EPA-R2-72-060
OCTOBER 1972 Environmental ProtectionTechnology
Low-Sulfur Char as a Co-product
in Coal Gasification
Office of Research and Monitoring
U.S. Environmental Protection Agency
Washington, D.C. 20460
-------
EPA-R2-72-060
LOW-SULFUR CHAR AS A CO-PRODUCT
IN COAL GASIFICATION
By
G.P. Curran, W.E. Clark,
Melvyn Pell, and Everett Gorin
Consolidation Coal Company, Inc.
Research Division
Library, Pennsylvania' 15129
Contract No. EHSD 71-15
Program. Element No: 1A2013
Project Officer: D. Bruce Henschel
Control Systems Division
National Environmental Research Center
Research Triangle Park, North Carolina 27711
Prepared for
OFFICE OF RESEARCH AND MONITORING
U.S. ENVIRONMENTAL PROTECTION AGENCY
WASHINGTON, D.C. 20460
October 1972
-------
EPA REVIEW NOTICE
This report has been reviewed by the Environmental
Protection Agency and approved for publication.
Approval does not signify that the contents necessarily
reflect the views and policies of the Agency, nor does
mention of trade names or commercial products constitute
endorsement or recommendation for use.
11
-------
ABSTRACT
A feasibility study is presented for the case where low-
sulfur char is produced as a co-product with low-sulfur producer gas
in a gasification-desulfurization operation with bituminous coal.
Calcium carbonate is used as a sulfur acceptor. Experimental data
are also presented to support the design feasibility study. These
data show that preoxidized coals are extremely responsive to
desulfurization under the conditions used in the feasibility study.
It is possible by this method to produce low-sulfur chars containing
0.5$ or less sulfur content. The economic evaluation shows that it
is possible to produce low-sulfur char at a lower Btu cost than would
be the case for complete gasification to low-sulfur producer gas.
The value of the char, however, as a boiler fuel is less than that of
the producer gas itself. It is, therefore, concluded that there is
no clear incentive to produce low-sulfur char as a co-product in a
gasification-desulfurization operation.
ACKNOWLE DGMENT
The principal investigator in this study was Everett Gorin
for Consolidation Coal Company and D. Bruce Henschel was Project
Officer for the Environmental Protection Agency in monitoring this
contract. Appreciation is expressed to Bruce Henschel for his many
helpful suggestions which were of considerable assistance in
fulfilling the goals of the research work under this contract.
iii
-------
TABLE OF CONTENTS
Page
I. SUMMARY 1
Table I - Preliminary Economic Comparison of Methods of
Producing Low-Sulfur Boiler Fuel 2
II. CONCLUSIONS AND RECOMMENDATIONS 3
III. INTRODUCTION " 4
IV. PROCESS DEFINITION 5
A. General Description 5
Figure 1 - Flow Diagram for Modified Process to
Produce Desulfurized Char Co-Product 6
B. Design Basis 7
C. Survey of Process Variables ' 9
D. Screening Evaluation - Results and Discussion 10
Table II - Variable Study for Desulfurized
Char Co-Product Processes 12
E. "Minimum" Gasification Case - Heat and Material Balance 13
Table III - Further Comparison of Cases 1 and 2 14
F. High Gasification Case - Heat and Material Balance 15
V. SUPPORTING EXPERIMENTAL DATA 16
A. General Background 16
B. Experimental 16
C. Results 17
Figure 2 - Schematic Drawing of Experimental Apparatus 18
Figure 3 - Differential Desulfurization of Preoxidized
Coals - Comparison with Data on LTC Chars 19
Figure 4 - Differential Desulfurization of Preoxidized
. Coal at 1500°F 21
VI. ECONOMICS 22
Table IV - Product Distribution 23
Table V~A - Investment, Manpower and Utility Summary
Minimum Gasification Case 24
Table V-B - Investment, Manpower and Utility Summary
High Gasification Case 25
Table VI - Expander Output, Power Output, and Air
Compressor Requirement . 26
-------
ii,
TABLE OF CONTENTS - Cont'd.
Table VII-A - Direct Operating Costs, Minimum
Gasification Case 27
Table VII-B - Direct Operating Costs, High
Gasification Case 28
Figure 5 - Cost of Fuel Gas as a Function of the Price
Receiyed for Char/Coal at 40^/MM Btu 29
VII. BIBLIOGRAPHY 31
VIII. APPENDICES 32
APPENDIX A 32
Figure A-l - Schematic Flow Diagram - Low-Sulfur Fuels - CO2 Acceptor
Process - Minimum Gasification Case, Dwg. XF-3255 33
Tables
A-I Mass and Heat Bala.nce - Preoxidizer - Gasifier 34
A-II Mass and Heat Balance - Reactor (Squires) 35
A-III Mass and Heat Balance - Heat Exchanger C-201 36
A-IV Mass and Heat Balance - Heat Exchanger C-202 37
A-V Mass and Heat Balance - Compressor JC-203 38
A-VI Mass and Heat Balance - Heat Exchanger C-301 39
A-VII Mass and Heat Balance - Product Gas at 206 psia 40
A-VIII Mass and Heat Bala.nce - Product Gas at 25.7 psia 41
A-IX Mass Balance ~- CO2 Removal System ( 42
A-X Mass Balance - Sulfur Recovery 43
APPENDIX B 44
Figure B-l - Schematic Flow Diagram - Low-Sulfur Fuels - C02 Acceptor
Process - High Gasification Case, Dwg. XF-3266 45
Tables
46
47
48
49
50
51
52
53
54
55
56
57
58
59
B-I
B-II
B-III
B-IV
B-V
B-VI
B-VII
B-VI I I
B-IX
B-X
B-XI
B-XII
B-XI I I
B-XIV
Mass
Mass
Mass
Mass
Mass
Mass
Mass
Mass
Mass
Mass
Mass
Mass
Mass
Mass
and
and
and
and
and
and
and
and
and
and
and
and
Heat
Heat
Heat
Heat
Heat
Heat
Heat
Heat
Heat
Heat
Heat
Heat
Balance
Balance
Balance
Balance
Balance
Balance
Balance
Balance
Balance
Balance
Balance
Balance
Balance
Balance
- Preoxidizer - Gasifier
- Reactor (Squires)
- Heat Exchanger
- Heat Exchanger
- Heat Exchangers
- Heat Exchanger
- Heat Exchanger
- Heat Exchanger
C-201
C-205
C-203 & C-204
C-202
C-206
C-207
- Compressor JC-203
- Heat Exchanger
- Product Gas at
- Product Gas at
C-302
2O6 psia
25.7 psia
- C02 Removal System
- Sulfur
Recovery
-------
1.
I. SUMMARY
The results of a feasibility study are presented here for the
case where low-sulfur char is produced as a co-product with low-sulfur
producer gas from bituminous coal in a gasification-desulfurization
operation in which CaC03 is used as a.sulfur acceptor.
The results of this study with costs escalated to 1976 and with
coal at 4O^/MM Btu are summarized in Table I. Where the total gas and
char product are used as fuel to a conventional boiler, the cost of the
desulfurized fuel increases with increasing extent of gasification. The
minimum incremental cost of the desulfurized fuel to the power plant is
20^/MM Btu above the cost of coal to the process.
If, instead, the gas component of the product mix is assumed to
have a higher value than the char, equal to the value for the total gasifi-
cation case, i.e., of 70.4^/MM Btu, the lowest cost of low-sulfur char is
still achieved with minimum gasification. In this case, the cost of the
low-sulfur char is 11^/MM Btu greater than that of the coal feed.
Experimental data are presented which show that very low-sulfur
levels in the product chars of the order of 0.5$ or less may be expected
in application of the process considered here. Preoxidized coals are
shown to be more readily desuifurized than other carbonaceous materials.
Experimental data indicate that a "flash" desuifurization process
using fine preoxidized coal may be feasible. Such a process would be lower
in cost than the processes considered, but the magnitude of the cost
reduction has not been-estimated.
-------
TABLE I
Preliminary Economic Comparison of
Methods of Producing Low-Sulfur Boiler Fuel
Total Fuel Burned in Conventional Boiler
Case
Plant Investment, $
Annual Costs. $ MM/yr (6132 hr/yr)
Direct Operating Costs, ex. Coal
Coal at 40^/MM'Btu
Capital Charges at 15%
Less Sulfur Credit at 15 $/L Ton
Less Power Credit at 9 mills/KWH
Net Annual Costs
This Studyv Char as Co-Product
Minimum High
Gasification Gasification
62.9
6 = 74
33.89
9.44
-1.71
-1.90
46.46
109.7
8.89
33.89
16.45
-1.77
-5.07
52.39
Annual Report to OAP
Total Gasification
(Case II)
118.0
10.42
33.89
17.70
-1.83
-1O.65
49.53
(Ref.l)
Fuel Product
HHV of Gas + HHV of Char, MM Btu/hr
Same, + Sensible Heat in Gas
% of Total as Gas
Fuel Cost. Delivered to Power Statiom 3)
5^/MM Btu (HHV)
j^/MM Btu (HHV + Sensible Heat)
:12,338
12,663
46.7
61.4
60. 0
11,5O9
12,358
73.8
74.3
69.1
1O,415
11,48O
100.0
77.6
70.4
MW Excess Power
Efficiency
Cold
OverallC2)
34.5
.893
.925
91.7
.833
..917
193.0
.754
.879
i) 1976 operation. Includes escalation and interest during construction at 7-1/2%/yr.
2) Overall efficiency = [HHV (gas + char) + gas sensible + 3413 (KW x's power)]/HHV in with coal.
(3) Gas and Char assumed to have the same value.
-------
II. CONCLUSIONS AND RECOMMENDATIONS
It is more economical in producing low-sulfur boiler fuel
to minimize the extent of gasification in the desulfurizing process.
It is not clear, however, that the magnitude of the cost reduction,
i.e., ca. lO^/MM Btu is sufficient to compensate for the disadvantages
of producing the low-sulfur char co-product. The disadvantages are:
ash handling and removal facilities must be provided at the boiler
plant, and the substantial 'control over NOX emissions is lost as
compared with combustion of an all-gaseous fuel. In addition, the
possibility of using the gas as a premium fuel in a combined cycle
power plant is largely lost because, if the char is to be burned in .a
conventional boiler, the gas will have to be fired simultaneously to
ensure stable combustion. Also, for the above reasons, the process
in which char is a co-product shows no clear incentive over flue gas
scrubbing. Total gasification is more expensive, but it can be
combined with cleaner combustion and more efficient power generation
techniques.
The major incentive for a char co-product process would come
about if fluid bed combustion were a developed process. The low-
sulfur char could then be used as fuel without need for supplementary
gas for firing and without further sulfur control in such a plant.
The gas could then be used as fuel for a combined cycle type power
plant.
Some cost savings over the cases considered here would be
possible by use of a modified "flash" desulfurization process. It is
recommended that experimental work be conducted in Consol's continuous
bench-scale, unit to establish the feasibility of such a process.
-------
III. INTRODUCTION
A studyV1/ had been prepared for the Control Systems Division of
the EPA on the adaptation of the C02 Acceptor Process to the problem of
producing clean low Btu producer gas from high-sulfur bituminous coals.
The study included a preliminary experimental evaluation of the process
concept as well as an economic feasibility study. The feasibility study was
directed solely to the case where complete gasification of the coal feed was
effected. This was done because it was felt desirable not only to eliminate
sulfur but also ash from the fuel product. Another reason was that subse-
quent combustion of an all gaseous fuel would be expected to yield a lower
and more controlled level of NOX emission.
The results of this study as well as some associated experimental
work are given in the Annual Reportv1) to the Control Systems Division of
EPA. This report will simply be referred to hereafter as the Annual Report.
It was pointed out in the Annual Report'1) that prior Consol
experimental data had shown that the process could easily be reoriented to
produce low-sulfur char as a co-product with the low-Btu gas. This fact was
reconfirmed by experimental data obtained during the aforementioned study.
It was quite clear also, that the total low-sulfur fuel mix could be produced
at a lower cost where char is a co-product.
