasification pilot plant operated by the Bureau of Mines in Morgantown, West Virginia.
The absorber was 8" in diameter and was loaded to a depth of 0.81m (3211) with
approximately k$.k kg (100 Ib) of limestone. Producer gas was fed to the reactor at
a rate of about 793 normal m3/h (2800 SCFH) for each run. Severe pressure drop
Increases, accompanied by corresponding decreases in flow rates, were experienced as
a result of coke formation from tars contained In the producer gas, which resulted
in termination of the experiments.
A conceptualized flow diagram prepared by the process developer is shown
in Figure 4. Producer gas generated by a coal gasifier is first cleaned of parti-
culates and then introduced to a fixed-bed limestone absorber at a temperature of
8?1°C (1600°F) and a pressure about atmospheric.
There are two absorbers operating in series, with the second functioning
as a guardcase. At 10 percent breakthrough, the first absorber is taken off stream
and cooled for regeneration, with the second absorber becoming the primary absorp-
tion vessel. The freshly regenerated absorber then becomes the second vessel. The
purified gas is combusted in a power plant to produce steam for electric power
generation.
-------
700
NOMINAL OPERATING CONDITIONS
760*C, 6100 PPM H->S IN FEED.
2x105 Pa.6000 (&SV
0 100 200 300 400 500 600 700 800
TIME ON STREAM. MINUTES
tOO 200 3OO 40 500 600 700 8OO
TIME ONSTREAM. MINUTES
FIGURE 3. MULTICYCLE ABSORPTION DATA
15
-------
RAW GAS
ABSORPTION
XI
in
3
I
a
HEATING UP
£ CALCINING
COOL DOWN &
REGENERATION
DUMPING LOADING &
STANDBY
j
AIR
HEATER
T
' SPENTLIMES TONE DISPOSAL
FVELMSJCW. ?(g
"orrT
GAS
CKGAS
DEOXt\
DAWN
POWER
PLANT
DISPOSAL
REACTORS
REGENERATION
GAS BLOWER
WATER* GAS
COMPRESSOR
I SULFUR
T
FIGURE 4. FLOW DIAGFAf^ OF THE AIR PRODUCTS PROCESS
-------
The fully-sulifded bed of absorbent is cooled to 566°C (1050°F) at a
pressure of approximately ^.8x10-* Pa (70 psia) and regenerated; the regeneration
gas is used for cooling. Four vessels in series are used for simultaneous cooling
and regeneration. Because the h^S concentration in the regeneration off-gas is
too low (about 1.8 percent on a wet basis and ^.8 percent on a dry basis), it
cannot be processed directly in a Claus plant. For this reason, after cooling and
removal of water, this gas is compressed to a pressure of 1x10° to 2x10° Pa (150
to 300 psia) and directed to the selective CQ.2 and h^S recovery section. The h^S
rich stream from this section is sent to a Claus plant, while the CQ-2 rich stream
is added to the power plant stack gas and used for regenerating the sulfided lime-
stone. The heat for calcination is provided from the combustion of clean fuel gas
with air, with the calcined limestone being employed for absorption. The useful
life of the limestone is ended after five absorption and four regeneration cycles.
The sulfided limestone is then soaked with water and treated with CO- enriched-
gas at 100°C (212°F) according to the Chance reaction:
CaS(s) +C02(g) +H20W) =CaC03(s) +H2S(g)>
The limestone slurry is pumped to a settling pond and the f^S containing gas is
directed to the selective C02 and I^S recovery unit.
An economic analysis of the process performed by the process developer
is presented below. The size of the desulfurization process is based on an 800 MW
power plant requiring 1.3x10" normal m3/h (^.5.5x10^ SCFH) of producer gas ^.8 x
106J/normal m3, i.e. (130 Btu/SCF; or 2.65x105 kg/h , i.e., 292 tons/h , of coal
at 2.88x!07j/kg, i.e. 12,^00 Btu/lb).
Process Economics (Excluding Power Plant and Coal Gasifier)
Capital Cost 60 $/kw
Operating Expense 0.88^/kg of coal
or 0.038
-------
gas, coking was caused by tars and heavy hydrocarbons. In addition, coke could
arise from carbon monoxide according to the following reaction:
2 CO = C + C02
(This reaction, however, is reversed at high temperatures and lower pressures.)
In the experiments, coke formation caused plugging of the bed, which resulted in
high pressure drops and equipment shut-downs. This problem could not be prevented
during the process study. The process developer indicates that an oxidative pro-
cess is an effective way to remove the coke. However, this method is damaging
to the sulfided limestone since it generates sulfates that form low melting eute-
citics with oxides and carbonates, causing sintering, loss of porosity, and loss
of activity.
Corrosion of some metallic parts of the equipment was also observed
during the experimental program. The reactor inlet line corroded at the point
where cold and pure h^S bottled gas mixed turbulently with the hot gases. Also,
the inlet gas distribution plate of the reactor was plugged in a few instances,
causing the cancellation of two experiments. An analysis of the residue of the
plugging substance revealed that it was comprised chiefly of iron and nickel sulfides,
originating from the corrosion. Since iron and nickel are more susceptible to
H2S attack than chrominum, the inlet lines and the gas distribution plate were
changed from Inconel 601 to SS 316. The process developer indicates that materials
of construction which can adequately withstand the corrosive atmosphere and condi-
tions in producer gas desulfurization have not been developed, and may require the
use of exotic alloys.
The Air Products Process consumes a large amount of energy. The process
developer estimates that the total energy input to the system is equal to 9.5 per-
cent of the clean producer gas heating value. This input is provided by electricity,
natural gas and steam.
Finally, the Air Product Process generates a large quantity of solids for
disposal. The process developer estimates that about 10 kg of limestone must be
discarded for every 100 kg of coal gasified.
3.1.1.2 Atlantic Refining Process -H : The Atlantic Refining Process -JJC inven-
ted by Leum, et al.(5), removes hydrogen sulfide from high temperature/high pressure
gaseous streams, either directly from a hydrodesulfurization process, or in the first
stage of a catalytic reforming process. In the case of the hydrodesulfurization
effluent, when the desulfurized stream is cooled to condense the liquefiable con-
stituents, a liquid phase (free of dissolved hydrogen sulfide) and a gaseous phase
(consisting of hydrogen sulfide-free effluent) would be produced. Alternatively,
the hydrogen sulfide-free effluent could be sent to further refining operations,
such as reforming or hydrocracking, without reducing either the temperature or pre-
sure of the effluent stream. In the case of a multiple stage catalytic reforming
process, such as one utilizing a platinum-containing catalyst, a hydrocarbon-hydro-
gen stream, free of hydrogen sulfide, can be produced without deviating from the
process conditions. Hot gas desulfurization provides savings in terms of energy and
18
-------
equipment costs.
The hydrogen sulfide containing gases are passed through a bed of alkaline
earth metal hydroxide, such as calcium hydroxide, at temperatures between 316° and
538°C and pressures between 7.0x10^ and 't.SxlO5 Pa (100 and 700 psi). The hydrogen
sulfide reacts with the alkaline earth metal hydroxide to produce the corresponding
alkaline earth metal sulfide and steam. If calcium hydroxide is used as a sorbent,
then the following reaction takes place:
Ca(OH)2 + H2S = CaS + 2H20
When the sorbent is saturated, the gaseous stream from the hydrodesulfur-
ization or catalytic reforming step is diverted to a second reactor containing a
bed of fresh alkaline earth metal hydroxide. The first bed of alkaline earth metal
sulfide is then regenerated in-place.
Regeneration is accomplished by treating the alkaline earth metal sulfide
(e.g., CaS) with an alkali metal hydroxide (e.g., NaOH, K.OH) or sulfide (e.g., Na2$,
«2$) in an aqueous solution, at concentrations varying between 1 percent and 20 per-
cent, while at the same time introducing an oxygen containing gas into the solution.
The reaction produces the alkali earth metal hydroxide and alkali metal sulfide
which, in the presence of oxygen, converts to water soluble alkali metal polysul-
fides. Based on the information given by Leum et al. (5), it is deduced that the
following regeneration reactions can take place. It is assumed that the alkali
metal is sodium and alkaline earth metal is calcium.
a) Regeneration with NaOH:
*» CaS + 2NaOH + 3 H20 + \ 02 —> 4 Ca(OH)2 + Na^
5 CaS + 2NaOH + k HZO + 2 QZ >5 Ca(OH)2 + Na^
b) Regeneration with Na2S:
3 CaS + Na2S + 3 H20 + |- QZ > 3 Ca(OH)2 + Na^
*4 CaS + Na2S + k H20 + 2 QZ > ^ Ca(OH)2 + Na2S
Regeneration can be carried out at any pressure below the boiling point of the aqueous
solution. However, regeneration at the absorption pressure eliminates the need for
depressuring and represurring the chamber containing the alkaline earth metal
compound. The reaction time is between k and 12 hours with/air, at rates between
10 and 300 standard liters per mole of alkali metal hydroxide. The insoluble alkaline
earth metal hydroxide pellets are separated from the solution and heated to desul-
furization temperature with high temperature steam or inert gas, before the next
absorption cycle is initiated.
A gas stream containing hydrocarbons, hydrogen and a small amount of
hydrogen sulfide was passed through a bed of calcium hydroxide cylindrical pellets
19
-------
of 48mm (3/16") diameter and 32mm (1/8") thickness at 482°C (900°F) until the
calcium hydroxide was completely converted to sulfide. Regeneration was achieved
by passing air through calcium sulfide pellets in sodium hydroxide solutions for
an hour. The amounts of sulfur removed in the regeneration were 34 percent, 66
percent, and 85 percent with I percent, 10 percent, and 20 percent sodium hydroxide
solutions, respectively. With 14 percent Na2$ solution, only 55 percent of the
sulfur was removed from the sulfided sorbent in two hours.
Further experiments were carried out to study successive desulfurization
and regeneration cycles. Absorption was carried out at A82°C (900°F) by passing a
gas containing hydrogen sulfide through a bed of calcium hydroxide pellets. Follow-
ing the absorption period, the resulting pellets of calcium sulfide were treated
with a 1 percent aqueous solution of sodium hydroxide at a temperature of 93°C
(200°F) and atmospheric pressure. Air was used as the oxidizing agent. Absorption
and regeneration efficiencies were maintained without loss during three cycles.
The regeneration efficiency was found to be independent of the polysulfide concen-
tration.
The Atlantic Refining Process -IT was terminated in the late 1950's,
and there are presently no plans for further process development. (A-l)
• Uncertainties and Problems: The limited available information on this
process was obtained from a patent (5), in which the experimental difficulties are
not mentioned.
Regeneration of the sulfided sorbent with an alkaline earth metal hydrox-
ide and an oxygen containing gas produces polysulfides. The concentration of poly-
sulfides increases with successive regeneration cycles. Although it was observed
that regeneration efficiency is independent of polysulfide concentration, no method
was proposed for the treatment of polysulfides to either recover sulfur or to dis-
pose of it in an environmentally acceptable manner.
3.1.1.3 CCNY Process; City College of New York developed a panel bed filter to
remove dust particles from a gas stream. The same device can, in principle, also be
employed to remove hydrogen sulfide from coal gases at elevated temperatures, if a
proper absorbent is used as the bed material. The panel bed filter has been tested
experimentally only for particulate removal. However, a mathematical model was
developed to study hot gas desulfurization in the panel bed filter.
Figure 17 illustrates the panel bed filter (6) which contains three col-
umns of louvers. At the left of Figure 5 "wishbone louvers" support gas-entry
surfaces of fine granular solid having sizes smaller than 20 mesh and larger than
100 mesh. This solid is supplied by gravity from an overhead bin to a tall, narrow
space between the wishbone louvers and the central column of the closely spaced
horizontal louvers. A coarse solid is held in place within a second tall, narrow
space, between the central louvers and the inclined louvers at the right of Figure 5.
When the panel bed filter is used for particulate removal, the dirty gas flowing
in the horizontal direction enters the bed through the spaces between the adjacent
wishbone louvers, and deposits a layer of dust at the entry surfaces. The clean
20
-------
gas leaves the other side of the panel bed filter (see Figure 6 a)- The dust
from the panel bed filter is removed by a reverse surge flow of a clean gas (puff-
back. A sharp puff of gas supplied to the clean face of the bed causes a mass
movement of the fine sol id toward the left. This causes the filter deposits to fall
away from the dirty face, together with a small amount of the fine bed solids (see
Figure 6 b). The central louvers are designed to prevent the coarse solids from
participating in the movement of the fine solids toward the left. Efficiencies
beyond 99.9 percent were observed in experiments for filtration of an aerosol of
1.1 micron particles through a sand bed.
The panel bed filter, with puffback cleaning, might also be used to ab-
sorb sulfur species from fuel gas or combustion products. A paper by Graff, et
al. (7) discusses the potential of panel bed filters for simultaneous removal of
fly-ash and sulfur dioxide with half-calcined dolomite at elevated temperatures.
Other papers (8, 9, 10) discuss the possible use of panel bed filters for hydro-
gen sulfide removal from fuel gases at high temperatures. Again, half-calcined
dolomite was considered. This process Is patented by Squires (11).
In the case of hydrogen sulfide removal from fuel gas, the gas would
flow horizontally across a bed of reactive solids. Each small space between the
adjacent wishbone louvers constitutes a small reactor with countercurrent flow of
granular solids and gas. The major advantages of panel bed filters for counter-
current contacting of gas and solids are predicted to be their compactness and
their capability to operate at exceptionally small pressure drops. For a fuel
gas having a face velocity of 0.15 m/s (30 ft/min), the panel bed can purify
183 m3 of gas per minute per square meter of ground area (600 ft3/min/ft^). The
ground area occupied by such a panel bed filter is comparable to that of a fluid-
ized bed operating at a superficial velocity beyond 3 m/s (10 ft/s). The rela-
tively low pressure drop across a panel bed filter is generally on the order of
2500 Pa (10 inches of water).
Figure 7 illustrates the CCNY Process (11). A combined flow of fuel
gas, produced by gasification of coal or residual oil, and calcium carbonate or
lime, is introduced into the panel bed filter. The fuel gas of reduced sulfur
content then leaves the other end of the panel bed filter. Solids spilled by
puffback pass to a separator, where coarse granular particles are separated from
the finer particles and returned to the panel bed filter via a hopper. Spent fine
particles containing calcium sulfide are reacted with water and carbon dioxide gas
to produce a stream rich in hydrogen sulfide. The regenerated calcium carbonate
particles are treated before being introduced into the hot fuel gas. If the fuel
gas is sufficiently hot and has low carbon dioxide partial pressure, so that cal-
cium carbonate is calcined when introduced in the gas, then the treatment consists
only of drying. However, if it is not sufficiently hot enough to calcine calcium
carbonate (above approximately 6A9°C, i.e. 1200°F), then the treatment includes
slaking as well as calcination. If the fuel gas is at a substantially elevated
pressure, greater than about l»x!05 Pa (k atm) , the sulfided particles are reacted
with steam and carbon dioxide at a temperature below about JOk°C (1300°F). Other
candidate solids for absorption of hydrogen sulfide from a fuel gas at elevated
pressure include a composite of iron oxide and fly-ash, fully-calcined dolomite,
21
-------
DIRTY
GAS
SPACE SPACE
FOR FOR
FINE COARSE
SOLID SOLID
CLEAN
GXIS
FIGURE 5. CROSS SECTION OF PANEL BED FILTER
22
-------
•" ".X "-"*-"•:'.-. -.
*•< i:
OFDIRT
6a.
GAS CLEANING
.~
»' « *4iiii>7 •"••.:;> .Vs S*.
* X.i 4' :• v/. •'.': •'•'•
W&fj
> !•:!;>.•.>;/
FIGURE 6b. PUFFBACK CLEANING OF PANFL
BED FILTF.t
SHARP PUFF
OF GAS
23
-------
HOT GAS
CONTAINING
HYDROGEN
SULFIDE
PARTICLES OF
FJNECALCIUM
CARBONATE
OR LIME OR
SLAKED LIME
PANEL
BED
SPENT
SOLIDS
SULFUR-FREE GAS
COARSE
GRANULAR
MEDIUM
SEPARATION
REGENERATION
-H20
DRYING/
CALCINATION/
SLAKING
GAS RICH
IN HYDROGEN
SULFIDE
CLAUS
PLANT
FIGURF 7.
CCNY PROCESS
2k
SULFUR
-------
and half-calcined dolomite (8).
Various design configurations for panel bed filters were studied through
computer simulations (6). The results show that granular solids used to filter
dust from gas might also advantageously serve as the active reagent for absorbing
sulfur from the gas at atmospheric pressure. However, Squires does not recommend
use of the panel bed filter for simultaneous sulfur and particulate removal in a
single panel bed, especially at high pressures (A-2) .
In addition to his involvement in the panel bed process, Squires has
also studied desul furizat ion of fuels with fully- or half-calcined dolomite.
A patent of Squires (12) describes a process whereby a fuel gas is desul furized
with fully- or half-calcined dolomite, and the sulfided dolomite is then reacted
with steam and carbon dioxide at an elevated temperature and pressure to yield
hydrogen sulfide and half-calcined dolomite. Hydrogen sulfide can then be con-
verted to sulfur in a Claus plant. The results of lab-scale kinetic studies of
sulfur absorption by dolomites are presented in several publications (13,1^,15,16).
The reasearch program on the panel bed filter for particulate removal has been
terminated at CCNY. Further development Is believed to be beyond the university
research level (A-2). There are no futher plans at CCNY to develop and test the
panel bed filter for sulfur removal (A- 3) •
• Uncertainties and Problems: Since no experimental testing was per-
formed with a panel bed for H2S removal, the operation of this device remains
highly uncertain for desul furizing hot gases. An obvious potential operational
problem involves the separation of sulfided absorbents from solid particles de-
posited on the bed. Other potential problems would include those observed under
other processes with similar solid absorbents, such as poor regenerabi 1 i ty and
decreasing activity of absorbent. Although panel beds as solid filters have
undergone considerable testing, they need to be experimentally proven for desul-
furizing fuel gases at elevated temperatures.
3.1.1.*t CONOCO Process; Development of the CONOCO Process (17,18) by the
Consilidated Coal Company was initiated in 1972 as an adaptation of the C02-Acceptor
Coal Gasification Process. This research project was funded by the Environmental
Protection Agency until it was terminated in 1976. There are no plans for further
process development due to lack of funding.
The CONOCO Process was initiated when it was concluded that there was no
real advantage in using the C02-Acceptor reaction together with a sulfur-acceptor
reaction in the gasifier, when low-Btu fuel gas was the desired product. Desul-
furization was then studied as a separate operation, outside the gasifier, by
passing the low-Btu gas through a fluidized bed of half-calcined dolomite (MgO.CaCO?)
The H/>S absorption reaction is:
+ H2S = MgO.CaS + H20 + C02.
The kinetics of this reaction were also studied by Squires (19).
25
-------
Figure 8 shows the variation of the equilibrium constant for the above
reaction with temperature. Since the reaction is endothermic, hydrogen sulfide
removal is favored at higher temperatures. On the other hand, the bed temperature
should be held below the calcination temperature of the CaCOo, which is dependent
on the carbon dioxide partial pressure in the low-Btu gas. The calcination reaction
is:
MgO.CaCO, = MgO.CaO + C02 .
The equilibrium relationship for the CaCQj-CaO-CC^ system is shown in Figure 1 of
Section 1 (Air Products Process). The regeneration reaction is the reverse of the
absorption reaction. It is favored by low temperature and high carbon dioxide and
steam partial pressure.
A schematic diagram of the CONOCO Process is shown in Figure 9. The
hot fuel gas from the gasifier is desulfurized in the HjS sorption bed, at 899OC
(1650°F) and 1.52x10° Pa (15 atm) , by the reaction discussed previously.
The sulfided acceptor is conveyed to the regenerator by continuously
recycling a stream of CO? and steam. The regenerator is maintained by 70A°C (1300°F)
and 1 .52x10^ Pa. The sulfided acceptor is then regenerated by the reverse of the
above desul furization reaction. The regenerated acceptor is returned by gravity to
the gas desul furizer. Spent acceptor from the regenerator is treated with water
and carbon dioxide in the Chance Reactors at about 93°C (200°F) and atmospheric
pressure before disposal. The Chance reaction is:
MgO.CaS + 2C02 + H20 - MgCO .CaCO. + H2$.
The H2S concentration in the regeneration off-gas is restricted by the equilibrium
conditions of the regeneration reaction. This gas, containing 3-6 percent h^S at
87 percent approach to equilibrium, is introduced into a liquid phase Claus Reactor
at about IS^C (310°F) and 1.52x10° Pa, where it is contacted with hydrosul furous
acid from an S02 absorption column. The liquid phase Claus reaction is:
2H2S + H2S03 = 3S + 3H20.
Liquid sulfur produced in the reaction is separated from the water. Part of the
product sulfur is burned to produce sulfur dioxide for absorption by water in the
S0£ absorption column. It should be noted that a conventional Claus plant is not
employed because the h^S concentration in the off-gas is too low.
Absorption and regeneration experiments were carried out in a bench-scale
apparatus, where regenerated solids were continusouly recycled to the absorber.
Bottled gases were used to simulate low-Btu producer gas, and the exit gas from the
absorber was recycled to reduce the gas consumption. Sorbent consisting of 28 x 35
and 35 x A8 Tyler mesh sizes was introduced into a 7-6 cm diameter, type 30A SS
absorber at a flow rate of 1.4-3.2 kg/h. The sorbent was fluidized with 8.5-17.7
normal m3/h (300-625 SCFH) low-Btu gas, producing a bed of 46 cm in height. The
operating temperature and pressure in the absorber were 871°C (1600°F) and 1.52x10^ pa
(15 atm), respectively. The sulfided sorbent was continuously drawn from the ab-
sorber and fed to a 5 cm diameter regenerator, where the bed wa« fluidized with
26
-------
(TJ
Q_
to
o
a:
C.Q
(103)I08
(102)107
(io])io6
(10°)105
CaC03 SOLID PHASE
(-)
CaS SOLID PHASE
PARTIAL PRESSURE, ATM.
(1 atm =r 105 Pa)
-. v - , 704 760 816 871 927
(1000) (1100) (1200) (1300) (HOO) (1500) (1600) (1700)
TEMPERATURE, °C (°F)
FIGURC 8. F.qUILIRP.IUC' CONSTANT FOR THf HHACTION OF H, 5> with CaCO
27
-------
CYCLONE
-o-
DE SULFUR IZED GAS^
G4S
OESULFUR/ZER
GASIFIER GAS
N)
00
CYCLONE
C0
MAKE-UP
,i
rLLU/Yt W.
•Ctua
REGENERATOR
3
RECYCLE GAS
H20
CHANCE
REACTOR
CO-
SLURRY
TO POND
LIQUID
PHASE
CLAUS
REACTOR
cup
H20
H2S
S02
H20
±
so->
ABSORPTION
COLUMN
SULFUR
BURNER
T
SULFUR
CO? AIR
S02
FIGURE 3. CONOCO PROCESS
-------
steam and carbon dioxide at 593-70^ (1100-1300°F) and 1 .52xl06Pa (15 atm) pressure.
The regenerator bed height was maintained at 1.17 m.
Four types of dolomites (Tymochtee, Canaan, Pennsylvania and Virginia)
and two types of limestones (Nebraska and 1691 from Seneca Falls, New York) were
tested and evaluated. The Nebraska limestone showed little activity. While the
1691 limestone gave good attrition resistance, it formed agglomerates. Among the
four dolomites tested, only Canaan stone, containing about 60 percent (by weight)
CaO and 40 percent MgO, on oxide basis, had attrition rates less than 1 percent,
acceptable activity, and long life. In the continous cycling studies (up to 37
cycles) with the Canaan dolomite, about 97 percent of the inlet h^S was removed in
the desulfurizer until most of the CaCQj was consumed. The decline in the regenera-
tion activity was slow, and the n^S concentration in the regeneration off-gas
approached 75 percent of the equilibrium value. Increased levels of h^S in the
regeneration off-gas were observed at higher temperatures (e.g., 760°C as opposed
to 70*» or 593°C) . The effect of solids make-up (about 5 percent under operating
conditions) was tested, and the result verified with a computer model based on
continuous runs with no make-up. Increase of particle size (20 x 23 vs. 35 x ^5
Tyler mesh) did not have any effect on the sorbent activity.
Canaan dolomite was also tested with once-through operation to provide
a basis for cost comparisons with process variations using regeneration. Hydrogen
sulfide absorption of up to 96 percent was obtained with an 85-90 percent solids
conversion. An attrition rate of about 1 percent of the solid feed was observed.
A hardening procedure was developed to reduce the attrition rates of
soft dolomites. This procedure consisted of partially sulfiding the dolomite,
oxidizing the sulfide to calcium sulfate, and introducing the sulfate to the ab-
sorber. The sulfate would then be reduced to sulfide under the operating conditions
of the absorber. Continuous runs with hardened Tymochtee and Virginia dolomites
gave attrition rates below 1 percent, but with some loss of regeneration activity.
Batch tests were conducted to develop a kinetic model of the overall
process, and to determine the effects of various operating variables on desulfur-
ization and regeneration. Regeneration was characterized as a "pseudo" first
order process. The following variables were concluded to Kave no effect on outlet
H2S concentrations in the desulfurized gas: age of stone, particle size over the
range 28 x 35 to 35 x A8 mesh, and inlet H£S concentration over the range of 0.3
percent to 1.1 percent. Deeper beds and higher temperatures reduced the h^S con-
centration and resulted in closer approaches to equilibrium. The percent regen-
eration of CaS decreased with increased stone age, and increased with increasing
temperature and bed depth; it was also adversely affected by hardening of the
stone. Cyclic studies revealed that high temperature regeneration was not useful
as a reactivation method, since a large portion of the CaCO-j produced was inactive
in the absrober. Also, reduced residence time in the absorber improved the
regeneration performance and cycle life.
Sulfur recovery from regeneration exit gas by the liquid phase Claus
reaction was studied in a Jerguson sight gauge reactor, with an internal cross
29
-------
section of 1.6 cm x 191 cm and a length of 137 cm. The reactor was filled with
glass beads of 3mm diameter. Experiments were carried out at 1.52x10" Pa pressure,
and the temperature was kept in the range of 15*<-160OC (310-320°F) to prevent an
increase In the viscosity of liquid sulfur above 160°C. Conversion of the feed
sulfur (about 5.5 mole percent, 2/3 of which is H2S and 1/3 S02) to elemental
sulfur was experimentally demonstrated. In a run where 66.7 percent of the inlet
sulfur was converted, the exit gas contained 1.2 percent H2$ and 0.1 percent S02,
on a dry basis. In other runs, a larger reactor was simulated by making runs with
inlet concentrations equal to outlet concentrations of previous runs.
The liquid phase Chance reaction was studied batch-wise in a glass
apparatus, In two steps. In the first step, called hydrosulfidation, a slurry of
CaS in water was treated with H2S:
CaS + H2S - Ca(HS)2.
In the second step, called carbonation, C02 was introduced:
Ca(tiS)2+ C02 + H20- CaC03 + 2H2S.
It was shown that, on a dry basis, about 95 percent sulfided sorbent, containing
26 percent sulfur, can be converted to a solid with 0.3 percent sulfur for disposal.
The reaction was run at 91°C (195°F) and atmospheric pressure with a residence time
of one hour.
• Uncertainties and Problems: The effect of temperature on sulfur
removal efficiency needs to be studied further. Desulfurization tests were per-
formed at 871°C (1600°F), with a hydrogen removal efficiency of about 97 percent
being obtained. However, at higher temperatures not above the sorbent calcination
temperature, improved sulfur removal efficiences would be expected because of the
desulfurlzation react ion thermodynamics (Figure 8).
During the batch cycling studies, reduced residence time in the desul-
furizer was found to improve the sorbentlife. This fact has yet to be confirmed
In continous runs. A reduced residence time would require smaller equipment, and
a longer sorbent life would reduce sorbent make-up requirements.
A sulfur conversion of about 90 percent wa-> estimated in the liquid phase
Claus Reactor by the process developer. This estimate was obtained by extrapolating
Information gathered from various experimental runs. A single run, simulating the
design conditions, would be more convincing. Furthermore, work is needed to define
the optimum system with respect to reactor volume, gas concentration, temperature
and mixing. There are also possibilities for improvements via the addition of
catalysts and buffers, and operation with nonaqueous solvents.
Hardening of the two types of soft dolomite reduced the sorbent attrition
rates with some loss of activity. Caution should be exercised in extrapolating
limited experiment results when evaluating the applicability of various types of
dolomites to the CONOCO Process. Similarly, the discouraging experimental results
30
-------
obtained with the two types of limestones should not be generalized for all types
of 1imestones.
Corrosion was another problem encountered during operation. The inlet
line to the desulfurizer and the body of the desulfurizer, which were originally
made of type 310 SS, corroded. The inlet line was then replaced with a type
kk6 SS, and the desulfurizer with type 310 SS. No corrosion problems were encoun-
tered in the regenerator, which had a body and internals made of type 310 SS.
Coke formation in the desulfurizer might be a potential problem for the
CONOCO Process. Although this problem was observed in the Air Products Process
and was, in fact, one of the reasons for the termination of that research program,
it was not observed in the CONOCO desulfurization studies. The desulfurizer in
the CONOCO Process was run at about the same temperature range as in the Air
Products Process, but at a higher pressure (1.5x10^ Pa vs. Ixlo5-2xl05 Pa), which
would normally favor coking. However, Air Products used higher CO and lower H20
concentrations than CONOCO, which would enhance coking.
