PB 204 711
EVALUATION OF SO2 - CONTROL PROCESSES
M0 W. Kellogg Company
Piscataway, New Jersey
15 October 1971
NATIONAL TECHNICAL INFORMATION SERVICE
Distributed ... 'to foster, serve
and promote the nation's
economic development
and technological
advancement.1
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THE M.W. KELLOGG COMPANY
Pullman Incorporated
A Unifion of .
EVALUATION OF SO2 - CONTROL PROCESSES
TASK #5 FINAL REPORT
SUBMITTED TO
ENVIRONMENTAL PROTECTION AGENCY
OFFICE OF AIR PROGRAMS
DIVISION OF CONTROL SYSTEMS
CONTRACT NO. CPA 7O-68
OCTOBER 15, 1971
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RESEARCH AND ENGINEERING DEVELOPMENT
EVALUATION OF S02-CONTROL PROCESSES
TASK #5 FINAL REPORT
Submitted to
ENVIRONMENTAL PROTECTION AGENCY
OFFICE OF AIR PROGRAMS
DIVISION OF CONTROL SYSTEMS
Contract No. CPA 70-68
APPROVED:
Project Direct
Manager
Chemical Engineering Development
(J Director
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K 4 ED 20 LAB. 3-68
THE M. W. KELLOGG COMPANY Page No
A DIVISION OF PULLMAN INCORPORATED r=^^-|
IKEUOBBI
Research I Engineering Development Report Ho. -SED-71-1257
EVALUATION OF S02-CONTROL PROCESSES
TASK #5 FINAL REPORT
EPA-OAP-DCS CONTRACT NO. CPA 70-68
October 15. 1971
Staff: See attached sheet
Period Covered: February to June 1971
L. 0. No. 4092-50
Distribution ; Copy No.
Office of Air Programs 1-25
L.C. Axelrod 26
M.C. Cambon 27
C.P. Chatfield 28
C.W. Crady 29
G.M. Drissel 30
J.B. Dwyer 31
S.E. Handman 32
A.N. Holmberg 33
R.H. Multhaup 34
J.J. O'Donnell 35
C.E. Scholer 36
W.c. Schreiner 37
F.H. Shipman 38
A.G. Sliger 39
M.J. Wall 40
T.H. Wasp 41
R.I.D. (4) 42-45
AUTHORS:
V.
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STAFF
W. Beck
M. Cambon
C. Chatfield
R. Christopher
R. Chute
R. Croke
W. Cronkright
G. Drissel
G. Frisone
P. Giambalvo
N. Gilbert
S. Handman
T. Huang
J. Jung
J. Kunz
W. Leddy
P. Lefrancois
H. Leftin
B. Mandelik
D. Masi
J. O'Donnell
J. Osborne
R. Paredes
E. Roberson
C. Royce
J. Salvati
L. Schneider
R. Schreiber
W. Schreiner
L. Scotti
A. Sliger
R. Smith
H. Terzian
C. Van Dijk
J. Van Hook
D. Vichi
T. Wasp
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EVALUATION OF S02-CONTROL PROCESSES
TASK #5 FINAL REPORT
CONTRACT NO, CPA 70-68
Submitted to
ENVIRONMENTAL PROTECTION AGENCY
OFFICE OF AIR PROGRAMS
DIVISION OF CONTROL SYSTEMS
by
THE M. W. KELLOGG COMPANY
RESEARCH & ENGINEERING DEVELOPMENT
PISCATAWAY, N.J.
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TABLE OF CONTENTS
I. PREFACE
II. INTRODUCTION
III. SUMMARY
IV. BASIS OF EVALUATION
A. Process
B. Estimating
V. RESULTS
A. Investment
B. Operating Costs
VI. IMPACT
VII. ADVANCED POWER CYCLES
VIII. REFERENCES
IX. APPENDIX
Page No.
1
8
11
11
16
20
20
30
48
52
57
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I. PREFACE
The major objective of the work described in this
report was to provide APCO with a consistent basis for com-
paring the feasibility of different SOjj- control processes,
both technically and economically. Neither time, funds,
nor input data were available to permit a detailed eval-
uation of each process to be made and the results of this
study should not, therefore, be used without considering
their limitations. The major use for these results is ex-
pected to be in process screening and program planning in
which case the relative rather than absolute merits of the
processes will be of prime importance. Since the accuracy
of the estimates is higher on a comparative basis than it
is on an absolute scale, it is felt that the results do,
in fact, provide a sound basis for program planning.
It should not be construed from the above discussion
that the individual estimates for the various processes
are so inaccurate as to be meaningless on an absolute basis.
The estimating accuracy itself, based on the flow sheets as
shown, is judged to be within about 30%. It is the lack of
detailed process design for each case that creates the most
uncertainty and this uncertainty will vary among the dif-
ferent processes evaluated according to the status of the
input data for each. Based on past experience, it has been
found, without exception, that for a first-of-a-kind plant,
the estimated investment increases as more data become
available up to the point where a definitive detailed
design can be made. Thus, in the present study, the estim-
ated investments of those processes for which the most data
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to the actual cost than those based on very little data
(e.g., potassium formate). In all cases, however, the in-
vestments presented in this report will be lower than the
actual cost of the first plant built for reasons given be-
low, excluding any increases owing to inflation.
The investments included in this report are
essentially for battery limits plants only
and therefore do not represent complete
plant costs. However, for those processes
producing byproduct acid, storage tanks
have been included.
The design basis assumed that technology was
well established in each case and safety
factors (e.g., gas bypasses) were not needed.
New power plants only were considered, no
retrofits.
Waste disposal was treated as an operating
cost (e.g., SOC/ton) whereas in some cases
it could increase the investment signifi-
cantly.
Items such as interest during construction,
start-up costs, and working capital were
not included.
No contingencies were included.
Detailed designs and cost estimates were not
prepared for a specific installation; rather
a general case was used.
The process design used in each case does not
necessarily reflect the latest data since
these were not always available at the time
of the study. In many cases, the data used
did not come directly from the developer of
the process but from other sources such as
published papers.
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Since many of the factors that influence complete
plant investment were not used in developing the present
costs, in no case should any of the investments reported
herein be compared with a definitive estimate prepared
from a detailed design for a specific installation.
As previously stated, the main objective of this study
was to provide a sound basis for program planning and
thereby to allow more selective funding of processes. A
second phase tentatively was planned whereby the most
promising candidates from this first study would be eval-
uated in greater depth to obtain investments for complete
plants, rather than battery limits only, and thus provide
costs that would more closely approach the actual costs.
These latter costs can, of course, be obtained only by
preparing a detailed design and definitive estimate for a
specific installation for each process considered. To
follow this latter procedure for a number of different
processes would require a very large expenditure of man-
hours and money and also would require the availability
of input data needed for a detailed process design.
Within the limitations imposed by the ground rules
of the study, the results reported herein do provide a
reasonably accurate comparison of the candidate SC>2-
control process. In addition, the values reported may
be used as indicative of the actual investments required
with the qualification noted above that they are low
since they represent battery limit designs rather than
complete plants.
Operating costs for each process are reported based
on a format derived during the course of the study. Since
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probably these also are lower than will be obtained in an
actual case, provided the other assumptions (chemical cost,
capital charge rate, depreciation, etc.) are correct. In
any event, it is extremely doubtful that the reported costs
will be lower than shown and more likely, will be higher.
More exact operating costs can be estimated after defini-
tive designs are prepared.
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II. INTRODUCTION
The work reported on herein was performed pursuant to
Contract No. CPA 70-68 with the Air Pollution Control Office
(APCO)*, Environmental Protection Agency (EPA) . The task speci-
fications were delineated by APCO in Task No. 5 Contract Task
Specification Request, dated January 27, 1971, Task Change No. 1,
February 4, 1971, and Task Change No. 2, March 12, 1971.