Accordingly, Consol Research, was requested by the Control Systems
Division to revise the previous feasibility study to include the case where
low-sulfur char is a co-product. This has now been done and the results are
reported herein.
A similar study(2) had previously been prepared for the Office of
Coal Research. In this instance, a feasibility study was prepared wherein
low-sulfur char was considered as a co-product with production of high Btu
pipeline gas.
Some new experimental data are also briefly presented which
demonstrate the ease of desulfurization of preoxidized coals at very mild
conditions.
-------
IV. PROCESS DEFINITION
A. General Description
The sulfur recovery system used in this process has been patterned
after that of Case II in the Annual Report.(1) This is because experimental
work carried out subsequent to the preparation of the Annual Report has shown
that the sulfur recovery system of Case I, though somewhat cheaper than that
of Case II, is of doubtful technical feasibility.
The process considered here differs from that discussed in the
Annual Report!1) in that with low-sulfur char as a co-product the need for a
C02 acceptor regenerator (CaC03 -* CaO) as the ultimate sink for ungasified
char disappears. The heat, formerly supplied by the acceptor action in the
gasifier is now supplied by additional combustion of air. In all the cases
discussed here there is no C02 acceptor action, only H2S acceptance. Sulfur
is removed in the gasifier by the reaction,
CaC03 + H2S = CaS + C02 + H2O.
The sulfided acceptor is regenerated by reversing the above reaction in a
separate vessel exactly as described in Case II of the Annual Report.(x)
This regeneration reaction is referred to elsewhere in this report as the
"Squires" reaction. In all cases, gasifier conditions are such that an
adequate C02 partial pressure exists to prevent calcination of CaC03 enter-
ing the gasifier.
1. Base Case
A simplified flow diagram is shown in Figure 1. After preoxidation
at adiabatic conditions, the coal and its preoxidation products are fed
directly to the bottom of the gasifier vessel along with additional air and
steam which have been preheated by exchange with the gasifier offgas. All
the heat duty for the gasifier is supplied by partial combustion of a portion
of the coal with.air.
Regenerated acceptor (CaCO3 + MgO) enters at the top of the gasifier,
showers down through the fluiclized bed of char, and segregates in a "boot"
below the entry points of the coal, air and steam. The acceptor in the boot
is fluidized by a stream of recycled product gas. The recycle gas is cooled
by exchange with the incoming air and boiler feed water, recompressed, and
then is reheated by exchange with the gasifier offgas.
All of the recirculating acceptor is picked up from the gasifier
boot by a stream of C02-steam and is carried to the regenerator vessel. The
regenerated acceptor then is returned to the gasifier. The H2S-bearing stream
from the regenerator is converted to elemental sulfur by the Wackenroder
reaction (liquid-phase Claus) as described on page' 27 of the Annual Report.x1)
-------
Figure 1
6.
TO GASIFIER
A
FINAL
CLEANUP |300°F GAS
BFW
7DO°F
AIR
RECYCLE
COMPRESSOR
GASIFIER
RE(
PREOXIDIZER
CaS
400 °F
AIR AIR STEAM
\
SULFUR
RECOVERY
ENERATOR
I200°F
OUT
- C02
_cg2
RECYCLE
\l
\
CHAR PRODUCT
-------
7.
2. Improved System - Use of Highly Caking Pittsburgh Seam Coal
The base case gasifier-desulfurizer system requires the use of a
relatively coarse coal feed to achieve a reasonable vessel cross-sectional
area. This is true because, the char is the continuous phase in the
gasifier-desulfurizer and its fluidization properties, i.e., average size and
particle density must be compatible with the fluidization velocity of 1.09
ft/sec used in the design cases discussed below.
Previous work presented in the Annual Reportv1) has shown that
adiabatic preoxidation of the-highly caking Pittsburgh Seam coals is not
effective at elevated system pressure in producing an operable gasifier feed
when relatively coarse (24 x 100 mesh) coal is used. Evidence is at hand,
however, that finer sized coal(say 65 mesh x 0) would be operable at adiabatic
preoxidation conditions.
For production of a desulfurized char co-product, the relatively
coarse particle size required to form a dense-phase fluidized bed is no longer
needed since gasification of fixed carbon and hence long residence time is not
required. In this improved case the acceptor would be the continuous phase in
the gasifier, fluidized at 2-3 ft/sec. Finely sized char would be fed at the
bottom of the bed, and would pass up through it. By controlling the acceptor
bed expansion, the mean retention time of the .finely sized char particles can
be held in the range of 10 to 20 minutes. The char particles ultimately would
be elutriated from the bed of coarse acceptor and recovered as product.
Thus, the improved case provides not only a more practical preoxi-
dation system than the base case for processing of Pittsburgh Seam coals, but
also has the additional advantage of reduced vessel costs.
At the time that this study was carried out, experimental data were
not at hand to show that the desired degree of desulfurization could be
achieved at the low residence times required for the improved case. Such
data have since been obtained which show that this is indeed feasible as will
be reported below.
B. Design Basis
A screening type evaluation of twelve specific cases around the flow
sheet of Figure 1 was made to select two specific cases for economic evaluation.
The computer program described in the Annual Report was used, with suitable
changes to fit the modified system of Figure 1, for this preliminary evaluation.
The major process restraints, i.e., those related to the fluidization properties
of the materials handled, the thermodynamics and kinetics of the acceptor
reactions, and of the carbon steam reactions, etc., remain the same as those set
forth previously A1)
Also, the design of the sulfur recovery section utilizing the
"Wackenroder reaction" is based entirely on the Case II system described in
the Annual Report.
The preliminary screening evaluation is aimed at determining the
overall heat and material balance, vessel sizes and process efficiency as a
function of the selected process variables for each case.
-------
The following design assumptions are common to all the cases
considered:
1. Coal
The same coal composition was used as in the original feasibility
study, i.e., high-sulfur eastern bituminous coal having 4.3% sulfur content.
The analysis is duplicated below for the convenience of the reader.
Feed Coal Analysis
Moisture (as received) 6.O wt $
Ultimate Analysis. Wt %
Hydrogen 4.8
Carbon 69.8
Nitrogen 1.2
Oxygen 7.6
Sulfur 4.3
Ash 12.3
Higher Heating Value 12,700 Btu/lb (MF basis)
2. Char Sulfur Content
The sulfur content of the product char, for the purpose of this
work, was conservatively taken as 0.75 wt % regardless of the level of gasi-
fication of fixed carbon. This assumption is based on previous Consol work
with low-temperature carbonized (LTC) char.
Work carried out since this design study was made shows, however,
that preoxidized coals desulfurize more readily than LTC char, and sulfur
contents well below 0.75 wt % should be achievable. These data as well as
other background data are discussed in Section V of this report.
3. Preoxidation Level
The extent of preoxidation was maintained within the adiabatic
restraints of the process, and as discussed previously,I1) amounts to about
8.6$ for a preoxidizer operated at 800°F. It was shown experimentally,!1)
that preoxidation to this level, was sufficient to render 24 x 100 mesh
Illinois No. 6 coal operable in the gasifier. However, when high sulfur
Pittsburgh Seam coals of the same size consist are used, it has been shown
that the extent of preoxidation required to establish operability is in excess
of 20%. Thus, with coarse coals of this type, a process modification is
required wherein heat is removed from the preoxidizer. The alternative is to
use finer coal and the improved process as discussed above.
The economic analysis for the base case, thus strictly applies only
to the use of the more weakly caking coals such as Illinois No. 6. The coal
analysis used in this study, however, actually is that of a high-sulfur
Pittsburgh Seam coal. It is felt, however, that the overall heat and
material balance would not change significantly if the substitution with
Illinois No. 6 coal were made.
Extent of preoxidation is expressed as a percentage and is defined by
100 x Ibs 02 consumed per Ib of moisture-free coal.
-------
9.
4. Acceptor Sulfur Reactions
It is essential, in order to achieve the required char sulfur
levels, that rapid and efficient removal of H2S inhibitor be effected in situ
by means of the reaction,
CaC03 + H2S = CaS + H20 + C02.
Prior laboratory data by Ruth, et al.,\3' have demonstrated that this reaction
is very rapid. It was also demonstrated in the operation of our continuous
unit(9) that efficient H2S removal was effected in the gasifier via the above
reaction.
5. Product Gas Cleanup
The assumption was made that particulate matter and alkali in the
product fuel gas, after cooling to 13OO°F, could be reduced sufficiently by
high-pressure drop cyclones to allow sustained operation of gas turbine engines
or expanders. The product gas is cooled to 130O°F in all cases by heat exhange
with the incoming air, steam and recycle gas streams as shown in Figure 1.
6. Preheating of Inlet Streams
In the cases where the gasifier temperature is 1600°F or higher, the
maximum preheat temperature was taken as 1200°F in order to avoid use of high
alloy tubing in the heat exchangers. In these cases, the recycle gas and
steam were preheated to.!200°F. The remaining available heat was used to
preheat the gasifier air stream.
For the 1500°F cases, the preheat temperature limit was raised to
1350°F in order to keep the maximum tube wall temperature roughly comparable
with that of the higher temperature cases.
In all the cases, the preoxidizer air stream temperature was that at
the compressor outlet (398°F for the 15 atm cases).
7. Acceptor Make-up Rate
As in the previous feasibility study the acceptor make-up rate was
0.5$, i.e., O.005 mols fresh MgC03-CaC03 added per. mol MgO-CaC03 circulated
through the gasifier. No data, are available as to whether this make-up rate
would be sufficient to maintain adequate acceptor activity. However, the
acceptor circulation rate is only about 1/10 that of the previous cases in
which C02 acceptor action was used. Thus, the incremental cost of a higher
make-up rate would be relatively small.
C. Survey of Process Variables
The specific conditions discussed here form the basis for the
twelve cases which were considered in the primary evaluation.
-------
10.
1. Percent of Fixed Carbon Gasified
By varying the gasifier temperature and the amount of inlet steam,
a wide range of fixed carbon gasification can be achieved which in turn
alters the ratio of energy available in the gas and char co-product streams.
Temperatures below 1500°F were not considered since our OCR work
has shown that this is the minimum temperature at which substantially
complete hydrocracking of tarry matter occurs. Temperatures above 17OO°F
were not used, primarily to avoid the problem of ash slagging at the gasifier
air inlet point. Also at temperatures above 17OO°F, in some situations the
C02 partial pressure would become insufficient to prevent calcination of CaCO3.
2. Extent of Conversion of CaCO3 to CaS
As noted above, no data on the effect of process cycling on
acceptor activity exist. For the purpose of this study, a base activity of
O.2O was used, i.e., 2O mols of CaS are formed per 100 mols total calcium
on each passage through the gasifier.
3. Regenerator Temperature
The base temperature for the acceptor regenerator ("Squires"
reactor) was taken as 13OO°F, the same as in the original feasibility study.
4. System Pressure
To be compatible with cases in the Annual Report,!1) the base
system pressure was taken as 15 atm (2O6 psia).
5. Fluidization of the Gasifier Boot
Air and steam cannot be used to fluidize the acceptor in the gasi-
fier boot because the CaS would be oxidized to CaS04, making regeneration
impossible via the "Squires" reaction. Therefore, recycle gas must be used.
The base amount of recycle gas corresponds to the ratio of boot cross-sectional
area to that of the gasifier of 1/6.
D. Screening Evaluation - Results and Discussion
The scope of this study does not include specification of how the
co-products, low-sulfur fuel gas and char, are to be used ultimately in the
generation of clean power. Nevertheless, some conceptual assumptions had to
be made to allow comparisons of the merits of the various cases.
Both fuels can be burned in a conventional boiler, with suitable
burner modifications. In this instance, the product gas is passed through
expanders to reduce the pressure to about 25 psia. The expanded gas is
delivered at about 700°F to the burners. The expanders drive compressors
for the process air and produce additional power which is sold. The
desulfurized char is cooled, depressured and delivered to the station silos.
-------
11,
The low-volatile char cannot be burned in a conventional boiler
without supplemental firing. By simultaneously burning the fuel gas co-
product stable combustion can be assured, even for the cases which have the
minimum fuel gas/char ratio, since the gas supplies about 45% of the heating
value of the total fuel.