Finally, COS removal with the half-calcined dolomite was not experimentally
studied by CONOCO, and thus remains an uncertainty.
3.1.1.5 U.S. Steel Process (A-*0 : U.S. Steel has been developing a process since
1973 to desulfurize coal-derived gas for use in the reduction of iron ore in a
gravitating bed furnace operating at 8l6-982°C (1500-l800°F) and atmospheric pres-
sure (1x105 Pa). Presently, natural gas is being used for this purpose. The
process is internally funded and proprietary, and published information is not
available. However, a patent application was recently submitted.
The U.S. Steel process employs fully-calcined dolomite (CaO.MgO) at
temperatures above 816°C and atmospheric pressure. Sulfided sorbent is regenerated
in a different method than the one used by Air Products and Chemicals, Inc. (which
uses steam and carbon dioxide). This regeneration method is proprietary.
Bench-scale experiments were initially carried out for about six months.
Bottled gases were used to simulate coal gas. The gas flow rate was varied
between 1.7 and 3-1*0 normal m3/h (60 and 120 SCFH) .
Later, a small pilot plant was operated with a packed bed absorber, 3m
high and 25 cm I.D., filled with fully-calcined dolomite. Coal gas was simulated
by reforming natural gas, and then adding hydrogen sulfide. The 85 m^/H at stan-
dard conditions (3000 SCFH), and the fixed bed was operated above 8l6°C and atmos-
pheric pressure. Hydrogen sulfide removal efficiencies above 98.5 percent were
achieved in these tests.
Further pilot plant tests are planned on gas from an existing gasifier.
• Uncertainties and Problems: Although the U.S. Steel Process is pro-
prietary and no information on the operating problems is available, it can be
anticipated that such problems as the deposition of tars and particualtes on the
absorption bed would occur. This would decrease the sorbent activity and limit
the type of dolomites that can be used for the process.
31
-------
3.1.2 HGC Processes Using Iron Based Absorbents
Five HGC processes using iron based absorbents were identified. These
processes are the Appleby-Frodingham, Babcock and Wilcox, Batelle Columbus Lab-
oratories, IMMR and MERC Processes. Among these only the Appleby-Frodingham
Process was ever used commercially. Currently, process research is being conducted
only on the IMMR and MERC Processes, with work on the other processes having been
terminated. The HGC processes using iron based absorbents are summarized below
and in Table 3 . Sections 3-1 -2.1 through 3-1.2.5 present detailed information on
these processes.
• Appleby-Frodingham Process: Hydrogen sulfide is absorbed with iron
oxide (Fe203) at temperatures up to 427°C (800°F). Regeneration is with air.
Three plants, two pilot and one full-scale, having daily capacities of 0.07x10°,
0.110x10^ and 0.906x10° m3 of inlet gas, respectively, were operated in the United
Kingdom in the 1950's and 1960's to purify coke oven and coal gases. Regeneration
off gases containing sulfur dioxide were used to produce sulfuric acid. These
plants were closed due to the availability of cheaper low-sulfur distillate fuels,
as well as some operating problems.
• Babcock and Wilcox Process: Oxidized carbon steel (FeO) rods are used
to absorb hydrogen sulfide from coal gases at temperatures between *»27 and 649°C
(800 and 120QOF). Regeneration is achieved with air at temperatures between 260 and
649QC (500 and 1200°F). Bench-scale and pilot plant tests were carried out to
study the process feasibility. A conceptual design of the absorption-regeneration
equipment was also developed. The research program was terminated in 1976 due to
lack of internal funding.
• Battelle Columbus Laboratories Process: A mixture of iron oxide
(Fe203), a proprietary material, and inerts (starch, bentonite, and sodium silicate)
was tested at 538 and 8I6°C (1000 and 1500°F), and atmospheric pressure. Regenera-
tion was achieved with air at 538°C (lOOO^C) . Bench-scale experiments were ter-
minated in 1975 due to lack of internal funding.
• IMMR Process: This process developed by the Institute for Mining and
Minerals Research at the University of Kentucky removes sulfur compounds at high
temperatures from fuel gases, with coal gasifier bottom ash (which contains iron
oxide Fe20^ as the active material) being employed as the absorbent. Regeneration
of the spent absorbent is achieved with air. Bench-scale experiments were carried
out at temperatures between 371 and 8l6°C (700 and 1500°F). This research program
is ongoing.
• MERC Process: The Morgantown Energy Research Center, with support
from Air Products and Chemicals, Inc. developed a process to desulfurize producer
gases at temperatures between 538 and 8l6°C (1000 and 1500°F). A pelletized mix-
ture of iron oxide (Fe20o) and flyash or silica is employed as absorbent. The
sulfided pellets are regenerated by oxidation with air at temperatures of 8l6°C.
The larger-scale experimental work at MERC was indefinitely suspended upon review
of the overall project by DOE. Instead, more basic research is being conducted.
32
-------
TABLE 3 - HGC PROCESSES USING IRON BASED
SOLID ABSORBENTS
HGC
PROCESSES
Appleby-
Frodingham
Bab cock 6.
Wi Icox
Batelle
Columbus
IMMR
MERC
STATUS ABSORBENT
Terminated Fe 0,
Z 3
Terminated FeO
Terminated Supported
Fe2°3
On-Going Gasifier
Ash
(contains
On-Going Supported
Fe.O,
2 3
ABSORP-
TION
TEMP .
360-
*27
l|27-
61)9
538
316
37»-
816
538-
816
PRES-
SURE
(Pa)
txlO5
IxlO5
r
^txtOp
9xl05
1x10^
2xl06
GAS METHOD OF
FLOW REGENERATION
RATE
norma 1
O.llxlO6 Air
170 Air
0.06 Air
1.22 Air
212 Air
REGENERATION
TEMPERATURE
593-
8)6
533-
6^9
533
^27-
6*19
538-
816
COMMENTS
Comme re i a 1
Experimental
work
suspended
-------
3.1.2.1 App1eby-Frodingham Process: With the increasing use of coke oven gas
in steel furnaces in the 1930 's and 19^0's, the Appleby-Frodingham Steel Company
of England developed a high temperature sulfur removal process for coke oven gases
for use in the production of low sulfur steels. At that time, the commercially
available process for removing hydrogen sulfide was through the use of iron oxide
purification boxes. This process operated at low temperatures (approximately 35°C
(95°F) and suffered from high capital costs, large space requirements for installa-
tion, dirty handling problems, and a lack of an available market for the spent
oxide. The process developed and tested.by the Appleby-Frodingham Steel Company
used iron oxide particles to absorb sulfur at temperatures between 3**3 and 399°C
(650 and 750°F).
Laboratory-scale experiments (18) at Appleby- Frodingham were carried out
between 19*»6 and 19^8, and were later resumed in 1953. Absorption of both hydrogen
sulfide and organic sulfur was studied using 16 mesh iron oxide particles at
temperatures up to 699°C (1290°F) and at space velocities above 3000 per hour. Re-
generation of the sulfided ore was achieved by means of air at temperatures up to
799°C (1470°F). The laboratory tests indicated high absorption rates and good
activity of the absorbent over hundreds of cycles.
As a result of these favorable laboratory tests, a pilot plant was in-
stalled at Scunthorpe, and the first hot desulfurization tests began in mid-1956.
The pilot plant purified approximately 0.071x106 m3 (2.5xl05ft3) of coke oven gas
per day in a continous operation. During operation, raw coke oven gas was heated
to about 2*»9°C (480°F) by heat exchange with th« purified coke oven gas, and was
admitted to the bottom of a two-stage fluidized-bed absorber which used iron oxide
particles as the absorbent. The coke oven gas was heated to 380°C (716°F) in the
absorber by means of heat exchange with the hot ore from the regenerator. The
sulfided ore was regenerated by fluidizing with air at temperatures between 600 and
800°C, yielding an effluent gas containing sulfur dioxide. The flow of solid
materials between the absorber and regenerator was achieved through fluidization of
the absorbent by means of air or an inert gas. The temperature of the pilot plant
was self-sustaining once the plant started up. The degree of desulfurization attained
in the pilot plant was only slightly lower (71 percent to 8*» percent) than that
achieved in the laboratory unit at similar space velocities and temperatures.
However, 96 percent to 99 percent desulfurization of the coke oven gas was achieved
by using two absoprtion beds in series. Removal of a portion of the organic sulfur
was also observed.
As a result of the performance of the laboratory and pilot plant tests,
a full scale desulfurization plant, with a daily capacity of 0.071 million cubic
meters (32xlo6ft3) of crude coke oven gas, was installed at Scunthorpe. This plant
was equipped with a unit for the manufacture of sulfuric acid from the sulfur removed
from the coke oven gas. The installation was designed to remove 90 percent of the
hydrogen sulfide in the crude gas in a two-bed absorber, with the final equivalent
production of 98 percent sulfuric acid.
Another plant was built to desulfurize 0.110 million normal m3/day (3.9
million SCF/day) of coal gas at the Exeter Works of the South Western Board of England.
-------
The anticipated performance was that approximately 99 percent of hydrogen sulfide
would be removed, together with 90 percent of all organic sulfur compounds other
than thiophene, of which 20 percent would be removed.
The schematic diagram in Figure 10 represents one of the three parellel
units of the Appelby-Frodingham desulfurization plant at Exeter. Two such units
were used in operation, keeping the third on standy-by.
The crude coke gas, from the main blower, is first heated by means of
purified coke oven gas in a shell and tube heat exchanger, after which it is intro-
duced to the base of the four-stage absorber. As the gas moves upward in the
absorber, it fluidizes iron particles at each stage. The iron oxide in the absor-
ber bed reacts with the hydrogen sulfide. The purified exit gas passes first through
cyclones, then the heat exchanger for cooling, and finally, the multicyclone for
dust removal. The iron sulfide from the absorber flows through two seals in series,
where it is fluidized with an inert gas produced from the combustion of towns gas.
It is then conveyed to the regenerator by means of S02-containing recycle gas from
the acid plant; air can also be used as an alternative. In the regenerator, iron
sulfide is converted to oxide by means of air. The oxide flows to the two lower
stages of the absorber by gravity through two seals in parallel. The proportion
of recycled oxide passing to either of the two lower beds is determined by tempei—
ature requirements and adjusted by solids control valves. Hot sulfur dioxide gases
at 421°C (790°F) pass from the regenerator through a cyclone for the removal of
coarse carryover particles, and then to an outlet manifold. The combined gases are
then passed through two cyclones for the removal of fine dust and are introduced,
finally, into the sulfuric acid plant. Fresh oxide for make-up is stored in an
overhead bunker and is diverted pneumatically to small hoppers located on the top
of each absorber, where it is introduced to the top stage of the absorber.
A.C. Bureau (21), et al, report that hydrogen sulfide removal at Exeter
was 99-9 percent, which exceeded the design value of 99.0 percent. The organic
sulfur removal, with the exception of thiophene, was the same as the design value
of 90 percent. About 20 percent of thiophene was removed. The removal efficiency
was strongly affected by the absorption temperature, with removal improving at
higher temperatures and decreasing at lower temperatures. The inlet gas hydrogen
sulfide concentration varied considerably between 7-16 and 14.64 g/m3 (3.13 and
6.40 grains/ft3) whereas the design value was 13.20 g/m3 (5.77 grains/ft3).
The capacity of each absorber was 0.110 million normal m /day (3.9 xlO
SCFD) , with the amount of iron oxide fed to the plant varying between 1044 and 1453
kg/day (2,300 and 3,200 Ibs/day) per absorber. Approximately 50 percent of the
particle sizes were between 40 and 80 mesh, the rest being bigger.
The absorber was operated below 399 C (750 F) . The temperature was
715-750°F at the bottom stage and 360-371°C (680-700°F) at the stage above the
bottom. These two lower stages removed most of the sulfur, whereas the third stage
contributed to sulfur removal only occasionally. The operating temperature in the
regenerator was between 699 and 749°C (1290 and 1380°F) and the plant was self-
supporting with respect to heat requirements. A slight loss of combustible con-
stituents during the passage through the absorber was observed, as can be seen from
Table 4.
35
-------
S02 OFF-GAS\ §
ABSORBENT^
INERT
GAS
STORAGE
BINS
GAS
CYCLONE
HEAT
EXCHANGER
CRUDE GAS
JNLET
SOL/DS
WATER
GAS COOLERS
WATER
MULT I
^CYCLONE
WATER
PURIFIED
GAS
FIGURE '0- /SP?LF,BY-FROniTiCilAJ'i PROCF.SS
-------
Table 4
TYPICAL GAS ANALYSIS ACROSS AN A3SQRBER
Inlet (vol percent) Exit (vol percent)
C02 4.8 5.0
CnHm 1.8 1.6
02 0.2 0.1
H2 ^5.0 44.3
CO 15.1 14.7
C2H6 2.3 2.5
CH/, 15.4 15.4
N2 15.4 16.4
About 90 percent of hydrocyanic acid present in the inlet gas, at 0.57-1-14 g/m
(0.25-0.50 grains/ ft^) concentration, was converted to ammonia, which was subsequently
removed at the exit of the desulfurization plant.
Loss of iron oxide from the plant varied between 7-6-10.9 g/m (474 and
678 lb/10"SCF) of coke oven gas. Loss of coke oven gas through the seals amounted
to 0.5-1 percent of the inlet gas.
Because the oxide particles were conveyed by means of air instead of
sulfur dioxide, the sulfur dioxide concentration in the gas to the acid plant was
reduced to 3-5-6 percent, as compared to the design value of 10 percent.
The closure of the plant in March 1965 was primarily due to economic rea-
sons such as the possibility of producing gas from light distillates at a lower
cost than from solid fuels and the substantial reduction in the anticipated revenue
from the sale of sulfuric acid. The problems involved with solids handling and
the loss of iron oxide absorbent were also significant factors.
• Uncertainties and Problems: The following operating problems were
observed at the desulfurization plants discussed above:
Absorbers: Loss of iron oxide was observed in the absorbers from the ends
of the joints between the separate sections of the distribution plate. This was
corrected by welding.
Indirect Coolers: Condensation took place in the hot gas section of the
indirect coolers. This was corrected by use of a water seal. Although condensation
in the cold gas section of the indirect coolers was minor, back-pressure built up
in the cooler tubes causing occasional shutdown of the plant. This was corrected
by a spray of water.
Gas Multicyclones: Condensation in the gas multi-cyclones caused
sludge build-up and blockage of tubes. This difficulty was solved by maintaining
the temperature of the inlet gas above the water dewpoint.
37
-------
Expansion Bellows: Expansion bellows were installed in the piping be-
tween the absorber, the regenerator, and the seals. These bellows failed and caused
shut down of the plant due to the temperature differences, the corrosion by sulfur
dioxide and trioxide, and the vibration of the fluidized equipments. This difficulty
was solved through an improved bellows design.
Seals: An expansion chamber and a baffle plate were installed after
carryover of fluidized particles in seals was observed.
Sulfur dioxide mains: The hot gas valves on the regenerator off gas main
were affected by high temperatures, presence of sulfur oxides, and iron oxide dust.
For these reasons, they did not operate properly.
Instruments: Some of the instruments, especially pressure and flow instru-
ments, did not work as desired because of fluctuations caused by fluidization and
the plugging of lines.
Oxide handling: _Oxide losses, due to the small particle sizes, were
greater than anticipated and resulted in increased make-up requirements. Dust caused
serious difficulties with blowers, hot gas valves and instruments. Plant shut down
was experienced due to the erosion of bends and rotary feed valves.
3.1.2.2 Babcock and Wilcox Process: The Babcock and Wilcox (B & W) Process uses
iron oxide to remove hydrogen sulfide from gases at high temperatures. Bench-scale
experiments and pilot plant tests were carried out to study the process feasibility.
A conceptual design of the absorption-regeneration equipment was also developed.
However, the need for larger scale experimentation and a lack of funding resulted in
termination of the research program in 197& (A-5,A-6).
Bench-scale tests were performed by passing about 0.11 normal m3/h bottle gas
(simulated synthesis gas) containing 1 percent H2$, 12 percent CO, 8 percent C02,
1 percent CHj, and 78 percent N£ through the reactor bed. Two absorbent materials
were tested: (1) sintered iron oxide, fabricated in the form of cyclindrical cells
of 6.35 mm I.D. and 26.7 cm length, used at temperatures of kk\ and 649°C (825 and
1200°F); and (2) carbon steel, used in annular (53 cm long and 9.5 mm threaded rod
in 12.7 mm I.D. pipe) and tubular (6.8 mm I.D. and 1.98 m long) arrangements at 6^9
and 871°c. The reactants were heated in air at temperatures to 760°C (1*»00°F) to
develop a thin surface layer of FeO :
/\
Fe/Fe + Air > Fe/FeO + N
J{ *m
and were used to absorb hydrogen sulfide from the synthesis gas:
Fe/FeO + xH,S * Fe/FeS + xH 0.
/\
-------
103
2
i
o
o
Uj -3
102
10
1*27
(800)
53^
(1000)
6'*9
(1200)
760
(1400)
87]
(1600)
982
(1800)
TEMPERATURE, °C (°F)
FIGURE II. EQUILIBRIUM CON'STANT FOK II r> ABSORPTION BY IRON OXTDF
39
-------
Sulfur removal above 97 percent efficiency was achieved in the absorption tests.
Regeneration was carried out with air at the absorption temperatures. Approximately
13 percent sulfur dioxide concentration was attained in the regeneration off-gas
and the absorption and regeneration reactions were found to be very rapid. However,
material problems were observed at temperatures around 871°C (1600°F).
Large scale tests (20) were carried out with about 182 kg/h (400 Ib/h)
of coal gas produced in a gasifier. This gas was first cleaned of particulates
(80 percent to 90 percent particulate removal efficiency), and then introduced into
refractory-lined absorbers. Three absorbers were employed in the tests, with two
usually used in parallel for absorption, while the third was regenerated with air.
Carbon steel rods, with a total surface area of 10.8 m2 (116 ft*), were arranged in
staggered grids. Each rod was 6.35 mm in diameter. The purified coal gas from the
absorbers was then introduced into a secondary combustor where the combustibles
were burned with air. The heat evolved was used to raise steam and to preheat the
inlet air. Figure 12 shows a schematic of the large scale test arrangement.
All tests were initiated with the oxidation cycle. During the absorption
cycle, the hydrogen sulfide concentration of the inlet gas varied from 1000 to 7500
ppm by volume, using sulfur addition. These concentrations would be obtained from
0.6 percent to 4. 5 percent sulfur coal (based on 50 percent theoretical air). The
temperature was varied between 427 and 649°C. Experimental results indicate that
higher temperatures and higher contact times increase hydrogen sulfide removal
efficiency. h^S removal efficiency for these pilot plant tests is sh'own in Figure 13,
The cells were regenerated with air at 500-650°F, with flow rates ranging
from A0.9 to 74.9 kilograms (90 to 165 pounds) of air per hour. The highest
"instantaneous" sulfur dioxide concentration observed in the regeneration off-gas
was 14 percent. However, based upon 40.9 kg/h of air flow, an "average" 3.3 percent
sulfur dioxide concentration was calculated for the regeneration cycle.
A more recent paper by Kertamus (2l) discusses results obtained from
bench-scale experiments performed at temperatures between 357 and 649°C (675 and
1200°F) in a 25.4 mm diameter reactor with simulated synthesis gas containing 1 per-
cent CHi,, 12 percent CO, 8 percent H2, 8 percent C02, 6 percent h^O, 1 percent 82$
and 6k percent N2. Plain carbon steel was used in these tests. When absorption
and regeneration tests were carried out at temperatures between 357 and 538°C
(675 and 1000°F), an instantaneous spike in sulfur concentration was observed in
the desulfurized gas. This spike was due to the evolution of sulfur dioxide formed
from the reaction between hydrogen sulfide and iron sulfate, which resulted from
low temperature regeneration. The reactions explaining the formation of the spike
are presented below:
Low temperature regeneration (<533°C) with air:
Fe/FeS -»- Air . Mlnor> Fe/FeOx + xSO, •*• N
X
-------
TO STACK
1 H2S CONVERSION
Y' GRIDS
RAW COAL
STORAGE
TRANSPORT
SECONDARY
COMBUSTION
-f=^ ^~
Vn rrl
SCREW
CONVEYOR
PRIMARY L
GASIFICATION
CHAMBER
COMBUSTOR
FIGURE 12. LARGE SCALE TKST CONFIGURATION
-------
100
90
x.-
£
fc
-j
5
p
Q:
-------
Sulfur concentration spike:
Fe/FeSO^ + A/3 H2$ ^Fe/FeS + A/3 ^0 + A/3 S02
The formation of the spike was eliminated by raising the regeneration
temperature above 538°C. Sulfur removal efficiencies in excess of 95 percent were
observed at temperatures between 357 and 538°C and space velocities of 2000-2500 per
hour. The absorption experiment was terminated when the H2S concentration in the
desulfurized gas reached 0.1 percent (end point). An average H2S concentration in
the desulfurized gas was then obtained at each absorption temperature. The results
are given in Figure lA. Under the same operating conditions, sulfur removal increased
linearly with temperature from 1.07 m3 H2S/100 m2 surface (3.5 SCF H2S/100 sq. ft.
surface) at 357°C to A.3 m3 H2S/100 m2 surface (1A SCF H2S/100 sq. ft. surface) at
538°C. This is shown in Figure 15.
The conceptual design of the absorption-regeneration equipment developed
by B & W Co. consisted of a stainless steel cylindrical unit segmented into sixteen
compartments. Each compartment is filled with carbon-steel plates oriented long-
itudinally with the gas flow, offering about 9.3 m2 (100 ft2) of surface area. Ab-
sorption takes place in thirteen of the sixteen compartments, while regeneration is
being carried out in the remaining compartments. The regeneration air passes in and
upward in the first compartment to a crossover, then downward for a second pass, and
upward for a third and final pass. At two revolutions per hour, each of the sixteen
compartments is regenerated twice per hour. The regeneration off-gas contains from
10 percent to 13 percent S02. The conceptual design is shown in Figure 16.
• Uncertainties and Problems: The process developer indicates that the
optimum operating temperature is between 593° and 6A9°C (1100° and 1200°F). Opera-
tion at a higher temperature, 871°C (1600°F), resulted in unsuccessful regeneration,
coupled with bending of the reaction tube due to the internal heat. However, no
test was performed at temperatures between 6A9 and 871°C; therefore, the upper
temperature limit is not clear. On the other hand, low temperature regeneration
resulted in the formation of FeSOz,, which, after reaction with the coal gas in the
desulfurizer, yielded sulfur dioxide gas. This was eliminated by operating above
538°C (1000°F).
Particle deposition on the grids can limit and even prevent the absorption
and regeneration reactions. In fact, such a problem was observed in the pilot plant
testing.
Carbonyl sulfide removal was not studied in these experiments and, there-
fore, remains an uncertainty.
The conceptual design of the absorption-regeneration equipment developed
by B & W was not tested, and the technical and economic feasibility of this concept
has yet to be demonstrated.
A3
-------
.05
o
o
13
.03
•"•'
.02
.01
\
\
\
SPA
ENL
\
X
CE VEH
i POINT:
"^
OCITY'21
010 PE
x^^
"^^.
500-2500
•RCENT
"^
V/V-HR
SULFUR
•
316 371 427
(600) (700) (300)
482 53R 593 6^9
(900) (1000) (1100) (1200)
TEMPERATURE,°C (°F)
FIGURE 14. SULFUR CONCENTPATION
VERSUS TEMPERATURE
(SCF H2S REMOVED/100 FT2 OF SURFACE)
mj H2S REMOVED/ 100 m2 OF SURFACE
-» —>. ^ **^ "
» *• O\ OO O N> !
2- S S S S S. «
3 — — fo W \Z J
r> 'r* bo '*• b ^j C
SRACi
END
/
f VELto
POINT:
/
CITY: 20 1
0.10 PEf
/
70-2500
?C£7v7S
/.
WV-HR.
'JLFUR
/
/
f
9
316 371 427 482 538 593 649
(600) (700) (800) (300) (1000) (1100) (1200)
TEMPERATURE, °C (°F)
FIGURE 15. SULFUR REMOVAL
-------
3.1.2.3 Sattelle Columbus Laboratories Process: Sattelle Columbus Laboratories
(3CL) developed iron-oxide sorbents (22) to remove hydrogen sulfide from coal gases
at elevated temperatures. Work at BCL began in 197^, as a result of funding under
the Battelle Energy Program, but was suspended in late 1975 due to termination of
this internal funding (A- 7) .
The sorbent contains a mixture of iron oxide and a proprietary material.
This material is believed to react, to a limited extent, with hydrogen sulfide, and
improves the overall reactivity of the sorbent. The absorption reaction of iron
oxide with hydrogen sulfide can be explained through either of two mechanisms:
(1) According to ICI (23), fresh iron oxide (F^O^) is converted to
Fe^Oi,, in the presence of hydrogen at temperatures above 1 77°C (350°F)
The Fe^O/j then reacts with hydrogen sulfide to produce FeS , as follows
Fe^ + H2 + 3H2S - 3 FeS + kHfl
(2) The U.S. Bureau of Mines (2k) reports that the following reaction
takes place:
Fe20 + 3H,S = FeS + FeS2 + 3^0.
The sorbent in the BCL tests was prepared by adding one gram each of
starch, bentonite, and sodium silicate to 100 grams of a mixture of iron oxide and
the proprietary material. The mixture was then either pelletized or extruded (the
extrudates being 3.17 mm in diameter. and 6.35 mm long), oven dried at 121°C (250°F),
sintered at 871-982°C (1600-1 800°F) , and screened to remove all minus 10 mesh
particles .
Laboratory-scale absorption tests were carried out in a fixed-bed tubular
glass reactor of 5.1 cm diameter at 538 and 8l6°C (1000 and 1500°F) temperatures
and atmospheric pressure. A glass (quartz) reactor was chosen after steel reactors
(types 30^t and kk6) were found to react severely with hydrogen sulfide. Bottled
gas containing 17 percent H2« 26 percent CO, 5 percent C02, 50.5 percent N2 and
1.5 percent H2S(21.7 g H2S/m* or 9.5 grains H2S/ft3 gas) was used to simulate coal
gas. Approximately 16.52 normal cm3/s (2.1 SCFH) of gas was passed over 100 grams of
sorbent in each absorption test. The space velocity was 600 vol. gas/vol .bed/hr .
Hydrogen sulfide removal efficiency was found to be above 97 percent, and was higher
at 538°C than at 8l6°C. However, lower desul fur izat ion efficiencies would be expected
under the same operating conditions if steam were present in the inlet fuel gas. The
sorbent developed by BCL proved to be superior to both pure iron oxide and iron
oxide-fly ash mixture (developed by the U.S. Bureau of Mines) with respect to sulfur
removal and attrition resistance.
Regeneration tests were performed at 533°C with an air flow rate of 16.52
normal cm3/s. However, experimental data did not provide conclusive evidence with
regard to the comparative regeneration characteristics of the various sorbents.
-------
REGENERANT
GAS
OUTLET
SOUR PRODUCER GAS
,,11111,,,
AIR
REGENERAM GAS
INLET
Illll"
SWEET PRODUCER GAS
FIGURE 16. 3 S W REGENERATIVE DESULFURIZER
-------
• Uncertainties and Problems: Severe corrosion was observed in the
reactors fabricated from types 30^ and kkC stainless steels. This experimental
problem was solved by employing a quartz reactor. However, the type of material
that can be used in larger scale applications remains uncertain.
The simulated coal gas used in the experiments did not contain any
particles, tars, or steam. Particles present in actual coal gases might deposit
on the absorbent surface, reduce the absorbent life, and even cause plugging of the
bed. Desulfurization at lower temperatures (e.g. 533°C) might cause the condensa-
tion of tars. Also, lower desulfurization efficiencies would generally be expected
than those obtained from these experimental runs, since steam is usually present in
coal-derived product gases.
In view of the limited data on the regeneration step, the regenerabi1ity
of the absorbent is uncertain.
3.1.2.^ IHMR Process: The process developed by the Institute for Mining and
Minerals Research (IMMR) of the University of Kentucky uses the iron contained in
coal ash as the absorbent for hot gas desulfurization. This research program,
which was initiated in 1973, is ongoing and is being funded by the Department of
Energy. Bench-scale tests were performed in fixed- and f1uidized-bed reactors to
develop and understand the basic chemistry of desulfurization, regeneration, and
sulfur extraction, and to determine the important parameters and optimum operating
condi t ions.
The primary desulfurization reactant in coal ash obtained from a high
temperature oxidizing atmosphere is ferric oxide (Fe203). Under the reducing
conditions of the desulfurizer, both reduction and desulfurization reactions take
place. The reduction reactions are between the reducing gases (hydrogen and carbon
monoxide) and iron oxides. These reduction reactions can be represented as follows:
3Fe203 + H2 = 2Fe3Oj, + H20 (l)
3Fe203 + CO = 2Fe30>4 + C02 (2)
Fe30lt + H2 = 3FeO + H20 (3)
Fe30it + CO = 3FeO + C02 CO
FeO + H2 + Fe + H20 (5)
FeO + CO = Fe + C02 (6)
The principal desulfurization reactions are between hydrogen sulfide and
ferric oxide, ferrous oxide and elemental iron:
Fe203 + 2H2S + H2 = 2FeS + 3H90 (7)
Fe203 + 4H2S = 2FeS2 + 3H20 + H2 (3)
FeO + H2S = FeS + H20 (9)
FeO + 2H2S = FeS2 + H20 + HZ (10)
-------
Fe + H2S = FeS + H2 (11)
Fe + 2H2S = FeS2 + 2H2 (12)
Temperature has an effect on the forms of reactants and products. At
680°C (1256°F) ferric oxide transforms from alpha to beta, and at 730°C (11»36°F)
from beta to gamma. Ferrous sulfide transforms from alpha to beta at 230°F, and
from beta to gamma at 325°C (617°F). Ferric disulfide (Fe$2) occurs in two crystal-
line structures: the cubic form (pyrite) and the rhombic form (marcasite). Marcasite
transforms to pyrite at temperatures above 360°C (680°F) .