Task Order No. 5, in short, specified that estimates of
plant investment were to be made for several different S02~
control processes. The processes to be evaluated were to be
specified by APCO and all input process data for preparing the
estimates likewise were to be furnished by APCO. In addition
to the process evaluations, other items to be studied were
applicability of these processes to power generating stations,
a time schedule relating to the construction of both power
plants and S0_- control processes, and the impact SO-- control
technology would have on emission reduction. APCO also re-
quested Kellogg to comment on advanced power cycle technology
based on a United Aircraft report.
The primary objective of this work was to develop a
sound basis for APCO to use in determining which of the many
proposed processes they should support. Specifically, APCO
had to prepare a five-year plan for SO-- control process
development and to insure that priorities were properly
assigned, they needed technical and economic evaluations
of the leading candidates on a consistent basis. To achieve
the desired objective within the time and funds allocated,
certain "ground rules" were established that were to be
followed throughout the study. These are tabulated below.
* Since changed to Office of Air Programs (OAP)
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A single size (1,000 MW) new power plant would
be assumed for all processes except Dry Limestone
Injection which was set at 350 MW.
All input data were to be supplied by APCO (see
Section IV A. 1. for a listing of the input data
actually received and used).
No process optimizations were to be made by
M. W. Kellogg; instead the process design in
each case would be used as received from APCO.
Only those changes would be made that were nece-
ssary to 1) convert the process to the standard
design bases; or 2) make the process feasible
if, based on either judgment or data, it appeared
the base design was inoperable.
Approximate plant investments were to be estimated
for each case. Since neither time nor input data
permitted a detailed design and estimate to be
made, it was recognized that the accuracy of the
individual estimates would suffer but the relative
comparative accuracy should be much higher.
Operating costs would be calculated for each case
with the various parameters (e.g., capital charge
rate, utility charges, etc.) to be specified during
the course of the task.
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The final results to include major equipment
costs and total investment with the operating
costs shown in detail. Additional discussion
of the estimating procedure is included in
Section II. A. 2. of this report.
The entire task, including the Applicability
Survey, initially was scheduled for completion
in three months (by April 26, 1971) but was
later extended to four months (to May 31, 1971
completion date).
In many cases the input data received were insufficient
for a process design and for these, additional data were obtained
from the general literature whenever possible. If literature
data were not available then assumptions were made as required
to complete the design. It was recognized that while this may
not be the most desirable"procedure, within the time and funds
allocated for the task, it was the only method that could be used.
It is tentatively planned to make a more detailed evaluation for
those processes which appeared to have the best technical and
economic potential based on the results of the present study.
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III. SUMMARY
Preliminary process designs and economic evaluations
have been prepared for twelve different S02- control pro-
cesses. The process economics are summarized in Table 1.
Investments are reported in millions of dollars for a
1,000 MW plant (except for dry limestone) so the investment
cost in dollars per kilowatt has the same value; i.e., a
20 million dollar investment for a 1,000 MW plant is equiv-
alent to 20 dollars per kilowatt. The actual investments
range from a low of $9.7/KW for dry limestone injection to
a high of $43.4/KW for the Cat-Ox process.
The operating costs are reported as dollars per ton of
sulfur not emitted, as mills per kilowatt-hour, and as a
percent increase of power cost. All of these costs are
shown as gross (excluding byproduct credit) and net (includ-
ing byproduct credit). In the "throwaway" processes there
is, of course, no difference in the gross and net since no
salable byproduct is produced. Further, the base case
ammonia scrubbing process is treated differently since the
ammonium sulfate formed in the scrubber ultimately is used
to produce fertilizer. This latter process is divided into
two sections, viz, gas scrubbing and fertilizer production.
The areas of primary interest for this study are the gas
scrubbing costs and these are shown as gross costs in
Table 1.
Since the primary purpose of the power plant is to
generate power (presumably at the lowest cost), it would
seem that the major concern with the SO2- control processes
would be the resulting increase in power cost. The cost of
the sulfur not emitted provides a convenient means of com-
paring processes but if the design conditions are not
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TABLE 1
SO2 REMOVAL PROCESSES - ECONOMICS SUMMARY SHEET
.f
Plant Size (MW)
Percent of Sulfur Removed From Flue Gas
Fixed Capital Investment (SMM) 10.8
Salable By-Product
By-Product Production (Tons/Day)
By-Product Credit (SMM/Yr)
Total Gross Production Cost ($MM/Yr) 5.14
S/Ton Sulfur Not Emitted 65.2
Mills/KWH 0.73
% Increase of Power Cost
Total Net Production Cost ($MM/Yr)
$/Ton Sulfur Not Emitted
Mills/KWH
% Increase of Power Cost
/ / /
<* *y ^ . &" fc> ^ r? x**" .^ & *y A v i? . *>
1000
95
34.6
Sulfur
261
1.52
12.32
148.4
1.76
26.1
10.80
130.1
1.54
22.8
1 *,
1000
90
34.9
Sulfur
245
1.43
13.32
169.3
1.90
28.1
11.89
150.7
1.70
25.2
^Cj v_
1000
90
10.2
*
9.98
126.8
1.43
21.2
_
-
-
-
«?*} ^
1000
90
16.2
(98.5%
(H2S04
570
2.00
7.86
99.9
1.12
16.6
5.86
74.3
0.84
12.4
^«5 V
1000
90
21.6
(98.5%
(H2S04
692
2.42
8.38
106.6
1.20
17.8
5.96
75.5
0.85
12.6
1000
87.5
29.8
(80%H2SO4
(60%HN03
III iia§?4
oiiii2-01
10.51
137.7
1.50
22.2
8.03
104.7
1.15
17.0
•** ° <&
1000
95
26.8
(98.5%
(H2S04
890
3.12
9.04
111.1
1.29
19.1
5.93
71.3
0.85
12.6
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identical among processes the cost of not emitting sulfur
may not indicate accurately the increase in power costs.
For example, if the cost of not emitting sulfur is plotted
against the percent increase in power cost, a straight
line can be drawn through all points except dry limestone.
This latter process is for a smaller plant and removes
only 50% of the sulfur from the flue gas whereas the
others remove 90%. Thus, caution should be exercised when
comparing different processes on a cost of sulfur not
emitted basis.
Of the twelve processes evaluated, none can be ranked
as commercially available at the present time. Wet lime-
stone has been tested in a 100 MW unit, dry limestone is
presently being evaluated in a 150 MW plant, Cat-Ox has
undergone pilot plant testing (=tl5 MW) , and it and several
others are presently being installed in power plants of
about 100 MW capacity or higher. However, the results
currently available to Kellogg for the present study are
insufficient to allow an assessment as to whether adequate
data exist to permit a full size commercial plant to be
designed and built with a high level of confidence. Some
processes have been developed more than others while some
that are more advanced could not be fully evaluated owing
to data unavailability.
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IV. BASIS OF EVALUATION
The results of the Applicability Survey have been
reported separately and therefore will not be repeated
here.
A. Process
1. Input Data
The twelve SO2- control processes for which APCO
requested evaluations are listed below along with the input
data received for each.
1) Wet Limestone Scrubbing - TVA Conceptual
Design and Cost Study Series Report,
Study No. 2, 1969.
2) Dry Limestone Injection - TVA Conceptual
Design and Cost Study Series Report,
Study No. 1, 1968.
3) Cat-Ox (Monsanto) - an SRI final report
covering their evaluation of the process
under Contract CPA 22-69-68.
4) Molten Salt (Atomics International) -
seven AI reports containing exploratory
data, the latest dated 1969. A
Singmaster & Breyer report showing the
design and estimate of a unit sized for
an 800 MW power plant.
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5) Formate Process (Consolidation Coal) -
Paper presented at the 1970 ACS
meeting in Houston.
6) Ammonia Scrubbing (Base Case - fertilizer
production) - TVA Conceptual Design and
Cost Study Series Report, Study No. 3, 1970.
7) Ammonia Scrubbing (Alternate A - Steam
Stripping) - same as 6. Plus some addi-
tional data from TVA's pilot plant.