For comparison among the various cases, an overall efficiency was
defined as:
[HHV (gas + char) + gas sensible heat + 3413 (KW excess power)]
Overall efficiency = ^ <• » LjL
HHV in with coal
The excess power is that available for sale after subtracting the power
required for compression of air and recycle gas and for miscellaneous drives
from the gross power output of the expander.
Another method of utilization is to burn the gas in a combined
cycle power plant. For the purpose of this study, overall station heat rates
were calculated for the supercharged boiler combined cycle described on page
54 of the Annual Report.(*/ In these instances, the desulfurized char co-
product would have to be burned in a fluidized bed boiler since supplemental
firing is not possible. For the char portion of the cycle, a thermal efficiency
of 38%, based on the HHV of the char was assumed, without specifying the exact
nature of the cycle.
Detailed heat and material balances were calculated for twelve
cases. Some of the pertinent results of the calculations are summarized in
Table II which also shows _the {same data for Case II in the original feasi-
bility study.C1)
Cases 1 through 7 show the impact of the extent of fixed carbon
gasification over the range of 0 to 65%. Gasifier temperature, amount of
inlet steam, and the gasification rate all are compatible with our available
data on gasification kinetics.
All the cases having more than zero percent fixed carbon gasified
show a decreased overall efficiency (conventional cycle) although the range is
not large. For the combined cycle, the overall station rates improve modestly
with increasing extent of fixed carbon gasification. However, when the gasi-
fier cross-sectional areas are compared, it is obvious that the lowest cost
per unit of total energy in the co-products occurs at zero percent fixed carbon
gasified. Direct operating costs and capital charges in these processes are
sensitive to the gasifier vessel volume and to the amount of process air which
is required, as study of the Economic Evaluation Section of the original
feasibility study presented in the Annual Report(1) will show.
Case 10 shows that by roughly doubling the acceptor circulation rate
only a slight penalty is involved in either efficiency or in vessel volume.
Case 11 shows that decreasing the regenerator temperature to 1200°F causes a
nearly negligible penalty in efficiency and vessel volume. If subsequent work
shows that the kinetics of the "Squires" reaction are adequate at 1200°F, the
total cost of the plant probably will be decreased because the sulfur recovery
section investment will be considerably lower.
-------
O.7S$ Sulfur in Product Char.
15 atm System Pressure Unless Noted.
Case
Gas ifier
Temperature, °F
CaS/E Ca
Regenerator Temperature, "F
% Fixed C Gasified
Bols to Gasifier
Steam
Air (includes preox. air) .
Recycle
Hols Product Gas
Hols MgO-CaCO, to Gasifier
R,^, Gasification Rate(B)
Heating Value of Product Gas,
Btu/ft', wet
Gasif ier Cross Section, % of Case II
PPU H2S in Product Gas
Sulfur Removed from Coal Feed
Cold Efficiency
$ of Total Product HHV in Gas
P at Top of Gasif ier, atm
Air preheat Temperature, *F
Recycle Preheat Temperature, °F
Btu HHV in Char
Btu HHV in Gas
Overall Station Rate, Btu/KSTH for
Combined Cycle plus Fluidized
Bed Boiler
(*) K» Excess Power
Btu Sensible Heat in Gas
Btu HHV + Sensible Heat in Gas
Overall Efficiency *
TABU II
Variable Study for Dasulfurized Char Co-Product Processes
s Noted.
1500
0
0
3.417
2.802
6.237
.598
O
216
38
279
89.1
.8933
45.38
1.67
1320
1350
619,700
514,830
9060
4.53
31,5OO
546,330
.930
Basis; ICO Ib
3
170O
52
1.970
8.272
6.646
14.793
.620
59
135
100
159
92.5
.8334
71.89
2.46
1200
1200
297,520
760,900
8815
11.05
66,900
827,800
.916
3
1700
65
2.710
9.540
7.669
17.070
.622
75
125
115
195
92.7
.8114
78.91
2.42
930
1200
217,330
813,150
8810
12.89
77,600
890,750
.907
1650
48
2.560
8.339
6.849
15.244
.607
56
124
1OO
307
90.5
.8219
68.96
2.41
750
12 OO
324,000
719,810
8930
11.73
7O,6OO
790,410
.909
dry coal fed to preoxldizer
1650
41
2. 113"
7.638
6.271
13.958
.608
48
130
92
274
9O.6
.8348
65.50
2.44
930
12OO
365,770
694,430
8920
10.67
64,6OO
759,030
.914
1650
34
1.616
6.832
5.612
12.490
.60S
39
140
82
232
90.7
.8497
61.95
2.46
1180
120O
410,610
668,510
890O
9.48
57,600
726,110
.920
16OO
20
1.322
5.616
4.741
10. 552
.500
22
150
68
337
89.3
.8649
55.07
2.40
1200
1200
493,520
604,900
8950
8.16
49,900
654,800
.926
1700
65
2.506
9.184
1.0
16.521
.624
75
132
81
174
93.1
.8225
79.19
2.44
54O
1200
217,330
827,200
8700
13.04
74,9OO
902,100
.916
1500
0
0
3.345
0
6.167
= 598
0
221
13
272
89.2
.8954
45. SO
1.66
1120
619,700
517,410
9030
4.75
31,200
548,610
.933
1500
.10
0
0
3.636
2.908
6.472
1. 194
O
205
39
297
89.0
.8870
44.99
1.67
1260
135O
619,700
506,800
9120
4.60
32,400
539,200
.925
11
15OO
.20
12 OO
O
O
3.535
2.859
6.364
.597
O
210
39
289
89.1
.8899
45.17
1.67
1280
135O
619,700
510,500
9O95
4! 56
32 ,OOO
542,500
.927
Ig(')
15OO
.20
13OO
0
O
3.143
2.627
5.847
.595
O
233
28
378
88.8
.8973
45.62
1.92
143O
1350
619,700
519,900
8990
4.83
27,300
547,200
.933
Case >t
1700
. O187
1300
65
3.92
6.59
23.00
5.878
75
115
ICO
240
96.2
.754
100
2.23
398
18OO
O
957,580
8870
17.75
98,70O
1,056,280
.878
* Overall efficiency = [HHV (gas + char) + gas sensible + 3413 (KW excess power)]/HHV in with coal.
(») Acceptor is continuous phase. Gas velocity is 2 x other cases.
(•) System pressure = 20 atm.
(i) Regenerator gas.
(4) D gasifier/D boot = 1/45.
(s) Pounds fixed carbon gasified/pounds fixed carbon in
bed/minute x 104.
(s) Entire product mix burned in conventional boiler.
-------
13.
The effect of system pressure is indicated by Case 12 which shows
that by increasing the system pressure from 15 to 20 atm, modest improvements
in efficiency and vessel volume occur. However, the vessel cost will not
decrease in direct proportion to the cross-sectional area. Note that the
smaller air and recycle gas requirements lead to violation of the 1350°F limit
on the air preheat temperature,
Case 8 (to be compared with Case 3) shows that appreciable improve-
ment in efficiency and in vessel costs would occur if the recycle gas flow to
the acceptor boot could be reduced considerably. In Case 8 the ratio of boot
cross-sectional area to that of the gasifier is about 1/45, compared with a
ratio of 1/6 in the other cases. How practical such a drastic reduction in
cross-sectional area would be is not known at this time.
Case 9 corresponds to the improved process where fine preoxidized
coal is passed through a bed of acceptor. The data of Table I show that such
a process would be very attractive from the point of view of investment and
operating cost since the vessel volume would be reduced drastically.
Case 1 was chosen for the more detailed economic study to exemplify
the minimum gasification case, in place of Case 9, since the experimental data
on the feasibility of achieving the sulfur levels required via Case 9 were not
available at the time the decision was made.
Because the vessel cross-section is much smaller, Case I clearly
will show a lower cost for the total low-sulfur fuel product than the other
cases given in Table II "where a substantial degree of fixed carbon gasifica-
tion is effected.
Case 2 was, however, also chosen for economic evaluation to cover
the eventuality that a high ratio of fuel gas/char may be desirable in some
instances, i.e., for a combined cycle power plant fuel. Case 3 corresponds to
an even higher degree of gasification of fixed carbon, but Case 2 was chosen
in preference to Case 3 on the basis of lower vessel costs. Case 2 is here-
after referred to as the "high" gasification case. Some further details of
Cases 1 and 2 are shown in Table III.
E. 'Minimum" Gasification Case - Heat and Material Balance
A schematic, but more detailed flow sheet than Figure 1 for this
case is given in the Appendix as Figure A-l. The streams are numerically
identified for the purposes of the detailed heat and material balances around
the different sections of the plant which are given in Appendix A as Tables
A-I through A-VIII, inclusive.
Certain minor changes were made relative to the balances and flow
sheet presented in Table II and Figure 1, respectively. In Figure A-l the gas
fludizing the boot is recycled directly at 1300°F without prior cooling as
given in Figure 1. If this design proves not to be feasible, the gas may be
cooled as shown in Figure 1 prior to recycle compression. ; The difference in
cost of the two procedures.is well within the precision.of the estimate.
-------
14.
TABLE III
Further Comparison of Cases 1 and 2
Basis: 100 Ib dry coal fed to preoxidizer
Case
Wt % Char Yield
H, wt % (dry basis)
C
N
0
s
Ash
of Coal HHV in
Gas
Char
Cold Efficiency, $
Sulfur Distribution. Ib -
in with coal
out with char
out with gas
out with acceptor
54.9
. 0.5
76.3
~ 0
~ 0
0.75
22.4
40.5
48.8
89.3
4.30
.412
.056
3.832
32.9
O.4
61.5
~ 0
~ 0
O.75
37.4
59.9
23.4
83.3
4.30
.247
.075
3.978
S Rejected
89.1
92.5
-------
15,
The product gas to the turbine expander is also available now at
a lower temperature than given in Figure 1, i.e., 969°F vs. 13OO°F and the
delivered expanded gas temperature is correspondingly reduced from the
previous value of 7OO°F to 462°F. The main reason for the reduction in
temperature is because of the need to blend in with the product gas that
portion of the gas from which C02 had been removed in the C02 recovery
system. The C02 recovery system operation is required to supply C02 to the
"Squires" reaction in the sulfur recovery system. Figure 1 is not highly
detailed, and the C02 recovery system is not included.
F. High Gasification Case - Heat and Material Balance
The flow sheet and material balances for this case are given in
Appendix B as Figure B-l and Tables B-I through B-XIV, respectively. Figure
B-l is essentially a more detailed version of the basic simplified flow sheet
of Figure 1. The major differences between the flow sheets for the high and
low gasification cases are thai; (l) in the former case gasifier steam is
generated from the offgas and in the latter case no steam is generated or
used, and (2) the offgas in the former case is cooled before recycle to the
gasifier boot.
In both flow sheets, the details of the sulfur recovery system are
not shown. The sulfur recovery flow sheet, however, is identical to the
corresponding section of Case II and is given in Figure 7 of the Annual
Report.(1)
-------
16.
V. SUPPORTING EXPERIMENTAL DATA
A. General Background
Experimental work carried out some time ago at Consol Research^4)
showed that low-temperature carbonized (LTC) chars undergo a remarkable degree
of desulfurization simultaneously with gasification by steam-hydrogen mixtures.
It was recognized, however, that the active desulfurizing agent was hydrogen
and not steam. Data were later published by Consol Research on both the
kinetic and equilibrium relationships in the desulfurization of LTC charsv5/
by hydrogen alone and devolatilization gases containing hydrogen.
Results obtained in the continuous desulfurization of LTC char were
also reported.'5/ A fluid bed bench-scale unit(5) was employed in which the
use of -both once-through hydrogen and recycle devolatilization gases was
studied. The extent of desulfurization achieved under these conditions was
shown to be limited by equilibrium and not kinetics.
It was recognized that in order to arrive at a satisfactory commer-
cial process it was desirable to use the "natural" hydrogen evolved by
devolatilization of the char(sa) on heating to desulfurization temperatures,
and it was essential to carry out the desulfurization process in the presence
of an H2S acceptor!6'3 tnru 6^) to eliminate the hydrogen sulfide inhibition
of the process.