Equilibrium curves for some of the above reduction and desulfurization
reactions are shown in Figure 17- It can be seen that hydrogen sulfide removal is
favored at lower temperatures. Total pressure has no effect upon the equilibrium
conversions, since an equal number of moles of gas appear on each side of all re-
duction and desulfurization reactions.
Regeneration of the spent sorbent can be achieved with air, according to
the following reactions:
+ 702 = 2Fe203 + *»S02 (13)
1102 = 2Fe203 + 8S02 0*0
Regeneration can also be achieved with sulfur dioxide, with sulfur being extracted.
These reactions are:
6FeS + 4S02 - 2Fe3Oj, + 5S2 (15)
3FeS2 + 2S02 = Fe^ + 4S2 (16)
Reactions (13) and (\k) are exothermic, and have very large equilibrium constants
(pC7l0100) at temperatures between 371 and 8lo°C (700 and 1500°F). On the other
hand, reactions (15) and (16) are endothermic. The equilibrium constants for
reactions (15) and (16) are very small, as shown in Table 5.
Table 5 - EQUILIBRIUM CONSTANTS FOR REACTIONS (15) and (16)
Temperature Equilibrium Constants
(K) Reaction (15) Reaction (16)
_1 3 _ln
650 1.9^x10 * 2.00x10
800 2.88xlo"U 2.21xlO"'3
950 1.21x!0"9 8.^3xlo"8
1100 2.0xlO"8 l.OlxlO"3
Fixed^bed experiments (25,26) were carried out in a 5 cm I.D., 1.22 m long,
stainless steel reactor with the bed height being varied between 79 cm and 122 cm.
The sorbent was a Western Kentucky No. 9 coal ash from a fixed-bed gas producer,
-------
CO
z
o
u
T.
ce
03
c-1
k _]
~T~ 'K
FIGURE 17. EQUILIBRIUM CONSTANTS FOR REDUCTION AND
DESULFURIZATION REACTIONS
-------
containing 23 percent iron as Fe20_j. The sorbent was crushed and screened to a
3x16 mesh size before being introduced into the reactor. The absorption reactions
were conducted at temperatures between 371 and 788°C (1^50°F), pressures between
6.5x10.5 and 1.22x10.6 Pa (80 and 150 psig), space velocities between 800 and 2,000h .
Bottled gases were used to simulate coal gases. The inlet gas composition was:
53 percent nitrogen, 8 percent carbon dioxide, 20 percent carbon monoxide, 15 percent
hydrogen, 3 percent methane and 0.5 to 1.5 percent hydrogen sulfide. Removal
efficiencies greater than 99 percent were achieved with inlet concentrations up to
1.25 percent hydrogen sulfide.
A large variation in ash absorption capacities was observed, ranging from
8.3 to 60.0 grams hydrogen sulfide per kilogram of ash (58-1*21 grains H2S/1b ash).
The following equation, derived from the experimental data, correlates the operating
variables to the absorption capacity:
CAP - (12.0 C) + (0.173 SV) + (0.097** T) + (1.42 P) - 3*»1.0
where
CAP » ash absorption capacity, grains f^S/lb ash
C - inlet h"2S concentration, grains h^S/SCF
SV = space velocity, volume gas/volume bed/hr
T » temperature, °F
P - pressure, psig
From the above equation, it can be seen that higher space velocity, temperature,
pressure, and inlet hydrogen sulfide concentration increase the bed capacity. The
highest absorption capacity, 60 grams h^S/kg ash or (?»21 grains H2$/lb ash) was
found to be 61 percent of the theoretical capacity, based on the available iron
in ash. Examination of the spent sorbent by X-Ray diffraction revealed the presence
of both ferrous and ferric sulfide throughout the structure of the particles.
Regeneration of the spent sorbent was carried out with air up to seven
cycles at 649°C (1200°F) (see equations 13 and I1*), and had a negligible effect on
the subsequent sorption tests. Regeneration of the spent sorbent was also studied
(27) by passing sulfur dioxide containing gas through a fixed bed in a 19.1 mm I.D.
glass reactor at temperatures from 693 to 316°C (1230 to 1500°F) and space velocities
from 430 to 1690 h~'. Sulfur dioxide concentration jn the inlet gas was varied between
22 and 100 percent (see equations 15 and 16). Experimental and computer analysis
suggested that this regeneration method would not be feasible for the production of
elemental sulfur. The primary reason is due to the low equilibrium constants, re-
quiring large quantities of sulfur dioxide for recycle. At 727°C (IS^F) and 20
percent S02 in the feed gas, a circulation load of 77 m3/kg (6x10$ ft3 per ton) of
sulfur produced would be required. In addition, the product gas stream would be
cooled for sulfur condensation and then reheated, resulting in an energy loss.
Furthermore, recycling of large gas volumes would cause increased pumping costs.
50
-------
Fluidized-bed desulfurlzation (26) was also investigated in a glass tube
reaction of 5 cm I.D. and 91 cm height, by passing a mixture of bottled gases
through the bed. The gas mixture contained hydroden, carbon monoxide, carbon
dioxide, methane, nitrogen, and hydrogen sulfide at 1.03x10$ Pa (0.25 psig) pressure
and three different temperatures: 427, 593 and 3l6°C (800, 1100, and 1500°F). The
ash sorbent, with 23 percent iron content, was the same as that used in fixed-bed
experiments, with particle size ranging from -60 to +80 mesh. The space velocity
was kept constant at 2297 n~' and the inlet hydrogen sulfide concentration ranged
between 1.00 and 1.25 percent. The test results showed an increase in ash absorp-
tion capacity after regeneration. The absorption capacity was 13 g H2S/kg ash
(92 grains H2S/lb ash) at 371°C and 1.03x!05 Pa, as compared to 3.6 g H2S/kg ash
(60 grains H2S/lb ash) at 427°C and 1.53x10$ Pa (80 psig) in the fixed-bed experi-
ments. Some tests were also carried out by injecting carbonyl sulfide (0.1 - 0.17
percent), carbon disulfide (0.05-0.10 percent) and tars into the feed gases (25).
Further tests were conducted in fixed bed quartz and Iconel 601 reactors
to determine the effect of operating parameters on the desulfurization process (28).
Four gasifier ash sorbents with Fe20? contents ranging from 5 to 22 percent were
used at Tyler Mesh particle sizes of 20x35, 10x20 and 4x8. The ash capacity was
found to increase slightly with higher temperatures and increased number of cycles
(A-8). The 4x8 size sorbent gave an efficiency of less than 65 percent at 15 per-
cent breakthrough of H2S, whereas smaller sizes had efficiencies of 80 to 90 per-
cent in the temperature range of 538 to 760°C. It was found that the sorbents
effectively removed COS and CS2 as well as H2S from simulated producer gases. How-
ever, the Inconel 601 was severely corroded by ^S. Corrosion tests made on 1020
C.S., 304 S.S., 316 S.S., Incoloy 800, Inconel and Armco 13 SR showed that metals
with low nickel content, such as Armco 18 SR, and chromium to nickel contents above
0.5 showed good corrosion resistance. Regeneration tests performed with air at
427-649°C gave 02 efficiencies of 80-95 percent at 15 percent breakthrough. Two
kinetic models were developed for the preliminary design of the commercial process.
• Uncertainties and Problems: In the cases of both the fixed and
fluidized bed experiments, steam was not present in the inlet gases. The presence
of steam would have had an adverse effect on desulfurization.
Regeneration with simultaneous sulfur extraction is not feasible, as
discussed previously.
3.1.2.5 MERC Process: The Morgantown Energy Research Center (MERC) is investigating
iron oxide (Fe20?)supported sorbents for desul furi zi ng producer gases at 538-8lS°C
(1000-1500°F). This research program was started in the late 1960's, and is still
on-going. It was first funded by the U.S. Bureau of Mines, later by U.S. ERDA, and
presently by the Department of Energy. Air Products and Chemicals, Inc. has assisted
MERC in developing the desulfurization process.
The MERC process removes H2S from raw producer gas at 538-8l6°C in fixed-
bed absorbers containing Fe20.j supported sorbents. The use of the support material
reduces the attrition and loss of the Fe203 sorbent, one of the major problems which
caused the shut-down of the Appleby-Frodingham Plant in England. In the absorber, H2S
51
-------
reacts with Fe203 to form FeS and FeS£, with the empirical composition approaching
FeSj c. The absorption reaction is mildly exothermic (A H = -91.3 kj/mole Fe2n7 or
-21. § kcal/mole Fe20) . and can be shown as:
Fe203 + 3H2S = ZFeS, 5 + 3H20.
This reaction can be obtained from the following two reactions, whose equilibrium
constants are discussed under the IMMR Process:
Fe203 + 2H2S + H2 = 2FeS + 3H2P
Fe203 + 4H2S = 2FeS2 + 3H20 + H2
Regeneration of the sorbent is achieved by passing a gas containing oxygen through
the sulfided bed. The regeneration reactions fs as follows:
This regeneration reaction is highly exothermic (A H = -1.466 kj/mole Fe203 or
-350.1 kcal/mole Fe203) . If during regeneration Fe^Oj, is formed, excess oxygen in
the regenerator converts the Fe,0/| to Fe20^. The reactions are:
Another regeneration scheme which is being considered at MERC produces
sulfur directly in the sulfided bed. According to this scheme, regeneration is
simultaneously accomplished in two beds of sulfided sorbents. An oxygen containing
gas, such as air, is supplied to the first bed, where the following reaction takes
place:
4FeS + 702 _ ^ 2Fe203 + 4S02
The S02 rich effluent gas from the first bed is then directed to the second bed,
where the following reaction at 816<>C produces sulfur:
3FeS + 2S02 - Fe^ H- 5S
The Fe^Oj, is then converted to Fe20, before returning to the absorption bed:
2Fe301| + 1/202 - > 3Fe203
Bench-scale tests (24,29,30,31) were intially carried out to screen
absorbents for removing h^S from producer gas at 538-8l6°C. These tests were con-
ducted In a 25 mm I.D. stainless steel reactor tube with a simulated producer gas
having the following composition: 5 percent C(>2, 26 percent CO, 1.5 percent ^S,
17 percent H£ and 50.5 percent t^2- T°e gas flow rate to the reactor was about
15 SCFH, and the space velocity was 2000 vol. gas/vol . bed/hr. Each absorption test
52
-------
was ended when the ^S concentration from the reactor reached 0 -15 percent (i.e.,
at 10 percent breakthrough). Forty-eight absorbents were tested. Among these,
commercial catalysts, alkalyzed alumina, and taconite disintegrated; illmenite,
open hearth slag, and blast furnace slag particles broke down and fused; sintered
mixtures of 75 percent fly ash and 25 percent metal oxides (13203, V20^, CoO^,
Sn02, ZnO) gave very low absorption capactities, pure sintered Fe^O^ snowed accepta-
ble absorption capacity, but disintegrated at 8l6°C. Pure fly ash had 25 percent
of the absorption capacity obtained with a sintered mixture, 75 percent fly ash
and 25 percent Fe203, but remained undamaged. The most promising absorbents were
found to be sintered mixtures of 75 percent fly ash and 25 percent Fe203, 75 percent
fly ash and 25 percent Jamaican red mud, *»0 percent pumice and 60 percent Fe203,
and pure red muds (Jamaican and Point Comfort),
Further investigations were performed to determine the optimum absorbent
compositions. Sintered mixtures of fly ash and Fe203 containing more than 37 percent
¥e2®3 showed pellet degradation. Sintered mixtures of 75 percent fly ash and 25
percent Fe2®3 and pure Jamaican red mud were found to have comparable absorption
capacities. The AO percent pumice and 60 percent Fe203 mixtures had similar
absorption capacity as the fly ash FeoO? mixtures and pure red mud at 677°C (1250°F),
but had lower absorption capacity at §lo°C.
Absorbent durability was tested through sfx sulfiding-regeneration cycles.
Regeneration was achieved with air at temperatures of 538-8l6°C. A reactor having
a 63.5 mm I.D. was used to reduce the exposed reator wall surface per volume of gas.
This change was made due to the reaction of H?S with the reactor wall in the 25 mm
I.D. reactor. Sintered red mud, pumice-Fe2Q3 mixtures, and fly ash-Jamaican red
mud mixtures suffered either pellet degradation or showed a tendency to fuse to-
gether. Sintered mixtures of 75 percent fly ash - 25 percent Fe20, did not exhibit
any of these problems below 8l6°C, but fused at 871°C (1603°F). The fly ash-Fe2Q3
sorbents were prepared by mixing 99 percent pure Fe?03 Pow^er» water and fly ash
(-*» to +6 mesh) containing *»7.9 percent Si02, 23,3 percent A1203, 15.7 percent
F6203, 0.6 percent P205, 2.3 percent Ti02, 3-6 percent CaO, 1.5 percent MgO, 1.9
percent Na20 and 2.2 percent l<20. Pellets of about 6,35 mm diameter were formed,
dried at 26QOC (500°F) , and finally sintered at 1093-1 1*»9°C (2000-2100°F) to increase
the!r hardness.
Long-term absorption-regeneration tests were performed up to 17^ cycles
at temperatures of 538-8l6°C and a space velocity of 1000 vol. gas/vol. bed/hr. using
sintered 75 percent fly ash and 25 percent Fe20T. No reduction in absorption activity
was observed in these tests with simulated dry and wet (containing 7 percent steam)
producer gas. However, absorption capacity increased with temperature, and was
found to be higher when the inlet gas did not contain steam. Figure 18 shows the
relationship between temperature and absorption capacity. Tests with actual pro-
ducer gases, which contained 8 g dust and 16 g tar per normal cubic meter (0.5 lb
dust and 1 lb tar per 1000 SCF) of gas showed 95-?7 percent H2S removal. Some
carbon deposition on the pellets, due to cracking of tars and accumulation of dust
in the bed, was observed.
53
-------
16
12
3
o
10
§
o>
o
•
en
O
538
(1000)
I
593
(1100)
649
(1200)
704
(1300)
760
(1400)
Pa)
816
(1500)
ABSORBENT TEMPERATURE, °C (°F)
FIGURE 18. ABSORPTION CAPACITY OF SINTERED PELLETS OF
75% FLY ASH - 25% Fe O
-------
Regeneration studies were conducted with both air and oxygen. With
air at a space velocity of 1000 vol. gas/vol. bed/hr, the effluent S02 concentration
reached about 10 percent. In the case of oxygen, more concentrated $62 effluents
were obtained. In all cases, the bed temperature was kept below 8l6°C to avoid
fusion of the pellets. Analysis of the sulfided sorbent indicated the empirical
composition to be FeS] 3. Figurel9 shows typical regeneration curves for air and/
or oxygen.
Further absorption tests (32) were carried out in a 15.2 011(6'.') I .D.stainless
steel reactor containing 33kg(73 lb) of 4.76mm (3/16") diameter sorbent pellets
(75 percent fly ash - 25 percent Fe20^, with 3 percent bentonite as binder). Low
Btu gas from a fixed-bed gasifier was first introduced into a cyclone at a flow rate
of 42.5 m^ (1,500 SCFH) for the removal of large particles and then to the absorber
at 593°C. The space velocity was 1,900 vol. gas/vol. bed/h • The absorption capa-
city of the sorbent averaged 8.25 percent. The h^S removal efficiency was between
90 and 9^ percent, and was not affected by tars and particulates in the inlet gas.
However, deposition of particulates was observed in the absorber bed. These particles
were burned out during regeneration with air.
Various binders were investigated to improve the physical strength of the
sorbents (75 percent fly ash - 25 percent Fe203), without reducing the sorption
capacity (33). Bentonite and sodium silicate were found to be satisfactory. The
crushing strength was increased up to 31.8 kg/cm (70 lb/cm) with 12.7 mm (0.5")
diamter and 9.65 mm (0.38") long extrudates. The sorption capacity was raised to
12-14 percent at 593-704°C by sintering at 982°C (1800°F) for 30 minutes in an
oxidizing atmosphere. Temperature control in the regenerator was also studied to
prolong the sorbent life. Such control was achieved by diluting the regenerator
inlet air with nitrogen.
Mixtures of silica and Fe20^ were also studied (34) since silica appeared
to be superior to fly ash with respect to having more uniform properties and a
higher melting point. Tests in the 50.3 mm I.D. absorber with 157 normal cm'/s
(20 SCFH) of simulated producer gas at temperatures around 982°C showed that mixtures
of Fe20o and silica, containing up to 45 percent Fe203, are promising. The addi-
tion of sodium silicate was found to increase both sorption capacity and crush
strength. A mixture of 55 percent silica and 45 percent Fe20o, prepared with the
addition of sodium silicate, had a crush strength of 30.9 kg/cm (68 lb/cm) and a
capacity of 22.8 g H2S/100 g sorbent.
Continous tests (35) were also performed in a larger-scale pilot plant.
Flow rates up to 212 normal m3/h (7500 SCFH) of coal gas were employed with absor-
bent containing 55 percent silica and 45 percent Fe20o. Regeneration was accomplished
with 5.7 normal m3/h (200 SCFH) of air plus 42.5 normal m3/h (1500 SCFH) of inert
gas. Experimental difficulties were observed during six-day-long continuous
testing.
Air Products and Chemicals, Inc. (APCl) provided assistance to MERC
from early 1975 to 1977 (36,37,38). Under the APCl research program, sorbents
were prepared by varying the source of iron oxide (7 sources), the iron oxide
content of sorbents (from 8 to 63 percent), the support materials (fly ash and
55
-------
REGENERATION COMPLETED, %
0 9093 91. 100
100
80
60-
Uj
Uj
40
20-
-------
silica), the type of binders (bentonite and kaolin), and the particle size of
sorbents (3-18 and 6.35 mm diameter pellets). These sorbents were tested for such
properties as crush strength, porosity, and surface area. They were evaluated after
undergoing a single absorption and regeneration cycle. The recommended sorbents
included 42 percent iron oxide on fly ash, 42 percent iron oxide on silica (for all
sorption pressure ranges), and 63 percent iron oxide on fly ash (for sorption pres-
sures around 400 psi). The 42 percent iron oxide sorbents had an absorption capacity
around 0.18 sulfur/g absorbent, while the 63 percent iron oxide sorbent had absorp-
tion capacities of 0.13 gram sulfur/cm3 absorbent at 20 ps i and 0.32 g sulfur/cm-*
absorbent at 2.76x10° Pa (400 psi). Bentonite was found to be the preferred binder,
and methocel was used as an extrusion aid. The effects of operating variables on
the absorption and regeneration characteristics were also experimentally studied
in the absorber. At a temperature of 649°C and a space velocity of 1800 vol. gas/vol.
bed/hr, about 70 percent of the fresh absorbent capacity was utilized at the 10 per-
cent breakthrough point (i.e., when the exit h^S concentration reaches 10 percent
of the inlet H2S concentration). In the regenerator, with an inlet gas containing
85 percent steam and 15 percent air, 85 percent of the sorbed sulfur was removed
after regeneration. Mathematical models were developed to fit the experimental
absorption and regeneration data.
A review of the overall MERC project by DOE in 1977 resulted in a decision
to suspend indefinitely the larger-scale process development work, and to emphasize
supporting research (39)- Under the revised program, the chemistry of hot gas
clean-up was studied for more than 20 possible absorbents (46).
• Uncertainties and Problems: The exothermic regeneration reaction
generates high temperatures, which can result in a loss of sorbent activity due to
sintering. This problem was observed in the bench-scale tests while operating with
both oxygen and air. Temperature control in the regenerator was better achieved by
diluting air with nitrogen. However, this would reduce the S02 concentration in
the regenerator effluent gas and increase the pumping cost and equipment size require-
ments due to the larger volumes of gases to be handled.
Corrosion of the stainless steel equipment was observed in the earlier
tests. Suitable materials of construction need to be investigated.
Plugging of the control equipment occured in the large-scale continuous
experiments. This problem might have occured due to the deposition of particulates
as well as the condensation of tars at lower temperatures. Similar problems might
also be expected in commercial scale operations.
3.1.3 HGC Processes Using Copper Based Absorbents
Four HGC processes using copper based absorbents have been identified.
These include the Atlantic Refining -X , ESSO, Johns Hopkins and Kennecott Processes.
None of these processes is currently being further developed. Among them, the most
extensive testing was performed on the Johns Hopkins Process in the 1930's. Sum-
marized information for these processes is presented below and in Table 6. Detailed
information for these processes is presented in Sections 3-1-3.1 through 3.1-3.3.
57
-------
TABLE 6 - HGC PROCESSES USING COPPER BASED
SOLID ABSORBENTS
vn
u>
HGC STATUS
PROCESSES
Atlantic Terminated
Refining -I
ESSO Terminated
Johns Terminated
Hopkins
Kennecott Terminated
ABSORBENT
Cu/Pb/Zn
alumi no-
si 1 icate
Supported
Cu
Supported
Cu/Cr/V
Cu/CuO/
Cu(OH),
ABSORP-
TION
TEMP.
(°C)
93-
538
299-
932
Up to
816
«i82
-------
• Atlantic Refining Process -I : Heavy metal al umi no-si 1 i cates such as
copper al umi nosi 1 i cate, lead al umi nos i 1 i cate or zinc al umi nos i 1 i cate , impregnated
upon inert carriers, are used to desulfurize petroleum distillates and gases at
temperatures between 93 and 530°C (200 and 1000°F) . Regeneration of the spent
sorbent is performed with steam and air at temperatures between 371 and 538°C
(700 and 1000°F). Bench-scale experiments were conducted in the 1940's. There is
no current research on this process.
• ESSO Process: Hydrogen sulfide from reducing gases is removed with
copper containing solids at temperatures up to 982°C (1800°F) in a fluidized bed
absorber. The cuprous sulfide is then reacted with cupric oxide at temperatures
around 732°C (1350°F) to regenerate the copper and obtain sulfur dioxide gas. The
cuprous oxide is produced by oxidizing copper in a separate vessel. Bench-scale
experiments were carried out to test the different components and operations of the
process. There is no current research on this process.
• Johns Hopkins Process: The Gas Engineering Department of Johns Hopkins
University, supported by two gas companies, conducted extensive testing of copper,
vanadium, chromium and uranium containing absorbents at temperatures up to 8l6°C
(1500°F) to desulfurize coal gases. Laboratory pilot-scale, and large-scale plant
tests were performed in the 1920's and 1930's. The research activities were termina-
ted due to the unfavorable economic conditions during the depression and the in-
creasing use of heavy oil.
• Kennecott Process: Hydrogen sulfide is absorbed by copper containing
materials such as copper, copper hydrate, cupric oxide, and copper concentrate di-
luted with sand in proportions varying between 75 and 87-5 percent at temperatures
between 482 and 496 C (900 and 925 F) . Regeneration is with air at 8l6°C (1500°F).
Bench-scale experiments were terminated in 1976 due to business related reasons.
3.1.3.1 Atlantic Refining Process— I : The Atlantic Refining Process - I , inven-
ted by Nachod, et al (41 ) , was developed to desulfurize petroleum distillates and
gases with heavy metal al umi no-s i 1 icates at temperatures ranging from 93 and 538°C
(200 to 1000°F). Regeneration of the spent sorbent is accomplished with steam and/
or air at temperatures between 371 and 538°C (700 and 1000°F) .
Impregnated treating materials such as metal salts or oxides, deposited
upon inert carriers (e.g., clay), usually are subject to attrition, with the active
material being separated from the carrier and removed with the gas to be desul fur i zed.
This difficulty can be overcome by fixing the heavy metal ion (e.g., lead, copper,
zinc) in such a way that it is still capable of effecting the desul furizing reaction,
yet is bound so tightly to the carrier that it cannot be removed, with resulting loss
in capacity and contamination of the gas under treatment. This is accomplished by
fixing the cation M in a cation exchanger AZ (such as sodium or potassium alumino-
silicate) in the following manner:
The cation exchanger may first be formed by reacting sodium aluminate, sodium sili-
cate and alum, in suitable proportions, in an aqueous solution at a pH between 3 and
59
-------
9. The ratio of Na20 to A120, to Si02 may range from 1:1:2 to 1:1:8. The re-
sulting alumino-si1icate is then washed, dried, and treated with an aqueous solution
of a heavy metal salt to replace the sodium with the heavy metal ions. The heavy
metal alumino-silicate is washed and finally dried.
Lead and copper alumino-si1icates were prepared by reacting sodium alumino-
si] icate and lead/copper acetate. The absorbents were then tested to remove sulfur
compounds from vaporized naphtha at temperatures between 20A and ^82°C (*tOO and 900°F)
and a space velocity of 2 vol. naphtha/vol. alumino-si1icate/hr. Sulfur compounds
in the naphtha were reduced up to 4 3 percent.
There is no current work or future plans for further development of the
process.
• Uncertainties and Problems: The available information on this process
was obtained from a patent, which does not discuss any operating problems. Although
the sorbent preparation method was described in the patent, and the use of some
sorbents for hot gas desulfurization were experimentally demonstrated, the superior-
ity of such prepared sorbents over impregnated materials was not proven.
The maximum sulfur removal efficiency reported in the patent was *»3 percent.
Such a low efficiency would only be useful for cleaning gaseous streams originating
from low sulfur fuels.
The proposed regeneration methods were not experimentally demonstrated.
3.1.3.2 ESSO Process: The Esso Process, invented by Lewis (k2) t was developed to
desulfurize reducing gases (such as coke oven gas, producer gas, water gas, hydro-
carbon gases, and gases synthesized from petroleum oils) with copper-containing
solids, and produce high quality sulfur dioxide gas from regeneration of the spent
sorbent. No information could be obtained as to why the process was terminated.
However, there is no current work or future plans to further develop this process.
The reducing gas fluidizes a bed of copper-containing solids at tempera-
tures between 299 and 999°C(570 and 183QOF). The desulfurization reaction is:
2 Cu + H2S = Cu2S + H2. (1)
The variation of the equilibrium constant with temperatures for the above
reaction is shown in Figure 20. As can be seen from Figure 20, hydrogen sulfide
removal is thermodynamically favored at lower temperatures. However, rapid desul-
furization is obtained at higher temperatures due to favorable kinetics.
Copper oxides are undesirable in the absorber since they undergo reduction
with the incoming fuel gas by forming steam and carbon dioxide at the expense of
hydrogen and carbon monoxide, thereby lowering the heating value of the desulfurized
gas. On the other hand, when the sorbent is copper rather than the copper oxides,
removal of hydrogen sulfide takes place with simultaneous generation of hydrogen.
60
-------
FIGURE 20. EQUILIBRIUM CONSTANT FOR THE REACTION:
2Cu + HS = CuS + H
61
-------
The sulfided copper is withdrawn from the absorption bed and transferred
to the regenerator, where it is reacted with cuprous oxide solids. These cuprous
oxide solids are produced in another vessel by oxidation of copper with air (or
oxygen) at temperatures between 399 and 593°C (750 and 1110°F), and then passed to
the regenerator. The regeneration off-gas contains sulfur dioxide. The regeneration
reaction in Equation 2 takes place in the temperature range from 399 and 593°C,
or higher:
Cu2S + 2Cu20 = 6Cu + S02. (2)
The pressure in the regenerator must be low enough to evolve S02 from the
Cu0 mixture.
A flow diagram of the Esso Process is shown in Figure 21 . Copper-contain-
ing materials are contacted with the reducing gas in the f luidized-bed desul furizer,
where reaction (l) takes place. Desulfurized gas leaves the top of the absorber,
while the spent sorbents (Cu£S containing solids) are continuously withdrawn for
regeneration. In the regenerator, Cu20 containing solids are mixed with a small
stoichiometric excess of the Cu£S containing solids. Intimate mixing can be obtained
by mechanical means (e.g., stirring) and by fluidizing the regenerator bed with a
sulfur dioxide recycle stream. The regenerated solids, containing mostly copper,
but with a minor amount of Cu2$ and some Ci^O, are directed into two separate com-
partments located at the top of the regenerator. In one of the compartments, a por-
tion of the solids from the main body of the regenerator is further treated to
convert any remaining copper oxides to copper. The solids which now contain mostly
copper and some Cu2$ are then withdrawn from the regenerator and returned to the
desulfurizer. In the other compartment, a portion of the regenerated solids is
exposed to an excess of C^O from the oxidizer, thereby ensuring that the solids
drawn off from this compartment into the oxidizer are practically free of Cu2S. A
portion of the regenerator effluent gas is recycled to the regenerator as a fluidizing
agent.
In the oxidizer, air or oxygen-containing gas is reacted with Cu to form
Cu20 according to the following reaction:
2Cu + 1/202 - Cu20 (3)
The Cu20 solids from the oxidizer are returned to the regenerator for reaction with
Cu2S.