8) Ammonia Scrubbing (Alternate B - Thermal
Decomposition) - same as 7.
9) Modified Chamber Process (TYCO) - TVA report
covering Preliminary Conceptual Design and
Cost Study; A.M. Kinney, Inc. report on heat
and material balance analysis; TYCO final
report, Contract CPA 70-59, April-October 1970,
10) Magnesium Oxide Scrubbing - Experimental
data from Babcock and Wilcox, Report 5153,
September 1970, and a block type flow sheet
of a conceptual design.
11) Zinc Oxide Scrubbing - Final report by the
Envirogenics Company (Div. of Aerojet-General
Corp.) for Contract PH 86-68-77, October 1970.
12) Citrate Process (USBM) - Status report from
the Bureau of Mines, January 1971.
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2. Design Bases
Since a major objective of the present study was to
compare all the different processes on a uniform basis, it was
imperative that the same design bases be used for the parameters
common to all processes. Accordingly, the bases shown in Table 2
were specified and used in this study. Design parameters peculiar
to each process are given in the individual reports included as
Section IX, the Appendix of this report.
A plant site in Cincinnati, Ohio was selected for the
evaluations since this general area contains a number of large
power plants and extensive coal deposits are nearby. (A map
showing the location of the major thermal power plants in the
United States is included in the Applicability Survey report.)
A coal sulfur content of 3.5% was desired as the power plant fuel
and a coal meeting this requirement was shown for Harrison County,
Ohio, in Table 1, Page 13, of USBM RI 7346, "Analyses of Tipple
and Delivered Samples of Coals," 1970. Since only the proximate
analysis is given, the composition was adjusted slightly to match
the reported heating value. Flue gas quantity was calculated using
20% excess air which was assumed to be at 80°F with a relative
humidity of 60%. A heat rate of 9,000 Btu/kwh was chosen as a
reasonable value for a new large power plant.
To simplify the process calculations, it was assumed
that all of the coal sulfur would appear in the flue gas (i.e.,
no sulfur left in the ash) with 98% forming SO- and the remaining
2% present as SO.,. A nominal S02~ removal of 90% was specified
for all processes but the actual removal varied slightly owing
to the inherent differences among the processes.
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Table 2
GENERAL PROCESS DESIGN BASES
Coal Composition
Component
Ash
S
«2
N
O
Water in Coal
Excess Air
Carbon Burned in Coal
Heating Value in Coal
Wet
Dry
Sulfur to SO-
Sulfur to SO^
SO- Removal
Flue Gas Temperature Out of Air
Preheater
Minimum Gas Temperature
to Stack After Treatment
Preferable Gas Temperature
to Stack After Treatment
Gas Velocities in Ducts
Plant Size
Heat Rate
Flue Gas Quantity
Flue Gas Molecular Weight
Flue Gas Weight
Flue Gas Composition
Component
SO
SO
CO
N2
°
Fly Ash
Dry Wt. %
15.2
3.5
5.0
67.2
1.6
7.5
4.8 Lbs/100 Lbs Wet
20%
100%
11,980 Btu/Lb
12,580 Btu/Lb
98%
2%
90%
300*F
200°F
250°F or Higher
60 Ft/Sec
1,000 MW
9,000 Btu/Kwhr
293,000 Moles/Hr
29.54
8,655,000 Lbs/Hr
Mole Fraction
0.00261
0.00005
0.13657
0.74087
0.03276
0.08714
77,900 Lbs/Hr
Water in Air (60% RH @ 80°F) 0.0212 Mols/Mol Dry Air
Coal Consumption Rate (Wet) 751,252 Lbs/Hr
Gas ducts 300°F or higher to be insulated.
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A preferable stack gas exit temperature of 250°F
was selected based on a TVA study, "Use of Limestone in Wet-
Scrubbing Process," Conceptual Design and Cost Study Series,
Study No. 2, 1969. While opinions vary widely among power
producers as to the optimum exit temperature, the 250°F value
seemed reasonable at the time this study was made. Also, since
all processes are designed using the same basis, the relative
comparisons should not change appreciably at other exit tempera-
tures. The effect of reheat temperature on process costs should
be studied in more detail when the process evaluations are ex-
tended.
The gas velocity in the ducts was specified as a
nominal 60 feet per second based on verbal information from
various power plant designers. The recommended velocities
ranged from about 40 to 120 feet/second with a nominal value
of 4,000 feet/minute frequently quoted. Thus the design figure
of 60 feet/second actually used falls within the normal range
and is a reasonable value for the present study. A more com-
plete study might require optimization of gas velocity —
balancing lower investment against higher pressure drop, hence
operating cost, as the velocity increases -- but this was beyond
the scope of the present evaluation.
The percentage of the total ash in the coal that
appears as fly ash generally falls within the range of 60-85.
For this study it was assumed that the quantity of fly ash
produced would fall about midrange.
The values shown in Table 2 formed the basis for
the process design in each case except as noted in the individual
discussions. Material balances which were received as part of
the input data were converted to these bases or, if none were
received, process balances were made using Table 2 values.
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B. Estimating
The present estimates are classified as approximate
owing to the limited data available and the special nature
of the equipment involved but are felt to represent the best
quality that could be prepared within the calendar time and
man-hours available. It should be noted that the estimates
contain many large cost items, the prices of which in some
cases have been established by an estimator's judgment only
since sufficient data and time were not available to develop
any better pricing. Further, the purchased cost of a large
item in many cases was obtained by contacting a single vendor
for a budget type quote and it is doubtful if the optimum cost
and/or design was always obtained. Also, since some process
items were ill-defined (data lacking or proprietary and un-
available) , they had to be treated as "black boxes" and the
resulting costs could be quite inaccurate.
Owing to the reasons cited above, it is difficult to
properly judge the accuracy of the present estimates. Never-
theless, some indication of accuracy is required and, therefore,
the following appraisal is offered. It is felt that the accuracy
of the present estimates is no better than 30% and could be
poorer. However, since all of the estimates were prepared
essentially simultaneously in a single department and some of
the equipment is similar, the relative comparative accuracy
probably is better than the accuracy of any one plant. It
should be noted that, based on past experience, costs invariably
increase as more data become available. It is reasonable to
assume, therefore, that the present costs, which are based on
processes not yet completely developed, would increase rather
than decrease as more - detailed evaluations are made. Accordingly,
v
the cited accuracy of 30% should not be considered as a plus or
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minus value since more likely it would be in the -5 to +25% range.
The degree of the expected cost increases would, of course, vary
from process to process generally as a function of the status
of the input data.
The estimates as reported in the next section of this
report do not include contingency nor undeveloped design allow-
ance (UDA). Neither do they represent complete plant investments
since a number of items are not included such as inventory, interest
during construction, working capital, cost of land, nor offsites
such as laboratories, site preparation, access roads, and incre-
mental costs of maintenance shops and warehouses. These items
have not been included since they are relatively small compared
to the total investment and would have no significant effect on
the comparisons. Therefore, it was decided that the time required
for including the items necessary to obtain a complete plant in-
vestment was not justified in view of the very limited overall
schedule. These items can be added when more detailed estimates
are made.
The general bases used in preparing the estimates are as
follows. The plant site selected was Cincinnati, Ohio, as pre-
viously mentioned, and labor productivity and rates for this area
were obtained. Thus the-total investments reported herein would
be somewhat different for other locations.
The general rationale followed was to charge the SO?-
control process for all costs that resulted, directly or indirectly,
from the installation of the process. Thus increased costs of
ductwork, flue gas fans, and dust removal equipment were considered
to be part of the control process. To insure that proper charges
were made in all cases, a "standard" power plant design was used to
determine the investment for the items cited above and allowances
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for these items were then credited to the control process.
(Basis for the standard power plant was TVA's Bull Run Steam
Plant which has a unit nameplate rating of 950 MW.) For
example, the weight of the standard plant ductwork was deter-
mined (by designing the ducts and making a layout) ar>d this
was subtracted from the total ductwork weight required when
an S02~ control process was added.