The effectiveness of the use of a "showering" lime acceptor in
desulfurizing a bed of ch'ar simultaneously undergoing partial gasification in
a continuous fluid bed bench-scale unit has also been reported.(7) Char
produced from a Pittsburgh Seam coal was reduced in sulfur to 0.29 wt $ by
this technique. Similar desulfurization results were also given in work
reported in the Annual Report.I1) The organic sulfur content of the char
gasification residue was reduced to 1% of that in the feed char.
B. Experimental
The previous work,(5) in which differential desulfurization of LTC
char was studied, was conducted by fluidizing a small batch of char with a
large excess of H2, i.e., H2/char ratios in the range of 140-400 SCFH/lb
were used. This was sufficient to substantially eliminate all but minor H2S
inhibition effects.
The present work was carried out with preoxidized coals previously
prepared. Details of the preparation are given in the Annual Report.I1)
-------
17.
A summary of the materials used is given below;
Sample Preoxidation, Preoxidation
No. Coal Feed to Preoxidizer Wt. 1a Temp., °F
6P Illinois No. 6, - 8.7 810
Hillsboro Mine
4P Pittsburgh No. 8, 27.9 . 750
Ireland Mine
A somewhat different experimental technique was used in this work.
The preoxidized coal was injected into a bed of calcined dolomite fluidized
with H2 at 1 atm pressure. This insured absence of H2S inhibition.
The reactor and apparatus for injection of preoxidized coal are
shown schematically in Figure 2.
The first step in a run was to calcine 35 grams of Tymochtee dolomite
(100 x 200 Tyler mesh) in air at 1650°F for 30 minutes. The reactor was
quenched and the air was replaced with hydrogen. The 4-way stopcock was cracked
so that hydrogen purged the reservoir but no coal came through.
The reactor was then immersed and equilibrated at temperature. Coal
feed was started by fully opening the stopcock. It took 30 seconds to one
minute to feed the coal charge of 3-5 grams. The timer was started when about
half of the coal had been fed.
Using this technique, the bed temperature never dropped more than
100°F and generally returned to temperature within 1-1/2 minutes.
Run 4P product could not be fed with either H2 or N2. It formed
agglomerates in the feed tube. This material was run by mixing it with the
calcined dolomite and then immersing the reactor containing both in the sand
bath.
The conversion of CaO to CaS was always less than 10$. A flow of
1.5 SCFH of H2 was standard. This gave a fluidizing velocity of 0.3-0.4 ft/sec.
C. Results
A comparison of results with preoxidized coals with previous data
with LTC chars is given in Figure 3. The weight percent sulfur elimination is
defined as follows:
Wt % Sulfur Elimination = (^ S in Feed Char - gms S in Product Char)
gms S in Feed Char
The data points for the preoxidized coals are shown in the graph, while the
line is the least squares line drawn under the "unproven" assumption that the
extent of sulfur elimination is independent of feedstock.
-------
18,
Figure
SCHEMATIC DRAWING OF EXPERIMENTAL APPARATUS
QUARTZ ^
REACTOR X
•":-v.^
'•,.::••!
» * „ , *.
."**.- *• **
v*rv'.
•5^-
l&i
• •' ''. y\
'-.•vS'
•' • . f '•
.>..'•.
. • " <•,
••• vc
* * . •
.
* f
-SI
t"».'*~J
••ll'lv
•-•'Vs:
• .* • r -.
.» • . • tj
*:.;V
f • •' •
•'.''':'•
•'• *, •..
'• w ' " '
• • _ .'
,' • *
• . ""'" ''•
/ -rf * •
* * t *
•••^."
^^
^-^^
SAND BATH
XS FURNACE
*••
FEED
RESERVOIR
GAS
4-WAY
STOPCOCK
-------
n .•
-------
20.
It is immediately obvious that the preoxidized coals desulfurize
more readily than the chars, and that the differential becomes more marked
as the temperature is reduced.
Figure 4 shows the reduction in sulfur content of preoxidized
Pittsburgh Seam coal versus time under process conditions, i.e., at 1500°F.
It is noted that the design value of 0.75 wt % sulfur is reached in only
10 minutes residence time.
The char residence time provided in the minimum gasification case
cannot be specified exactly since it depends on the average particle size
and density of the bed solids which is not known at this time. The order of
magnitude is, however, two hours while the average H2 partial pressure in
Case I is slightly higher, i.e., 1.1 atm, than that used in the experimental
work. The residence times and the hydrogen partial pressure in the high
gasification cases are even higher. It is, therefore, clear that the sulfur
content of the product char should be well below the level of O.75 wt $
specified.
As was mentioned above, it is likely that to establish operability
with Pittsburgh Seam coals within the framework of adiabatic preoxidation,
it will be necessary to use a fine grind of the coal feed. This means that
one must resort to the improved case discussed in Section IV-A2 to effect
desulfurization. The results shown in Figure 4 suggest that adequate
desulfurization can be achieved at the short residence times required by the
improved case.
Figure 3 suggests that adequate desulfurization can be accomplished
at temperatures well below the 1500°F temperature used in the min-imum gasifi-
cation case. Operating at lower temperatures would produce a higher ratio of
char to gas but would likely produce a lower cost net product. It should be
pointed out, however, that in this instance one could not use CaC03 as the
acceptor since the equilibrium in the acceptor reaction,
CaCO.s + H2S = CaS + H20 + C02
becomes unfavorable. It would be necessary to use a metal oxide acceptor in
this instance such as MnO as had been proposed previously.!5)
-------
0 I
-------
22.
VI. ECONOMICS
The investment and operating costs for the cases considered here
were not arrived at by a detailed estimation procedure but are only approxi-
mate. They were arrived at by ratioing unit costs as detailed in the Annual
ReportC1) to 'the equipment sizes required here.
A summary of the economics of the two cases considered here is
given in comparison with the total gasification case, i.e., Case II of the
Annual Report, in Table I of this report (cf. Summary). The figures are
based on 1976 operation and include escalation in labor and materials, and
interest during construction at the rate of 7-1/2$ per year. The operating
factor of the plant is taken at 70$. The product distributions for the cases
are summarized in Table IV.
/
The investment break down by section, and manpower and utilities
for two cases are given in Tables V~A and V-B. The power balances around the
expander and air compressor are given in Table VI. Operating cost breakdowns
are given in Tables VII-A and VII-B, respectively.
Figure 5 shows that the cost of the combined fuel product decreases
progressively from 70j^/MM Btu to 60c'/MM Btu, with coal input at 40jz(/MM Btu,
as one goes from the total to minimum gasification case. The process thermal
efficiency also increases in going through the same progression.
Where the total product is used as conventional boiler fuel, it does
not appear that the process considered here is attractive relative to' flue gas
scrubbing. The estimated cost of sulfur removal for the minimum gasification
case cited here is 20p/MM Btu of fuel burned versus 15jzf/MM Btu for the formate
scrubbing case cited in the Annual Report.C1) It should also be noted, in
making the comparison with flue gas scrubbing, that burning the char has an
additional cost penalty associated with it of particulate removal. This must
also be considered in comparing the merits of the total versus minimum gasifi-
cation case. Total gasification lends itself to combustion processes which
are both more efficient and cleaner.
The revised improved case cited earlier would effect some further
cost improvements, however, the magnitude has not been estimated.
The major incentive for low-sulfur char co-product cases would be
where a separate use existed for the low-sulfur char and where the gas would
have a premium value. Such a case would comprise use of the char in a
fluidized boiler and the gas in a combined cycle power plant.
Figure 5 has relevance to such a case and shows the cost of the low-
sulfur char for the cases given in Table III as a function of the value of the
gas. It is clear here that the minimum gasification case also produces the
lowest cost char as long as the gas value is below about 75.5^/Ml Btu.
-------
TABLE IV
Product Distribution
23.
Case
Coal Consumed
Lb/hr
MM Btu/hr (HHV)
Products
Gas
Mols/hr
MM Btu/hr, HHV
Temperature, °F
Pressure, psia
Btu/SCF, HHV
Ghar
Lb/hr
MM Btu/hr, HHV
Sulfur Content, Wt,
This Study,
Char as Co-Product
Minimum
Gasification
«-
<-
604,300
(3,740
0.75
High
Gasification
1,157,000
13,812
365,500
3,236
0.75
Annual Report to OAP
Case II
Total Gasification
67,277
5,598
462
219
159,320
8,273
512
r>c n
137
239,267
10,415
665
115
O
Sulfur, Ib/hr
41,500
43,100
44,700
-------
TABLE V-A
Investment, Manpower, and Utility Summary
Minimum Gasification Case
24.
Sulfur
Recovery-
Section
ISBL Investment, $ MM
Operators, Men/Shift
Utilities
Electricity, KW
Cooling Water, GPM
OSBL Investment. $ MM
Offsites at 6.7% of ISBL
Investment
Miscellaneous Utilities at-
l/o of ISBL Investment
Electrical
Cooling Water at 33 $/GPM
Sub-Total, OSBL
Installed Plant Cost, ISBL + OSBL, $ MM
Coal
Preparation
5.0
3
4. ,200
Gasification
20.3
7
4,250(0
5,510
Solids
Disposal
13.9
6
15,520
11,260
Totals
39.2
16
23,97O
16,770
Escalation
Interest During Construction
Total
2.6
0.4
2.6
0.6
6.2
45.4
8.5
53.9
9.0
62.9
) Does not include the main air compressors, which
are driven directly by expanders.
-------
TABLE V-B
Investment, Manpower, and Utility Summary
High Gasification Case
25.
Section
ISBL Investment, $ MM
Operators, Men/Shift
Utilities
Electricity, KW
Cooling Water, GPM
Boiler Feed Water, GPM
OSBL Investment. $ MM
Offsites at 6.7% of ISBL
Investment
Miscellaneous Utilities at
l/o of ISBL Investment
Electrical
Cooling Water at 33 $/GPM
Boiler Feed Water at 380 $/GPM
Sub-Total, OSBL
Installed Plant Cost, ISBL + OSBL, $ MM
Coal
Preparation
5.0
3
4,200
Gasification
48.6
9
7,316(1)
13,35O
813
Sulfur
Recovery-
Solids
Disposal
14.2
6
15,420
11,700
Totals
67.8
18
26,936
25,O50
813
Escalation
4.5
0.7
5.0
0.8
0.3
11.3
79.1
14.8
93.9
Interest During Construction 15.8
Total 109.7
(i) Does not include main air compressors,.which
are driven directly by expanders.
-------
TABLE VI
Expander Output, Power Output, and
Air Compressor Requirement
26,
Case
Output from Expanders, MM Btu/hr
Less 5% Mechanical Inefficiency,
MM Btu/hr
Less Main Air Compressor Requirement,
MM Btu/hr
Net Power Generated, MM Btu/hr
Equivalent MW
Minimum
Gasification
286
.0(0
-14.3
271.7
Maximum
Gasification
721.3(2)
-36.1
685.2
-153.8
117.9
34.5
-372.1
313.1
91.7
(i) From Table A-VIII.
(2) From Table B-VII.