Bench-scale experiments were performed to test the desul furization and
regeneration reactions. Copper-containing solids were prepared by impregnating sil-
ica gel with a copper nitrate solution. The mix was then dried, roasted in air,
crushed to sizes finer than 100 mesh, and reduced with hydrogen at a temperature of
149 to 204°C (300 to 40QQF) . A hydrogen gas stream, with 1.0 - 1.5 percent H2S,
was contacted with copper impregnated silica gel containing 2 to 20 percent Cu.
flow rates greater than 3.12 cnP/3/g Cu (130 SCFH gas/lb Cu) in the bed were em-
ployed. Space velocities were greater than 500 vol .gas/vol .bed/h . Hydrogen sulfide
removal was determined at different temperatures. Within the optimum temperature
62
-------
DESULFURIZED
DtSULFURIZER-
I
f — '
»
\»_
1
IMP
REDl
GA
_~-^
URE
7C/A
S
MENTS <
<£ i
i! i
Juy^xJl
3L
Cu8 Cu2S
COPPER SOLIDS
REGENERATOR-
SULFIDED
COPPER SOLIDS
a
i
i
i
i
• \
1
I
j
^
-.
i
I
1
|
I
i
L*<
II
Ii
ii
,,,11.
ii
1
O/20
SOLIDS
. . • • >
SOLIDS
OXIDIZER —
Cujo
"~^r OUL/L/O
>^
T
AIR
c
>
r
FIGURE 21. ESSO PROCESS
63
-------
range of 3^9 to 593°C (660 to 1110°F), the removal efficiency was about 98.5 percent.
At higher temperatures, in the range of 593 to 927°C (1100 to 1700°F), the removal
efficiency was found to be between 80 and 95 percent.
Regeneration of the spent sorbent was first accomplished in a bed fluidized
with air at temperatures between 399 and 704°C. The air rate was controlled to
provide only enough oxygen to convert the Cu2$ to Cu by the reaction:
Cu2S + 02 = 2Cu + S02 (k)
However, some Cu£0 was observed as a reaction product. The CuoO solids were con-
tinuously removed from the bed and reacted with Cu^S in excess of 1 mole per 2 moles
of Cu£0 at temperatures ranging from 449 to 733°C (840 to 1360°F) . Sulfur dioxide
gas evolved from the reaction zone, and the solids were free of Cu20 and contained
about 90 percent Cu and 10 percent
In another experimental run, regeneration was carried out in a stirred
reactor by supplying Cu20 at a rate of a little less than 2 moles per mole of Cu2S,
at temperatures between 449 and 738°C, without any addition of air, oxygen, or
nitrogen. While the evolution of S02 was proceeding, the solids mixture contained
copper, primarily in the form of Cu, with much smaller amounts of Cu2S, and very
1 ittle Cu20.
• Uncertainties and Problems: The available information on the Esso
Process is limited. There are no data on the level of mixing or solids handling.
Although fluidization with S02 is recommended as an alternative method for mixing
in the regenerator, no experiments were performed for confirmation.
A small amount of copper, impregnated on the carrier, improves temperature
control in the desulfurizer. However, this requires larger quantities of solids to
be circulated and larger treating vessels.
3.1.3-3 Johns Hopkins Process; The Gas Engineering Department of the Johns Hopkins
University (J.H.U.) carried out extensive research and test programs into hot gas
desulfurization during the late 1920's and early 1930's to develop suitable absor-
bents that might reduce gas purification costs. This work was supported by several
gas companies, including the Consolidated Gas Electric Light and Power Company of
Baltimore and the Consolidated Edison Company of New York. The activities carried
out by J.H.U., which are described by Huff and Logan (43), included laboratory-
scale research, and both small and large-scale plant tests.
Laboratory Research: Laboratory tests were performed to develop suitable
absorbents for the simultaneous removal of both H2$ and organic sulfur compounds at
elevated temperatures. Blue gas, containing about 2.29 grams of H2$ and 0.09-0.l8g
of organic sulfur per m3 of gas, (100 grains of H2S and 4-8 grains of organic sulfur
per 100 ft^ of gas) was used for test purposes. Regeneration on the sulfided
absorbent was achieved by blowing air through the sulfided absorbent at a flow rate
of 15.7 cm3/s (2 f t3/h) . The first experiments made on copper oxide and iron oxide
showed that neither copper nor iron alone was effective for removing sulfur. Copper
did not absorb sulfur at 349 and 44g°C (660 and 840°F) , whereas iron oxide plugged
64
-------
quickly with carbon at 449 C, and could not be regenerated at 3i»9°C. Other absorbents
tested in the range of 249 and 449°C (480 to 840°F) with commercial blue gas in-
cluded 85 percent copper - 15 percent uranium on pumice, 50 percent copper - 50
percent uranium, 30 percent copper - 10 percent uranium - 10 percent chromium, 80
percent tin - 20 percent uranium, 60 percent tin - 40 percent uranium, 80 percent
antimony - 20 percent uranium, 80 percent uranium - 20 percent cerium, 80 percent
bismuth - 20 percent uranium on pumice, compressed lime, compressed molybdenum, 85
percent iron - 10 percent uranium - 5 percent manganese, and compressed vanadium.
Of these materials, the copper-uranium-chromium mixture gave the most promising re-
sults.
Next, continuous tests were carried out at 450°C (842°F) with 15 cm of
absorbent consisting of 80 percent copper, 10 percent uranium and 10 percent chromium.
Blue gas, with the following composition, was used in these tests: 13-2 percent C02,
2.1 percent 02, 14.9 percent ^2, 16 percent CO, 0.1 percent CHi,, 0.2 percent C2Hg,
and 53-5 percent N2- The amounts of t^S and organic sulfur added to 1 m' of this
gas were 1.56-3.09 and 0.09-0.18 grams (73-135 grains H2S/100 ft3 gas and 4-3 grains
organic sulfur/100 ft3 gas), respectively. The space velocity was approximately
3,000 per hour. About 0.03 m^ of air was used to regenerate the absorbent. Hydro-
gen sulfide removal was above 97.3 percent in all cases. The addition of 30 percent
steam to the inlet gas improved the efficiency of hydrogen sulfide removal to over
99 percent and resulted in almost complete removal of organic sulfur.
While the mechanism of the reactions was not investigated, Huff and Logan
concluded that uranium or chromium is primarily responsible for the conversion of
organic sulfur to hydrogen sulfide, and copper for the absorption of sulfur as
hydrogen sulfide.
Additional tests were made to develop a suitable carrier or binder for
the active material, to enable it to withstand plant use without crumbling. These
tests included such treatments as fusing the active material (copper, vanadium,
chromium, uranium) either alone or with aluminum, and baking with alundum cement or
clay. These tests indicated that copper-vanadium-clay mixtures were the most pro-
mising. A mixture of 70 percent clay - 30 percent oxides (80 percent copper -
20 percent vanadium as metals), baked at 899-949°C (1650-1740°F), removed hydrogen
sulfide and organic sulfur completely from a carburetted water gas containing 1.97
to 20.0 g H2S/m3 (86 to 875 grains H2S/100 ft3) gas at 449°C. The absorption-
regeneration cycle consisted of 3.5 minutes for absorption, followed by 3 minutes
for regeneration. The space velocity was about 5,600 per hour.
Small-Scale Plant Tests: On the basis of the encouraging laboratory tests
described above, a test apparatus was installed in 1928 at the Spring Gardens Plant
of the Consolidated Gas Electric Light and Power Company of Baltimore. This apparatus
was attached to the top of a superheater for backrun carburetted water gas produced
by a water gas machine, and heated independently to maintain the desired experimental
temperatures. The operating conditions were more adverse than those prevailing in
the laboratory, especially with respect to the presence of dust and tar in the raw
carburetted water gas. The ability to control the temperature of the purification
process, and the concentration of the oxygen in the blast gas during regeneration,
were maintained during these tests.
65
-------
An absorbent of 80 percent copper - 20 percent vanadium mixed with clay
(70 percent clay and 30 percent oxides) was molded into small rings (1.45 mm long
x 13.5 mm O.D. x 7.1 mm I.D.), and tested at temperatures between 482 and 510°C.
In one experiment of 33-5 hours duration, 7.96 m3 (281 ft3) of gas containing 2.47g
H2S/m3 (108 grains H2S/100 ft') was completely desulfurized with 0.64 Ib of puri-
fying material. Regeneration was achieved by means of air.
These promising results led to the installation of a similar, but larger,
apparatus operating on both uprun and backrun gas which contained 3-00 and 5.1 g
H2S/m3 (131 and 222 grains H2S/100 ft3), respectively, and about 0.39 g organic
sulfur/m3 (17 grains organic sulfur/100 ft3). Average removal efficiencies of
93.2 percent for H2S and 74 percent for organic sulfur were obtained. Short tests
also indicated copper-chromium-clay mixtures to be promising.
Small plant tests were continued in an attempt to reduce the cost of
producing the absorbent by coating clay forms with a layer composed of 60 percent
clay and 40 percent oxides.
In a 1500 hour experiment conducted at around 538 C, hydrogen sulfide
removal efficiency was 79-5 percent during the first 200 hours and 66.7 percent
during the last 200 hours. Removal of organic sulfur averaged about 65 percent.
The space velocity in this experiment was 2,000 per hour. The gas flow varied be-
tween 3.1 to 10.2 m3/hr (282 to 360 ft^/hr), and H£$ concentration in the inlet gas
was between 2.4 and 3-4 g/m3 (106 and 147 grains/100 ft3). Another experiment was
conducted to determine the effect of increased temperature. Raising the temperature
from 5IO°C to above 8l6°C, and then lowering it to 527°C, did not affect the activity
of the material and the removal of hydrogen sulfide remained in the neighborhood
of 85 percent.
Large-Scale Plant Tests: Large-scale plant tests were performed by
placing the coated absorbent forms in the top of a superheater, instead of placing
them in a separate chamber outside the superheater, to save the cost of additional
plant equipment.
The first test was carried out at the gas plant in Baltimore discussed
previously. Nine thousand small, circular, finned forms (with copper-vanadium
mixtures as active material) were used, occupying a space of about 1.82 m in diame-
ter and 0.91 m in height. This test lasted 442 hours, during which the removal
efficiency decreased from 76.4 percent to 41 percent. The number of forms was then
increased to 23,000 for the next test. However, an initial desulfurization efficiency
of 75-80 percent declined to 24 percent after 3,700 hours of operation. Regeneration
was achieved by introducing air at the bottom of the superheater.
Two additional plant-scale experimental programs, one in Baltimore and the
other in New York, were intiated during 1930. Results were generally similar to the
initial tests, with absorbent activity decreasing dramatically with time. The
facility at the Central Union Plant of the Consolidated Edison Company in New York
was first run with 2,100 forms of the kind used previously in the test in Baltimore.
These forms occupied the upper third of the superheater. The removal efficiency
66
-------
dropped from ^9-5 to 39 percent after 1300 hours of operation. The number of forms
was then increased to 60,000 and an efficiency of 77-5 percent was obtained with
carburetted water gas for the first TOO hours. However, the efficiency dropped to
67 percent after about 800 hours of operation, and then to below 50 percent. Lab-
oratory tests showed a loss of active chemicals, particularly copper, from the
absorbent coatings.
Tests at Baltimore were performed with coated rectangular grid forms,
employing copper-vanadium as the active material. An initial efficiency of 91 per-
cent for the first 600 hours dropped to about 20 percent after 2,200 hours of oper-
ation. Loss of active material was also experienced during these tests.
The large-scale plant tests gave lower removal efficiencies than the
earlier laboratory and small-scale plant tests. This was due to the high tempera-
tures (about 760-3l6°C in the superheaters, as compared to 538°C in small-scale
tests), the presence of excessive chloride impurities, high ash, and frequent
leaching with steam when treating the raw carburetted water gas.
Investigation of the lowering of activity of the coated forms led to the
development of a modified form in which the active chemicals, copper and chromium,
were heavily incorporated throughout the material.
The formula was 55 parts clay, 20 parts stoneware grog,, and 25 parts
oxides (18.87 Cu20, 6.13 C^O^) . Tests performed at Spring Gardens, Baltimore,
showed removal efficiencies of 60.7 percent the first month and 15 percent at the
end of 7300 hours of testing. As was the case with the coated forms, considerable
loss of copper was again experienced.
Laboratory tests suggested that the loss of activity might be due to a
slow chemical reaction between the copper and the silica of the clay. Preliminary
tests were carried out to develop an absorbent that would maintin its activity during
plant use. Studies were made to determine the effects of adding an alkaline (MgO)
substance to the absorbent mixture. While preliminary results were encouraging, no
plant data was obtained with the new absorbent forms because research activites
were suspended. The apparent reasons for terminating this test program involve the
economic conditions during the period of "the depression", and the increasing use
of heavy oil during the 1330's, which obviated the need for coal-derived water gas.
Problems experienced at the various stages of development of the Johns
Hopkins Process are incorporated in the discussions pertinent to each stage.
3.1.3.^ Kennecott Process: Kennecott Copper Corporation (KCC) investigated the
applicability of copper-containing absorbents for hot gas desulfurization. Research
at the Ledgemont Laboratory of KCC started in 1971, and was terminated in 1976 due
to business related reasons. The research program was internally funded (A-9).
Bench-scale experiments were carried out in a reactor of 25 mm I.D. and
15 cm height (A-10). Operationwith Lurgi fuel gas at 2.0xl06 - 2.5xlO& Pa (20-25
atm) pressure was simulated by using bottled gases with the equivalent hydrogen
67
-------
sulfide partial pressure at 1.0x10 Pa (1 atm) total pressure (1*7) . The inlet bottle-
gas composition was: 23 percent HpS, 30 percent CO, and k7 percent H_. Four
different copper containing absorbents were tested: copper metal (Cu;, copper
hydrate (Cu(OH)2), cupric oxide (CuO) and copper concentrate (2*t.3 percent Cu,
26 percent S, 21.5 percent Fe). These absorbents were diluted with sand in propor-
tions varying between 75 percent and 87.5 percent of sand (by weight). Successive
absorption and regeneration cycles (3 or 4 cycles per experiment) were conducted
at absorption temperatures of 482 and 510°C (900 and 950°F) and a regeneration
temperature of 8l6°C (1500°F). Each of the four absorbents was found to be capable
of removing in excess of 90 percent of the hydrogen sulfide, with a gas residence
time of about 0.5 seconds. Each of the sulfided absorbents was regenerated with air
at 816°C with a similar reaction time. The regeneration off-gas contained 12-14
percent sulfur dioxide for 99 percent regeneration of the sulfided absorbent.
The equilibrium constants for reactions involved in the absorption cycle
are given in Figure 22 as a function of temperature. This figure indicates that
cuprous sulfide is the absorption product, even when cupric oxide is used as an
absorbent under reducing atmosphere.
• Uncertainties and Problems: The scale of these tests was too small to
draw any definite conclusions regarding the technical or economic feasibility of
the process.
Hot gas desulfurization was studied over the temperature range between
482 and 510°C. No testing was performed beyond this range. Regeneration tests were
not conducted at any other temperature but 8l6°C.
The simulated fuel gas contained only hydrogen sulfide, carbon monoxide
and hydrogen. However, the presence of carbon dioxide and steam would reduce the
sulfur removal efficiency. Also, the presence of particulates and tars in the
fuel gas might cause operational problems.
Removal of carbonyl sulfide and any other organic sulfur was not demon-
strated experimentally.
An initial period of incomplete hydrogen sulfide absorption was observed
for all the absorbents tested. This might be due to the evolution of sulfur
dioxide resulting from the reduction of copper sulfate formed in the regeneration
cycle.
The temperature difference between absorption and regeneration steps is
more than 260°C (500°F). This would reduce the overall efficiency of the process.
The absorption reactions are thermodynamically favored at lower temperatures, whereas
the regeneration reaction is favored at higher temperatures. Regeneration at low
temperatures results in the formation of copper sulfate.
3.1.4 HGC Processes Using Zinc Based Solid Absorbents
Three HGC processes using zinc based solid absorbents were identified,
including the Atlantic Refining -I , Catalysts and Chemicals, and IFP Processes.
68
-------
s
1
Uj
FIGURE 22. EOUILIBRIUiM CONSTANTS FOR DF.SULFURIZATION WITH
COPPER-CONTAINING ABSORBENTS
-------
Among these only the IFF Process is on-going, with the research being conducted in
France. Table 7 shows the process characteristics for these three processes. The
Atlantic Refining Process -I is described under the "HGC Processes Using Copper
Based Solid Absorbents" in Sections 3-1-3 and 3.1-3-1 and, therefore, is not dis-
cussed further in this section. Summarized information on the Catalysts and
Chemicals and IFP Processes is presented below. Detailed information for these two
processes is presented in Sections 3.1-4,1 and 3.1.4.2.
• Catalysts and Chemicals Process: Zinc oxide having a minimum surface
area of 30 square meters per gram is employed to remove sulfur compounds from gases
containing steam at temperatures up to 650°C. Bench-scale experiments were termina-
ted a few years ago, and there are no plans to continue the research program.
• IFP Process: Supported zinc oxide is used to remove hydrogen sulfide
from low-Btu gases at temperatures between 400 and 600°C. Regeneration is achieved
with a gas containing oxygen at temperatures between 600 and 900°C. Bench-scale
experiments were conducted in France with both fixed and fluidized beds. The
research program is on-gojng.
3.1-4.1 Catalysts and Chemicals Process; The Catalyst and Chemicals Process
described by Gutmann, et al. (45), employes zinc oxide to remove sulfur compounds
(hydrogen sulfide, carbonyl sulfide and mercaptans) from gaseous streams containing
steam at temperatures between 149 and 649°C (300 and 1200°F). The sorbent is not
recommended for use at temperatures above 649°C because of marked reductions in
sorbent activity (A-ll).
Zinc oxide is a very effective absorbent at temperatures between 21 and
427°C (70 and 800°F) for removing sulfur compounds from gases which do not contain
steam, and is widely used in refinery operations. Information on absorbents con-
taining zinc oxide, e.g. ICI zinc oxide catalyst 32-4, can be obtained from the
open literature (23).
Zinc oxide is the predominant form of zinc in h^S absorption, mainly since
neither h^ nor CO can reduce zinc oxide to zinc. The variation of the equilibrium
constant with temperature for the following desulfurization reaction is shown in
Figure 23-
ZnO + H2S - ZnS + H20.
The kinetics for the desulfurlzation reaction was studied by Westmoreland et al.
(46), and the rate was found to be slower than h^S absorption using either Fe70_,
MnO or CaO. 5
The presence of steam in gas streams reduces the effectiveness of the
zinc oxide and, for this reason, zinc oxide has not been recommended for treatment
of such streams. However, Gutmann, et al., found that at temperatures above 148°C
(300°F), there is a direct relationship between the surface area of zinc oxide and
Its affinity for sulfur compounds in gaseous streams containing steam. The recom-
mended surface area of zinc oxide is generally between 30 and 100 square meters
70
-------
TABLE 7 - HGC PROCESSES USING ZINC BASES
SOLID ABSORBENTS
HGC
PROCESSES
Atlantic
— Ref tni ng ""I
Catalysts &
Chemi ca Is
IFP
STATUS ABSORBENT
Terminated Zn/Cu/Pb
A1 umi no-
s i 1 icate
Terminated ZnO
On-Going Supported
ZnO
ABSORP-
TION
TEMP.
(°C)
93 -
538
150-
650
400-
600
PRES-
SURE
(Pa)
6xl05-,
2.6x10°
METHOD OF
REGENERATION
HO
Air
02 Con-
taining
REGENERATION
TEMPERATURE
/O-v
( c>
370 -
538
600-
900
COMMENTS
Regenerat ion
not studied
Proprietary
process
Gas
-------
JO1
10
JO
s
§
2
JQl
T
(Kf
FIGURE 23. EOUILIERIUM DATA FOR THE REACTION;
ZnO + '
H S =
+ HO
72
-------
per gram. In the case of effluent streams from steam reforming or from the first
stage of shift conversion processes, and area between 30 and 60 square meters per
gram is recommended.
A set of bench-scale absorption experiments (A5) were performed by passing
a "dry" gas stream containing 99.7 percent nitrogen and 0.3 percent hydrogen sulfide
through a fixed bed containing 48 mm ZnO pellets and extrudates at a temperature of
i»00°C and 7.8x105 Pa (100 psig) pressure. The sorbent capacities were 15 percent
and 18.2 percent when the surface areas of zinc oxide were 5 and 30-^0 square meters
per gram, respectively. These experiments show that, in removing hydrogen sulfide
from a gas stream which contains no steam, there is little difference in the capa-
city of high and low surface area zinc oxide for hydrogen sulfide.
Another set of absorption experiments was conducted with a gas mixture
containing 3-5 percent carbon monoxide, 22.5 percent carbon dioxide, 55.0 percent
hydrogen, and 19.0 percent nitrogen at 232°C (i»50°F) and 2.8x10 Pa (kOO psig)
pressure. The hydrogen sulfide concentration in the inlet gas was 50 ppm, and the
steam to gas ratio 0.6. Sorbent capacity was 1 percent at 9 square meters per gram
of zinc oxide, and 6 percent at 30 square meters per gram of zinc oxide. These
experiments show that when steam is present in the gas stream, the affinity of zinc
oxide for sulfur compounds is increased when higher surface area zinc oxides are
employed.
Regeneration of the sulfided sorbent was not attempted. Further experi-
ments were also carried out to remove carbonyl sulfide (COS) with zinc oxide.
However, the tests results are not available.
• Uncertainties and Problems: The information concerning the Catalyst
and Chemicals Process is limited, and the experimental difficulties are not dis-
cussed in the patent. However, it is. known that the process is restricted to the
removal of "trace" quantities of hydrogen sulfide, due to process economics. For
the same reason, regeneration of the sulfided absrobent is not feasible. Operation
below 6^9°C (1200°F) is recommended to prevent the loss of activity of the sorbent.
3.1.^.2 IFP Process: Institut Fran9ais du Petrole (IFP) is developing a process
using supported zinc oxide (ZnO) for the desulfurization of low-Btu gases at temper-
atures between ^00 and 600°C (A-12).
Laboratory tests were performed by passing a synthetic low-Btu gas through
both fixed and fluidized beds containing 100 grams of absorbent in an electrically
heated quartz reactor. In the fixed bed experiments, various metal oxides were
tested, and zinc oxide was found to be active as well as thermally stable. The
chemistry of the absorption reaction is discussed under the Catalyst and Chemicals
Process.
In fixed bed runs of about 100 cycles each, the optimum operating condi-
tions—based upon desulfurization efficiency, absorption capacity and absorbed
life— were obtained at temperatures between *fOO and 600°C. Regeneration of the
sulfided sorbent was achieved with a gas containing oxygen. A low regeneration
73
-------
temperature favored the formation of zinc sulfate, which resulted in the emission of
SQ-2 in the subsequent absorption cycle, whereas a high regeneration temperature re-
duced the absorbent capacity due to crumbling of the bed particles. The temperature
rise in the regenerator was controlled by limiting the oxygen concentration in the
inlet gas.
Fluidized bed experiments were also carried out with a synthetic low-Btu
gas containing 1 percent h^S, at desul furization temperatures of AOO-600°C. The
sorbent particles were 0.3-0.5 mm in diameter and contained 70 percent zinc oxide
by weight. The sulfided sorbent was regenerated with a gas containing k percent
oxygen at temperatures between 600 and 900°C. Table 3 shows the results of a run
which lasted 2200 hours, during which the sorbent was cycled 615 times.
TABLE 0. FLUIDIZED BED ABSORPTION-REGENERATION EXPERIMENT
Sorbent Capacity for h^S exit cone. H2S + S02 Cone, from
Cycle No. 1 500 ppm (mo 1 HgS/kg Sorfaent) Absorber (ppm) _
17' 70
50 3.5 90
200 2.2 100
600 2
• Uncertainties and Problems: The available information on the I FP
Process Is very limited. This process seems 'to be at an early stage of development.
There is no information on the nature of the support material. The use of low-Btu
gas containing tars might severely effect the sorbent characteristics.
3.1-5 HGC Processes Using Other Solid Absorbents
The processes described In this section are those which could not be
included in the categories of solid absorbents discussed in Sections 3-1.1 through
S-l.'i. Two such processes, consisting of the Exxon and Foster Wheeler Processes,
have both been terminated. There is no indication for further development. Table 9
presents the process characteristics of these two processes. Summarized information
Is presented below, while Sections 3.1-5.' and 3.1.5-2 provide more detailed dis-
cussions.
• EXXON Process: Lanthanum oxide deposited on steam-treated and calcined
alumina removes hydrogen sulfide at temperatures and pressures up to 927°C (1700°F)
and 7.8x10? Pa (100 pslg). Regeneration at these conditions is achieved by treating
the sulfided sorbent first with steam and then with an oxidizing gas. Bench-scale
work at Exxon Research and Engineering Company was terminated in the early 1970 's.
• Foster Wheeler Process: Hydrogen sulfide from coal gases is removed
with nickel containing absorbents at temperatures between 538 and 788°C (1000 and
)450°F). The sulfided sorbent is regenerated with air at temperatures between
538 and 732°C (1000 and 1350°F). Bench-scale experiments were terminated in 1976 in
favor of higher priority projects.
-------
TABLE. 9 - HGC PROCESSES USING OTHER
SOI. ID ABSORBENTS
HGC
PROCESSES
Exxon
Foster
Wheeler
STATUS ABSORBENT
Terminated Supported
Terminated Supported
Ni
ABSORP-
TION
TEMP.
149-
927
538-
788
PRES-
SURE
(Pa)
Ixiof-
6x!05
GAS
FLOW
RATE
no rma )
mVh
0.10
METHOD OF
REGENERATION
Steam
Air/02
Air
REGENERATION
TEMPERATURE
927
538-
788
-------
3.1.5.1 Exxon Process: The Exxon Process, invented by Wheelock, et al .
removes hydrogen sulfide and organic sulfur by means of a rare earth metal component
deposited on steam-treated and calcined alumina at elevated temperatures. This pro-
cess was developed as an improvement over an earlier process (k&) that employed un-
steamed support material. Suitable rare earth metal components include metals having
an atomic number 57 through 7' inclusive, or mixtures thereof. The preferred metal
component is lanthanum. The inorganic oxide support is subjected to steam treatment
at temperatures from 38 to 149°C (100 to 300°F) higher than the temperature at which
the process is to be operated. Steam treatment reduces the surface area of the sup-
port by about 20 percent relative to unsteamed support, and requires lesser amounts
of lanthanum sorbent in the sorbent composite, without loss of sorption capacity.
After the impregnation of a lanthanum compound (e.g., La(OH)3) on the support, the
solid is first dried and then calcined at the process operating temperatures. Pro-
moters such as sodium and potassium can also be incorporated into the support (about
1.3 atoms of promoter per atom of sorption active metal) to increase the sorbent
capaci ty .
A schematic diagram of the Exxon Process is shown in Figure 2k. First,
absorption is carried out in reactor A and regeneration in reactor B. The feed
gas is introduced into the fixed-bed absorber at temperatures between 1^9° and 927°C
(300 and 1700°F) and pressures between 105 and 7-8x105 Pa (0 and 100 psig). The
space velocity is maintained between 300 and 1000 vol. gas/vol . supported sorbent/
hour. The supported sorbent removes hydrogen sulfide and acts as a catalyst for the
conversion of carbonyl sulfide to hydrogen sulfide with the following "possible"
mechanism, based on the formation ancj decomposition of thio acids:
COS + H20 = H2C02S
If lanthanum oxide is used as the active material, then the desul furization reaction
is:
Regeneration in reactor 8 is achieved in two steps. In the first step called desorp
tion, the sulfided sorbent is treated with steam or a gas containing steam at absorp
tion temperatures and pressures. The reverse of the absorption reaction takes place.
The amount of steam is about 20-^0 moles per mole of sorbed hydrogen sulfide. The
second step consists of restoring the activity of the sorbent by contacting it with
an oxidizing gas, such as air or oxygen, at temperatures used in the sorption cycle.
The space velocity is between 500 and 2500 vol. air/vol. supported sorbent/hour.
The off-gas from this second step may contain small amounts of sulfur dioxide and,
therefore, be environmentally objectionable for discharging to the atmosphere. In-
stead, they can be used in the process, possibly in the gasifier (A-13) .
When the supported sorbent in reactor A has reached capacity, the regenera
tion (desorption/activation) ©Deration in reactor B is substantially complete and
the reactors are switched; that is, reactor 3 will go on the absorption cycle while
reactor A is on the desorption/activation cycle. Various sorbents and support mater-
76
-------
STEAM AIR
FEED
GAS
1 ' \ r
A
B
CLEAN GAS
REGENERATOR
OFF GAS
FICURF, 2». EXXON PROCESS
77
-------
ials were tested with a bottled gas mixture containing 0.96 percent H?S- These tests
were carried out at temperatures between 538 and 927°C, pressure of 0 psig and space
velocities between 700 and 1000 vol. gas/vol . supported sorbent/hour. The lanthanum
oxide supported sorbents, prepared according to the method proposed by Wheelock, et
al., gave a higher absorption capacity than any other sorbent tested. Hydrogen
sulfide removal efficiency was observed to be around 98 percent.
These experiments were dropped in the early 1970's due to the lack of
internal funding. With the exception of information presented in the referenced
patents, data obtained in these experiments are proprietary (A-13).
• Uncertainties and Problems: The Exxon Process is proprietary, and
available data on this process are very limited. There is no information on the
experimental difficulties encountered during the bench-scale testing. However, it is
indicated that the process is effective for desulfurizing gases with hydrogen sulfide
concentrations less than 3 mole percent. Since coal gases contain much lower H2$
concentrations, the Exxon Process may be effective for cleaning coal gases.
The experiments reported in the patents were performed with lanthanum as
the active material, although the patents claim that any earth metal component with
an atomic number between 57 and 7' can be used for desul furization. The technical
and economic feasibility of such absorbents remains to be demonstrated.