The incremental fan horsepower required by the control
process was calculated from the total gas flow and increased
pressure drop which is caused by the additional ducts and
equipment.
The incremental cost of the dust removal equipment,
assumed to be an electrostatic precipitator (ESP), required
that a "standard" be specified for the power plant. Since the
requirements were not well-defined, two different base case
ESP's were estimated: one at 95% overall dust removal effi-
ciency and one at 99%. The effect of each of these is shown
in the process investments. If the process itself removes the
fly ash then a credit is taken for eliminating the ESP. On the
other hand, if the process requires an unusually high dust re-
moval efficiency (e.g. 99.5% minimum), then the increased cost
of the more efficient unit is added to the process investment.
These conditions apply only to a new plant since no credit could
be taken for an existing ESP nor would the incremental costs of
increasing efficiency be the same in an existing plant.
Equipment needed for by-product recovery, i.e., acid
plants or Glaus plants, are included as part of the process in-
vestment. When acid is produced, a nominal 30 days storage has
been provided to insure continuity of operation even when faced
with short-term delays such as temporary loss of tank cars or
trucks, rail strikes lasting a few days, etc.
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One important point in the design rationale that
should be noted is that no by-passes have been provided
around the process units. If these were required, the cost
would increase in all cases. Provisions have been made,
however, for shutting off one complete train of process
equipment which provides some flexibility since multiple
trains are required in all cases.
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V. RESULTS
A« Investment
The estimated investment for each of the twelve processes
evaluated is tabulated below. Note that the dry limestone pro-
cess is for a 350 MW plant but all the others are based on a
1,000 MW plant. A detailed breakdown of the investment costs is
shown in Table 3. For convenience the major assumptions and
major changes to process designs received as input data are sum-
marized in Table 4. For complete bases see Table 2 and the in-
dividual discussions for each process given in the Appendix of
this report.
PROCESS INVESTMENT, $ $/KW
1. Wet Limestone Scrubbing 10,800,000 10.8
2. Dry Limestone Injection 3,400,000 9.7
3. Cat-Ox 43,400,000 43.4
4. Molten Salt 34,600,000 34.6
5. Formate 34,900,000 34.9
6. Ammonia Scrubbing (Fertilizer) 38,200,000 38.2
7. Ammonia Scrubbing (Steam Stripping) 16,200,000 16.2
8. Ammonia Scrubbing (Thermal Decomposition) 21,600,000 21.6
9. Tyco 29,800,000 29.8
10. Magnesium Oxide Scrubbing 26,800,000 26.8
11. Zinc Oxide Scrubbing 21,800,000 21.8
12. Citrate 21,200,000 21.2
As previously mentioned, all of the wet scrubbing processes
have a gas reheat system designed to produce a stack gas tempera-*
ture of 250°F. The system selected was based on a TVA study and
for their conditions gave the lowest costs. However, it appears
that the carbon steel material of construction specified for the
exchanger downstream of the air preheater (i.e., upstream of the
gas scrubber) is inadequate since the gas will be cooled below
-------
TABLE 3
SO2-CONTROL PROCESSES J?
COST BREAKDOWN >
Major Equipment Cost
Direct Materials
Sub-Contracts
Ductwork, Duct Valves, and
Expansion Joints
Piping, Electrical & Instruments
Foundations, Concrete & Steel
Structures, Buildings, Paint,
and Insulation
Direct Materials
Sub-Contracts
Subtotal - Materials & Sub-Contracts
Other Costs
Construction Forces,
Home Office Engineering,
Procurement, Insurance,
Start-up and Contractor's Fee
Investment Cost for Reforming Unit,
Oxygen Plant & K2SO3 Recovery
Subtotal Investment Cost
for Process Plant
Other Process Units Investment Costs
Claus Sulfur Plant
Sulfuric Acid Plant Including
Sulfuric Acid Storage
Fertilizer Plant Including
Storage and Nitric Acid Plant
Electrostatic Precipitator
Deduct Standard rower Plant
Induced Draft Fans and Ducts
Deduct Standard Power Plant
Electrostatic Precipitator (300°F)
NET COST INVESTMENT
*y -O *?
(2,300) (700)
(2,100) (900)
10,800 3,400
(Costs are in M$)
1,360
4,231(a)
1,940
440
340
21,836
17,064
38,900
l,400
(2,300)
(2,100)
34,600
«P«?
11,329
120
1,106
3,145
475
552
16,727
13,173
6,700
36,600
2,700
(2,300)
(2,100)
34,900
//£
/# ""
5,436
89
936
1,418
247
160
8,286
6,314
14,600
28,000
(2,300)
(2,100)
38,200
x
•/<£•&
5,997
282
936
1,689
258
265
9,427
7,473
16,900
3,700
(2,300)
(2,100)
16,200
0*"^
/$£
8,093
85
936
2,066
342
440
11,962
9,638
21,600
4,400
(2,300)
(2,100)
21,600
^$
&
9,754
6,949(d)
811
2,360
410
290
20,574
10,226
30,800
l,300
(2,300)
0
29,800
/;*/
«••»/'
/
-------
TABLE 4
Process
Wet Limestone Scrubbing
Major Assumptions
i
K>
10
Dry Limestone Injection
Basis: TVA 1969 Conceptual Design
•Limestone £ 110% stoichiometry
•250°F gas reheat required
•3-stage UOP TCA scrubber vessel
•100% calcination of limestone and
25% conversion of CaO in boiler
•90% S(>2 removal
•99+% solids removal
•12% solids slurry to disposal pond
•No deleterious effect on boiler
operability from limestone injection
•No plugging problems in slurry circuit
Basis: TVA 1968 Conceptual Design
•350 MW maximum size boiler than can
be used
•200% stoichiometric limestone
addition
• 50% SC>2 removal (25% conversion of
CaO)
•No deleterious effects on boiler
operability
•Stack emission particulate level of
0.1 gr/SCF
Major Changes to Process
•Scaled from 200 to 1000 MW, heat,
material balances, AP changed as
needed
•Higher AP in reheat system
•Increased liquid circulation rate
in scrubbers and added 1 extra
stage of contact
•Recycle waste disposal pond over-
flow liquid back to process
•Heat, material balances, AP
changed to reflect MWK bases
•Increased grinding from 70%-200
mesh to 99%-325 mesh
•15.2% ash in coal vs TVA's 12.0
•Increased size of electrostatic
precipitator from TVA's 2.8 to
3.1-fold over standard power plant
also increased efficiency from
-------
TABLE 4 (Cont'd)
Process
Major Assumptions
Cat-Ox
i
(O
<*> Molten Carbonate
Basis: SRI Report, Contract CPA22-69-78
•Converter temperature - 900°F
•S02 conversion - 90%
•Flue gas from absorber - 225°F
•Product acid strength - 78%
•Maximum available size Teflon T.M.