-------
27,
TABLE VII-A
Direct Operating Costs
Ex Coal Costs
Minimum Gasification Case
Direct Operating Costs (6.132 hr/yr) $ MM/yr
1. Direct Operating Labor
16 men/shift at (40,OOO $/yr)/man/shift) 0.640
2. Maintenance Labor
1.6$ of Investment O.726
Sub-Total, Direct Labor 1.366
3. Direct Supervision, 15$ of 1 + 2 0.205
4. Indirect Overhead, 5O$ of 1 + 2 + 3 0.786
5. Payroll Overhead, 15$ of 1+2+3+4 0.353
6. Maintenance Materials, 2.4$ of Investment 1.090
7. Miscellaneous Supplies,
15$ of Maintenance Materials 0.164
8. Utilities, Chemicals and Catalyst
Electricity, 23,970 KW at 9 mills/KWH 1.512
Cooling Water, 16,770 GPM at 2^/1000 gal 0.123
Make-up Acceptor, 6,60O Ib/hr at 3.5 $/Ton 0.071
K2C03, 7.1 MM SCFH Gas Treated
at (1,000 $/yr)/MM SCFH 0.006
Total Direct Operating Costs 5.676
Escalation 1.064
Total 6.740
-------
28,
TABLE VII-B
Direct Operating Costs
Ex Coal Cost
High ,.asification Case
Direct Operating Costs (6.132 'hr/vr)
1. Direct Operating Labor
18 men/shift at (40,000 $/yr)/(man/shift)
2. Maintenance Labor
1.6$ of Investment
Sub-Total, Direct Labor
3. Direct Supervision, 15$ of 1+2
4. Indirect Overhead, 50$ of 1 + 2 + 3
5. Payroll Overhead, 15$ of 1+2+3+4
6. Maintenance Materials, 2.4$ of Investment
7. Miscellaneous Supplies,
15$ of Maintenance Materials
8. Utilities, Chemicals and Catalyst
Electricity, 26,936 KW at 9 mills/KWH
Cooling Water, 25,O5O GPM at 2^/1000 gal
Boiler Feed Water, 813 GPM at 25^/1000 gal
Make-up Acceptor, 6,80O Ib/hr at 3.5 $/Ton
K2C03, 6.7 MM SCFH Gas Treated
at (l,OOO $/yr)/MM SCFH
Total Direct Operating Costs
Escalation
Total
$ MM/yr
0.72O
1.085
1.805
0.271
1.O38
0.467
1.627
0.244
1.699
0.185
0.074
0.073
P. 006
7.489
1.404
8.893
-------
FIGURE 5
COST OF FUEL GAS AS A FUNCTION OF THE
PRICE RECEIVED FOR CHAR:COAL AT 40y(/MM BTU
29.
3
•P 4J
ca a
0)
s
0)
10 M
01 K
O V
OT
§ *P
-------
30.
It is noted in the cases reported here, as well as in Case II of
the Annual Report,V1) that a major part of the total cost of the operation
is in the sulfur recovery operation.
It is of interest to compare the costs given here, for the liquid-
phase Glaus system, with the more "conventional" once-through Claus system
as costed in a recent report by Shell. (8)
To get the cost of the liquid-phase Claus system, one must exclude
from the recovery system costs, the "Squires" reactors, the C02 recovery and
handling system and the acceptor disposal systems uniquely associated with
our system. Thus, the investment in the liquid-phase Cl^us system itself is
only about 54% of the total recovery cost. The installed cost for the liquid-
phase Claus unit to produce 450 long tons of sulfur/stream day is thus about
$7.5 MM for the minimum gasification case or about $17 M/long ton of
sulfur/stream day. The "comparable" Shell casev8) at 10 atm handling 2% H2S
gas in a "conventional" Claus has an installed cost, excluding utilities and
offsites of $3.05 MM for 116 long tons of sulfur/stream day or about $26
M/long ton sulfur/stream day. The Shell system seems more expensive, but the
bulk of the cost is tied up in the incinerator system which is not required
in our particular case. The "conventional" Claus, as expected, is cheaper
without the incinerator system, i.e., about $1O M/long ton of sulfur/stream
day. The "conventional" Claus cannot be used in the "Squires" system because
of the high steam content of the gas treated.
-------
31,
VII. BIBLIOGRAPHY
1. "Development of the CO2 Acceptor Process Directed Towards Low-Sulfur
Boiler Fuel." Curran, G.P.; Clancey, J.T.; Fink, C.E.;
Pasek, Bedrich; Pell, Melvyn and Gorin, Everett.
Annual Report from Consolidation Coal Co. to Control Systems
Division, Office of Air Programs, Environmental Protection Agency,
Under GAP Contract EHSD-71-15 for Period
Sept. 1, 1970 to Nov. 1, 1971
NTIS Accession No. PB 210-840.
2. Consolidation Coal Co., Research & Development Report No. 16 to
The Office of Coal Research, U.S. Dept. of the Interior
Under Contract No. 14-01-0001-415, Interim Report No. 2,
"Low-Sulfur Boiler Fuel Using the Consol C02 Acceptor Process."
Nov. 1967. U.S. Dept. of Commerce, National
Tech. Information Service PB-17610.
3. Ruth, L.A.; Squires, A.M.; and Graff, R.A.; "Desulfurization of Fuels
With Calcined Dolomite: Part III - First Results for Reaction
of CaCO3 with H2S." Paper presented at ACS Meeting,
Los Angeles, March 31, 1971.
4. Zielke, C.W.; Curran, G.P.j Gorin, Everett; and Goring, G.E.;
Ind.' Eng. Chem., 46. 53 (1954).
5. Batchelor, J.D.; Gorin, Everett; and Zielke, C.W.,
Ind. Eng. Chem., 52. 161 (i960).
6. United States Patents Issued to Consolidation Coal Company:
a. U.S. 2,717,868 - Sept. 13, 1955
b. U.S. 2,824,047 - Feb. 18, 1958
c.
d.
e.
f.
U.S. iJ,71Y,8bH - Sept. J.3, iyt)t)
U.S. 2,824,047 - Feb. 18, 1958
U.S. 2,927,063 - Aug. 23, 1960
U.S. 2,950,229 - Aug. 23, 1960
U.S. 2,950,231 - Aug. 23, 1960
U.S. 3,101,303 - Aug. 20, 1963
7. Curran, G.P.; Fink, C.E.;, and Gorin, Everett; Production of Low-Sulfur
Boiler Fuel - Application of C02 Acceptor Process." Paper
presented before the Second International Conference on
Fluidized Combustion, Hueston Woods, Ohio, Oct. 4-7, 1970.
8. Shpall, R.T., "Claus Technical and Economic Study." Shell Development
Co. to Control Systems Division, Office of Air Programs,
Environmental Protection Agency,
Contract EHSD-71-15, April, 1972.
9. Consolidation Coal Company unpublished work.
-------
32.
VIII. APPENDICES
Appendix A
Mass and Heat Balances
Minimum Gasification Case
Figure A-l
Table A-I
A-II
A-III
A-IV
A-V
A-VI
A-VII
A-VIII
A-IX
A-X
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
34.
Table A-I
Mans and Heat Balance
Prcoxidizor - Gaslflor
Minimum Gasification Case: See Figure A-l
aals: 1 hour Datum: 60° F, H2<> CA)
Stream
nput
Feed Coal 1
MF Coal
Moisture
Sub-Total
Air to Prcoxidizer 2
°2
Kn
Moisture
Sub-Total
Air to Gasifler 5
°2
N2
Moisture
Sub-Total
Makeup Acceptor 3
MgCOs-CaCOs
Inert . .
Sub-Total
Recycle Acceptor .4
JIgO-CaC03
Inert
Sub-Total
Recycle Gas ' . 6
Beat of Reaction:
Totals
C'
7
..-t Gas
CH4
H2
CO
C02
"2
NHs .
H2S
H20 (v)
Sub-Total (Net Gas)
Recycle Gas
Sub-Total (Stream 7)
Char 8
Acceptor 9 .
HeO'CaCO-
MgO-CaS
Inert
Sub-Total
Heat of Reaction:
Pounds
1,087,600
69,400
1,157,000
92,400
304,500
1,100
396,000
157,200
517,900
1,800
675,900
6,100
500
6,600
912,700
110,600
1,023,300
747,400
Mols Mol %
-
3,832
2,888 20.9
10,869 78.7
60 0.4
13,817 10O.O
4,911 20.9
18,482 78.7
101 0.4
23,494 10O.O
33
-
6,500
- ' .-
30,477
H
52,200
7,700
59,900
-
-
100
100
-
-
200
200
-•
-
-
-
-
-
25,600
C
759,200
-
759,200
.
—
.
<•
"
-
-
•-
800
.-
800
78,100
-
78,100
143,300
HHV of Coal - (HHV of Net Gas + HHV of Ch
4,009,200
129,200
15,200
344,200
274,800
822,300
8,700
600
. 70,400
1,665,400
747,400
2,412,800
604,300
734,800
146,200
111,100
992,100
-
8,055 11.9
7,539 11.1
12,287 18.1
6,243 9.2
• 29,347 43.2
511 0.7
19
3,906 5.8
C7.907 100.0
30,477
98,384
-
5,233
1,300
- -
6,533
1DIV of HoS which reacted
H2b + MgO
•CiiCOs > HKO-
85,800
32,500
15,200
-
• -
1,500
-
7,90O
57,100
25,600
82,700
3,100
.
-
-
-
981,400
96,700
. -
147,600
75,000
-
. -
- .
319,300
143,300
462 , 600
456,000
62,800
-
-
62,800
with Acceptor:
CaS + CO2 + H^O
Elemental
N
13,000
-
13,000
'
304,500
-
304,500
-
517,900
517,900
-
-
-
- ,
-
-
372,300
Balance, Pounds
0 S
82,600 46,800
61,700
144,300 46,800
92,400
— . —
1,000
93,400
157,200
-
1,600
158,800
2,600 . -
-
2-, 600
312,000
-. -
312,000 -
206,000 200
ar) = (13,812.4 - (5,595.3 +
1,207,700
-
•
-
•
822,300
7,200
-
•
829,500
372,300
1,201,800
5,900
-
.
-
-
1,300 mol x
(A): 1,300
Calcination of Makeup ; ccoptor, MRCOs'CaCOs » M^O
Heat Loss
Totals
4,009,200
_
83,800
981,400
1,207,700
917,100 47,000
.
.
196,600
199,800
-
600
62,500
458,900 600
206,000 200
664,900 800
1,000 4,500
251,200
41,700
-
251,200 41,700
242,100 Btu/tnol
Ash or
Inert
133,800
-
133,800
-
—
-
- •
•
-
-
-
-
500
500
-
110,600
.110,600
-
6,740.0)
244,900
-
- •
• -
-
-
-
-
-
-
133,800
-
-
111.100
111,100
X
Temp.
UflO'Ca =F
60
s - 60
-
398
398
398
-
1,318
1,318
1,318
-
2,700 60
. 60
2,700
522,600 1,300
1,300
522,600
1,350
) MM Btu -
329,300
1,500
,500
,500
,500
,500
,500
,500
,500
-
1,500
-
1,500
420,800 1,425
104,500 1,423
1,425
525,300
Enthalpy
Ot
/SH '
Btu/lb MM Btu
0
0
0
73.9
84.4
1,213.5
305.5
327.6
1,672.7
0
0
0
326.0
300.0
488.2
1
2
1,190
5,063
382.8
377.6
378.7
994
409.2
1,772.8
1
499
364.5
275.7
334.4
mol x 30,620 Btu /mol =
•CaCOs + CO2: 33
917,100 47,000
mol x 51,
244,900
200 Btu/raol m
325,300
2
0
0
0
6.8
25.7
1.3
33.8
48.0
169.7
3.0
220.7
0
0
0
297.5
33.2
330.7
364.9
,477.1
,427.2
153.7
77.0
131.8
103.8
311.4
8.6
0.2
124.8
911.3
409.0
,320.3
301.5
267.8
40.3
37.1
345.2
314.7
39.8
1.7
104.0
,427.2
Hoat (HHV)
of Combustion
MM Btu
13,812.4
3,082.7
928.7
1,495.5
-
-
84.1
4.3
-
5,595.3
6,740.0
-------
Table A-II
Mass and Heat Balance
Reactor (Squires)
Minimum Gasification Case: See Figure A-l
35,
Basis: 1 hour
Input
Acceptor
MgO • CaCOa
MgO • CaS
Inert
Sub-Total
Gas
CO 2
H20 (v)
Datum: 60°F, H;jO (,£)
Temp.