3.1.5.2 Foster Wheeler Process: Steiner (4g) . describes a continuous process for
the selective removal of hydrogen sulfide from coal -derived gases at temperatures
between 538°C and 783°C (1000°F and H50°F) with a nickel -containing material. The
absorbent can be supported or unsupported nickel, nickel alloy (e.g., nickel-alumi-
num), or nickel oxide mixed with other oxides. As the nickel content of the sorbent
material becomes spent through transformation to solid nickel sulfide compounds,
the sulfided sorbent is withdrawn from the bed and regenerated with an oxygen-con-
taining gas such as air in a f luidized-bed reactor at temperatures between 538 and
732°C (1000 and 1350°F). The regenerated sorbent is continuously recycled into the
absorber.
The equilibrium constants for the absorption reactions are shown in Figure
25. Since the reactions are exothermic, lower temperatures yield higher sulfur
removal. A reaction occurring in the regenerator can be represented as:
NixSy + y02 - xNi + yS02.
Bench-scale experiments that started in 197^ were performed by passing
0.10 normal m3/h (3.5 SCFH) of a bottled gas mixture through a sorbent bed in a
22 mm (0.8$") I.D. quartz reactor (A-14,15). The inlet gas contained 4500 ppm hydro-
gen sulfide. The bed removed 85 to 95 percent of the hydrogen sulfide from the inlet
gas, and the reaction was rapid (less than one second). The optimum desul furization
temperature was found to be around 704°C (1300°F). The experiments were later dis-
continued in 1976, and there are no plans for further research.
78
-------
7000
S
o
1
8
FIGURE 25. EQUILIBRIUM CONSTANTS FOR Ni-CONTAINING
SORBENTS
79
-------
• Uncertainties and Problems: No information was available concerning
the problems encountered during the experimental program. It appears that the
process was abandoned at a very early stage. Uncertainties about the process would
include the desulfurization and regeneration rates, activity of the sorbent, sulfur
dioxide concentration in the regeneration off-gases, and corrosion of equipment.
3.2 HGC PROCESSES USING MOLTEN SALT ABSORBENTS
Three HGC processes using molten salt absorbents were identified. All
three processes use either an alkali or alkali-earth metal carbonate to remove sul-
fur compounds, especially H-S, from gases. Unlike solid absorbents, these molten
salts are not subject to attrition and, therefore, should require less make-up.
Also, they provide good contact with the gas. However, because of the corrosive
nature of molten carbonates, special and expensive materials of construction would be
required. In addition, molten salt processes involve certain inherent complications,
such as those associated with materials handling and transfer, process start up and
shut down, and sodium and potassium removal from the desulfurized gases to protect
gas turbine and other components.
The three HGC processes using molten carbonate absorbents are summarized
below and in Table 10. Sections 3-2.1, 3-2.2 and 3.2.3 provide more detailed informa-
tion on these processes.
• Battelle Northwest Process: Sulfur compounds from producer gas are
removed by a mixture of molten sodium potassium, lithium and calcium carbonates at
temperatures above 600°C in a scrubber. Regeneration is achieved with steam and
carbon dioxide. Research was initiated in 1372, and the process was developed
through the PDU level. Process development was terminated in October 1973 due to
the curtailment of funding.
• HRI Process: Hydrogen sulfide from fuel gases is removed by a molten
alkali-metal carbonate (e.g., I^CO?) impregnated within the pores of a refractory
material with a Si02 content less than 2.0 wt. percent. The absorption reaction
takes place at temperatures between 927 and 1093°C (1700 and 2000°F), and pressures
from atmospheric to A.18x10° Pa (600 psig). Regeneration is by means of carbon
dioxide and steam at temperatures above the melting point of the sulfide. Bench-
scale tests at HRI were terminated in 1976 after a patent was issued.
• Pullman Process: A flowing film of molten sodium, potassium and/or
lithium carbonate is contacted with industrial waste gases containing sulfur com-
pounds (H2S, S02, 503) at temperatures between 816 and 1093°C (1500 and 2000°F)
and pressures between 1x105 and 5*10^ Pa (l and 5 atm). Regeneration consists of
solidification, dissolution in water, filtration and dehydration of the sulfided
mel t.
3.2.1 Battelle Northwest Process
The Battelle research program, which started in November 1972 under a
contract with the Office of Coal Research, focused on cleaning hot fuel gas by molten
salt scrubbing. The project was unfunded between December 197^ and October 1§75.
80
-------
TABLE 10 - HGC PROCESSES USING MOLTEN
SALT ABSORBENTS
HGC
PROCESSES
Battel le
Northwes t
HRI
Pul Iman
STATUS ABSORBENT
Terminated Li, Na , K
Ca Carbon-
ates
Terminated Supported
,K2C03/
Na2C03
Terminated Li , Na, K
carbonates
ABSORP-
TION
TEMP.
(°c)
600-
910
907-
1093
816-
1093
PRES-
SURE
(Pa)
IxlO5
1.2xl06
*ixlO°
IxlO5
5x10?
CAS
FLOW
RATE
normal
m3/h
127
0.01
METHOD OF
REGENERATION
H20-C02
H 0-CO
£ t-
Metal
Bicarbonate
- CO,
REGENERATION
TEMPERATURE
(°c)
500-
600
907-
1093
COMMENTS
PDU
Operated
-------
The U.S. Department of Energy then supported this research program until October 1978,
at which time the process development work was terminated. There are no plans for
further research.
The development of the Battelle Process has undergone three phases. First,
bench-scale tests were carried out to study various selected ^S absorbents. Second,
PDU tests were conducted to demonstrate the H2S absorption characteristics under
batch mode of operation. Finally, a PDU was constructed for continuous operation to
demonstrate both the absorption and regeneration steps of the process.
The Battelle Northwest Process is patented by Moore (50).
Bench-Scale Experiments (51): Bench-scale experiments were performed in an
alumina-lined InconelSOO bubbler contactor having a capacity of about I liter of
molten salt and accepting flows up to 10 £ /min. The synthetic produces gas in the
tests was slightly above 10^ Pa (1 atm) pressure and contained 20 percent (by volume)
H2, 20 percent CO, 15 percent C02 , 45 percent N2, and 0.5 percent H2S. Early exper-
iments conducted with a binary eutectic of CaCl2:NaCl as a solvent for CaCOs (or CaO)
at temperatures up to 85QOC, where CaO solubility is 7-3 mole percent, exhibited good
absorption and regeneration performance with respect to H2S . However, the vapor
pressure of NaCl is too high to permit its use in turbine applications.
Tests with molten Na2C03 at around 850°C showed almost complete extraction
of H2S from moist synthetic fuel gas, although corrosion of the Inconel 600 scrubber
was observed. Efficient regeneration could not be achieved at temperatures above
the melting point of Na2C03 (854°C) since chemical equilibria do not favor reversal
of the extraction reaction at such temperatures.
A ternary eutectic mixture of equal weights of Li2C03, Na2C03, and
with a melting point of 393°C (74l°F) also showed excellent extraction of H2S (above
95 percent). Figure 26 shows the dependency of the equilibrium constant, calculated
from experimental results, on temperature for the ternary mixture. The data for
the absorption of H2S with ^2^3 are also included in Figure 12 for comparison.
Regeneration of the spent ternary mixture was achieved with a mixture of 75 percent
H20 and 25 percent C02, at a salt temperature of 854°C. In the first two hours, 84.5
percent of the total H2$ evolved at an average steam utilization of 31-4 percent.
Again, severe corrosion was observed.
Separate corrosion tests were conducted with various alloys in the ternary
eutectic mixture discussed above, sparged with h^S. About 1.5 percent of the carbon-
ate was converted to sulfide. The corrosion test results of two 50-hour exposures
are presented below:
Corrosion Rate, mils/month (1
347SS
Haynes 25
50 Cr-50 Ni
HAPO-20
Haste Hoy F
82
Period 1
0.5
1.5
0.2
0.2
1.1
Period 2 *
1.6
1.2
1.9
1.4
0.6
-------
Another set of experiments were performed with 15 percent (by weight)
mixed with equal parts (by weight) of potassium, lithium and sodium carbonates.
Approximately 50 percent of the CaCO^ was converted to CaS by sparging with n^S.
High density A^O^ ceramics, type 406 SS , and some alonized steels were found to be
corrosion resistant. With alonized 30^ SS , the penetration rate was found to be
2.5AyHm/month (0 . 1 mil /month).
Further bench-scale absorption tests were performed with a molten salt
mixture containing 27.3 percent (by weight) 1(2003, 26.5 percent Na2C03, 27.5 percent
K£C03, and 18.3 percent CaC03, at temperatures between 600 and 910°C and atmospheric
pressure. The experimental data showed that the addition of CaCO^ substantially
improved the removal of hSS. Regeneration of the spent salt was almost complete,
although steam utilization was poor. In two hours, 71-5 percent regeneration was
achieved with 10 percent steam utilization. Further tests were performed with a
reduced concentration of 1)2003 (about 13 percent by weight), since Li2C03 is the
most expensive of the salt ingredients. The result was an absorption of H-S in the
range of 90 to 99 percent, with a maximum of 63 percent consumption of the avail-
able
It was concluded from these bench-scale tests that the mixed alkali
carbonates, containing about 15-20 mole percent CaC03, are excellent sorbents for
H2$. The equilibrium constant calculated from experimental results for the quaternary
mixed carbonates is given in Figure26as a function of temperature. In the salt
mixture, CaC03 is the most reactive ingredient, with the absorption reaction occur-
ring between CaCC>3 and H2S until most of the CaC03 is consumed.
The general reaction between h^S and the molten carbonate can be expressed
as follows:
MC03 (t) + H2S(g) = MS (/) + C02(g) + H20(g),
where M represents the metal. Regeneration can be accomplished by reversing this
reaction. If both the liquid and gas phases are ideal, then the equilibrium constant
for the above reaction can be expressed as follows:
where X; is the mole fraction of compound i in the liquid gas,
YJ is the mole fraction of compound i in the gas phase, and
P is the total pressure.
Batch-Mode PDU Experiments: The design of the PDU for batch-mode operation
(50) is based upon a once-through flow of the gas and cyclic use of the quaternary
salt mixture. The experiment was designed to purify 1.^2-2.12 normal m'/h (50-75
SCFM) of producer gas, generated from Dattelle Northwest's fixed-bed gasifier. The
83
-------
10
Q.
Z
«t
o
u
z:
(J03)108
(io')io6
(10°)105
(io"2)ioj
E \
\
\
\
\
\
\
\
\
\
\
aCQ3 (s)
\
\
\
QUATERNARY MIXED CARBONATES
\
\
\
(Li. No. /O>CO,
k
\
as as ;.o ;.; 12 1.3 i.t
FIGURE 26. TFMPERATURF: DEPENDENCE OF EOUILIBRIUM CONSTANTS FOR
REACTIONS BETWEEN H S AND MOLTEN CARDOMATPS
84
-------
temoerature of the exit gas from the gasifier was in the range of 20A-*»27 C v-"ju -
800°F), and the pressure was 10^ pa (1 atm). Since the producer gas was not quite
as hot as desired, an electrically powered heater was provided to preheat the gas.
The h^S concentration of the gas was adjusted to 0.65 percent (by volume). The
inventory of the salt was 75.7-9^.6 dm' (20-25 gallons) enough to allow k-(> hours
operation. Alonized type 30*» SS was selected as the material of construction. A
block diagram of the PDU is shown in Figure 27.
The gas heater is adjusted to heat the gasifier exit gas from *»27°C up to
900°C. The molten salt mixture, from Pot Furnace 1, flows to the vertically oriented
venturi scrubber, where it is atomized and entrained in the gas. The venturi scrubber
is designed to yield above 99 percent collection efficiency of particles V m or
larger. The salt and gas mixture is carried to the deentrainment and demister sec-
tion located above Pot Furnace 1, under a gravity induced flow. Clean gas flows to
the burner where it is combusted and then vented. For regeneration, the gas flow
is stopped and all salt is collected in Pot Furnace 2, where the salt is cooled to
550-580°C, and reacted with carbon dioxide and steam. The h^S rich off-gas from
Pot Furnace 2 is then contacted with a sodium carbonate solution in a packed bed
scrubber. A sodium hydrosulfide solution is formed according to the following
react ion:
Na2C03(>0 + 2H£ S(g) = NaHSC?) + C02(g) + H20(g) ,
and is then disposed. The regenerated salt is lifted back into Pot Furnace 1 with
ni trogen.
The salt composition was 13-0 percent (by weight) l^CO^, 36.0 percent
^2^03, 37-3 percent ^03, and 13.7 percent CaC03 in these batch PDU tests. At
7k3°C (1380°F), the H2S recovery was between 9*» and 97 percent. This recovery rate
agrees with data developed during the. bench-scale testing. The CaC03 utilization
ranged between 16.7 and 29.5 percent. In a set of runs at higher temperatures
(around 820°C), CaC03 dissociated, forming CaO as a precipitate, after its solubil-
ity was exceeded. In these tests, 99 percent H2S removal was achieved. However,
the presence of COS in the product gas limited overall sulfur removal. No corrosion
of the alonized steel equipment and piping was observed. The particle removal
efficiency, as well as the particle size distribution, was obtained through sampling.
The product gas contained approximately 12 ppm salt, which is about an order of
magnitude higher than the acceptable level for turbine fuel.
Continuous-Mode PDU Tests: The continuous mode PDU (52) was designed to
demonstrate that the salt can be continuously regenerated, simultaneous with the
extraction of sulfur compounds and particles from producer gas. The equipment design
is shown in Figure 28. As in the batch design, the venturi scrubber is used to lift
the gas-salt dispersion to the deentrainer; therefore, no mechanical pump is required.
The inlet gas produced in the gasifier flows at 85-212 normal m3/h (50-75 SCFM) and
its H2S concentration is adjusted to 1.3 percent. The design of the venturi scrubber
is such that the gas velocity at the throat is 76 m/s (250 ft/s). The salt contains
a mixture of lithium, sodium and potassium carbonates. Calcium carbonate was not
incorporated in this mixture because of its tendency to decompose and precipitate
at elevated temperatures.
85
-------
SALT TRANSFER
LINE
FROM _
GASIFIER
**•
«0I
PURGE
CMS
VENTUW
INJECTION
GAS
TO BURNER
*• WASTE TO DRUMS
OR DRAIN
FIGURE 27. BATCH-MODS TEST APPARATUS
-------
EXHAUST TO BURNER
Jj—M—CQ2
GASJFIER
(i)
(2)
(3)
(4)
(5)
(6)
(7)
(8)
(9)
(10)
(ID
(12)
Venturi
Baffled De-Entrainer
Heat Exchanger
Packed De-Entrainer
Extraction Column
Venturi Feed Tank
Emergency Salt Storage Tank
Air Preheator
Regenerator
Salt Make-Up Tank and Reheater
Steam Preheater
Steam and CO Preheater
F1GURF 28. CONTIMdCUS-MODE TEST APPARATUS
37
-------
A portion of the salt removed from the deentrainer by inertial impaction
flows at 3 drnVmin through a cooler and is then introduced into the regenerator,
where it is contacted with H20 and C02 countercurrently. The regenerator is a
152 mm (6-inch) diameter column containing eight bubble cap trays. It operates be-
tween 500 and 600°C and produces a C02-H2S gas mixture, following steam removal, con-
taining 30 percent f^S. The regenerated salt mixture is drawn into the salt make-up
tank, where it is heated to 700°C , and then introduced into the top tray of a 25*» mm
(10-inch) diameter bubble cap extraction column. This column has two trays and is
designed to further desulfurize gas from the venturi scrubber so that the i^S con-
centration in the exit gas is 0 .Qk$ percent. Above the extraction column are a
deentrainer, packed with the alumina particles, demister, and filter. The salt
from the extraction column is returned to the venturi feed tank. Particles are re-
moved from the salt by means of a bleed stream extracted from the venturi feed
tank.
Figure 29 shows a schematic flow diagram of salt bleed recovery system (53).
This proposed process removes solids and recovers t^CQ,, because of its higher cost
than the other salt components, by taking advantage of the differences in solubility
of carbonates and bicarbonates. The molten salt bleed stream is added to an aqueous
phase, containing water and KHCOj, in a quencher-dissolver. The effluent stream
from this unit is filtered and the filter cake is treated with C02 in a carbonator,
where solid 112^03 is converted to soluble LiHC03, and thus separated from the solid
waste. The LJHC03 solution is then stripped with air or steam to recover 112003
as a precipitate.
The liquid from the filter, rich in Na2C03 and K2C03, is carbonated to
NaHC03 and KHC03 (which are much less soluble than the carbonates), and removed as
precipitates. The liquid phase is then recycled to the quencher-dissolver.
Start-up of the continuous-mode PDU (A-lS)occurred in November 1977.
Plugging, due to corrosion of alonized steel pipe and vessels, was experienced during
the initial shake-down tests. A decision was then made to use CaCO-j in the salt
mixture as a corrosion inhibitor. Sixteen runs were performed with the four-component
system. However, experimental problems were observed during the continuous-mode of
operation of the PDU.
The research program was terminated in October 1973 due to the curtailment
of funding by DOE. There are no furture plans for further process development.
• Uncertainties and Problems: The use of a pump was avoided by incorpora-
ting a venturi scrubber which atomizes the molten liquid by means of the product gas
and raises the liquid-gas mixture to a higher level. This process is the first in
which the use of a molten liquid was attempted in a venturi scrubber. In the con-
tinuous PDU shake-down tests, erosion of the throat of the venturi scrubber was
observed. The eroded materials, along with the corrosion materials from the alonized
steel equipment and piping, deposited and plugged the valve and the venturi feed
line. To prevent this problem, a filter trap was designed, but has not yet been
tested. Battelle decided to conduct further PDU testing with a quaternary mixture
of carbonate, introducing CaCO, into the lithium, sodium, and potassium carbonate
38
-------
CO
MOLTEN SALT
BLEED
QUENCHER
DISSOLVER
MAKE-UP WATER
STEAM
•** BYPRODUCT STEAM CW.
-b-
r^WASTE GAS
c
(ORFLUEGASl
->
mi
'CARBONATOR)
"6
WASTE GAS
_COp
(OR FLUE GAS)
(PRECIPITATE)
WASTE GAS
DECARBONATION
STRIPPER/LiCO*
PRECIPITATO
iy*)
R
WASTE
SOLIDS
(PRECIP)
FIGURE 29. SCHEMATIC FLOW DIAGRAM OF SALT BLF.ED RECOVERy
-------
mixture. It was believed that CaCOj would act as a corrosion inhibitor. This, how-
ever, has not yet been experimentally proven while the addition of CaCOj would im-
prove the recovery of I^S, it would make regeneration more difficult and would limit
the maximum operating temperature of the process. In the batch mode PDU operation
at 820°C, CaCO^ dissociated into CaO precipitate, after its solubility in the
quaternary salt mixture was exceeded.
Separation of the molten liquid from the gas is another problem, due to
the nature of the process. In the batch mode PDU operation, the amount of salt re-
maining in the clean gas was about an order of magnitude higher than the acceptable
level for turbine applications. Also, it is assumed that tars and oils originating
from the gasifier will be collected in a tar trap. Again, this assumption has
yet to be tested.
The Battelle Northwest Process has only been tested at 10 Pa pressure.
However, a conceptual design study was carried out for operation at 2xlO^Pa pressure.
It was calculated that desul furization can be accomplished with three to four times
as many stages as required at lO-'Pa pressure, and that regeneration can be achieved
with half the number of stages as required at lO^Pa pressure. The salt bleed section
of the process, proposed by Battelle, was not studied in detail. The molten liquid
and the aqueous solution are mixed in the quencher-dissolver. The design of this
equipment should take into account the serious problem of superheat limit explosions
which can occur when a molten salt is added to an aqueous phase.
In the continuous -mode PDU shakedown test with a three component mixture,
problems such as power failure, malfunction of the coal conveyor, and others
associated with the start-up of the extraction column were observed. These problems
were eventually corrected. Corrosion was also observed during these shakedown tests.
Calcium carbonate was then added to remedy this problem. With the four-component
system, the major difficulty was observed in operating the regeneration column.
Other problems involved controlling flows at low rates and preventing particulates
from plugging the valves. Again, some corrosion and erosion were observed. Finally
during the testing, the ceramic lining of the venturi scrubber was damaged due to
thermal stresses (A-16) .
3.2.2 HR1 Process
The HRI Process (5^,55) is based upon contacting H.S and COS containing
fuel gases with an abrasion resistant, porous solid having molten sodium or potassium
carbonate within its pores. Sulfur removal Is accomplished at temperatures above
the melting point of the carbonate according to the endothermic reaction:
M2C03(0 +H2S(g) -M2S(1) +H2°(g) + C02(g)
where M stands for sodium or potassium. The CQ-2 partial pressure is maintained at
such a level as to prevent the endothermic decomposition of carbonates to oxides.
Regeneration of the spent sorbent is accomplished by reversing the above absorption
reaction.
90
-------
The HRI Process, shown in Figure 30, is described by Wolk, et al (55).
Producer gases generated by the reaction of steam and oxygen with coal or oil are
introduced into a f1uidized-bed absorber for desulfurization. The absorber contains
a strong, abrasion resistant, porous particulate material, such as aluminum oxide,
with potassium carbonate dispersed within its pores. The operating conditions for
the absorber are 927-1093°C (1700-200QOF) and about 1 .5xloM .2x10° Pa (200-600 psig)
pressure. The sulfur compounds in the hot fuel gas are reacted with the molten po-
tassium carbonate within the pores of the particulate absorbent to form potassium
sulfide. The substantially sulfur-free fuel gas is then passed through a gas-solid
separation unit, in which particulates are removed. The resulting fuel gas can then
be combusted with compressed air and expanded in a turbine for power generation.
As the absorption capacity of potassium carbonate is substantially depleted,
the absorbent is regenerated by withdrawing a portion of it from the fluidized-bed
absorber and passing it to the regenerator. There, the absorbent is regenerated in
a fluidized bed with pressurized steam and carbon dioxide. The regeneration tempera-
ture is sufficiently high so as to maintain the potassium sulfide in a molten condi-
tion, such as 927-1093°C. The regenerated absorbent is then returned to the fluid-
ized-bed absorber. The gas stream from the regenerator is further treated to re-
cover sulfur. The process can employ moving-bed or dual-fixed beds instead of
fluidized-bed reactors.
Impregnation of the liquid carbonates into the pores of a strong solid
offers several advantages. First, it reduces the corrosive effects of the melt.
Secondly, absorption and regeneration take place without involving the crystalline
structure of the solid support; therefore, attrition of the support material is not
promoted. Thirdly, this system allows for good gas sorbent contact, and provides
high heat transfer rates to sustain the rapid endothermic desulfurization reaction.
Laboratory experiments were initially performed with unsupported molten
potassium carbonate. These experiments revealed complete removal of h"2S from a gas
containing 5 percent ^S and 95 percent H£ at 927-I093°C, 105 Pa pressure, and with
a space velocity of 460 vol. gas/vol. melt/h . Complete regeneration of the melt
was not achieved with equimolar steam and carbon dioxide because of the presence
of oxygen entering the system with the steam, thereby converting the sulfide to sulfate.
An experiment was also made in a 22 mm (7/3") I.D. alumina reactor at 927 C.
The reactor bed was packed with 6.4 mm (1/4") diameter alumina and silica supported
sorbents containing 12.6 percent I^CO-,. About 11.3 dm3/h (0.4 ft3/h) of gas contain-
ing 5 percent hydrogen sulfide in hydrogen was passed through the reactor at a space
velocity of 460 vol. gas/vol. bed/h . There was no evidence of physical degradation
of solids when the absorption period was over. However, in an experiment in which
the supported sorbent contained 16.3 percent I^COo, physical degradation of the
support material was observed. Formation of the water soluble metasilicate was sus-
pected from the reaction of silica in the support with potassium carbonate:
Si02(s) + K2C03«) -K2O.SI02(s) +C02(gK
Another experiment was performed in the same fixed bed reactor at 954°C (1750°F)
by contacting 18 wt. percent I^CO^ in the same support material with an inlet gas
mixture containing 1.2 percent hydrogen sulfide in hydrogen. More than 95 percent
of the H-S was removed. Regeneration was achieved with equimolar amounts of steam
91
-------
IO
co
AIR/OXYGEN-*
ITrAM ..... I-.....*
rnA! »
G4S
S£
GASIFIER
\
ySOLID
WRATOR
k
/
GAS/SOLID
SEPARATOR
JGLFX
' UAt
ABSORBEN
MAKE-UP
ABSORBER
i
\ ' '
.SPEW
ABSORBENT
\N
"•
r
REGENERATOR
1 1
SULFUR
RECOVERY
SULFUR
ASH
W; 0 C02
FIGURE 30. HRI pnocnss
-------
and carbon dioxide. However, the formation of sulfate could not be prevented and,
during the two following absorption cycles, removal efficiencies dropped to 63 per-
cent and 39 percent. Upon examination of the inside of the reactor after the experi-
ment, migration of the melt from the silica, as well as fusion of particles to each
other (formation of metasi1icate), were observed.
Experimental work was terminated when the patent was issued. There are
presently no plans to do further work on this process due to a lack of funding and
the uncertain potential for commercialization of hot gas cleanup processes in
general (A-17) .
• Uncertainties and Problems: Although both references state that removal
of H£S can be achieved with either potassium or sodium carbonate, or a mixture of the
two, experimental work was carried out only with potassium carbonate. Thus, sodium
carbonate and mixtures of the two carbonates were not tested.
Complete regeneration was never achieved. The reason was apparently due
to the formation of potassium sulfate by the oxygen dissolved in the water employed
to produce the steam used for regneration. This sulfate was then converted to
sulfide by hydrogen:
K2SVf)+"H2(g) =K2S(C) +/*H2°(g).
Formation of metasi1icate due to the reaction between silica and potassium carbonate
was a problem when the support material contained 13.*» percent silica:
Si02 + K2C03 = K2O.Si02 + C02 .
This was prevented by using a support material containing high purity alumina with
a very low silica content (0.4 wt percent).
The problem of migration of melt from the pores of the alumina was never
resolved, but the inventors believe that it can be prevented if both desulfurization
and regeneration are carried out in a f1uidized-bed reactor.
3.2.3 Pullman Process
The Pullman Process, invented be Lefran9ois, et al. $6), was developed to
remove sulfur compounds (F^S, S02, SQj) , nitrogen compounds, and fly-ash from indus-
trial waste gases by means of molten sodium, potassium and/or lithium carbonates at
temperatures between 816 and 1093°C (1500 and 2000°F) and pressure between 1x105 and
5x10^ Pa (1 and 5 atm). The gas and liquid are contacted in chambers, where a moving
film of liquid is deposited on alumina trays, and the gas is directed by means of
baffles. Pressure drop is maintained less than 1x10^ Pa (1.5 psi) to eliminate the
entrainment of liquid droplets into the gas stream. The absorption scheme is shown
in Figure 31 .
The preferred regeneration procedure is described in U.S. Pat. 3,567,377-
According to this procedure, the sulfided melt is diluted with an aqueous solution
of the metal bicarbonate. If necessary, it is filtered to remove solids from the
93
-------
MELT
wo
CLEAN
GAS
REGENERATION
FRESH
MELT
GAS
FIGURE 31. PULLMAN PROCESS
-------
resulting solution. Then, this solution is passed into a two stage carbonation zone.
In the first stage, any carbonate in the solution is converted to bicarbonate preci-
pitate by contact with a carbon dioxide gas. This treatment results in a more viscous
and less alkaline solution. The solution is then passed to the second stage, where
it is contacted with carbon dioxide gas for removal of sulfur values as hydrogen
sulfide. Finally, a portion of the bicarbonate is converted to carbonate, which is
heated and recycled for absorption.
An experiment was performed at about 9^8 C (17^0°F) by bubbling a gas mix-
ture of 1 percent hydrogen sulfide in nitrogen through molten sodium carbonate at
a velocity of 15 cm/s (0.5 ft/s) and a pressure drop of 10 Pa. The exit nitrogen
from the demister zone was initially found to contain less than 5 ppm of carbonate
and hydrogen sulfide.
• Uncertainties and Problems: The information obtained on the Pullman
Process is very limited. A corrosion problem was observed during the testing. After
30 minutes of operation, the stainless steel demister filter required replacement
due to heavy corrosion. Experimental data concerning the regeneration of the sulfided
melt is not given in the patent from which this information was obtained.
3.3 HGC PROCESSES USING MOLTEN METAL ABSORBENTS
Only the IGT-Meissner Process has been identified in this group. This
process, under development by the Institute of Gas Technology, is ongoing at the
bench-scale level. In this process, molten lead is sprayed into a reducing gas con-
taining hydrogen sulfide at temperatures between 327 and 1200°C (621 and 2192°F).
The lead - lead sulfide slurry formed from the absorption reaction is electrolyzed
to regenerate lead and obtain sulfur. In general, this process has similar advantages
and di sadvantages of processes using molten salt absorbents. A detailed description
of the IGT Process is presented in Section 3.3-1-
3.3.1 IGT-Heissner Process
The IGT-Meissner Process employs molten lead to remove hydrogen sulfide from
hot reducing gases. This proprietary process is being developed and funded by the
Institute of Gas Technology (IGT) in conjunction with its U-Gas Process (57). Inform-
ation from IGT (A-18) reveals that the present lab-scale research is basically along
the same lines as the process described in the patent, issued to Meissner (58).