exchanger used in acid cooling
system is 8"0 x 13'
•Catalyst removal, cleaning and re-
charging facilities are "black box"
items
•Electrostatic precipitators available
for 900°F service @ 99.5% efficiency
•Product acid storage capacity for
30 days production
Basis: Singmaster & Breyer Design,
Atomics International reports contain-
ing exploratory data
•Electrostatic precipitator available
with efficiency of 99.5% 0 850°F
•0.03% Cl in coal
•95% SO2 removal
•95% reduction of M2SO4 by carbon
•Reducer effluent melt can be filtered
to remove solids
•88% recovery of lithium salts contained
in fly ash and coke filter cakes
•No recovery of sodium and potassium
salts from filter cakes
•92% recovery of sulfur in Glaus plant
Major Changes to Process
•Heat, material balances, AP
changed as needed to reflect
MWK bases
• Equipment sized to reflect MV7K
bases
•Material and heat balances scaled
to reflect MWK bases
•Gas velocity in absorber reduced
from 25 to 10 ft/sec. (8 spray
towers instead of 4)
-------
TABLE 4 (Cont'd)
Process
Molten Carbonate (Cont'd)
Potassium Formate
i
N>
Major Assumptions
•Bra^-rr^^^ES?* -.-w .~'^*>s*i.-=sj-im*BeJ2:iatr'.-3rr~rz>-*,.sfa
•Assume reducers will work as shown§
two compartment vessels, oxidation and
reduction zones separated by bridge wall
with melt circulating between zones,
3 ft/sec gas velocity, 20 min. liquid holdup
•Claus plant off-gas recycled to absorbers
•Make-up carbonates are anhydrous
•No plugging of spray nozzles by fly ash
•Demisters will perform satisfactorily
•Suitable materials of construction are
available
•Carryover of salts into flue gas stream can
be reduced to practical level
Basis: ACS paper presented at 1970 meeting
in Houston
•80% fly ash removed in cyclones
•0 back pressure of SC>2 in formate solution
•90% SO2 removal
•Absorber is 3-stage spray tower
•Absorbers are in heat balance - no enthalpy
data available
•Fly ash escaping cyclones removed in
scrubbers and separated for disposal in
centrifuges
•Stripper in heat balance - no enthalpy
data available
•97.5% recovery of sulfur in Claus
•02 used in Claus instead of air
•Natural gas- C(>2 reforming used to generate
CO
•"Black box" processing step used to recover
potassium CO2 scrubber liquids
•Residence time in regen-
erator lowered from 2 hrs
to 25 min.
•Total pressure in regen-
erator increased from 700
psi to 2000 psi to maintain
same CO partial pressure r
in ACS paper
•Detailed flow sheet not
given in input data - MWK
design developed from the
data given. Supplemented
as required by judgment
factors and assumptions
see discussion in Append;^-
-------
TABLE 4 (Cont'd)
Process
Ammonia Scrubbing
(Base Case - Fertilizer
Production)
Major Assumptions
Ammonia Scrubbing
(Case A - Steam
Stripping Regeneration)
i
to
en
Ammonia Scrubbing
(Case B - Thermal Decom-
position Regeneration)
Basis: TVA 1970 Conceptual Design
•90% S02 removal
•4-stage sieve tray (impinjet) scrubbers
•Oxidation of sulfite and bisulfite takes
place in separate vessel
•No fly ash removal upstream of scrubbers,
99.5% in scrubbers
•Scrubber effluent is 50 wt. % salt solution
•Ammonium sulfate solution produced and fed
to fertilizer plant
• Standard fertilizer plant, CaSC>4 waste
disposed of as 20% slurry
Basis: Same as Base Case plus additional data
from TVA's pilot plant; also literature data
•Packed column under vacuum used to regen-
erate scrubber effluent with steam-heated
thermosiphon reboiler
•68.8% recovery of sulfur removed from
flue gas
• 90% SC>2 removal from flue gas
•15% of sulfur absorbed makes sulfate
•5/1 sulfate/sulfite-bisulfite ratio used
•Weak SC>2-water solution condensed from
stripper overhead vapors air-stripped to
improve SO2 recovery
Basis: Same as Base Case and Alt. A
•90% S02 removed from flue gas
•83.7% sulfur removed from flue gas is
recovered
•Effluent S/NH3 mole ratio, S/C, = 0.85
•80% conversion (NH4>2S04 in elec. furnace
•1025 x stoichiometric CaCO3 used in sulfate
reactor
•Utilities consumption - 0»2f Steam/* (NH4)2S04
in feed and 350 kwh/fcon(NH4) 2^04 in feed
Major Changes to Process
•Scaled heat, material
balances, AP, to reflect
MWK design conditions
(size, coal, etc)
•Flow Sheet heat, materia
balances, AP, equipment
selection/design largely
MWK since no detailed
process design available
-------
TABLE 4 (Cont'd)
Process
TYCO
Major Assumptions
to
MgO
Basis: TVA Conceptual Design & Cost Study
(Preliminary), A.M. Kinney, Inc. Report,
Tyco Final Report
•80% NOX, 90% SOX removed
•Filtering system for removing fly ash from
H2S04 stream is "black box" item
•Plug flow conditions in SO2-oxidizer-reactor,
with 15 sec. residence time (direct scale-up
from bench scale)
•Same height of packing in N203 scrubber as
indicated by extrapolating miniplant data
•99% stripping efficiency @ 272°F with 99%
of NOx as NO2, no N20 formed
•Overhead vapor from stripper contains 10% NO2
.Flooding characteristics for granular carbon
same as for 1/4-inch rings
•60% HN03 and 80% H2SO4 produced as byproducts
•0.005 wt. % ash can be filtered from 272OF
80% H2S04 on commercial scale
Basis: Experimental data from Babcock and
Wilcox, Report 5153, 1970, and block type
flow sheet of a conceptual design
•Venturi scrubbers remove 98% of fly ash in
flue gas
•3-stage TCA scrubbers, 95% SO2 removal,
0.762 wt. % bisulfite concentration in
scrubber outlet slurry
•Slurry thickened to 30% solids & centrifuged
•Rotary dryers used to dry solids and expel
water of hydration
•Rotary kiln calciners used to regenerate MgO
by reduction with carbon
•Acid plant included as part of process
•30 days' storage of byproduct acid, 98% H2SO4
•Mg(OH>2 losses reduced by converting to
soluble sulfate and recovering
Major Changes to Process
•Mill I iMM i I i tr,- J^ffl«&*rt^=;•«a*yjt^-.E:.'--VT^i-i. ijV¥.
-------
TABLE 4 (Cont'd)
Process
ZnO
Major Assumptions
i
K)
-J
t
Citrate
Basis: Final Report by The Envirogenics Co.,
(Div, of Aerojet General Corp.) for Contract
PH 86-68-77, Oct. 1970
•Scaled rates and material balance given in
report to MWK size and coal; in some cases,
equipment scaled
•Absorbers in heat balance, 90% S(>2 removal,
4-stage Impinjet,scrubbers used
•Sulfite to sulfate ratio = 3/1
•Clarifier scaled from report - no settling
data given
•Process equipment for formation, crystalli-
zation and thickening of ZnSO3 scaled from
Envirogenics report
•Superdecanting centrifuges are capable of
separating ZnS03 crystals from lean solu-
tion of sulfite-bisulfite
•All calcined particles are larger than
40 microns
•Indirect kilns used to calcine hydrated zinc
sulfite
•30 days' storage of 98.5% H2SO4 byproduct
acid provided
•Filters on thickened waste Blurry (fly ash
+ CaSC>4) sized based on gypsum filtering rates
Basis: Status reports and summaries of lab-
oratory and pilot plant testing by USBM, Jan,
1971; general literature for H2S production
•90% SO2 removal § 105°F
•Liquid sulfur-fly ash slurry centrifuged to
remove fly ash
•4.3 gm. S02/liter solution, equivalent to
about 80% approach to equilibrium
Major Changes to Process
•Add mechanical separators
for fly ash removal
•Add H2S04 plant
•Substitute grinding mill
for hammer mill on dryer
effluent solids
•Higher sodium to citric
acid ratio used, 583 vs 1
c
-------
Process
Citrate (Cont'd)
TABLE 4 (Cont'd)
Major Assumptions
'Citric acid concentration, 0.87 gm mol/
liter
•Na/citric acid ratio of 5.3 atoms/mol
•100% conversion of SO2 0 5% excess H2S
and 10-min. residence time
•Extrapolation of absorption data to low
concentration of SO2 (most available data
on smelter gas)
•Two-step system to produce I^S:
CH4 + S — > CS2 + H2S, followed by hydrolysis
of CS2 to H2S
Major Changes to Process
i
N>
CO
-------
its dew point, thereby forming H-SO. on the exchanger surface.
Although carbon steel has been used in the present estimates,
corrosion resistant materials (e.g. Alloy 20) are actually
required and this would significantly increase the cost. A
new optimization study is required to determine the best method
of reheat including several different levels of gas temperature.