Stream Pounds Mols °F
9
734,800 5,233 1425
146,200 1,300 1425
111,100 - 1425
992,100 6,533
10
805,400 18,300 1147
329,700 18,300 1147
Sub-Total 1,135,100 36,600
Heat of Reaction
MgO • CaS + C02 + H20 (v) > MgO . CaCOa + H2S
1.300 mols x 30.620 Btu/mol =
Totals
Output
Recycle Acceptor
MgO . CaCOa
Inert
Sub -Total
Spent Acceptor
MgO • CaCOa
Inert
Sub-Total
Gas
C02
H2S
H20 (v)
Sub-Total
Heat Loss
Totals
2,127,200
4
912,700 6,500 1300
110,600 - 1300
1,023,300
11
4,600 33 1300
500 - 1300
5,100
12
748,200 17,000 1300
44,300 1,300 1300
306,300 17,000 1300
1,098,800 35,300
2,127,200
Enthalpy
Ah AH
Btu/lb MM Btu
x "364.5 267.8
275.7 40.3
334.4 37.1
345.2
273.6 220.4
1,581.3 521.4
741.8
39.8
1,126.8
326.0 297.5
300.0 33.2
330.7
326.0 1.5
300.0 0.1
1.6
318.1 238.0
341.7 15.1
1,661.8 509.0
762.1
32.4
1,126.8
-------
Table A-III
Mass and Hoot Balance
Heat Exchanger C-201
Minimum Gasification Case: See Figure A-l
36.
Basis: 1 hour
Datum: 60°F,
Stream
Pounds
Hols
Enthalpy
' /ih.
Btu/lb
MM Btu
High Temperature Side;
Input
Oas
Output
Gas
CH4
H2
CO
CO 2
H2
MH3
H2S
H:P
Sub-Total
Oas
CH4
H2
CO
C02
H2
NH3
H2S
21
23
Sub-Total
Beat Exch
Heat Loss
Totals
ange
UJ
2,412,800
98,384
1500
S47.2
1,320.3
1
1
2
93
11
249
199
596
6
31
,207
93
11
249
198
593
6
30
,204
,412
,700
,000
,600
,300
,500 .
,300
500
,000
,900
,500
,000
,000
,700
,000
,300
500
,900
,900
,800
5
5
8
4
21
2
49
5
5
8
4
21
2
49
98
,842'
,468
,912
,528 .
,285
371
14
,833
,253
,828
,454
,890
,517
,232
370
14
,826
,131
,384
1327
1327
1327
1327
1327
1327
1327
1327
1327
1327 '
1327
1327
1327
1327
1327
1327
1,003
4,439
333
326
330
852
353
1,677
1,003
4,439
333
326
330
852
353
1,677
.4
.1
.1
.3
.O
.5
.4
.1
.1
.3
.0
.3
94
48
83
65
196
5
0
85
579
93
48
83
64
196
5
0
83
577
163
0
1,320
.0
.8
.2
.0
.9
.4
.2
.6
.1
.8
.8
.0
.8
.4
.4
.2
.4
.8
.2
.2
.3
Low Temperature Side;
Input
Air
N2
H20 (v)
Sub-Total
Beat Exchange
Totals
Output
Air
°2
Totals
(v)
137,200
517,900
1,800
676,900
398
398
398
73.9
84.4
1,213.5
157,200
517,900
1,800
676,900
4,911
18,482
101
23,494
1318
1318
1318
JOS. 5
327.6
1,672.7
48.0
169.7
3.0
220.7
(1) By difference, to force the heat balance.
-------
Table A-IV
Mass and Hoat Balance
Boat Exchanger C-202
Minimum Gasification Case; Seo Figure A-l
37.
Basis: 1 hour
Datum: 60°F, H2O
Enthalpy
Stream
Pounds
Mo Is
High Temperature Side;
jnput
Gas
Output
Oas
CH4
H2
CO
CO2
N2
NH3
H2S
HzO (v)
Sub-Total
Beat Exchange
Heat
Totals
21
22
1,207,900
49,253
1327
379.1
1
1
93
11
249
199
596
6
51
,207
,207
,700
,000
,600
,300
,500
,300
500
,000
,900
,900
5
5
8
4
21
2
49
49
,842
,468
,912
,528
,285
371
14
,833 .
,253
,253
904
904
904
904
904
904
904
. 904
597
2,936
216
205
215
531
223
1,456
.4
.5
.1
.1
.9
.3
.2
56
32
54
40
128
3
0
74
389
189
0
579
.0
.3
.0
.9
.3
.4
.1
.3
.3
.5
.3
.1
Low Temperature Side;
Input
Gas
CO2
HzO (v)
Sub-Total
Gas
O>2
H20 (v)
Sub-Total
Heat Exchange
Totals
15
16
57,200
700
57,900
748,200
329,000
1,077,200
1,135,100
1,300
40
1,340
200
200
676
676
?9.2
1,124.9
143.9
1,343.9
1.7
0.8
2.5
107.7
442.1
10
805,400
329,700
1,135,100
18,300
18,300
36,600
1147
1147
273.6
1,581.3
(1) By difference, to force the heat balance.
-------
Table A-V
Mass and Heat Balance
Compressor JC-203
Minimum Gasification Case; See Figure A-l
38,
Basis: 1 hour
Input
Cos
Datum: 60°F, HjO
Stream
23
Pounds
1,204,900
Mo Is
49,131
I son tropic Work ='Hp TI J(P2/Pl) -lj (49,131 mols) =
Inefficiency•
(9.4/0.89) - 9.4
Totals
Gas to C02 Plant
CH4
H2
CO
C02
N2
NH3
H2S
H20
Sub-Total
Totals
13
Temp.
op
1327
1,204,900
58,000
6,800
154,500
123,200
369,100
3,900
300
31,600
747,400.
35,500
4,200
94,500
75,500
225,900
2,400
200
19,300
457,500
49,131
3,615
' 3,383
5,515
2,802 ,
13,170
230
9
1,753
30,477
2,213
2,071
3,375
1,715
8,062
140
• 5 ..
1,073
18,654
-
1350
1350
1350
1350
1350
1350
1350
1350
1350
1350
1350
1350
1350
1350
. 1350
1350
Enthalpy
1,204,900
49,131
Btu/lb
-\
-
1,027
4,521
339.9
332.8
' 336.5
870 . 8
360.3
1,690.1
1,027
4,521
339.9
332.8
336.5
870.8
360.3
1,690.1
•• ^ >
MM Btu
377.8
9.4«>
1.2(«
588.4
59.8
30.7
52.5
41.0
124.2
3.4
0.1
53.4
364.9
36.5
19.0
32.1
25.1
76.0
2.1
0.1
32.6
223.5
588.4
(1) 'cp = 9.318 Btu/lb nol °B
TI = 1787°R
» = Cp/(cp-1.99) = 1.2731
Pa/Pi •> 1,0550 Iff ~ 12 psl)
(2) 89% efficiency.
-------
Table A-VI
Mass and Heat Dnlnnce
Hoat Exchanger C-301
Minimum Gniiification Caac; See Figure A-j.
39.
Basis: 1 hour
Datum: 60°F, H2O
Stream
Pounds
Mols
Temp.
oF
Enthalpy
Ah
Btu/lb
AH
MM Btu
High Temperature Side!
Input
Gas to CO2 Plant
13
457,500
18,654
1350
223.5
Output
Gas to CO2 Plant
CH4
HZ
CO
C02
»2
»H3
H2S
HzO
Sub-Total
• Heat Exchange
Heat Loss*1'
Totals
24
35,900
4,200
94,500
75,500
225,900
2,400
200
19,300
457,500
2,213
2,071
3,375
1,715
8,062
140
5
1,073
18,654
524
524
524
524
524
524
• 524
524
293.5
1,605
116.6
105.1
116.3
275.5
117.0
1,271.7
10.4
6.7
11.0
7.9
26.3
0.7
•
24.5
87.5
135.9
0.1
457,500
18,654
223.5
Low Temperature Side;
Input
Gas
from CO2 Plant
26
CH4
H2
CO
C02
"2
NH3
H2S
H20
Sub-Total
Heat Exchange
Totals
35,500
4,200
94,500
18,300
225,900
2,400
200
: 31,400
412,400
2,213
2,071
3,375
415
8,062
140
5
1,743
18,024
230
230
230 .
230 :
230
230
230
230
97.5
385.1 .
42.3
35.8
42.3
97.3 .
41.3
1,138.1
3.5
2.5
4.0
0.7
9.6
0.2
-
35.7
56.2
' 135.9
412,400
18,024
192.1
14
35,500
4,200
94,500
18,300
225,900
2,400
200
31,400
412,400
2,213
2,071
3,375
415
8,062
140
5
1,743
18,024
1150
1150
1150
1150
1150 .
1150
1150
1150
824.8
3,806
. 283.8
274.5
281.2
713.7
297.3
1,582.9
29.3
16.0
26.8
. 5.0
63.5
1.7
0.1
49.7
192.1
(1) By difference, to force the heat balance.
-------
Table A-VIII
Mass and Heat Balance for Expander
Product Gas at 25.7 psia
Minimum Gasification Case; See Figure A-l
Basis: 1 hour Datum: 60°F, H20 (J
Stream
Input
Gas 22
Gas from C02 Plant 14
Totals
Output
Product Gas 17
CH4
H2
CO
C02
N2
NH3
H2S
H20 (v)
Sub-Total . ' .
Heat Loss*1'
Totals
!-)
Pounds
1,207,900
412,400
1,620,300
129,200
15 , 200
344,100
217,600
822,400
8,700
700
82,400
1,620,300
1,620,300
Mols
49,253
18,024
67,277
8,055
7,539
12,287
4,943
29,347
511
19
4,576
67 , 277
67 , 277
Temp.
OF
904
1150
969
969
969
969
969
969
969
969
Enthalpy
Ah
Btu/lb
655.1
3,165
234.1
223.2
232.3
578.7
242.4
1,489.2
MM Btu
389.3
192.1
581.4
84.6
48.1
80.6
48.6
191.0
5.0
0.2
122.7
580.8
0.6
581.4
(1) By difference, to force the heat balance.
-------
Table A-VIII
Mass and Heat Balance for Expander
Product Gas at 25.7 psia
Minimum Gasification Case; See Figure A-l
Basis: 1 hour Datura: 60°F, H2Q
Input
Product Gas at 206 psia
Stream
17
Pounds.
1,620,300
Mols
67,277
969
Enthalpy
580.8
Output
Work= -0.91
(P2/Pl)
(N-D/N
Product Gas at 25.7 psia
CH4
H2
CO
C02
N2
NH3
H2S
H20
Sub-Total
Totals
-l] (67,
277 mols) =
1,620,300
67,277
286.0
(1)
129,200
15,200
344,100
217,600
822,400
8,700
700
82,400
1,620,300
8,055
7,539
12,287
4,943
29,347
511
19
4,576
67,277
462
462
462
462
462
462
462
462
249.1
1,389
100.8
89.8
100.5
236.3
100.0
1,242.9
32.2
21.1
34.7
19.5
82.7
2.1
0.1
102.4
294.8
580.8
(1) 91% efficiency
Cp = 8.385 Btu/lb mol °R
Tx = 1429 °R
N = c /Ccn-1.99) = 1.3112;
-------
42.
Table A-IX
Mass Balance
CC>2 Removal System
Minimum Gasification Case: See Figure A-l
Basis: 1 hr.
Input
Gas to CO9 Plant
CH4
H2
CO
C02
N2
NH3
H2S
H20 (v)
Sub-Total
Make-up Water
Totals
Output
Gas from C02 Plant
CH4
H2
CO
C02
N2
NH3
H2S
H20 (v)
Sub-Total
Make-up CO?.
C02
H20 (v)
Sub-Total
Stream Pounds
2E>
35,500
4,200
94, 500
75,500
225,900
2,400
200
19,300
457,500
20 12,800
470,300
26
35,500
4,200
94,500
18,300
225,900
2,400
200
31,400
412, 400
15
57,200
700
57,900
•
Mols
2,213
2,071 >
3,375
1,715
8,062
140
5
1,073
18,654
710
19,364
2,213
2,071
3,375
415
8,062
140
5
1,743
18,024
1,300
40
1,340
230
230
200
Totals 470,300 19,364
-------
Table -A-X
43.