The desulfurization reaction between the molten lead and the hydrogen
sulfide in reducing gases can be represented as follows:
Pb + H2S = PbS + H2 (1)
Figure 32 shows the variation of the equilibrium constant for this reaction with
respect to temperature. Whether or not the incoming gas is reducing depends upon
the mole ratios of C02/CO and H^O/hL. These ratios are important in minimizing the
formation of lead oxide according to the following reactions:
PbO + CO = Pb +C02 (2)
PbO + H, = Pb +H90 (3)
L L 95
-------
o
o
o
Ui
vt
10
S
FIGURE 32. EQUILIBRIUM CONSTANTS FOR DFSULPURI7.ATION WITH
LEAD ABSORBP:NTS
96
-------
The two ratios are related through the water gas shift reaction:
H20 + CO = H2 + C02 CO
Equilibrium information on reactions (2) and (3) is also included in Figure 32.
A schmatic flow diagram of the IGT-Meissner Process is shown in Figure 33.
Desul furizat ion can be achieved by various contact methods at temperatures
between 327°C (621<>F, melting point of lead) and 1 200°C (2192°F). The preferred
method, described in the patent and shown in Figure 33 consists of spraying a shower
of droplets through the flowing gases passing through the reaction chamber. A
rotating disk or paddle wheel, partially immersed in the metal pool in the bottom
of the chamber, can be used to spray the droplets through the gas. Providing that
the H20/H2 and C0.2/CO ratios are of proper magnitude, undesirable oxides do not form
and molten lead retains its ability to react with sulfur. The lead sulfide which
is formed dissolves in the liquid until the saturation point is reached, at which
time additional sul f ide preci pi tates . The amount of sulfide formed is important
with respect to the pumpability of the slurry. Figure J>k shows a temperature/
composition diagram for the lead-lead sulfide system.
Regeneration can be achieved by various methods. One method involves
roasting the lead sulfide to oxide, and then reacting the oxidized lead compound with
more lead sulfide to form metallic lead and sulfur dioxide. Lead oxide may also be
reduced to lead when reacted with carbon, hydrogen or carbon monoxide. The regenera-
tion method discussed in detail in the patent and shown in Figure 33 utilizes
electrolysis. The cell uses molten lead as the cathode and a carbon rod as the anode.
The carbon rod is dipped in an electrolyte layer of lead chloride or an alkali metal
salt (such as potassium chloride or sodium sulfide), in which lead sulfide is sol-
uble. Sulfur forms at the anode, either in liquid or gaseous form, depending upon
the operating conditions. Molten lead in the melt is returned for absorption in
amounts sufficient not only to react with all the sulfur compounds, but also provide
a pumpable slurry with lead sulfide present.
An experiment was performed at A26°C (800°F) by passing fuel gas containing
0.3 percent h^S and 0.1 percent COS through a spray of molten lead. The amount of
lead recycled to the desul furi zation chamber was 0.82 kg/nr3 (51 lb/1000 ft^) of gas
treated. Sulfur removal efficiency was above 50 percent in this experiment. Re-
generation was achieved by electrolysis in a chloride solution, with sulfur being
released in gaseous form.
• Uncertainties and Problems: Since the IGT-Meissner Process is proprie-
tary, the available information on this process is very limited. The major anticipated
problems include corrosion of vessels, transfer lines, pumps and valves. Difficulties
in equipment shut-downs due to the solidication of the molten metal and plugging of
the transfer lines, pump and valves might also be expected.
97
-------
ABSORBER
REDUCING GAS
PURIFIED GAS
CATHODE /
REGENERATOR
VOLUGE
SOURCE
FIGUPF 33. IGT-MEISSNER PROCESS
-------
ATOMIC %S
FIGURE 31*- TEMPERATURE/COMPOSITION DIAGRAM
FOR LEAD-LEAD SULFIDE SYSTEM
99
-------
SECTION k
APPLICABILITY OF HGC PROCESSES TO COAL GASIFICATION SYSTEMS
Conversion of coal into clean fuel gas requires several process steps as
illustrated in the simplified block diagram presented below.
COAL
GASIFICATION
PARTICULATE
REMOVAL
SULFUR
REMOVAL
CLEAN
FUEL GAS
The first step is the coal gasification process itself. The gasification
of coal can produce low, intermediate or high Btu gas which contains impurities
such as particulates and sulfur compounds. These impurities can be objectionable
either because of environmental regulations or process restrictions. The removal
of these impurities can be achieved either at low or high temperature. Low temp-
erature cleanup requires quenching of the gasifier product gas, as well as removal
of sulfur compounds through commercially available processes. On the other hand,
high temperature removal processes are designed to operate at temperatures above
(800°F).
The applicability of HGC processes to a variety of coal gasification
systems is assessed in this section. Various coal gasification processes are
reviewed, possible end uses for low and intermediate Btu gases are evaluated, the
types of gasifiers and their potential end uses are matched, and finally the appli-
cability of the hot gas cleanup processes described in Section 3.0 to the gas-
ifier-end use pairs is evaluated. The use of HGC processes to clean high Btu gases
is not considered in this study, since it would not be economically possible to
transport these gases at high temperatures over long distances.
4.1
GASIFIERS
Many gasifiers have been developed, and some have been commercially used
to produce low and intermediate Btu gases from coal. Table 11 shows a list of
promising low and intermediate Btu gasifiers and their characteristics. Low and
intermediate Btu gasifiers can be classified as:
100
-------
TABLE 11
PROMISING LOW AND INTERMEDIATE 3TU GASIFIERS
TYPE OF GASIFIER
Fixed Bed
Lurgi
Wei 1 man-Galusha
Chapman (Wilputte)
Woodal1-Duckham/Gas Integrale
Foster Wheeler/Stoic
MERC
BGC/Lurgi Slagging
GFERC Slagging
Riley Morgan
GASIFICATION MEDIUM
(LCM Btu) (Int. Btu)
Air/H20 02/H20
x
x
X
X
X
X
X
X
X
X
X
X
X
X
PRESSURE
Low H i gh
x
X
X
X
X
X
X
Fluidized Bed
Winkler
Entrained Bed
Koppe rs-Totzek
Coalex
Bi-Gas
Texaco
Foster Wheeler
x
x
x
x
x
x
x
x
x
x
x
X
101
-------
o fixed bed gasifiers,
o fluid!zed bed gasifiers, and
o entrained bed gasifiers
Below is a brief description of each generic gasifier type, including the
characteristics of the pollutants present in the raw product gas.
4.1.1 Fixed Bed Gasifiers
In the fixed or moving bed gasifiers, coal is introduced at the top of
the gasifier and contacted countercurrently with a mixture of air or oxygen and
steam. This gas mixture is admitted to the gasifier beneath the grate supporting
the coal bed. The bed is maintained at a fixed level and the flow of solids can
be described as essentially plug flow. The gas solid contact results in various
zones in the bed. At the top zone coal is heated and devolatilized. Below that
is the gasification zone where carbon in the coal reacts with steam and carbon
dioxide to produce carboji monoxide and hydrogen. The third and bottom zone is the
combustion zone, in which a portion of coal is burned to provide the heat necessary
for the endothermic reactions in the gasification and devolati1ization zones.
Fixed bed gasifiers have the advantage of providing high carbon conversion with
high energy efficiencies, but are generally limited to non-coking coals at speci-
fied size ranges.
Several fixed bed gasifiers including the Wei 1 man-Gal usha, Chapman
(WJlputte), Woodall Duckham/Gas Integrale and Lurgi have been commercially used.
Among these, the Lurgi gasifier is the only one employed at high pressures.
The distribution of sulfur compounds in the exit streams from a Lurgi
gasifier tested on American coal was determined as a percentage of the sulfur
present in the feed coal. The sulfur balance is shown below (59):
Percent of inlet sulfur
H2S in gas stream 92.0
Organic sulfur compounds in gas stream 1.5
Sulfides in condensate 2.5
Tars/oils 2.5
Gasifier bottom ash 1.5
The gasifier bottom ash showed the widest variation, containing up to
6% of the feed sulfur in coal in some cases. About 1 to 2% of the feed sulfur was
in the form of organic sulfur compounds.
Fixed bed gasifiers also produce tars and oils which condense at low
temperatures. The amount of these hydrocarbons and organic compounds in the raw
102
-------
gas depends upon the type of coal feed and gasifier operating conditions. Lower
gasifier temperatures and pressures enhance the formation of heavier hydrocarbons.
About *»0 to 75 grams of condensable hydrocarbons per kilogram (80-150 Ib/ton) of
moisture and ash free (MAP) coal, or 16-3^ g/normal m^ (7-15 grains/SCF) of raw
product gases, is reported from tests with American coals in a Lurgi gasifier (60).
Fixed bed gasifiers generate lower particulate loadings in the raw pro-
duct gas than entrained and fluidized bed gasifiers. Particulate loadings from
fixed bed gasifiers range from about 0.9 to 5.7 g/normal nr (0.4 to 2.5 grains/SCF).
The size distribution data in the Lurgi runs on American coals indicate that 40-80%
of particulates is under lOOyMm (60) . Other data obtained from the MERC fixed bed
gasifier (61) tests show that 50-70% of the particles collected in cyclones are
under 150/Wm in size. These particles contain about 75 to 80%, and sometimes up
to 90% of devolati1ized coal char, the remainder being coal ash. The Lurgi data
show that coal ash in the raw gas represents 0.5% of the ash in the feed coal (60).
The fate of nitrogen in the feed coal was also determined from tests on
American coals in a Lurgi gasifier (60). About 60% of the nitrogen in the feed
coal appeared in the raw gas as NH,, 35% as N2 and the rest in tars and oils. It
was also found that the raw gas contained 5-50 ppm by volume HCN and that the
bottom ash from the gasifier did not retain any nitrogen.
J».1.2 Fluidized Bed Gasifiers
In a fluidized bed gasifier, the gas flows upward through the coal
particles at a high enough rate to maintain the bed in suspension. The vigorous
solids movement in the gasifier enables intimate solid-gas mixing, which provides a
uniform bed temperature. Gasification reactions are primarily a combination of
combustion and water-gas reactions at temperatures of 816 to 1093°C (1500 to 2000°F).
Substantial quantities of char are produced when gasification is not achieved to
completion. The advantages of fluidized bed gasifiers are their high gasification
rates and excellent heat transfer characteristics which prevent the formation of
clinkers, and ability to handle a large variety of coals with respect to quality
and size.
Winkler is the only fluidized bed gasifier which has been commerically
used. However, much more information is available from fluid bed gasifiers under
development than from the Winkler gasifier.
There are limited data available concerning the distribution of sulfur
compounds in various streams from fluidized bed gasifiers. The percentage of
sulfur gasified is reported to be greater than the percentage of carbon gasified (62)
Data from the Winkler gasifier show that more than 96-97% of the feed sulfur is
present in the raw gas (63) (64). The sulfur compounds in the raw gas are primarily
in the form of ^S, with some organic sulfur compounds (carbonyl sulfide, methyl
mercaptan and ethyl mercaptan) also present (65)•
Due to its high operating temperature (about 1700°F), the Winkler gasifier
does not produce any tars and heavy hydrocarbons, and the presence of phenols and
light hydrocarbons in the raw gas has not been reported (63) (66). Similar results
were obtained from tests with the bench-scale Hygas fluidized bed gasifier (65).
Data indicate that about 38 grams of light oil (benzene, toluene) per kilogram of
dry coal feed may be present in the raw gas. However, in tests with the bench-scale
103
-------
Synthane fluidized bed gasifier (steam/oxygen gasification), it was found that coal
type and coal injection geometry significantly affect the amount of tars produced.
For example, 8-38 g tar/kg MAP coal was obtained with North Dakota lignite,
whereas 45-80 g tar/kg MAP coal was obtained with Illinois #6 and Pittsburgh
coals (67) .
Fluidized bed gasifiers produce high particle loadings in the gas streams
due to the high gas velocities of 1.5-3-' m/s. Data from tests with Winkler gasifier
show particle loadings of about Il4g/normal m3(50 grains/SCF). These particles contain
about 70% ash and 30% carbon (63). Similar results are reported from tests.with
the Hygas gasifier (68). However, particle loadings of 4.6-11.4 g/normal m
(2-5 grains/SCF) are reported from tests with Synthane gasifier (69), which does
not gasify coal to completion but produces a char by-product.
No information is available on nitrogen compounds from the Winkler
gasifier. Conversion of up to 70% of the feed nitrogen from coal to NH. is
reported for Hygas (70) and Synthane (67) processes.
4.1.3 Entrained Bed Gasifiers
Entrained bed gasiffers use pulverized coal, and are characterized by low
inventory of fuel in the reaction zone. The temperature and, therefore, the gas-
ification rate decrease from inlet to outlet of the gasifier. For this reason it
is not economically feasible to gasify more than 85 to 90% of the carbon in a
single pass. However, higher carbon conversions can be achieved by recycling the
sol id residue.
The major advantage of entrained bed gasifiers is their ability to handle
any type of coal with no concern about the particle sizes and caking characteristics.
Also, the overall energy production rate per reactor volume is higher than for other
types of gasifiers. However, the need for recovery of heat and chars from these
gases to achieve high overall thermal efficiency are drawbacks for the entrained bed
gasi fiers.
Among the various entrained bed gasifiers, the Koppers-Totzek gasifier
is the only one that has been used commercially. Gasification has been carried out
for the production of hydrogen in ammonia synthesis. There are no commercial
installations of the Koppers-Totzek gasifier in the United States.
Information on impurities from entrained bed gasifiers is very limited.
For the Koppers-Totzek gasifier (71), sulfur in the feed coal appears mainly in
the raw gas stream as h^S with some COS and S02- Condensable hydrocarbons are not
present since they are cracked to gaseous constituents at high gasification temp-
eratures above 1093°C (2000°F). The particulate loading in the raw gas from the
Koppers-Totzek gasifier is reported to be 27.5-55.0 g/normal nr (12-24 grains/SCF).
It is also reported that the raw gas contains both NHj and HCN, although thermo-
dynamics favors the decomposition of ammonia into nitrogen and hydrogen at high
temperatures.
104
-------
k.2
POTENTIAL END USES FOR LOW AND INTERMEDIATE BTU GASES
In 1976, the electric utility and industrial sectors together contributed
55 percent of the energy consumed in the United States. Natural gas and petroleum
supplied 80 percent of the energy input to industry and 30 percent to the electric
uti 1 ities(72l The use of coal-derived gases to replace natural gas and petroleum in
these areas could have important economic benefits in the United States, in addition
to reducing the nation's dependency on foreign, unreliable sources of energy. The
following sections examine how these two energy sources are used in the industrial
and electric utility sectors, and whether they can be replaced by low and inter-
mediate Btu gases.
Jt.2.1 I ndus trial Appl icat ions
Natural gas and petroleum provide SO percent of the energy input to
industry, primarily for use as:
o fuel for direct process heat,
o fuel for small and industrial size boilers to raise steam,
o fuel for gas turbines, and
o reducing gas for the reduction of iron ore and regeneration gas
for some catalysts and sorbent materials; synthesis gas.
It.2.1.1 Pi rect Process Heat; Low and medium Btu coal gases appear to be appli-
cable as fuel for direct process heat in the petroleum refining, chemical and
allied products, primary metals, glass, cement, lime, ceramics, and possibly other
i ndustries.
In the petroleum refining industry, low and intermediate Btu gases have
potential application in some areas where natural gas or oil is presently fired.
They can be especially used in tube-type furnace heat exchangers which are designed
to heat the feedstocks for various operations such as distillation, cracking and
reforming.
In the chemical and allied products industry, the bulk of process heat
from fossil fuels is used in firing furnaces. These furnaces have tube-type
heat exchangers fired externally, with chemical feed inside the tube in much the
same manner as watertube boilers, and would be adaptable to either low or inter-
mediate Btu gas firing.
In the iron and steel industry coal-derived gas was used in open hearth
furnaces about fifty years ago to provide the heat required for the production
of steel from iron. Today, natural gas and oil are used for this purpose.
Shifting to coal gases again probably would not cause any major difficulty.
Another area where coal gases can easily be substituted for natural gas is in pre-
heating the air combusted in blast furnaces for the production of pig iron from
iron ore. This represents about 10% of the total energy used for preheating the
105
-------
air (73). Also, about half of the energy required to heat the steel ingots prior
to being forged, drawn or extruded into desired shapes, is currently supplied with
natural gas, which can be replaced with coal gases with no major difficulty.
About 80 percent of the energy input to ferrous foundries is used in
melting ferrous charges in furnaces. These furnaces include cupolas (coke, fired),
open hearths (gas and oil fired), air furnaces (pulverized coal fired), rever-
beratory furnaces (gas and oil fired) and electric furnaces. Open-hearth and
reverberatory furnaces would be the two potential applications for low and inter-
mediate Btu coal gases, which represents about 15? of the total energy usage in
ferrous foundries.
Among the primary nonferrous metal industries, the aluminum industry con-
sumes the most energy, primarily in the form of electricity and natural gas to
calcine alumina trihydrate to anhydrous alumina in rotary kilns at about 1200°C.
Other major industries with extensive use of oil and natural gas include the pri-
mary copper, zinc and lead industries. Conversion from natural gas and/or oil to
coal gases would not present great difficulties.
In the glass industry, about 75-80? of the process energy is consumed
for glass melting and is provided by either natural gas or electricity. Conversion
of the melters to low energy gas from air-blown gasifiers would be difficult, and
the trend is towards electric heating.
In the cement industry, most of the energy is consumed in rotary kilns
for the production of clinker, and is provided" by pulverized coal, oil or natural
gas. Conversion from natural gas or oil to low or intermediate Btu coal gases
would require readjustment of the combustion process in the kiln, to maintain a
desired flame pattern.
In the lime industry, low or intermediate Btu coal gases could be used
in rotary-hearth kilns to calcine limestone or dolomite. Retrofitting of these
kilns would require readjustment of the combustion process, which would be less
critical than in the cement industry, since combustion takes place outside the
kiIns.
The ceramics industry produces a variety of products, including brick,
clay refractories, ceramic tile and pottery. Low or intermediate Btu gases could
be used for firing kilns and various baking and dryi.ig operations. Conversion
from the present fuel sources (natural gas and oil) to low and intermediate Btu
gases from coal should not present any serious difficulty.
k.2.}.2 Industrial Boilers: Industrial boilers represent another potential
application of low and intermediate Btu gases, since at least half of the energy
input to the industrial sector is used for steam generation, either for process
use or on-site electricity generation (59). Industrial boilers can be classified
into two types: (1) fire-tube boilers, where combustion gases are passed through
tubes immersed in water, and steam is raised outside of the tubing; and (2) water-
tube boilers, where steam is raised by passing water through the tubes, with com-
bustion taking place outside of the tubing. Fire-tube boilers have sizes of
159-12,530 kg steam/h (350-27,600 Ib/h), and can withstand pressures up to
106
-------
2 x 10 Pa (300 psi). Water-tube boilers have sizes ranging from 4,540 to over
45^,000 kg steam/h (10,000 1 000,000 Ib/h) , and are suited for pressures between
several hundred and 1.36 x IO/ Pa (2,000 psi).
Conversion of industrial boilers burning natural gas or oil to low or
intermediate Btu gases would require development of suitable burners. Also, due
to the increase in gas volumes, especially for low Btu gases, pressure drops would
be high and may require redesign of the boiler.
4.2.1.3 Gas Turbines: Gas turbines burn fossils fuels in the combustor section
of the turbine to generate electricity. Gas turbines used in industry are more
rugged and can handle lower grade fuels than those in a!rcrafts, which operate with
fuel gases at pressures between 7.8 x 10^ - 1.6 x 10^ Pa (100 and 200 psig) and
have very stringent fuel gas specifications to prevent corrosion and erosion of
turbine materials. Fuel specifications for industrial gas turbine applications are
in the nature of engineering estimates. Westinghouse Electric Corporation projects
a maximum allowance of 0.041 g/normal m^ (0.018 gr/SCF) for particle loading (in
the inlet fuel gas to the turbine), with no particles greater than 6/*m (74) . The
maximum allowable amount of sulfur, as estimated by the United Aircraft Corpo-
ration (75), is 0.18 mole percent in the fuel gas. In the presence of sulfur
compounds, sodium and potassium in the inlet fuel gas form sulfates, which promote
corrosion of the turbine blades at temperatures above 1093°C (1200°F). Vanadium
oxidizes to vanadium pentoxide, which catalyzes the corrosion of the turbine at
temperatures above 8l6°C (1500°F). Particulates not only cause corrosion and
erosion, but also form hard deposits on the turbine blades and reduce the energy
conversion efficiency.
Retrofitting of combustor-turbine systems would be easier for intermediate
Btu than low Btu gases. For gases having heating values less than 7.45 MJ/m'
(200 Btu/SCF) new combustors may be necessary, whereas for gases with heating
values less than 5.591 MJ/m^ (150 Btu/SCF), the turbine would be derated due to the
lack of sufficient volume to handle the combustion products (73). Furthermore, high
pressure fuel gases from coal gasification units would probably require considerable
treatment to meet the turbine specifications.
4.2.1.4 Other Industrial Applications: Low and intermediate Btu gases can also
be used as a reducing agent in some industries. In the iron and steel industry,
iron ore is converted to pig iron in blast furnaces, with coke being employed as the
reducing agent. Natural gas is also injected to reduce the coke requirement and
represents about half the energy needed in blast furnaces. If low or intermediate
Btu gases are to be used instead of natural gas, they should be at a high temp-
erature and contain practically no C02 or water vapor to minimize endothermic
effects.
Low and intermediate Btu gases can be used in ammonia and methanol
synthesis. For ammonia synthesis, a product gas rich in carbon monoxide and hydro-
gen is required. For methanol synthesis, a nitrogen free gas with a hydrogen to
carbon monoxide mole ratio of 2 is desirable.
Low and intermediate Btu gases can be used in the regeneration of cat-
alysts. In the petroleum industry, catalysts employed in catalytic cracking
107
-------
operations are regenerated by burning off the carbon deposited from the oil at a
temperature of around 593°C OlOO°F) and pressure of 1.7 x 10$ Pa (10 psig).
For some cases in which additional fuel is required to maintain these conditions,
low or intermediate Btu coal gases could have potential as a fuel. Low and inter-
mediate Btu gases can similarly be used to regenerate absorbent materials used in
flue gas desulfurization systems.
k.2.2 Electric Utilities
Low and intermediate Btu gases can be used in:
o Base and intermediate load power plants to raise steam in boilers,
o Peak load power plants to generate electricity by means of gas
turbines, and
o Combined cycle power plants.
A.2.2.1 Base and Intermediate Load Power Plants: In a base or intermediate
load power plant, the fossil fuel (coal, natural gas or oil) is burned with air at
atmospheric pressure in a boiler, with the heat of combustion being used to produce
high-pressure steam which is expanded in a turbine. The boilers are generally
similar to the industrial watertube boilers previously discussed, but are more
complicated, much larger and operate at high pressures. Heat transfer in the boilers
is provided by two mechanisms: radiation and convection. Radiant heat is transferred
directly from the hot flue gases to the banks of tubes which are placed downstream
from the combustion zone. The balance between these two heat transfer mechanisms
is an important parameter in determining the outlet temperature of steam.
Retrofitting of existing fossil-fueled electric plants to low or inter-
mediate Btu gases appear to have very limited, if any, potential. First,
intermediate load plants are operated with frequent shut-downs and start-ups,
which would require that the gasifier output be closely controlled to match the
power plant demand. For this reason, base load plants which operate continuously
are more suited for retrofitting. Second, switching from fossil fuels to low or
intermediate Btu gases would create an imbalance in the heat transfer mechanisms in
the boiler. This is due to the fact that low and intermediate Btu gases have lower
flame luminosity than fossil fuels, and therefore would provide less radiant
heat transfer. However, convective heat transfer would be improved. Third, flue
gases generated from the combustion of low and intermediate Btu gases are 5 to 12
percent and 30 to 55 percent greater, respectively, than those resulting from
fossil fuels. The existing induced and forced draft fans might not be able to
handle the increased volumes of flue gases. Since boilers designed for coal are
less compact and have wider tube spacings than those designed for natural gas or
oil, they would be more attractive for retrofitting to low or intermediate Btu gases
than boilers designed for other fuels. However, a study performed by the Tennessee
Valley Authority (76) for EPRI shows that firing low Btu gases in retrofitted
power plants is not economically competitive with burriinq coal in conventional
power plants and then cleaning up the stack gases. Therefore, the economic in-
centive appears to favor coal fired conventional power plants, rather than low/
intermediate Btu gas-fired power plants.
108
-------
It.2.2.2
peak electrical
downs than the
Peak Load Power Plants:
Peak load power plants use gas turbines to meet
demands. They are subject to more frequent start-ups and shut-
intermediate load turbines previously discussed.
Gas turbines used in utilities require gases at pressures greater than
apply
low or
able in coal
Pa (200 psig). The fuel specifications given for industrial gas turbines
r gas turbines in utilities as well. Retrofitting of these gas turbines to
ntermediate Stu gases would require sophisticated controls not yet avail-
gas i ficat ion.
if.2.2.3
to have good
Comb ined Cycle Power Plants:
Combined cycle power generation appears
potential for low and intermediate Btu gas use. In a combined cycle,
electricity is produced both by expansion of pressurized gases in a gas turbine and
by expansion of steam in a steam turbine. A simplified combined gas-steam turbine
system is shown in Figure 35- Low or intermediate Btu gases are burned and then
Low/Intermediate Btu Gas
1
Air.
Compressor
Gas
Turb ine
Electric
Generator
Steam
Boiler
Flue Gas
Electric
Generator
FIGURE 35. COMBINED CYCLE WITH UNFIRED WASTE HEAT RECOVERY SYSTEM
109
-------
expanded in the gas turbine, which generates electricity and also drives the air
compressor. The flue gases from the turbine are used to deliver heat to the steam
boiler, and then released to the atmosphere via the stack. Electricity is also
generated from the steam cycle by the expansion of steam in the steam turbine.
Combined cycle power plants (CCPP) are generally more efficient than
comparable conventional power plants. A recent study (77) showed that a complete
coal gasification - CCPP system would cost up to 15 percent less to build and
20 percent less to operate than conventional coal fired plants. Combined cycle
plants have been built and operated in Europe, some of which are the NEWAG plant
(55 MW) in Korneuburg, Austria, the EBV-Krafwerk Anna plant and the Hohe Wand
plant (75 MV). In addition, there are at least six other combined gas/stream tur-
bine plants in West Germany (78). However, the only CCPP using coal-derived
gases is the KDV plant in Lunen, West Germany. This plant gasifies coal at about
2 x 10 Pa (20 atm) in a Lurgi gasifier, and uses a supercharged boiler before
expanding the gas in a gas turbine.
k.3 GASIFIER-END USE COMPATIBILITY
As a result of the discussions from the previous section, it appears
that low and intermediate Btu gases have their greatest potential for the following
applicat ions:
o Combustion fuel
o Combined cycle fuel, and
o Synthesis/reducing gas
Low and intermediate Btu gases to be used as combustion fuel can be pro-
duced at a site within a relatively short distance from the industrial user. This
is usually referred to as "on-site combustion". For this application, space and
facilities are required for coal storage, preparation, gasification and gas puri-
fication. The selection of the gasification pressure depends upon the user's
requirements as well as the economics of the system. In order to keep the scope as
general as possible, gasifiers operating at both atmospheric and high pressures are
considered in this study for on-site combustion applications. The gasifiers con-
sidered for on-site combustion applications are shown in Table 12.
Off-site combustion refers to a direct combustion process which consumes
the gas at a site distant from the gasifier (e.g., 50 or 150 km). This might re-
quire a central gasification plant connected to various users via pipelines. For
off-site combustion applications, pressurized gasification seems to be more suitable
than gasification at atmospheric pressure due to less compression requirements of
the product gas from the gasification plant to the industrial users' sites. Also,
oxygen blown gasification and transportation systems seem to be more economical than
air blown gasification and transportation systems because of the high cost of
transporting large volumes of gas. However, since the economics depends upon the
transportation distance, both atmospheric and oxygen blown gasifiers are considered
in this study for off-site applications. Table 12shows that Lurgi, MERC, BGC/Lurgi
110
-------
TABLE 12
GASIFIER - END USE COMPATIBILITY
L: atmospheric pressure
P: high pressure
A : a i r b 1 own
0: oxygen blown
— ~~~~— -__ ^n<^ ^ses
~-— -~-__^ 0 n - S i te
Gas i f i e rs Comb us t i on
«/el 1 man-Gal usha LA
LO
Chapman (Wilputte) LA
LO
iVoodal 1-Duckham/ 1 LA
Gas Integrale 1 LO
.urgi PA
PO
i
Foster Wheeler/Stoic LA
LO
MERC J PA
3GC/Lurgi Slagging PO
GFERC Slagging PO
Riley Morgan LA
LO
Winkler LA
LO
Koppers Totzek •;
Coal ex i LA
1
Bi-Gas PA
i PO
Texaco ;
Foster Wheeler ', PA
j PO
i
Off-Site
Combust ion
PA
PO
PA
PO
PO
i
PA
i PO
PA
j PO
Comb i ned
Cycle
PA
PO
PA
PO
PO
(
i PA
PO
PA
PO
Reducing/
Synthesis
Gas
LO
PA
PO
111
-------
Slagging, GFERC Slagging, Bi-Gas and Foster Wheeler Gasifiers can be used for
off-site applications.
Combined cycle power generation requires expansion of high pressure fuel
gas in the gas turbine. For this reason pressurized gasifiers are most appli-
cable. Lurgi, MERC, BGC/Lurgi Slagging, GFERC Slagginq, Bi-Gas, Texaco and
Foster Wheeler Gasifiers can be used for this application. Both oxygen and air
blown gasifiers are considered for combined cycle applications.