It is felt, however, that the present allowance for the reheat,
system investment is adequate even though the method of reheating
is still uncertain. As a point of interest, the cost of Alloy 20
exchangers, which was received too late to incorporate in the
estimates, would add about $3-4 MM to the present investment.
Some "black box" items were required where process re-
quirements were not sufficiently defined to permit a normal
estimate to be made. Specific examples of this is the fly ash
filter in the Tyco process and the solution regeneration operation
in the Formate process. In the Tyco case, a brief discussion was
held with a vendor and a very approximate type investment was ob-
tained for the fly ash filter but additional data are needed to
firm up this number. For the Formate solution regeneration operation
a purely arbitrary allowance was included. Obviously, additional
process definition is required to eliminate the "black box" items
from all the flow sheets but the overall effect on process invest-
ment should be small.
In those processes which require the use of a very effi-
cient mist eliminator (Tyco/ Cat-Ox, Formate), a Brink type has
been specified. The vendor for these items was: contacted and a
price was obtained which supposedly included the containment
vessels. After the estimates had been completed, however, the
vendor sent a revision which indicated that the: original quote
did not include the containment vessels. If the cost of these
vessels does have to be added then the present estimates will
-------
increase by several million dollars. However, since the
cost presently included for the Brink mist eliminators is
several million dollars, it is likely that some other
equally efficient type might be found for this price and
the present costs would, therefore, be unchanged. Con-
sequently, the additional cost of the containment vessels
was not included in the present estimates. Since mist elimi-
nation in these processes is of major importance, extensive
experimental data in this area eventually will be required
to provide a sound basis for both process design and cost
estimating.
Waste disposal is a major problem in some processes,
notably those using limestone, both wet scrubbing and dry
injection. Since the method of disposal will vary from site
to site (e.g., ocean dumping, land fill, removal by outside
contractor), it was decided not to attempt to estimate the
cost of waste disposal as an investment but rather to include
it as an operating cost. Waste disposal would have to be con-
sidered in detail, however, for a specific installation, includ-
ing such factors as potential water pollution.
B. Operating Costs
Operating costs have been calculated in the manner re-
quested by APCO, viz, as mills per kilowatt hour and as dollars
per ton of sulfur not emitted, considering both credit and no
credit for by-products. A breakdown of these costs for each
-process studied is given in Tables 6-17. Since the individual
items shown in the tables are generally self-explanatory, no
detailed discussion of these will be given. However, a separate
listing, and in some cases a brief rationale, of the specific
values used for the various items is included below for convenience.
-------
1. Plant Site Location - Cincinnati, Ohio.
2. Fixed Cost - Annual % of capital investment.
a. 15-year plant life, sinking fund method with
cost of money at 8%; depreciation is 3.68%.
Interim replacements of 0.35%; total is 4.03%.
b. Taxes - Federal tax of 3.61%, local taxes of
2.33%, total is 5.94%; basis FPC Power Supply
Area 12.
c. Insurance - 0.25%.
The sum of a + b + c is 10.22% which combined with
the cost of money at 8% gives a total of 18.22% for
fixed charges. Other combinations of plant life and
cost of money would, of course, produce other fixed
charge rates (e.g., 30-year life and 6% money gives
13.8% while 10-year life and 9% money results in
22.12% charges) but the one selected appears reason-
able for the present evaluation.
3. Direct Costs
a. Labor Rate - $4.50/hour based on U. S. Bureau
of Labor statistics for stationary engineers,
boiler firemen, and average production worker
in the Cincinnati area. Supervision at 15% of
labor.
b. Maintenance - Will vary with the process. A
range of 4-5% of capital investment has been
specified for the present evaluations with the
actual value used based on judgment factors
applied to each process.
-------
c. Cost of Raw Materials, Plant Supplies -
Plant supplies will> be 15% of maintenance
costs. A listing of chemical prices is
shown in Table 5.
d. Waste Disposal (Excluding Fly Ash) -
5DC/Ton.
e. Utilities -*•_
Steam ( = 150 psia) 40C/M Lb
Process Water 20£/M Gal
Treated Process Water 40£/M Gal
(BFW Quality)
Cooling Water 2.5£/M Gal
Coal . . $8.00/Ton *
Fuel Gas 40£/MM Btu
Electricity ,,6.75 mills/Kwh
(This electricity cost was derived from coal- . •
fired steam-electric plant production expense
data available from the FPC and the National
.Coal Association for the plant site area.
Projected escalation for the 1971 power costs
is included.
The value shown is the total production cost
and assumes that the power used by the SO~-
control process will cost as much to generate
as does power sold to outside customers. An
alternative approach would be to consider the ,,
power requirements of the SO--control process
on an incremental basis such that the power
cost is equal only to the additional fuel burned.
This latter method would produce power costs about
one-half those actually used. Since there is not
a large difference in the power consumed among
the processes studied, it was not considered justi-
fied to spend the time needed to prepare a separate
study of the proper method of charging power to
SO-- control processes. Therefore, the total
production costs were used in the present evalua-
tions .) .
-------
TABLE 5
CHEMICAL AND RAW MATERIAL PRICES
Material
Price
Ammonia
Catalysts:
Activated Carbon
Nitric Acid Catalyst
Sulfuric Acid Catalyst
Citric Acid
Coal
Coke (Petroleum)
Lime
Limestone
Lithium Carbonate
Magnesia
Nitric Acid (60%)
Oxygen ("over-the-fence")
Phosphate Rock
Potassium Carbonate
Sodium Carbonate
Sodium Hydroxide
Zinc Oxide
$50/Ton (Anhyd,)
$.38/Lb.
$120/troy oz.
$1.09/Lb.
$.30/Lb.
$8/Ton
$8/Ton
$15/Ton
$2/Ton
$.42/Lb.
$40/Ton
$25/Ton
$12/Ton
$11.88/Ton
$.05/Lb.
$.02/Lb.
$.03/Lb. (dry)
$.15/Lb.
-------
4. Indirect Costs
a. Payroll Burden - 20% of operating labor and
supervision.
b. Plant Overhead - 50% of operating labor,
supervision, maintenance, and supplies,
5. By-Product Credit
Sulfuric Acid
98% acid @ $12/short ton of 100% acid,
80% acid @ $ 6/short ton of 100% acid.
Nitric Acid (57-60%) at $25/short ton of 100% acid.
As previously indicated, the power station thermal efficiency
used in the present study is about 38% (9,000 Btu/Kwh) and the load
factor is about 80% (7,000 hours/year). The same load factor (i.e.,
7,000 hours/year) was used for the full fifteen-year assumed life
of the SO-- control plant which differs somewhat from TVA's recom-
mended figures of 7,000 hours/year the first ten years and 5,000
hours/year for the next five. However, since there does not appear
to be universal agreement on what load factor should be used for
power plants and since the load factor will not have a major in-
fluence on the comparisons of the various S02~ control processes
studied, use of an 80% load factor appears reasonable.