Mass Balance
Sulfur Recovery
Minimum Gasification Case;
See Figure A-l(1)
Basis: 1 hour
Input
Gas to Sulfur Recovery
co2
H2S
H20 (v)
Sub-Total
Combustion Air
02
•^2
H20
Sub-Total
Totals
Output
Gas
C02
H2O (v)
Sub-Total
Vent Gas
S02
N2
H2O (v)
Sub-Total
Product Sulfur
Make H20
Stream Pounds
12
748,200
44,300
306,300
1,098,800
21,000
69,200
200
90, 400
1,189,200
16
748,200
329,000
1,077,200
18
400
69,200
100
69, 700
19 41,500
800
Mols
17, 000
. 1,300
17,"000
35,300
657
2,471
14
3,142
17,000
18,260
35,260
7
2,471
8
2,486
1,293
46
Temp,
op,
1300
676
90
Totals
1,189,200
(1) Small streams to and from the Spent Acceptor Stripping Column
have been neglected.
-------
44.
VIII. APPENDICES - Cont'd.
Appendix B
Mass and Heat Balances
High Gasification Case
Figure B-l
Table B-I
B-II
B-III
B-IV
B-V
B-VI
B-VII
B-VIII
B-IX
B-X
B-XI
B-XII
B-XIII
B-XIV
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
Mass and Heat Balance
Preoxidizer - Gpsiflor
Basis: 1 hour Datum: 60° F, HjO (i) ' •
• ' . Elemental Balance, Pounds '
Input
Peed Coal
Uolstun
Sub-Total
Air to Preoxldlzer
02
»2
Moisture
Sub-Total
Air to Gaslfler
02
Moisture
Sub-Total
Steam to Gaalfler
U k A
Inert
Sub-Total
Recvcle Acceptor
i!gO-CaC03
Inert
Sub-Total
Recycle Gas
Beat of Reaction:
Straao Pounds
1
1,087.100
60,400
1,157,000
t
82,400
304,500
1,000
397,900
a
511,800
1,680,200
6,000
2,20-1,000
27 386,000
3
6,200
600
6,800
4
947,400
114.800
1,062,200
Hols Hal % a C
82,200 759,200
3,852 - 7.700
59,900 .759,200
2,888 20.9
10,869 78.7
59 0.4 100
13,816 100.0 100
15,994 20.9
60,136 78.7
329 0.4 700
70,500 100.0 700
21,428 > 43,200
34 - - 800
800
8,747 - - 81,000
81,000
8 1,714,900 72,283 - 46,100 249,300
HHV of Coal - (HHV of Net Gas + HHV of Char)
Ash or '
H OB Inert UgO*Ca
13,000 82,600 46,800 133,800
81.700 ...
13,000 144,300 '-iS.SJO 133,800 -
92,400. ...
304,300 - - ...
. 900
304,500 • 93,300 ...
311,800 ...
1,686,200 -
5,300 ...
1, 686,200 317,100 ~ . - =
342,800 - -
- ' 2,700 - - 2,700
. . «nn
2,700 - 600 2,700
323,900 - - 842,300
- - 114,800
323,900 . 114,800 342,300
897,200 521,900 400 -
• (13,812.4 - (8,272.8 -f 3,236.4)) MM Btu •
Trap,
•P
«O
60
398
398
398
1,200
1,200
1,200
1,200
60
an
1,300
1,300
1,200
EnthnlDT Beat (HHV)
Bhj/lb
0
0
73.9
84.4
1,213.3
275.0
294.7
1,609.3
1,809.3
0
S
0
326.0
300.0
438.9
£U of Combustion
tni Btu ID! Btu
0 13,812.4
0
0
23.7
1.2
140.7
496.9
9.7
84773
821.2
0
0 . . .
308.9
34.4
343.3
787.4
2,303.2
6,928,800
150,000 1,090,300 2,900,900 1,946,000 47,200 249,200 545,200
4,736.1
EV
S°2
"2
SH3
H-)S
HaO (v)
Sub-Total (Met Gas)
Recycle Gas
Sub-Total (Stream 7)
Char
Acceptor
l!gO'CaC03
BgO'CaS
Inert
Sub-Total
Beat of Reaction: .
Beat Loss
Totals
55,500
53,200
830,400
577,500
1,990,600
8,700
900
301,800
3,818,600
1,714,900
3,533,500
365,500
762,600
151,800
115,400
1,029, £00
Hire of n.,3
H.,S + UgO-
3,459
26,387
29,646
13,123
71,053
511
26
16,751
1G0.956
72,283
233,239
2.1
16.4
18.4
8.2
44.2
0.3
-
10.4
100.0
-
-
14,000
53,200
-
-
-
1,500
100
33,800
102,600
46,100
1-18,700
41,500
.
356,100
157,600
-
-
-
-
555,200
2-19,300
b'0-1,500
1,300 . 220,600
5,431
1,350
.
which
CaCO-j— »
reacted
MgO-CaS
.
-
_
-
65,200
-
-
65,200
-
-
-
.
1,990,600
7,200
.
-
1,997,800
897,200
2,895,000
5,900
_
-
_
-
.
-
474,300
419,900
-
.
.
268,000
1,162,200
521,900
1,684.100
1,200
260,700
-
-
260,700
...
.. . .
. _
. .
. . - -
. . .
800
-
800
400
1,200
2,700 133,800
436,700
43,300 - 108,500
115,400
43,300 115,400 545,200
1,700
1,700
1,700
1,700
1,700
1,700
1,700
1,700
1,700
1,700
1,700
1,700
1,700
1,442
3,774
440.8
438.3
' 435.9
1,127
, 472.2
1,884.9
343.8
432.0
333.3
411.6
' •
with Acceptor: 1,350 mol x 242,100 Btu/mol =
+ CO,
Calcination of Uakoup Acceptor,
+ I!;,o CU: 1
MecOj-cacOj-*
,.150 mol x
HuO-CaC03
30,620 Btu/mol
•*• CO2: 34 mol
=
x 51,200 Dtu/mol =
80.0
307.2
366.0
253.1
867.7
9.8
0.4
368.9
2,453.1
1,101.7
3,554.8
199.3
344.7
30.9
47.3
443.1
326.8
41.3
1.7
168.9
1,323.8
3,251.1
3,607.8
-
.
83.9
6.2
-
8,272.8
3.236.4
6,928,800
150,000 1,090,300 2,900,900 1,940,000 47,200 249,200 545,200
4.736.1
-------
Table B-II
Mass and Heat Balance
Reactor (Squires)
High Gasification Case: See Figure B-l
47,
Basis: 1 hour
Input
Acceptor
MgO.CaCOa
MgO- CaS
Inert
Sub-Total
Gas
C02
H20 (v)
Sub-Total
Heat of Reaction
MgO. CaS + CO2
1,350 mols x
Totals
Output
Recycle Acceptor
MgO- CaC03
Inert
Sub-Total
Spent Acceptor
MgO- CaCOs
Inert
Sub-Total
Gas
C02
H2S
H20 (v)
Sub-Total
Heat Loss
Totals
Datum: 60° F, H20 (I)
Stream Pounds Mols
9
762,600 5,431
151,800 1,350
115,400
1,029,800
10
836,600 19,009
342,500 19,009
1,179, 100
+ H20(v)-*~MgO.CaCO3 + H2S
30,620 Btu/mol =
2,208,900
4
947,, 400 6, 747
114,, 800
1,062,, 200
11
4,, 700 34
600
5,, 300
12
777,200 17,659
46,000 1,350
318,200 17,659
1,141,400 36,668
2,208,900
. Enthalpy
Temp. Ah AH
°F Btu/lb MM Btu
1700 452.0 344.7
1700 335.3 50.9
1700 411.6 47.5
443.1
943 215.9 180.6
943 1,476.0 505.5
686.1
41.3
1,170.5
1300 326.0 308.9
1300 300.0 34.4
343.3
1300 326.0 1.5
1300 300.0 0.2
1.7
1300 318.1 247.2
1300 341.7 15.7
1300 1,661.8 528.8
791.7
33.8
1,170.5
-------
Table B-III
48.
Mass and Heat Balance
Heat Exchanger C-201
High Gasification Case: See Figure B-l
Basis: 1 hour Datum:
High Temperature Side:
Input
Gas
Output
CH4
H2
CO
C02
N2
NH3
H2S
H20
Sub-Total
Heat Exchange
Heat Loss^5
Totals
Low Temperature Side:
Input
Air
02
N2
H20 (v)
Sub-Total
Heat Exchange
Totals
Output
Air
02
N2
H20 (v)
Totals
60°F, H20
Stream
7 5
28
1
2
5
5
31
1
2
2
5
1
(.£)
Pounds
,533,500
80, 400
77,100
,203,300
836,900
,884,600
12,600
1,300
437, 300
,533,500
, 533, 500
511,800
,686,200
6,000
, 204, 000
, 204, 000
511,800
, 686, 200
6,000
Mols
233,239
5,012
38,237
42,960
19,016
102,962
740
38
24,274
233,239
233,239
15,994
60,186
329
76,509
76,509
15,994
60,186
329
Temp.
oF
1700
1627
1627
1627
1627
1627
1627
1627
1627
946
946
946
1200
1200
1200
Enthalpy
Ah
Btu/lb
1,334
5,526
419.5
415.9
414.9
1,102
451.5
1,844.6
209.8
226.2
1,477.5
275.0
294.7
1,609.3
AH
MM Btu
3,554.8
107.3
426.1
504.8
348.1
1,196.8
13.9
0.6
806.6
3,404.2
149.6
1.0
3,554.8
2,204,000
76,509
107.4
381.4
8.9
497.7
149.6
647.3
140.7
496.9
9.7
647.3
(1) By difference, to force the heat balance.
-------
49.
Table B-IV
Mass and Heat Balance
Heal; Exchanger C-205
High Gasification Case: See Figure B-l
Basis: 1 hour Datum: 60°F H20 (I)
Stream
High Temperature Side:
Input
Gas 28
Output
Gas 29
CH4
H2
CO
C02
N2
NH3
H2S
H20
Sub-Total
Heat Exchange
Heat Loss*1?
Totals
Low Temperature Side:
Input
Recycle Gas 30
Heat Exchange
Totals
Output
Recycle Gas 6
CH4
H2
CO
C02
N2
NH3
^s
H20
Totals
]?ounds
5,533,500
80, 400
77,100
1,203,300
836, 900
2,884,600
12,600
1,300
437,300
5,533,500
5,533,500
1,714,900
1,714,900
24,900
23,900
372,900
259, 400
894, 000
3,900
400
135,500
1,714,900
Mols
223,239
5,012
38,237
42,960
19,016
102,962
740
38
24,274
233,239
233,239
72,283
72,283
1,553
11,850
13,314
5,893
31,909
229
12
7,523
72,283
Temp.
OF
1627-
1483
1483
1483
1483
1483
1483
1483
1483
700
1200
1200
1200
1200
1200
1200
1200
1200
Enthalpy
Ah
Btu/lb
~
1,171
5,001
377.9
372.4
373.9
979.7
403.6
1,763.4
•. -
-
873.9
3,983
297.6
289.0
294.7
752.2
312.7
1,609.3
AH '
MM Btu
3,404.2
94.1
385.6
454.7
311.7
1,078.6
12.3
0.5
771.1
3,108.6
294.6
1.0
3,404.2
492.8
294.6
787.4
21.8
95.2
111.0
74.9
263.4
2.9
0.1
218.1
787.4
(1) By difference, to force the heeit balance.
-------
Table B-V
50.
Mass and Heat Balance
Heat Exchangers C-203 and C-204
High Gasification Case: See Figure B-l
Basis: 1 hour Datum: 60°F, H20
Stream
Pounds
Mols
Enthalpy
Ah
Btu/lb
AH
MM Btu
High Temperature Side;
Input
Gas 29
Output
Gas to C-202 21
CH4
H2
CO
C02
N2
NH3
H2S
H20
Sub-Total
Gas to C-206 23
CH4
H2
CO
C02
N2.