For such applications as reducing or synthesis gas, the Koppers-Totzek
and Texaco Gasifiers are best suited, since their product gases contain large
concentrations of carbon monoxide and hydrogen and small concentrations of hydro-
carbons .
A. k APPLICABILITY OF HGC PROCESSES FOR SELECTED GAS I Fl ER-ENF) USE PAIRS
Seven gasifiers from Table H were selected for the purpose of assessing
the applicability of the HGC processes over a wide range of gasifier-end use pairs.
The selected gasifiers include Lurgi, WeiIman-Galusha, MERC, BGC/Slagging, Winkler,
Bi-Gas and Foster Wheeler. The first four of these are fixed bed gasifiers, whereas
Winkler is fluidized bed and the last two are entrained bed gasifiers. Operation
with both air and oxygen were considered for the Lurgi and Foster-Wheeler gasifiers,
thereby increasing the number of gasification options to nine. Table 13 shows the
operating characteristics of each gasification option, including the effluent wet
gas composition.
Particulates present in the raw product gases are assumed to be removed
at the gasifier exit conditions. The product gases leaving the particulate
collection equipment are then admitted to the HGC equipment for desulfurfzation.
The inlet gas to the HGC equipment can, therefore, be assumed to have the same
characteristics as the raw product gases, as presented in Table 13 .
Among the four end uses considered in Table 12 , HGC processes would not
seem to couple with off-site combustion applications due to the consideration of
transporting the hot gases over long distances via pipelines. For this reason,
off-site combustion end uses are not further considered. The clean-up requirements
of low or intermediate Btu gases as reducing or synthetic gases depend upon the
specific industrial process. For this reason, this end use is not further
evaluated. The following discussion is, therefore, pertinent to the applicability
of HGC processes for on-site combustion and combined cycle pairs.
The amount of sulfur that must be removed for each gasifier-end use pair
was calculated. For on-site combustion applications, the EPA New Source Perform-
ance Standards (NSPS) of 0.52 g S02/MJ (1.2 Ib S02/10° Btu) heat input was
employed as the governing limitation on the sulfur content of the product gases.
While it is recognized that some states have more restrictive limits than this for
new plants and that a less restrictive regulation might apply in many retrofit
situations, this NSPS is representative of the regulation that would generally
limit the sulfur contents of the product gases. For the combined-cycle appli-
cations, both EPA's proposed emission standard and process restrictions were
112
-------
TABLE 13
CHARACTERISTICS OF SELECTED GASIFICATION OPTIONS
Gasification
Medium
Raw Cos
C02
CO
H2
CHj,
"2
H2S
H20
Others
Tcmperature(°C)
(OF)
Pressure'Pa)
(atm)
Coal Type
Coal HHV(MJ/kg)
(Btu/lb)
vt. -;S in Coal
Cas Prod'n Rate:
normal tn^/kj
coal**
SCr/lb coal
* Assumed
Lurgi(79)
15.8
7.5
20.9
4.3
0.2
0.62
50.3
0.4
621
1150
2.0xl05
20.4
III. #6
28.5
12.235
3.3
4.33
69.0
Lurgl(79)
Ai r/H20
9-7
11-9
17.1
2.8
29.6
0.54
27.7
0.7
549
1020
2.0xl06
20.4
111. 16
28.5
12,235
3.8
4.83
77-5
We 11 man
Galusha(SO)
Air/H20
3.0
26.0
13.9
2.5
45.6
0.70
8.3
587
1038
1x105
1 +
Bituminous
27.9*
12,300*
•»4.0
4.95
79.3
MERC(8l)
Air/H,0
7.8
16.8
'13.6
2.3
41.8
0.73
16.7
0.2
538
1000
8.5xl05
8.5
11 1 . #6
27.3
11.750
3.7
3.65
53.5
BGC/
02/H20
1.8
54.4
26.3
7.9
0.4
1.38
4-9
2.3
427
800
2.0xl06
20.4
III. 16
28.5
12,235
3.8
1.37
29.9
Winkler(6j)
02/H20
16.3
2fi.4
28.7
1.3
0.8
1.10
25.4
927
1700
lx!05
1 +
German Coal
21.7
9,320
3.3
2.19
35-1
BI-Gas(82)
Air/H20
8.3
13.1
13-1
3-6
45.8
0.59
10.2
0.3
982
1800
3-IxlO6
31.0
Not Avail.
28.4
12,200
3.8
4.82
77.3
Foster
Wheel erf 79)
02/H20
12.6
34.5
30.1
6.0
0.4
1.07
14.4
927
1700
2.5xl06
24.5
111. 16
23.5
12,235
3.8
2.56
41.1
Foster
Whee!er(79)
Air/H20
3.4
29.2
13.9
3.4
47-9
0.65
2.2
0.4
927
1700
2.5xl06
24.5
111. *6
28.5
12,235
3.3
3-99
63.9
** Coal; as received
-------
considered. EPA's proposed emission standard for stationary gas turbines is 0.8
weight percent sulfur in the fuel gas (dry basis) burned by turbines (83) (A-19).
The gaseous fuel specifications for turbines, based on equipment considerations, is
reported to be 0.18 mole percent hydrogen sulfide (SM . Table lA'shows the minimum
amount of sulfur that must be removed from the product gases to meet the appli-
cable EPA standards or turbine specification. For combustion purposes, at least
79 percent sulfur removal is required for each selected gasifier-end use pair.
For combined cycle applications, the turbine specification is more stringent than
the proposed EPA standard. Thus the turbine specification was considered as the
minimum HGC requirement in assessing the use of HGC processes for combined cycle
application. Accordingly, a minimum of 67 to 87 percent of the sulfur in the fuel
must be removed, depending upon the type of the gasification process.
Nine candidate HGC processes were selected for the applicability assessments.
These HGC processes and their active materials are the Air Products (CaO), Babcock
and Wilcox (FeO), Battelle Northwest (molten salt mixture) CONOCO (CaCOJ , Foster
Wieeler (Ni), IFP (ZnO), IGT-Meissner (molten Pfa), Kennecott (Cu), and HERC (Fe^O,)
processes. The selection was made to generate a large number of representative
HGC processes each using~a different type of active material in the absorbent.
However, it was not intended to show a preference of one HGC process over another,
where both use the same active material (e.g., MERC and IMMR processes both employ
Fe20, as active material).
The H2S removal efficiencies of the above HGC processes were estimated
from thermodynamic equilibrium relationships for the respective absorbents, and
compared with the minimum removal requirements presented in Table ]k to assess the
applicability of each HGC process for the gasifier-end use pairs. Identical results
were obtained for both the pressurized on-site combustion and combined cycle appli-
cations for the product gas characteristics as shown in Table 13. These results
of these applicability assessments are presented in Table 15. It should be noted
that for the applicability of the Battelle-Northwest process, the insufficient data
are due to the uncertainties in the physical properties of the absorbent such as
the temperature and pressure at which the carbonate mixture decomposes into solid
ox i des.
The applicability of HGC was then generalized to coals with different
sulfur contents, by treating the hydrogen sulfide concentration in the fuel gas as
a variable while keeping all the other parameters constant. The coal characteristics
considered herein were assumed to be identical to those in Table 13, except for the
variable sulfur content. As a point of reference for these assessments, the
maximum H2$ concentrations in product gases that would not require any sulfur
removal were estimated based on both the emission regulations and turbine specifi-
cations previously discussed. Table 16 shows the hydrogen sulfide concentrations
below which no clean-up is required.
Applicability of the HGC processes for use in on-site combustion and
combined cycle systems was then assessed between the H.S concentration limits
indicated in Tables 13 and 14 for the respective gasification options. For any
product gas concentration in this range, the same result given in Table 15 was
-------
TA3LE 1*4 MINIMUM SULFUR REMOVAL FOR
GASIFIER/END USE PAIRS*
Percent Sulfur Removal
^^•^-^^ End Use
Gas from ^^-^^
Gasifier ^>>>-\^^
Lurgi (0^)
Lurgi (Air)
Wei Iman-Galusha
MERC
BGC/Slagging
Winkler
Bi-Gas
Foster Wheeler (02)
Foster Wheeler (Air)
On-s i te
Combust ion
EPA's std.
80
80
85
80
81
83
81
79
79
Combined Cycle
EPA's std.
57
19
20
29
65
64
4
60
8
Turbine spec.
71
67
74
75
87
33
70
83
72
Based on product gas characteristics in Table 13
115
-------
TABLE 15- ASSESSMENT OF HGC PROCESS APPLICABILITY*
Gasifiers
HGC Processes
Air Products
Bab cock £ W 11 cox
Bat te lie Northwest
CONOCO
Foster Wheeler
IFP
IGT-Meissner
Kennecott
MERC
LurgI
02/H20
k
k
3
k
\
k
1
2
1
Lurgi
Air/H20
k
1
3
4
1
1
1
2
1
Wei Iman Gal usha
Air/H20
k
1
1
2
1
1
1
2
1
MERC
Air/H20
1*
1
3
k
1
1
1
2
1
Explanation of numbering system:
1. Thermodynamically favorable, and gasifier product gas temperature is within range of HGC
experimental test conditions.
2. Thermodynamically favorable, but gasifier product gas temperature is outside range of
HGC experimental test conditions.
3. Insufficient data.
4. Unfavorable thermodynamics or absorbent properties.
* Results apply to both on-site combustion processes and combined cycles.
(continued)
-------
TABLE 15 (continued)*
Gasi f iers
HGC Processes
Ai r Products
Babcock & Wi 1 cox
Battel le Northwest
CONOCO
Foster Wheeler
IFP
IGT-Meissner
Kennecott
MERC
BGC/Slagging
02/H20
k
1
3
it
2
J
1
2
2
Winkler
02/H20
2
k
3
k
It
It
It
2
2
Bi-Gas
Air/H20
It
k
3
2
It
4
It
2
2
Foster Wheeler
0,/H,0
£- i.
2
k
3
k
it
k
k
2
2
Fos te r
Air/H
2
k
3
k
k
k
k
2
2
Wheeler
2°
Explanation of numbering system:
I. Thermodynamically favorable and gasifier product gas temperature is within range of HGC
experimental test conditions.
2. Thermodynamically favorable, but gasifier product gas temperature is outside range of
HGC experimental test conditions.
3. Insufficient data.
^. Unfavorable thermodynamics or absorbent properties.
* Results apply to both on-s i te combust ion processes and combined cycles.
-------
TABLE IS H2S PRODUCT-GAS CONCENTRATIONS BELOW
WHICH NO CLEAN-UP IS REQUIRED
^£nd Use
Gasifier -^^^^
Lurgi (02)
Lurgi (Air)
We 11 man Galusha
MERC
BGC/Slagging
Winkler
BI-Gas
Foster Wheeler (0.)
Foster Wheeler (Air)
82$ concentration (wet mole %)
On-site
Combustion
0.12
0.11
0.11
0.14
0.30
0.19
0.11
0.21
0.14
Combined Cycle
0.18
0.18
0.18
0.18
0.18
0.18
0.18
0.18
0.18
118
-------
obtained for both applications. This interesting result is due to a reduction in
the H2S removal efficiency at decreasing concentrations of H-S in the product
gas.
Combining results presented in Tables ]2and 15 , the applicability of the
HGC processes for each selected gasifier-end use combination is obtained. These
results are presented in Table 17 . It should be noted that processes using
Fe.^Q-3 (such as MERC, IMMR) and Cu (such as Kennecott) are applicable for all
gasifier-end use combinations under consideration.
119
-------
TABLE 17- HGC PROCESSES APPLICABLE FOR SELECTED GASIFIER-END USE COMBINATIONS
N)
O
Gasif iers
End Lurgi
Use (02)
On-slte
Combustion
Combined
Cycle
Foster Wheeler*
IGT-Meissner*
Kennecott**
MERC*
Foster Wheeler*
IGT-Meissner*
Kennecott**
MERC*
Lurgi Wei Iman-Gal usha
(Air) (Air)
Babcock 6
Wilcox*
Foster Wheeler*
IFP*
IGT-Meissner*
Kennecott**
MERC*
Babcock &
Wilcox*
Foster Wheeler*
IFP*
IGT-Meissner*
Kennecott**
MERC*
Babcock &
Wilcox*
Battelle
Northwest*
CONOCO**
Foster Wheeler*
IFP*
IGT-Meissner*
Kennecott**
MERC*
MERC
(Air)
Babcock &
Wilcox*
Fos te r
Wheeler*
IFP*
IGT-Meissner*
Kennecott*
MERC**
Babcock £
Wi Icox*
Fos te r
Wheeler*
IFP*
IGT-Meissner*
Kennecott**
MERC*
BGC/Slagging
Babcock &
Wilcox*
Foster
Wheeler**
IFP*
IGT-Meissner*
Kennecott**
MERC**
Babcock &
Wilcox*
Fos te r
Wheeler**
IFP*
IGT-Meissner*
Kennecott**
MERC*
* Thermodynamlcally favorable, and gasifier product gas temperature is within range of HGC experi
mental test conditions.
**Thermodynamically favorable, but gasifter product gas temperature is outside range of HGC
experimental test conditions.
(continued)
-------
TABLE 17 (continued)
Gas i f iers
End
Use
On-s i te
Combustion
Comb i ned
Cycle
Winkler
(oj
Air Products**
Kennecott**
MERC**
Bi-Gas
(Air)
CONOCO**
Kennecott**
MERC**
CONOCO**
Kennecott**
MERC**
Foster Wheeler
(Oj
Air Products**
Kennecott**
MERC**
Air Products**
Kennecott**
MERC**
Foster Wheeler
(Air)
Air Products**
Kennecott**
MERC**
Ai r Products**
Kennecott**
MERC**
* Thermodynamical ly favorable, and gasifier product gas temperature is within range of HGC
experimental test conditions.
** Thermodynamically favorable, but gasifier product gas temperature is outside range of HGC
experimental test conditions.
-------
SECTION 5
ADVANTAGES AND DISADVANTAGES OF HGC
The advantages and disadvantages of HGC are assessed generically rel-
ative to low temperature cleanup systems. Sections 5.1 through 5.5 present
this comparison in terms of thermal efficiency, tars, particulates, NOX, and
corrosion. The uncertainties and problems of each particular HGC process are
addressed in Section 3.
Economic comparisons between hot gas cleanup and low temperature
desulfurization have been made by a number of investigators. The results of these
studies are summarized in Section 5.6.
5.1 THERMAL EFFICIENCY
Cooling of the raw product gas for- low temperature cleanup results in
the condensation and removal of tars and the loss of the sensible heat of the gas.
This reduces the overall thermal efficiency of the coal conversion processes used
with low temperature cleanup, as compared to HGC processes.
In a recent study performed by Jones et al. ( 79 ) low and high tempera-
ture desulfurization processes are compared in terms of their thermal efficiencies.
Two desulfurization processes (Benfield Process for low temperature desulfurization
and MERC process for HGC) were coupled with five gasification options for com-
bined cycle application. The process thermal efficiencies were then compared for
gas turbine inlet temperatures of 1066oc (1950QF) amd 1316°C (2kQO°F). The results
are given In Table 18.
As expected for all gasification options higher thermal efficiencies are
obtained with HGC. The difference in thermal efficiency for the two cleanup op-
tions is greatest for Lurgi gasification. The reason is two-fold. First, Lurgi
gasifiers produce tars, whose heating value is ultimately converted to electricity
in a combined cycle application when HGC is used. When a low temperature desul-
furization process is used, these tars are removed from the raw product gas by
water scrubbing. In addition, Lurgi gasifiers require large amounts of steam to
prevent clinckering of ash In the gasifier. This steam then appears in the raw
product gas. When water scrubbing is employed, most of the sensible heat in the
steam can not be used for steam generation. In the case of most other gasifiers,
however, most of the steam fed to the gasifier is converted to hydrogen and carbon
122
-------
TABLE 18
SUMMARY OF ESTIMATED THERMAL EFFICIENCICS (79)
Cold HGC Cold HGC
Purification Process Benfield MERC Benfield MERC
Gas Turbine Inlet
Temperature, °C 1066 1066 1316 1316
Thermal Efficiency, %*
Lurgi (02) 29.
-------
monoxide in the gasifier. Thus, the associated efficiency loss due to the water
scrubbing step is less than for the Lurgi gasifier.
The conversion of sulfur compounds from HGC regenerator to product sulfur
also results in a loss of thermal energy. This loss in thermal efficiency is
estimated to be 1.5 percent in the Jones study.
Ashworth, et al.(85) concluded that HGC should always yield greater power
plant thermal efficiencies when compared with low temperature desulfurization. The
incentive for HGC was found to be greater when tars are present in the fuel gas.
5.2 TARS
Tars consist of a mix of polycyclic and heterocyclic aromatic compounds
produced from gasification of coal. The types of coal and gasification process
affect the types and amounts of tars in raw product gases from coal gasifiers. For
example, gasification of bituminous coals produces heavier tars than those generated
from gasification of lignite. Also, the formation of heavier tars is favored at
lower temperatures and pressures. The amounts of tars emitted from various types
of gasifiers are discussed in Section 4.
If a low temperature cleanup process is used to purify a tar-laden gas,
cooling of the gas stream in heat exchangers could cause condensation of tars at
temperatures below 538°C (1000°F), resulting in fouling of the heat exchangers.
As an alternative, water scrubbing can be used to remove the tars from the raw pro-
duct gases upstream of any heat exchange or pol 1 ution cojitrol equi pment. However,
no environmentally acceptable method for treatment or disposal of tar-laden water
stream has yet been demonstrated ( 40 ). In addition, the removal of tars implies
a loss of some heating value (about 0.04-0.9 MJ/normal m^ or 1-25 Btu/SCF) from
the product gas resulting in reduced energy efficiency.
If desulfurization of tar-containing coal gases is carried out at temper-
atures above 538°C, then condensation of tars can theoretically be prevented. The
tars would then pass through the absorber as vapors along with the product gas,
and be combusted either in the boiler or the combustor of the gas turbine. Thus,
the heating value of the tars would be conserved and utilized efficiently. If some
tars should condense in cold spots of the hot gas desulfurizer, then they might
cause operating problems such as a reduction in sorbent activity or plugging of
the bed. If such tar deposition does not hinder the operation of the absorber, then
the tars can be burned off in the regenerator (assuming that an oxygen containing
gas is used for regeneration).
5.3 PARTICULATES
Gas streams from coal gasifiers contain particulates at various concentra-
tions depending upon the type and operating conditions of the gasifier. These
particulates are composed of carbon and mineral matter present in the coal .
Sections 4. 1.1 through 4.1.3 discuss particulate emissions from coal gasifiers.
124
-------
The participates may be objectionable because of environmenta/ and/or
process considerations. The New Source Performance Standard established by EPA
for fossil fuel steam generators limits participate emissions to 0.1*3 9/MJ (0.10
lb/106 Btu) heat input (83 )• For stationary gas turbines, no such standard has
yet been established. However, in the case of gas turbines, particulates are
objectionable because of process considerations, since they cause erosion and
corrosion in the turbine. The erosion is due to the impact of the particulates on
the turbine blades, while the corrosion results from the formation of sulfates and
oxides, and the subsequent reaction of these compounds with the turbine metal. The
particulate-related specifications for these turbines are based on engineering
estimates, since gas turbines have not been commercially applied to coal gases.
Westinghouse Electric Corporation projects a maximum allowance of O.O1*! g/normal m*
(0.018 gr/scf) for particle loading (at the turbine inlet), with no particles
greater than 6/1 m (7k). For a typical particulate emission of 3-0 g/normal nr
from a fixed bed gasifier - a minimum removal efficiency of 9?.63 percent is re-
quired. For fluid and entrained bed gasifiers, the corresponding figures would
be 26 and 110 g/m3, and 99.^ and 99-96 percent removal efficiencies respectively.
The larger size particulates (i.e.,& 6^/Km) from coal gasifiers can be
efficiently removed by means of cyclones at temperatures up to Bl6°C (1500°F). If
low temperature desulfurization is to be used, the effluent gas stream from the
cyclones can then be cooled and water scrubbed to remove the smaller particulates
Particulate free gas can then be directed to the sulfur removal equipment. If high
temperature desulfurization is to be used, both the larger and smaller particles
need to be removed at high temperatures.
The types of equipment that can be used to collect particulates include
inertial separators, surface filters, scrubbers, granular bed filters, and
electrostatic precipitators. The inertial separators (e.g., cyclones) do not re-
move fine particulate effectively. If such equipment were coupled with a HGC
process that employes solid absorbents, then the particles from the gas stream
would deposit on the HGC absorption bed and reduce the activity of the sorbent. If
the HGC process uses a molten metal or salt absorbent, then the particulates would
have to be separated from the molten absorbent. Such a separation might be very
difficult.
Filters generally remove small particulates with a very high collection
efficiency. Among them, the commerci al lyavailable baghouse filters have not been
used above 288°C (550°F). Ceramic fabrics and porous ceramic filters, on the other
hand, can operate at temperatures up to 1650°C (3000°F), but are not commercially
available (they are under development). Developmental work on a metallic filter
by Brunswick Corporation has been discontinued (A-20).
Molten salt scrubbers can also be used at high temperatures. However,
molten salt systems create additional problems, as discussed in Section *». The
major anticipated problems include corrosion caused by the molten salt, which
necessitates the use of expensive construction materials, and the entrainment of
salt particles in the gas stream from the scrubber.
125
-------
Granular bed filters can remove small particulates with high collection
L.ficiency, but have not been commercially used at high temperatures (above i»30°C) .
electrostatic precipitators have been tested at 1IOO°C, but have not been employed
commercially above approximately 500°C.
It can be seen from the above discussion that no particulate control de-
vice is yet available to efficiently remove the particles from product gases of
coal gasifiers.
5.A NOX
The combustion of a fuel gas produces nitrogen oxides such as NO, N02,
and N20. These nitrogen oxides are usually referred as NOX. The New Source Stan- ,
dard of Performance for coal-fired steam generators limits NOX emissions (0.7 lb/10
Btu) to a maximum of 0.30 grams NOX (expressed as N02) per megajoule of heat input.
This standard does not apply to lignite-fired steam generators, for which the pro-
posed limit is 0.60 Ib NOX/10° Btu heat input ( 83 ). For stationary gas turbines,
the NOX emission standard depends upon the heat rate (at peak load) of the gas
turbine and/or the nitrogen content of the fuel.
NOX emissions are generally in the range of 500-1000 ppm for coal-fired
boilers, 100-1000 ppm for gas-fired boilers, and 100-500 ppm for oil-fired boilers
( 86 )• More than 90 percent of these emissions are in the form of NO, a relatively
unreactive nitrogen oxide. For this reason, NOX removal is a difficult problem,
and the most effective method of control is to limit its formation during the
combustion process.
The sources of nitrogen in a coal derived fuel gas consist of the atmos-
pheric nitrogen and the nitrogen in the coal. The NOX emissions due to the atmos-
pheric nitrogen can be controlled by the design and operating conditions, flame
temperature, excess air, and residence time in the combus tor. The gasification of
coal also produces nitrogen compounds primarily in the form of ammonia and some
hydrogen cyanide and pyridene. The concentrations of these nitrogen compounds in
the fuel gas vary with the gasifier operating conditions, with lower temperatures
and higher pressures favoring the formation of such nitrogen compounds. These
nitrogen compounds are readily combustible to product NOX.
When low temperature desulfurization is used, the fuel gas from the
gasifier is scrubbed with water. This scrubbing process removes the ammonia and
other nitrogen compounds, and, therefore, results in reduced NOX emissions. On
the other hand, when HGC is used, the raw gas from the gasifier is first introduced
to the particulate removal equipment and then to the high temperature desulfurization
unit. Thus, the nitrogen compounds are not removed and, therefore, contribute
to the NOX emissions from combustion in the boiler or turbine.
Decompostion of ammonia into stable nitrogen and hydrogen may be a sol-
ution to the above NOX emission problem (57 ). The decomposition reaction is
as follows:
126
-------
NH .« . » 1/2 N2 * 3/2 H2
The decomposition of ammonia is thermodynamically favored at high temperatures and
low pressures. .The effect of pressure is very small, increasing the partial pressure
equilibrium constant (K ) by 3 percent for a reduction in pressure from 30 to 10
atmospheres (temperatureprange 3^9-399 C. On the other hand, K is strongly dependent
upon temperature; its value being 38 at 3*»9°C (66o°F) , 7& at 339°C (750°F) and
152 at ^9°C (8*tO°F), all at 105 Pa (10 atm) pressure. However, the kinetics for
reaction are slow and a catalyst Is needed to speed the rate of this reaction. This
problem of converting NOX to ammonia and then decomposing the latter via a catalyst
has already been addressed by the automobile manufacturing companies, such as
General Motors(87) and Ford Motor Company (38). Various developed catalysts con-
taining Ni, Pt, W, Mo and Ru all have the deficiency of getting poisoned with sul-
fur compounds remaining in the gas after the cleanup. Another problem associated
with the use of these catalysts is sintering, which would result in reduced
activity. The only commerical catalysts reported in the literature that is poten-
tially capable of decomposing ammonia contains 5 percent F6203 in inert alumina.
Since this catalyst contains the same active material as some HGC processes (such
as MERC and IMMR processes), it might be possible to combine HGC and NH3 removal
into one step. Further research is needed to determine the feasibility of such
an operation.
5.5 CORROSION
Certain chemical compounds that originate from coal or the absorbent
material can cause corrosion of the equipment. Corrosion problems appear to be
much more severe for HGC than low temperature desulfurization processes.
In the low temperature processes, corrosion may take place during the
cooling of the raw product gas and the treatment of the liquid stream. During the
cooling of the raw product gas, condensation and deposition of KC1 and NaCl as
solid particles on the waste heat boiler tubes may cause corrosion. Iron base
alloys experience oxidation at temperatures above *»82°C (300°F), and nickel base
alloys above 6i»9°C (1200°F). Corrosion by the condensates has been observed in
gas cleanup systems used in by-product coke oven installations and in Koppers
systems ( 89 ).
In the case of HGC processes, various process equipment is exposed to
corrosion with H^S at high temperatures. Tests conducted by the Westinghouse
Research Laboratory ( 89 ) at 1000°C (1332°F) showed that Incoloy 800, aluminized
Incoloy 800 and L605 (a cobalt-based superalloy) are more corrosion resistant
than other metals. However, none of these materials was found to be satisfactory
at an H2$ concentration of 0.5 mole percent. Corrosion tests, performed by IMMR
on six metal specimens (1020 C.S., 30*t S.S., 316 S.S., Incoloy 600, Inconel 800
and Armco 18) at 6^3°C (1200°F) with a gas containing 1 percent H2S and 3 percent
H20 showed that Armco 13 withstood attack best, and that metals with low nickel
content and chromium to nickel ratios above 0.5 were most corrosion resistant (28 ).
More information on the H£S corrosion problem is presented in Section 3-
127
-------
In addition to H~S, alkali metals can also cause corrosion. These
alkali metals orginate from coal, water, and sometimes the absorbent material.
Dolomites contain sodium and potassium. The absorbents in the three molten salt
processes described in Section 3 also contain these alkalis. Upon combustion
of the fuel gas, alkali metal sulfides convert to sulfates, and then condense on
the turbine blades ( if the end use is combined-cycle), which results in corrosion
and failure of the turbine blades. Similarly, alkali chlorides from the fuel gas
may condense and cause corrosion. Information on corrosion problems due to molten
salt absorbents can be found in Section 3.
5.6 ECONOMIC ANALYSIS
A comparative economic analysis of HGC versus low temperature desulfur-
ization processes was performed by various investigators.
A study by Robson, et al,(82,90 ) compares HGC with low temperature de-
sulfurization of coal gases from the Bi-Gas gasifier for the combined cycle appli-
cation. The Selexol Process was selected for low temperature cleanup, whereas the
CONOCO Process was chosen for HGC. This study shows that the overall thermal
efficiency improves with HGC (45.0 percent for HGC versus 43.5 percent for low
temperature clean-up), and that the costs of generating electricity are 24.2 mill/
kWh with HGC and 25.4 mills/kWh with low temperature cleanup.
In a more recent study by Jones, et al, (79 ) coal gasification-combined
cycle power plants of 1,000 MW capacity were considered for two possible values of
gas turbine inlet temperatures (1066°C and 13l6°C). The MERC Process was chosen
as the representative of HGC processes, since it was judged to be at a more ad-
vanced stage than the CONOCO, Air Products or Battelle Northwest Processes. The
Benfield Process was chosen as the representative of low temperature desulfurization
processes. Five gasification options were considered, including the air blown
Lurgi (fixed bed), oxygen blown Lurgi (fixed bed), oxygen blown BGC (slagging,
fixed bed), and air blown Foster Wheeler (entrained bed) gasification processes.
In the combined cycle power plant designs with HGC, the raw product gas
from the gasifier passes through the particulate removal equipment and the MERC
H2S absorber, and is then combusted before being expanded in the gas turbine.
Sulfur dioxide in the MERC regenerator effluent gas is converted to sulfur with the
Applied Chemical S02 reduction technology, which uses the fuel gas as the reducing
agent (79 ) .
In the combined-cycle power plant designs with low temperature desulfuri-
zation, the raw product gas from the gasifier is quenched with water, desulfurized
in the Benfield Plant, and then introduced into the combustor of the gas turbine.
The liquid stream from the quench operation is treated, and tars are separated.
The gas stream from the Benfield Plant is treated in a Claus Plant and tail gas
incinerator ( 79 ).