-------
TABLE 6
ANNUAL OPERATING COSTS
WET LIMESTONE SCRUBBING
Plant Size (MW) 1000
Fixed Capital Investment (PCI) - $10,800,000
Stream Time (hrs/yr) 7000
DIRECT COST $/YEAR
1. Operating Labor (2 men/shift @ $4.50/hr) 78,840
2. Supervision - 1555 of Item 1 11,830
3. Maintenance Labor and Materials - 4$ of FCI 432,000
4. Plant Supplies - 15% of Item 3 64,800
5. Utilities
A. Cooling Water - — GPM at $.025/M gal
B. Process Water - — GPM at $.20/M gal
C. Electric Power - 25,310 KW at 6.75 mills/KWH 1,195,860
D. Natural Gas - — MSCFH at $.40/MSCF
E. Steam — M Ibs/hr at $.45/M Ibs '
o. Chemicals and Raw Materials
A. Limestone - 1,237 TPD g $2/TON 721,700
B. Extra Coal* - 3 TPH @ $8/TON ' 168,000
6a. Solids Disposal (ex fly ash) - 57 TPH § $.50/TON 199,500
7. Subtotal Direct Cost (ex credits) $ 2,872,530
8. Credits
None ' ——
9. Total Direct Cost $ 2,872,530
INDIRECT COST
10. Payroll Overhead - 20% of (1+2) 18,130
11. Plant Overhead - 50% of (1+2+3+4) 293,740
12. Total Indirect Cost $ 301 870
FIXED COST
13- Capital Charges - 18.22$ of PCI $ 1,967,760
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13) $ 5,142,160
UNIT PRODUCTION COST
15. Gross - Items (7+12+13) $ 5 1/12 160
A. Mills/KWH .73
B. $/Ton S not emitted 55 2
16. Net - Items (9+12+13) $ 5 1/12 160
A. Mills/KWH .73
B. $/Ton S not emitted . 65.2
* Extra coal needed for limestone calcination In boiler
-------
TABLE 7
ANNUAL OPERATING COSTS
DRY LIMESTONE INJECTION
Plant Size (MW) 350
Fixed Capital Investment (FCI) - $3,400,000
Stream Time (hrs/yr) 7000
$/YEAR
1. :\rer;iti.n[: Labor (1 man/shift @ $4.50/hr) 39,420
r/uporvision - 15£ of Item 1 5,910
3. Maintenance Labor and Materials - >\% of FCI 136,000
•i. Plant Supplies - 15* of Item 3 20,400
5. Utilities
A. Cooline Water - — GPM at $.025/M gal
R. Procosc Water - — GPM at $.20/M gal
C. KJectrJc Power - 4,420 KW at 6.75 mills/KWH 208,850
u. Natural Gas - — MSCFH at $.40/MSCF
E. Steam - — M Ibs/hr at $.45/M Ibs
6. Chemicals and Raw Materials
A. Limestone - 716 TPD @ $2/TON 417,700
U. Extra Coal* - 1.05 TPH 6 $8/TON 58,800
6a. Gnlids Disposal (ex fly ash) - 29 TPH @ $.50/TON 101,500
7. Subtotal Direct Cost (ex credits) $ 98« ,580
3. Credits
None '
9. Total Direct Cost $ 9«« ,5bO
INDIRECT COST
10. Payroll Overhead - 20$ of (1+2) 9,070
11. Plant Overhead - 50$ of (1+2+3+4) 100,870
12. Total Indirect Cost $ 109,940
FIXED COST
13. Capital Charges - 18.22? of FCI $ 619,480
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13) $ 1,717,800
UNIT PRODUCTION COST
15. Gross - Items (7+12+13) $ 1,718,000
A. Mills/KWH .70
B. $/Ton S not emitted 11?'. 2
16. Net - Items (9+12+13) $ 1,718 000
A. Mills/KWH .70
B. $/Ton S not emitted 112.2
* Extra coal needed for limestone calcination in boiler
-------
TABLE 8
ANNUAL OPERATING COSTS
CAT-OX
1
2
3
6
Plant Size (MW)
Fixed Capital Investment (PCI)
Stream Time (hrs/yr)
1000
$43, 400, 000
7000
JTHKCT COST
Operating Labor (4 men/shift § $4.50/hr)
Supervision - 152 of Item 1
Maintenance Labor and Materials - 4.5.2 of FCI
Plant Supplies - 152 of Item 3
Utilities
A. Cooling Water - 548 GPM at $.025/M sal
P.. Process Water - — GPM at $.20/M f$al
<"• , Electric Power - 17,800 KW at 6.75 mills/KWH
, Natural Gas - — MSCFH at $.40/HSCF
, Steam — M Ibs/hr at $.45/M Ibs
Chemicals and Raw Materials
A. Sulfuric Acid Catalyst - 1330 ft3/yr 6 $38.23/ft3
Subtotal Direct Cost (ex credits)
Credits
A. Sulfuric Acid (782) - 1061 TPD @ $6/TON 1002
B. Stack Gas Coolinc* - 162 MM Etu/hr § $.40/MM Btu/hr
Total Direct Cost
KKCT COST
10. P;iyroll Overhead - 20^ of (1 + 2)
11. Plant Overhead - 50% of (1+2 + 3 + 4
12. Total Indirect Cost
FIXED COST
13. Capital Charges - l8.22£ of FCI
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13)
UNIT PRODUCTION COST
15. Cross - Items (7+12+13)
A . i-ijlls/KWH
B. .p/Ton S riot emitted . '
1C. ,'Ict - Items (9 + 12+13) ; -
A. Mills/KWH
B. $/Ton S not emitted .
* Credit for stack gas at 225°F vs base case 300°F
$/YF,AR
157,6^0
?3,6r;0
1,953,000
29?,9r,0
5,760
-
841,050
50,850
$ 3,324,940
$ 1,424,090
36,270
1,213,6^0
$ 1,2^9^910
$
7,907,^80
$10,581,460
$12,482,330
] .78
158.0
$10,581,460
1.5!
133.9
-------
TABLE 9
ANNUAL OPERATING COSTS
MOLTEN SALT
Plant Size (MW) 1000
Fixed Capital Investment (PCI) - $34,600,000
Stream Time (hrs/yr) 7000
DIRECT COST
1. Operating Labor- (5 men/shift fi $4.50/hr)
I-.'. Supervision - 15% of Item 1
3. Maintenance Labor and Materials - 5% of PCI
'I. Plant Supplies - 15% of Item 3
b. Utilities
A. Cooling Water - 650 GPM at $.025/M gal
B. Process Water - — GPM at $.20/M gal
C. P.lcctrlc Power - 14,575 KW at 6.75 mills/KWH
n. Natural Oas - — MSCFH at $.40/MSCF
E. Steam — M Ibs/hr at $.45/M Ibs
6. Chemicals and Raw Materials
A. Lithium Carbonate - 115 Ibs/hr @ $.42/lb
B. Sodium Carbonate - 990 Ibs/hr § $.02/lb
C. Potassium Carbonate - 1684 Ibs/hr § $.05/lb
D. Petroleum Coke - 15.75 TPH @ $8/TON
6a. Solids Disposal (ex fly ash) - 1 TPH 6 $.50/TON
7. Subtotal Direct Cost (ex credits)
8. Credits
A. Sulfur - 261 TPD g $20/Short Ton
9. Total Direct Cost
INDIRECT COST
10. Payroll Overhead - 20% of (1+2)
11. Plant Overhead - 50% of (1+2+3+4)
12. Total Indirect Cost
FIXED COST
13. Capital Charges - 18.22% of FCI
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13)
UNIT PRODUCTION COST
15- Gross - Items (7+12+13)
A. Mills/KWH
B. $/Ton S not emitted
16. Net - Items (9+12+13)
A. Mills/KWH
B. $/Ton S not emitted
$/YEAR
197,100
29,570
1,730,000
259,500
6,830
688,670
338,100
138 ,600
589,400
882,000
3,500
$. 4,«63,270
(1,522,500)
$ 3,340,770
45,330
1,108,090
$ 1,153,420
$ 6,304,120
$10,798,310
$12,320,910
1.76
148.4
$10,798,310
1.54
130.1
-------
TABLE 10
ANNUAL OPERATING COSTS
POTASSIUM FORMATE
Plant Size (MW) 1000
Fixed Capital Investment (PCI) - $34,900,000