NH3
H2S
H20
Sub-Total '
Heat Exchange
Heat Loss(1>
Totals
Low Temperature Side:
Input
Boiler Feed Water 35
Heat Exchange
Totals
5,533,500 233,239
406,300
406,300
1483
5,533,500 233,239
22,554
22,554
214
154
3,108.6
1,
3,
1,
2,
49,
47,
739,
514,
773,
7,
268,
402,
31,
29,
463,
322,
110,
4,
168,
131,
400
400
900
600
700
700
800
900
400
000
700
400
300
900
900
500
400
100
3,
23,
26,
11,
63,
14,
143,
1,
14,
16,
7,
39,
9,
89,
082
512
416
693
310
455
23
926
417
930
725
544
323
652
285
15
348
822
1200
1200
1200
1200
1200
1200
1200
1200
1200
1200
1200
1200
1200
1200
1200
1200
873.
3,983
297.
289.
294.
752.
312.
1,609.
873.
3,983
297.
289.
294.
752.
312.
1,609.
9
6
0
7
2
7
3
9
6
0
7
2
7
3
43.
188.
220.
148.
522.
5.
0.
432.
1,562.
27.
118.
137.
93.
327.
3.
0.
271.
978.
566.
1.
2
8
2
7
7
8
3
8
5
1
3
9
1
4
7
2
0
7
0
4
3,108.6
62.1
566.0
628.1
Output
Slowdown
Steam
Totals
34
27
20,300
386,000
406,300
1,128
21.426
22,554
394
1200
340.2
1,609.3
6.9
621.2
628.1
(1) By difference, to force the heat balance.
-------
Table B-VI
51,
Mass and Heat Balance
Heat Exchanger C-202
High Gasification Case: See Figure B-l
Basis: 1 hour
High Temperature Side:
Input
Gas
Output
Gas
CH4
H2
CO
C02
NH3
H2S
H20 (v)
Sub-Total
Heat Exchange
Heat*1)
Totals
Low Temperature Side
Input
Gas
C02
H20 (v)
Sub- Total
Gas
C02
H20 (v)
Sub-Total
Heat Exchange
Totals
Output
Gas
C02
H20 (v)
Totals
itum: 60°F, H20 (I)
Stream Pounds
21 3, 402, 400
22
49,400
47, 400
739,900
514,600
1,773,700
7,700
800
268,900
3,402,400
3,402,400
15
59,400
800
60,200
16
777, 200
341,700
1,118,900
1,179,100
10
836, 600
342, 500
Mols
143,417
3,082
23,512
26, 416
11,693
63,310
455
23
14, 926
143,417
143,417
1,350
42
1,392
17,659
18,967
36,626
38,018
19,009
19,009
Temp.
°F
1200^
1104
1104
1104
1104
1104
1104
1104
1104
200
200
676
676
943
943
Enthalpy
Ah
Btu/lb
780.3
3,643
271.0
261.3
268.7
678.8
283.1
1,558.8
29.2
1,124.9
143.9
1,343.9
215.9
1,476.0
AH
MM Btu
1,562.5
38.5
172.7
200.5
134.5
476.6
5.2
0.2
419. 2
1,447.4
112,5
2.6
1,562.5
1,179,100
38,018
1.7
0.9
2.6
111.8
459.2
571.0
112.5
686.1
180.6
505.5
686.1
(1) By difference, to force the heat balance.
-------
Table B-VII
52.
Basis: 1 hour
Mass and Heat Balance
Heat Exchanger C-206
High Gasification Case: See Figure B-l
Datum: 60°F, H20
Stream
Pounds
Mols
Enthalpy
Temp.
Ah
Btu/lb
AH
MM Btu
High Temperature Side:
Input
Gas 23
Output
Gas 13
CH4
CO
C02
N2
NH3
H2S
H20
Sub-Total
Heat Exchange
Totals
Low Temperature Side:
Input
Gas from C02 Plant 26
CH4
H2
CO
C02
N2
NH3
H2S
H20
Sub-Total
Heat Exchange
Totals
Output
Gas from C02 Plant 14
CH4
H2
CO
C02
N2
NH3
H2S
H20
2,131,100 89,822
1200
31,000
29, 700
463, 400
322,300
1,110,900
4,900
500
168,400
2,131,100
1,930
14,725
16,544
7,323
39,652
285
15
9,348
89,822
1099
1099
1099
1099
1099
1099
1099
1099
775.6
3,625
269.6
259.9
267.4
675.0
281.6
1,556.2
2,, 131,100 89,822
6,100
5,800
90,500
3,500
216,900
1,000
100
27, 700
351,600
377
2,875
3,230
80
7,743
56
3
1,539
15,903
230
230
230
230
230
230
230
230
97.5
585.1
42.3
35.8
42.3
97.3
41.3
1,138.1
351,600 15,903
6,100
5,800
90,500
3,500
216,900
1,000
100
27,700
377
2,875
3,230
80
7,743
56
3
1,539
870
870
870
870
870
870
870
870
978.7
24.0
107.7
124.9
83.8
297.1
3.3
0.1
262.1
903.0
75.7
978.7
124.4
3.5
16.3
18.8
0.7
44.7
0.5
0.0
39.9
Totals
351,600 15,903
124.4
-------
53.
Table' B-VIII
Mass and Heat Balance
Heat Exchanger C-207
High Gasification Case: See Figure B-l
Basis: 1 hour Datum: 60° F, H20 (.&)
High Temperature Side:
Input
Gas
Output
Gas
CH4
H2
CO
C02
N2
NH3
H2S
H20
Sub-Total
Heat Exchange
Heat Loss*1'
Totals
Low Temperature Side:
Input
Air
02
N2
H20 (v)
Sub-Total
Heat Exchange
Totals
Out put
Air
Stream Pounds
13 2,131,100
24
31,000
29,700
463, 400
322,300
1,110,900
4,900
500
168, 400
2,131,100
2,131,100
511,800
1,686,200
6,000
2,204,000
2,204,000
31 2,204,000
Mols
89,822
1,930
14,725
16,544
7,323
39,652
285
15
9,348
89,822
89,822
15,994
60,186
329
76, 509
76,509
76, 509
Enthalpy
Temp. " Ah
°F Btu/lb
\
1099
671 404. 1
671 2,118
671 154.7
671 142. 6
671 154.0
671 371.0
671 157.0
671 1,341.5
398 73.9
398 84.4
398 1,213.5
946
AH '
MM Btu
903.0
12.5
62.9
71.7
46.0
171.1
1.8
0.1
225.9
592.0
310.3
0.7
903.0
37.8
142.3
7.3
187.4
310.3
497.7
497.7
(1) By difference, to force the heiat balance.
-------
54.
Table B-IX
Mass and Heat Balance
Compressor JO 203
High Gasification Case; See Figure B-l
Basis: 1 hour Datum: 60°F, H20
Enthalpy
Temp. Ah AH
Stream Pounds Mo Is °F Btu/lb MM Btu
_ 24 2,131,100 89,822 671 - 592.0
Isentropic Work = cp T1 [(P2/Pi) N" -lj (89,822 mols) = * 18.2(1)
Inefficiency = (18.2/0.89) - 18.2 =
Gas to CQ2 Plant 32
CH4
Totals 2,131,100 89,822 612.4
Output
Recycle Gas 30
CH4
H2
CO
co2
N2
NH3
H2S
H20
Sub-Total
24,900
23,900
372,900
259,400
894,000
3,900
400
135,500
1,714,900
6,100
5,800
90,500
62,900
216,900
1,000
100
32,900
416,200
1,553
11,850
13,314
5,893
31,909
229
12
7,523
72,283
377
2,875
3,230
1,430
7,743
56
3
1,825
17,539
700
700
700
700
700
700
700
700
700
700
700
700
700
700
700
700
426.9
2,219
162.3
150.2
161.5
390.4
165.0
1,355.5
426.9
2,219
162.3
150.2
161.5
390.4
165.0
1,355.5
10.6
53.0
60.5
39.0
144.4
1.5
0.1
183.7
492.8
2.6
12.9
14.7
9.4
35.0
0.4
-
44.6
119.6
H2
CO
co2
N2
NH3
H2S
H20
Sub-Total
Totals 2,131,100 89,822 612.4
(1) "c = 7.832 Btu/lb mol °R
TI = 1131 °R
N = c /(c -1.99) = 1.3406
P2/P1 = 1.0927 (AP~19 psi)
(2) 89% efficiency
-------
55.
Table B-X
Mass and Heat Balance
Heat Exchanger C-302
High Gasification Case; See Figure B-l
Basis: 1 hour
High Temperature Side;
Input
Gas to CO2 Plant
Output
Gas to C02 Plant
CH4
H2
CO
C02
N2
NH3
H2S
H20
Sub-Total
Heat Exchange
Totals
Low Temperature Side;
Input
Boiler Feed Water
Heat Exchange
Totals
Output
Boiler Feed Water
)atum: 60°F, H20 U)
Stream Pounds
32 416,200
25
6,100
5,800
90,500
62,900
216,900
1,000
100
32,900
416,200
416,200
33 406,300
406,300
Temp.
Mols °F
17,539 700
377 236
2,875 236
3,230 236
1,430 236
7,743 236
56 236
3 236
1,825 236
17,539
17,539
22,554 60
22,554
Enthalpy
Ah AH
Btu/lb MM Btu
„
119.6
101.1 0.6
605.9 3.5
43.8 4.0
37.1 2.3
43.7 9.5
100.7 0.1
42.8
1,140.8 37.5
57.5
62.1
119.6
0 0.0
62.1
62.1
35
406,300 22,554 213
153
62.1
-------
56.
Table B-XI
Mass and Heat Balance
Product Gas at 206 psia
High Gasification Case; See Figure B-l
Basis: 1 hour Datum: 60°F, EyO
-------
57.
Table B-XII
Mass and Heat Balance for Expander
Product Gas at 25.7 psia
High Gasification Case; See Figure B-l
Basis: 1 hour
Datum: 60°F, H^O CO
Stream
Pounds
Enthalpy
Mols
Ah
Btu/lb
AH
MM Btu
Input
Product Gas at
206 psia
17
3,754,000 159,320 1081
1,570.7
Output
Work = -0.91 C
Product Gas at
25.7 psia
CH4
H2
CO
C0£
N2
KP2/Pi) ~ -1J (159,320 mols) =
H2S
H20
Sub-Total
Totals
55,500
53,200
830,400
518,100
1,990,600
8,700
900
296,600
3,754,000
3,459
26,387
29,646
11,773
71,053
511
26
16,465
159,320
512
512
512
512
512
512
512
512
721.3
(1)
284.6
1,563
113.6
102.1
113.2
267.7
113.8
1,266.1
15.8
83.2
94.3
52.9
225.3
2.3
0.1
375.5
3,754,000 159,320
849.4
1,570.7
(1) 91% efficiency
Cp = 7957 Btu/lb mol °R
TI '= 1541°R
N = Cp/Ccp-1.99) =1.3335; (N-1)/N =0.2501
P2/Pl = 25.7/206 = 0.1248
-------
58.
Table B-XIII
Mass Balance
COj) Removal System
High Gasification Case; See Figure B-l
Basis: 1 hour
Stream
Input
Gas to CO2 Plant 25
CH4
H2
CO
C02
N2
NHs
H2S
H20 (v)
Totals
Output
Gas from C02 Plant 26
CH4
H2
CO
C02
N2
NH3
H2S
H20
-------
59.
Table B-XIV
Mass Balance
Sulfur Recovery
High Gasification Case; See Figure B-l
Basis: 1 hour
Stream
Input
Gas to Sulfur Recovery
CO 2
H2S
H20 (v)
Sub-Total
Combustion Air
~02
N2
H20 (v)
Sub-Total
12
Pounds
777,200
46,000
318,200
1,141,400
21,800
71,900
200
93,900
Mo Is
17,659
1,350
17,659
36,668
682
2,566
14
3,262
Temp.
°F
1300
Totals
1,235,300
Output
Gas
CO2
H20
Sub-Total
Vent Gas
SO 2
N2
H20 (v)
Sub-Total
Product Sulfur
Make H20
Totals
16
18
19
777,200 17,659
341,700 18,967
1,118,900 36,626
400
71,900
100
72,400
43,100
900
1,235,300
7
2,566
8
2,581
1,343
48
676
90
(1) Small streams to and from the spent acceptor stripping column
have been neglected.
------- |