The economic assumptions used by Jones, et al (79 ), can be summarized
as follows:
128
-------
o Mid-1975 dollars with no escalation,
o 36 - month construction period,
o 10 percent construction loan interest, compounded quarterly,
o Coal cost of $1.00/MM Btu (Illinois # 6, HHV: 12,235 Btu/lb as
received),
o 70 percent operating load factor,
o 25 - year plant life,
o 50:50 debt to equity ratio,
o Eight percent bond interest, compounded semiannually,
o 12 percent return on equity after taxes, compounded semiannually, and
o 15 percent contingency in plant investment costs.
The estimated capital requirements and costs of electricity for coal
gasification-combined cycle power plants are presented in Table 19 and 20.
Table 21 summarizes the estimated capital requirements for the purification sys-
tem components ( 79 ).
Tables 19 and 20 show that there is a considerable incentive in cleaning
the Lurgi gas at high temperatures. The reductions in costs for HGC are 3k percent
for capital requirements and 30 percent for the cost of electricity. However, the
incentive for using HGC for the other types of gasification options is relatively
smal1.
The attractiveness of HGC for Lurgi gasification is due to the presence
of large amounts of steam and tars.in the gasifier effluent gas. Low temperature
desulfurization of the Lurgi gas requires a large water treatment facility and a
tar-fired boiler with stack gas scrubber, which would result in a substantial
increase in costs. For the other gasification systems, such a large increase does
not occur due to the small amount of steam and the absence of tars in the raw
product gases.
129
-------
TABLE 19
SUMMARY OF ESTIMATED CAPITAL REQUIREMENTS FOR 1,000 MW
GASIFICATION - COMBINED CYCLE POWER PLANTS (79)
Capital Requirements $/kW*
-^^^^^ Purification
Gas i f i cat ion -~-^JJrocess
Process " "—• — ^^
Lurgf (02/Dry Ash)
Lurgi (Air, Dry Ash)
BGC Slagging (02)
Foster Wheeler (Air)
Foster Wheeler (02)
Gas Turbine Inlet Temp.
1066°C
Cold
Benfield
1,117
1.000
643
619
670
HGC
MERC
739
667
629
. 616
679
Gas Turbine Inlet Temp.
1316°C
Cold
Benfield
1.046
936
629
604
6^8
HGC
MERC
703
642
606
597
657
* Capital Requirements ($/kW) correspond to Plant Investment with adjustments for
Illinois Sales Tax, Preproduction Costs, Royalty Payments, Initial Catalyst and
Chemicals, Construction Loan Interest and Working Capital.
The Capital Requirements tabulated above represent an increase of 34 percent
over the Plant Investment estimates due to inclusion of these allowances.
130
-------
TABLE 20
SUMMARY OF ESTIMATED COSTS OF ELECTRICITY FOR
GASIFICATION-COMBINED CYCLE POWER PLANTS OF 1,000 MW CAPACITY (79)
Power Costs, MMIs/kWh*
^~~~~~-~-^^^ Purification
Gas i f i catTorT^-^^Process
Process "~-~-^
Lurgi (02)
Lurgi (Air)
BGC Slagging (0«)
Foster Wheeler (Air)
Foster Wheeler (02)
Gas Turbine Inlet Temp.
1066°C
Cold
Benf iel d
48.3 (60.2)**
43.8 (55.0)
30.4 (39.9)
29.2 (38.4)
31.5 (41.4)
HGC
MERC
33. B (43.7)
31-1 (40.5)
29.7 (39.0)
29.0 (38.1)
31.5 (41.1)
Gas Turbine Inlet Temp.
1316°C
Cold
Benf ield
44.9 (55.5)
40.6 (50.7)
29.2 (38.0)
28.1 (36.7)
30.5 (39.8)
HGC
MERC |
31.6 (40.4);
29.3 (37.7)1
28.2 (36.8)
27-7 (36.1)
30.2 (39.1)
* Based on: Delivered coal costs of $1.00/MMBtu. Mid-1975 dollars with no
escalation included and an operating load factor of 70 percent.
** Numbers in parentheses represent Power Costs related to a delivered coal cost
of $2.00/MM Btu.
131
-------
TABLE 21
SUMMARY OF ESTIMATED CAPITAL REQUIREMENTS FOR
THE PURIFICATION SYSTEM COMPONENTS OF EACH CONFIGURATION INVESTIGATED (79)
Capital Requirements $/kw*
*""•" — ^^^ Purification
Gas if icat iorr^-^^Process
Process ^ ~~^^_^
Lurgi (0-)
Lurgi (Air)
BGC Slagging (02)
Foster Wheeler (Air)
Foster Wheeler (0-)
Gas Turbine Inlet Temp.
1066°C
Cold
Benfield**
141
126
52
102
86
HGC
MERC***
115
121
87
106
99
Gas Turbine Inlet Temp.
1316°C
Cold
Benfield**
129
114
48
95
80
HGC
MERC***
103
109
80
99
93
* Based on mid-1975 dollars with no escalation included.
**Cold purification costs include the costs of the HjS scrubbing system, the Claus
plant, the tail gas treating system as well as crude gas cooling equipment.
***Hot purification costs include the costs of the iron oxide system, the S02 recycle
and sulfur recovery systems, and the high temperature particulate removal devices
upstream and downstream of the iron oxide system.
132
-------
REFERENCES
1. Cornelius, E.B., et al., "Hot Stage Desulfurization of Gasified Coal - Bench
Scale Limestone Fixed Bed Adsorption and Regeneration Studies", Air Products
and Chemicals, Inc., Under Contract to the Bureau of Coal Research, Department
of the Interior, Interim Report No. 1, October 1972 to April 1972*, NTIS No. FE-
1510-T-l.
2. O'Brien, W.G., Jr., et al., "Hot Stage Desulfurization of Gasified Coal - Bench
Scale Limestone Fixed Bed Adsorption and Regeneration Studies", Final Report,
September 1973 to January 1976, Air Products and Chemicals, Inc., Under Contract
to the Energy Research and Development Administration, NTIS No. FE-1510-T-2.
3. Squires, A.M., "A Reaction Which Permits the Cyclic Use of Calcined Dolomite
to Desulfurize Fuels Undergoing Gasification", presented at the 152nd National
Meeting of the American Chemical Society - Division of Fuel Chemistry, Symposium
on Gasification, New York, New York, Vol. 10, No. k, September 11-16, 1966.
k. Squires, A.M., et al., "Desulfurization of Fuels with Calcined Dolomite. I.
Introduction and First Kinetic Results", Chemical Engineering Progress Symposium
Series, 115, Vol. o7: "Important Chemical Reactions in Air Pollution Control",
edited by Butt, J.B. and R.W. Qoughlin, 1971.
5. U.S. Patent 2,8^5,382, Leonard N. Leum and Paul M. Pitts (Atlantic Refining
Company), "Cyclic Process for the Removal of Hydrogen Sulfide from High Temper-
ature Gaseous Streams without Reduction in Temperature", July 29, '968.
6. Lee, K.C., I. Rodon, M.S. Wu, R. Pfeffer, A.M. Squires, "The Panel Bed Filter",
May 1977.
7. Squires, A.M., R.A. Graff, "Panel Bed Filters for Simultaneous Removal of Fly
Ash and Sulfur Dioxide III. Reaction of Sulfur Dioxide with Half Calcined
Dolomite", Journal of the Air Pollution Control Association, Vol. 21, No. 5,
May 1971.
8. Squires, A.M., "The City College Clean Fuels Institute: Program for (l) Gasi-
fication of Coal in High-Velocity Fluidized Beds and (II) Hot Gas Cleaning",
Clean Fuels from Coal Symposium II papers (sponsored by 1GT), Chicago, Illinois,
June 23-27, 1975.
133
-------
REFERENCES (cont'd)
9. Squires, A.M., "Clean Power from Dirty Fuels", Scientific Ameri'can. 227.
pp. 26-35, October 1972. ~
10. Graff, R.A., S.I. Dobner, A.M. Squires, "Conversion of Coal to Clean Power",
presented at the 7th Intersociety Energy Conversion Engineering Conference
(Proceedings published by the American Chemical Society), 1972.
11. U.S. Patent 3,928,532, A.M. Squires, "Treating Gas with Chemically Reactive
Dust in Panel Bed", December 23, 1975.
12. U.S. Patent 3,402,998, A.M. Squires, "Processes for Desulfurizing Fuels",
September 2k, 1968.
13. Ruth, L.A., A.M. Squires, R.A. Graff, "Desulfurization of Fuels with Half-
Calcined Dolomite: First Kinetic Data", Environmental Science and Technology,
Vol. 6, No. 12, pp. 1009-1014, November 1972.
14. Kan, G.L., A.M. Squires, R.A. Graff, "High Pressure TGA Studies on the Cyclic
Use of Half-Calcined Dolomite to Remove Hydrogen Sulfide", presented at the
172nd National Meeting of the American Chemical Society - Division of Fuel
Chemistry, San Francisco, California, Vol. 21, No. 4, August 29 - September 3,
1976.
15. Curran, P., J. Clancey, B. Pasek, P. Mell, et al., "Production of Clean Fuel
Gas from Bituminous Coal", December 1973, NTIS No. PB-232-695.
16. Curran, P., B.J. Koch, B. Pasek, P. Mell, et al., "High Temperature Desulfur-
zation of Low Btu Gas", April 1977, NTIS No. PB-271-008.
17. Ruth, L.A., A.M. Squires, R.A. Graff, "Desulfurization of Fuels with Half-
Calcined Dolomite: First Kinetic Data", Environmental Science and Technology,
Vol. 6, No. 12, pp. 1009-1014, November 1972.
18. Reeve, L., "Desulphurlzation of Coke-Oven Gas at Appleby-Frodingham", Journal
of the Institute of Fuel. 31, pp. 319-324, July 1958.
19. Bureau, A.C. and J.J.F. Olden, "The Operation of the Fordingham Desulphurizing
Plant at Exter", The Chemical Engineer, pp. 55-62, March 1967.
20. Bhada, R.K. and W.L. Sage (Babcock & Wilcox Co.), "Desulfurizing Fuel via
Metal Oxides", paper presented at the American Chemical Society Meeting,
Chicago, September 1970, ACS Preprint; 14(4), PP- 121-34.
21. Kertamus, N.J. (Babcock & Wilcox Co.), "Removal of H2S on Oxidized Iron",
paper presented at the American Chemical Society Meeting, Dallas, April 1973,
ACS Preprint 18(2), pp. 131-40.
134
-------
REFERENCES (cont'd)
22. Murthy, K.S. (Battelle Columbus Laboratories), "Investigations on the Removal
of Hydrogen Sulfide at High Temperature from Coal Gas", paper presented at
the 170th National Meeting of the American Chemical Society, Division of Fuel
Chemistry, Vol. 20, No. 4, August 24-29, 1975.
23. "Catalyst Handbook", Chapter k: "Desulphurization", Wolfe Scientific Books,
London, U.K.
2k. Abel, W.T., F.G. Schultz, P.F. Langdon, "Removal of Hydrogen Sulfide from Hot
Producer Gas by Solid Sorbents", Rl 73^7, Morgantown Energy Research Center,
1974.
25. Schrodt, J.T., G.N. Hilton, and C.A. Rogge, "High Temperature Desulfurization
of Low-CV Fuel Gas", Fuel, pp. 269-272, 5^,October 1975.
26. Schrodt, J.T., "Hot Fuel Gas Desulfurization", NTIS No. PB-257-036, May 1976.
27. Schrodt, J.T., J.E. Best, "Sulfur Recovery from Fuel Gas Desulfurization
Sorbents", Department of Chemical Engineering, University of Kentucky, May 1976,
28. Schrodt, J.T., "Hot Gas Desulfurization", Progress Report, Department of
Chemical Engineering, University of Kentucky, ORO-5076-4, December 1977.
29. Schultz, F.G., J.S. Berber, "Hydrogen Sulfide Removal from Hot Producer Gas
with Sintered Absorbents", Journal of the Air Pollution Control Association,
Vol. 20, No. 2, pp. 93-96, February 1970.
30. Lewis, P.S., F.G. Schultz, W.E. Wallace, Jr., "Sulfur Removal from Hot Pro-
ducer Gas", presented at the 166th National Meeting of the American Chemical
Society - Division of Fuel Chemistry, Chicago, Illinois, Vol. 13, No. 4,
August 26-31 , 1973.
31. Schultz, F.G., P.S. Lewis, "Hot Sulfur Removal from Producer Gas", presented
at the Proceedings of the Third International Conference on Fluidized Bed
Combustion, Hueston Woods Lodge, Ohio, December 1973, NTIS No. P8 231-977.
32. Oldaker, E.C., A. Poston, Jr., W.L. Farrior, Jr., "Removal of Hydrogen Sulfide
from Hot Low-Btu Gas with Iron Oxide - Fly Ash Sorbents", February 1975,
NTIS No. MERC/TPR-57/1.
33. Oldaker, E.C., A.M. Poston, Jr., W.L. Farrior, Jr., "Hydrogen Sulfide Removal
from Hot Producer Gas with a Solid Fly Ash - Iron Oxide Sorbent", May 1975,
NTIS No. MERC/TPR-75/2.
34. Farrior, Jr., W.L., A.M. Poston, Jr., E.C. Oldaker, "Regenerable Iron Oxide-
Silica Sorbents for the Removal of 'r^S from Hot Producer Gas", presented at
the Energy Resources Conference, January 7-8, 1976.
135
-------
REFERENCES (cont'd)
35. Pitrolo, A. A., "Quarterly Technical Progress Report", January 1 - March 31,
1977, NTIS No. MERC-77/1.
36. Jenkins, R.J., "Hot Low Producer Gas Desulfuri zat ion in Fixed Bed of Iron
Oxide - Fly Ash, Final Report: Preparation and Properties of h^S Absorbents",
July 1976. NTIS No. FE 2033-14.
37. Leuenberger, E.L., "Hot Low Btu Producer Gas Desul furizat ion in Fixed Bed of
Iron Oxide - Fly Ash", Annual Report No. 1, July 1, 1975 - June 30, 1976,
October 7, 1976 , NTIS No. FE-2033-17.
38. Joshi, O.K., E.L. Leuenberqer, "Hot Low Btu Producer Gas Desul furizat ion in
Fixed Bed of Iron Oxide - Fly Ash", Final Report - July 1, 1975 - April 30,
1977, Vols. 1 and 2, September 1977, NTIS No. FE-2033-19.
39. Pitrolo, A. A., "Quarterly Technical Progress Report, April 1 - June 30, 1977,
NTIS No. MERC-77/2.
40. Case, G.D., "Chemistry of Hot Gas Clean-up in Coal Gasification and Combustion",
November 1977-
41. U.S. Patent 2,442,982, Nachod, C.F. and N.J. Haddonfield (Atlantic Refining
Corporation), "Desulfurization of Hydrocarbons", June 8, 1948.
42. U.S. Patent, 3,079,223, Warren K. Lewis (Esso Research and Engineering Company),
"Desulfurizinq Reducing Gases", February 26, 1963.
43. Huff, .W.J. and L. Logan, "The Purification of Commercial Gases at Elevated
Temperatures", AGA Proceedings, 18, pp. 724-759, 1936.
44. "Hot Gas Desulfurization Process" by Ledgemont Laboratory, Kennecott Copper
Corporation, April 1976.
45. U.S. Patent 3, Ml ,370, W.R. Gutmann, R.H. Wright (Catalysts and Chemicals, Inc.),
"Method of Removing Sulfur Compounds from Gases", April 29, 1969.
46. Westmoreland, P.R., J.B. Gibson, D.P. Harrison, "Comparative Kinetics of
High -Temperature Reaction Between H2S and Selected Metal Oxides", Environmental
Science and Technology, Vol. 11, No. 4, pp. 483-1*91, May 1977.
47. U.S. Patent 4,002,720, K.S. Wheelock, R.S. Say (Exxon Research and Engineering
Co.), "Gas Desulfurization", January 11, 1977-
48. U.S. Patent 3,974,256, K.S. Wheelock, C.L. Aldridge (Exxon Research and
Engineering Co), "Sulfide Removal Process", August 10, 1976.
136
-------
REFERENCES (cont'd)
A9. U.S. Patent ^,039,619, Peter Steiner (Foster Wheeler Energy Corporation),
"Process for the Selective Removal of Hydroqen Sulftde from Hot Coal
Gasification Gases", Auqust 2, 1977-
50. U.S. Patent 3,919,390, Raymond Moore, "Process and Composition for Cleaning
Hot Fuel Gas", November 11, 1975.
51. Moore, R.H., C.H. Allen, G.F. Schiefelbein, R.F. Maness, "A Process for
Cleaning and Removal of Sulfur Compounds from Low BTU Gases", Interim Report
for the period October 1972 to Auqust 197*», NTIS Mo. PB-236-522.
52. Moore, R.H., G.F. Schiefelbein, G.E. Stegen, D.G. Ham, "Molten Salt Scrubbing
for Removal of Particles and Sulfur from Producer Gas", September 1977,
NTIS No. BNWL-SA-6365.
53- Moore, R.H., et al., "Process for Cleaning and Removal of Sulfur Compounds
from Low BTU Gases", Quarterly Summary Report, January-March 1978, April 1978,
NTIS No. PNL-20^0-9.
5^- Maruhnic, P., W.C. Rovesti, and R.H. Wolk, "Investigation of a Supported Molten-
Carbonate, Hot-Gas Clean-Up Process—A Status Report", Hydrocarbon Research,
Inc., Trenton, New Jersey, June 1973-
55. U.S. Patent 3,996,335, R.H. Wolk, W.C. Rovesti, and P. Maruhnic, "Desulfuri-
zation of Fuel Gas at High Temperature Using Supported Molten Metal Carbonate
Absorbent", December 7, 1976.
56. U.S. Patent 3,671,185, P.A. Lefrancois, K.M. Barclay (Pullman Inc.),
"Purification Waste Gases", June 20, 1972.
57. Robson, F.L., W.A. Blecher, C.B., Col ton, "Fuel Gas Environmental Impact",
United Technologies Research Center, Hittman Associates, prepared for U.S.
Environmental Protection Agency, p. 26 EPA-600/2-76-153, June 1956.
58. U.S. Patent 3,95*»,938, H.P. Meissner, (Institute of Gas Technology), "Removal
of Hydrogen Sulfide from Reducing Gases", May *», 1976.
59. Becker, D.F., et al., "Feasibility of Reducing Fuel Gas Clean-Up Needs",
Gilbert Associates Inc., NTIS No. FE-1236-5, June 1976.
60. "Trials of Americal Coals in a Lurgi Gasifier at Westfield, Scotland",
Woodhall-Duckham Ltd. for U.S. ERDA, R & D Report #105, 1975.
61. Rahfuse, R.V., et al., "Gasification of Caking-Type Bituminous Coal at 75 to
150 psig in a Stirred-Bed Gas Producer", NTIS No. MERC/TPR-75-3, Morqantown
Energy Research Center, July 1975.
137
-------
REFERENCES (cont'd)
62. "Hygas: 1964 to 1972, Pipeline Gas From Coal-Hydrogenation", prepared for
ERDA by IGT, R & 0 Report #22: Final Report Vol. 1, July 1975-
63. Jahnig, C.E., "Evaluation of Pollution Control in Fossil Fuel Conversion
Processes, Gasification: Section 3. Winkler Process", Exxon Research and
Engineerinq Co., NTIS No. PB-249 846, September 1975-
6k. Anastasia, L.J., and W.G. Bair, "The HYGAS Process", paper presented at Clean
Fuels From Coal Symposium If, IGT, June 23-27, 1975.
65. "Hygas" 1972 to 1974, Pipeline Gas from Coal -Hydroqenat ion, Report #110:
Interim Report #1", prepared for ERDA by IGT, July 1975.
66. Banchik, I.M., "The Winkler Process for the Production of Low-Btu Gas from
Coal", presented at Clean Fuels From Coal Symposium I, IGT, Chicago, Illinois,
September 10-14, I973r
67. Nakles, D.V., et al., "Influence of Synthane Gasifier Conditions on Effluent
and Product Gas Production: Pittsburgh Energy Research Center, PERC/RI-75/6,
December 1975-
68. "Pipeline Gas From Coal - Hydrogenation", Report #130, prepared for ERDA and
AGA by IGT, July 1975-
69. Forney, A.J., et al., "Trace Element and Major Component Balances Around the
Synthane PDU Gasifier", Pittsburgh Energy Research Center, PERC/TPR-75/1,
August, 1975.
70. Jahnig, C.E., "Evaluation of Pollution Control in Fossil Fuel Conversion
Processes; Gasification: Section 6. Hygas Process", Exxon Research and
Engineering Co., NTIS No. PB-247 225, August 1975.
71. Farnsworth, J.F., et al., ^Clean Environment with K-T Process", presented
at the EPA meeting on Environmental Aspects of Fuel Conversion Technology,
St. Louis, Missouri, May 13-16, 1974.
72. Patterson, D.R., and C.A. Bolez, "Production and Use of Low and Medium BTU
Gas", Gilbert Associates, paper presented at 5th Energy Technology Conference
and Exposition, Washington, D.C., February 1978.
73. Ball, D., et al., "Study of Potential Problems and Optimum Opportunities in
Retrofitting Industrial Processes to Low and Intermediate Energy Gas From Coal",
Battelle Columbus Labs, NTIS No. PB-237 116, May 1974.
74. Meyer, J.P., and M.S. Edwards, "Survey of Industrial Coal Conversion Equipment
Capabilities: High Temperature, High Pressure Gas Purification", Oak Ridge
National Laboratory, ORNL/TM-6072, June 1978.
138
-------
REFERENCES (cont'd)
75. Fulton, R.W., et al . , "Survey of High Temperature Clean-Up Technology for Low
BTU Fuel Gas Processes", Aerotherm Report 75-1 3*» , January 1975.
76. Waitzman, D.A., H.L. Faucett, and E.E. Kindahl, "Evaluation of Fixed-Bed
Low-3tu Coal Gasification Systems for Retrofitting Power Plants", Tennessee
Valley Authority, NTIS No. PB 2^1 672, February 1975-
77. Chandra, K. , et al . , "Economic Studies of Coal Gasification Combined Cycle
Systems for Electric Power Generation", Fluor Engineers and Constructors, Inc.,
EPR! AF-6A2, January 1978.
73. Kahrwez, H.M., "Combined Gas/Steam Turbine Power Stations with Coal Pressure
Gasification Unit Operating to the STEAG-LURGI System, STEAG", Proceedings
of the 55th Annual Convention of Gas Processors Association, March
79. Jones, C.H., and J.M. Donohue, "Comparative Evaluation of High and Low Tempera-
ture Gas Cleaning For Coal Gasification - Combined Cycle Power Systems",
Stone 6 Webster Engineering Corporation, EPRI AF-**l6, April 1977.
80. Hoffer, F.D., W.Y. Soung, and S.E. Stover, "An Overview of Control Technology
for Industrial Fuel Gas From Coal", Hydrocarbon Research, Inc., report prepared
for EPA - Fuel Process Branch under contract No. 68-02-2601.
81. Moore, A.S., Jr., "Cleaning Producer Gas From MERC Gasifier", Morgantown
Energy Research Center, May 1977.
82. Robson, F.L., et al., "Fuel Gas Environmental Impact: Phase Report", United
Technologies Research Center - Foster Wheeler Energy Corporation, EPA 600/2-
75-078, November 1975.
83. Webber, O.K., and D.E. Whittaker, "Federal and State Standards Report", (Draft),
Pullman Kellogg, February 1978.
8^. Becker, D.E., and B.N. Murthy, "Feasibility of Reducing Fuel Gas Clean-up Needs,
Phase I, Survey of the Effect of Gasification Process Conditions on the Entrain-
ment of Impurities in the Fuel Gas", Gilbert Associates, Inc., NTIS No. FE 1236-
15, June 20, 1976.
85. Ashworth, R.A., et al . , "Low Btu Gasification High Temperature - Low Temperature
H2$ Removal Comparison Effect on Overall Thermal Efficiency in a Combined Cycle
Power Plant", Gilbert Associates, Inc., NTIS No. PB-235780, January 197^.
86. Consodine, D.M., "Energy Technology Handbook", McGraw-Hill Book Company, New
York, 1977-
87. Klimisch, R.L., and K.C. Taylor, "Ammonia Intermediacy as a Basis for Catalyst
Selection for Nitric Oxide Reduction", Environmental Science and Technology,
Vol. 7, No. 2, February 1973.
139
-------
REFERENCES (cont'd)
88. Shelef, M. and H.S. Ganhi , "Ammonia Formation in Catalytic Reduction of Nitric
Oxide by Molecular Hydrogen", Industrial Engineering Chemistry, Vol. 11, No. I
1972.
89. "High Temperature Turbine Technology Program, Phase 1. Program and System
Definition. Topical Report: Fuels Cleanup and Turbine Tolerance", Westinghouse
Electric Corporation, (Lester, Pa.), NTIS No. FE-2290-27, February 1977.
90. Robson, F.L. and W.A. Blecher, "Combined-Cycle Power Systems Burning Low-Btu
Gas", United Technologies Research Center, paper presented at the Environmental
Aspects of Fuel Conversion Technology III Symposium, sponsored by Industrial
Environmental Research Laboratory, EPA/RTP, 1977.
140
-------
APPENDIX A
PERSONAL COMMUNICATIONS
A-l. Robert Rieve, Atlantic Richfield Corporation, December, 1978.
A-2. A.M. Squires, Virginia Polytechnic Institute, January, 1978.
A-3. R. Pfeffer, City College of New York, December, 1978.
A-*». Edward Nemith, Bureau of Research, U.S. Steel Corporation, Pittsburh,
Pennsylvania, April, 1978.
A-5. R. K. Bhada, Babcock & Wilcox Co., January, 1978.
A-6. H.P. Markant, University of Kentucky (formerly with Babcock & Wilcox),
January, 1978.
A-7- Joseph Oxley, Battelle Columbus Laboratories, April, 1973.
A-8. J.T. Schrodt, University of Kentucky, August, 1978.
A-9. J.C. Agarwal, Kennecott Copper Corporation, Ledgement Laboratory,
January, 1978.
A-10. J.R. Sinek, Kennecott Copper Corporation, Ledgement Laboratory,
January, 1978.
A-ll. W. R. Gutmann, Catalysts and Chemicals, Inc., December, 1978.
A-l 2. A. Deschamps, Institute Francais du Pltrole, August, 1978.
A-13. K.S. Wheelock, Exxon Research and Engineering Co., April, 1973.
A-14. Peter Steiner, Foster Wheeler Energy Corporation, April, 1978.
A-15. Carl Gutterman, Foster Wheeler Energy Corporation, May, 1978.
A^. Raymond Moore, Battelle Pacific Northwest Laboratories, December, 197?
A-l 7 . Wen Soung, Hydrocarbon Research, Inc. Trenton, New Jersey,
December, 1977.
-------
APPENDIX (cont'd)
A-18. Dennis Duncan, Institute of Gas Technology, Chicago, Illinois,
January, 1978.
A-19. Douglas Bell, U.S. Environmental Protection Agency, IERL-RTP, N.C.,
December, 1978.
A-20. K. Mills, Brunswick Corporation, August, 1978.
-------
TECHNICAL REPORT DATA
(Please read Inttructions on the reverse before completing)
REPORT NO.
EPA-600/7-79-169
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
Hot Gas Cleanup Process
. REPORT DATE
July 1979
6. PERFORMING ORGANIZATION CODE
7. AUTHORIS)
A. Bekir Onursal
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME ANO ADDRESS
Dynalectron Corporation/Applied Research Division
6410 Rockledge Drive
Bethesda, Maryland 20034
10. PROGRAM ELEMENT NO.
EHE623A
11. CONTRACT/GRANT NO.
68-02-2601
12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
Industrial Environmental Research Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD COVERED
Final; 1/77 - 3/79
14. SPONSORING AGENCY CODE
EPA/600/13
15.SUPPLEMENTARY NOTES T£RL-RTP project officer is Robert A. McAllister, Mail Drop
61, 919/541-2134.
16. ABSTRACT
report gives results of a study to identify and classify 22 hot gas clean-
up (HGC) processes for desulfurizing reducing gases at above 430 C according to
absorbent type into groups employing solid, molten salt, and molten metal absor-
bents. It describes each process in terms of its status, chemistry, operating charac-
teristics, problems, and uncertainties . It assesses the applicability of nine HGC
processes to a variety of coal gasification systems for several end uses for the
duct gases. It evaluates advantages and disadvantages of HGC relative to conven-
tional low temperature cleanup systems with respect to thermal efficiency, the pre-
sence and/or emissions of tars , particulates , and NOx, and corrosion. It also pre-
sents economic comparisons between HGC and low temperature desulfurization. HGC
processes are best suited for combined- cycle and on-site combustion applications
coupled with low- or intermediate -Btu gasifiers. The Kennecott and MERC processes
are applicable for desulfurizing gases at high temperatures. HGC processes provide
greater overall efficiencies than low temperature desulfurization. Processing gases
at high temperatures result in increased NOx emissions. Particulate removal at
high temperatures is inefficient and corrosion problems increase. HGC processes
complicate and add uncertainties which economically offset some thermal efficiency.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
Pollution Control
Stationary Sources
Hot Gas Cleanup
Reducing Gases
Particulate
c. COS AT I Field/Group
Pollution Corrosion
Coal Gasification Gases
Desulfurization
Reduction (Chemistry)
Tars
Dust
Nitrogen Oxides
T3E
13H
07A,07D
07B,07C
11G
18. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (ThisReport)
Unclassified
21. NO. OF PAGES
20. SECURITY CLASS (This page)
Unclassified
-153.
22. PRICE
EPA Form 2220-1 (9-73)
143
------- |