Stream Time (hrs/yr) 7000
DTHKC'P COST $/YEAR
1. Operating Labor (4 men/shift 0. $4.50/hr) 157,6fiO
J. Supervision - 15£ of Item 1 ' 23,650
3. Maintenance Labor and Materials - '-1.5% of PCI 1,570,500
;i . Plant Supplies - 15% of I ten 3 235,580
•,'."*• i 1. it ies '"•
A. Cool In,- W.itor - 30,000 GPM at $..025/*T p;al 315,000
B. Process Water - 1,840 GPM at $.20/M nal 154,560
C. Electric Power - 37,540 KW at 6.75 mills/KWH 1,680,890
1). Natural Gas - 553 MSCPH at $.40/MSCF 1,5-48,800
F,. Steam - 116 M Ibs/hr at $.45/M Ibs 157,500
(-.. Chemicals and ;iaw Materials
A. Potassium Carbonate - 250 Ibs/hr @ $.05/lb 87,500
7. Subtotal Direct Cost (ex credits) $ 5,931,660
8. Credits ' . '
A. Sulfur - 2^5 TPD (3 $20/Short Ton (1,429,200)
9. Total Direct Cost $ 4,502,460
INDIRECT COST
10. Payroll Overhead - 20% of (1+2) . 36,270
11. Plant Overhead - 50# of (1+2+3+4) 993,710
12. Total Indirect Cost • $ 1,029^980
FIXED COST
13. Capital Charges - 18.22$ of FCI $ 6,358,780
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13) $11,891,220
UNIT PRODUCTION COST
15. Gross - Items (7+12+13) $13,320,420
A. Mills/KWH ' 1.90
B. $/Ton S not emitted ' . 169.3
16. Net - Items (9+12+13) $11,891,220
A. Mills/KWH 1.70
B. $/Ton S not emitted' 150.7
-------
TABLE 11A
ANNUAL OPERATING COSTS
AMMONIA SCRUBBING (BASE)
(Nil,,) 2 SO. PRODUCTION
(324,800 TPY)
7
Plant Si?.e (MW)
Fixed Capital Investment (PCI)
Stream 'I'iMe (hrs/yr)
',rT COST
1000
$ 10,200,000
7000
OpRT'-itlrif; Labor (3 men/shift fl $'i.50/hr)
Supervision - lri? of Item 1
Maintenance Labor and Materials - 4.5$ of FCI
Plr-nt Supplies - 15* of Item 3
U t ! 1 .i 1 1 e o '
A. Coc.li.nc; Water - 8,800 GPM at $.025/M cal
B. Process \vator - r>,000 GPM at $.20/M gal
C. El oc trie Power - 24,64? KW at 6.75 mills/KWH
• -) . Natural Has - — MSCFH at $.40/MSCF
'•:. Steam - 390 M Ihs/hr at $.45/M Ibs
Chemicals and Haw Materials
A. Ammonia - 287 TPD @ $50/TON
Subtotal L'irect Cost (ex credits)
Credits
Total IHrect Cost
$/YEAR
118,260
17,740
459,000
68,850
92,400
420,000
1,164,600
1,228,500
4,190,200
$7,759,550
TNI)I;-!HOT COST
10. Payroll Overhead - 20?, of (1 + 2)
11. Plant Overhead - 50£ of (1+2+3+4)
I-'1. Total Indirect Cost
FIXHP COST
13- Capital Charges - 18.227, of FCI
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13)
UNIT PRODUCTION COST
l;.. 'Iross - Items (7 + 12 + 13)
A. Mills/KWH
H. ^./Ton S not emitted
C. $/Ton (NII4)2 SO 14
27 ,200
331,930
$ 359,130
$1,858,440
$9,977,120
$9,977,120
] .43
126.8
30.80
-------
TABLE 11B
ANNUAL OPERATING COSTS
AMMONIA SCRUBBING (BASE)
HN03/FERTILIZER PRODUCTION
(691,250 TPY FERTILIZER)
Plant Size (MW) - 1000
Fixed Capital Investment (FCI) - $28,000,000
Stream Time (hrs/yr) ' - 7000
CT COST $/YF:AR
1. Operating Labor '(14 men/shift % $4.50/hr) 551,880
P. "upervislon - 15£ of Item 1 82,780
;. , •";!•' ntenanoe Labor and Materials'.- 4.5$ of FCI 1,260,000
': . Pi ant Supplies - 15# of Item 3. ,,..., : 1.89,000
S Utilities
A. OooHnr loiter - — GPM at $.025/M gal •• • -
P.. rrocer,:-, Water - 25,000* GPM at $.40/M gal ••' 4,200,000
C. Electric Power - .6,135 KW at 6.75 mills/KWH 289,900
D. Natural Gas - — .MSCFII at $.40/MSCF -
F.. raean - 398 M Ibs/hr at $.45/M Ihs 1,253,700
r . " helicals and Hav; Materials
A. (1111^)2304 Transfer Price - 1112 TPD @ $30.80/TON 10,800,880
3. Ammonia . - 517 TPD @ $50/TON '. 7,5^8,200
C. Pliosphate Rock - 1128 TPD @ $11.88/TON .'.. 3,908,500
D. Anti-Foam Ar;ent - QOO Ibs/day '% $.15/lb 39,420
F,. Nitric Acid'Cat - 119.4 tr.oz @ $120/oz • ,. 14,330
:'. Conditioner — 48 TPD § $46.6/TON 653,150
('••:. Solids Disponal -.. 72 TPH @ $.50/TON . 252,000
7. Subtotal Dir-ect Cor.t (ex credits) ',; $30,2'I3,740
Ii . Credits . _ - -
9. Total Direct Cost . , $ =
INDIiU'CT COST / ' '
10. Payroll Overhead - 20? of (1+2) . 126,930
11. Plant Overhead - 50^ of (1+2+3+4) 1,041,830
12. Total Indirect Cost $ l,lbd,760
FIXED COST
13. Capital Charces - 18.222 of FCI " $ 5,101,600
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OFEHATING COST
14. ;jet Production Cost - Items (9 + 12 + 13) $36,514,100
UNIT PRODUCTION COST
15. Gross - Items (7+12+13) $36,514,100
A. $/Ton Fertilizer . $ 52.82
D. $/Ton Fertilizer w. Ammonia & $30/TON $ 46.03
* Chloride free water of BFW quality @ $.40/M gal ..
-------
TABLE 12
ANNUAL OPERATING COSTS
AMMONIA SCRUBBING (A)
(STEAM STRIPPING)
Plant Size (MW) 1000
Fixed Capital Investment (PCI) - $16,200,000
Stream Time (hrs/yr) 7000
DIRECT COST $/YEAR
]. Operating Labor (4 men/shift @ $4.50/hr) ' 157,680
2. Supervision - 15% of Item 1 23,650
3. Maintenance Labor and Materials - 1.5JS of PCI 729,000
';. Plant Supplier, - 155? of Item 3 109,350
';. Utilities
A. Ccolinc Water - 24,640 GPM at $.025/H gal 258,830
B. Proces? Water - 3,000 GPM at $.20/M gal 252,000
C. Electric Power - 18,298 KW at 6.75 mills/KWH 864,COO
D. Natural Gas - — MSCFII at $.40/MSCF
K. Steam - 500 M Ibs/hr at $.45/M Ibs 1,575,000
C. Ciiemlcals and Raw Materials
A. Limestone - 250 TPD g $2/TON 291,700
'-. Ammonia - 300 Ibs/hr @ $50/TON 52,?nn
6u. oclids Disposal (ex fly ash) - 13.3 TPH @ $.50/TON ;l6,550
7. Subtotal Direct Cost (ex credits) $ 4,360,^60
3. Credits
A. Sulfuric Acid (98.5£) - 570 TPD @ $12/TON (1.995,000)
9. Total Direct Cost $ 2,365,^60
INDIRECT COST
10. Payroll Overhead - 20^ of (1+2) 36,270
11. Plant Overhead - 50£ of (1+2+3+^) 509,850
12. Total Indirect Cost $ 5^6,120
FIXED COST
13. Capital Charges - 18.22?, of FCI $ 2,951,640
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13) $ 5,863,620
UNIT PRODUCTION COST
15. Gross - Items (7+12+13) / $ 7,858,620
A. Mills/KWH 1 1.12
B. :p/Von S not emitted ' 99-88
16. Net - Items (9+12+13) $ 5,863,620
A. Mills/KWH 0.84
B. .p/Ton S not emitted ' 74.25
-------
TABLE 13
ANNUAL OPERATING COSTS
AMMONIA SCRUBBING (B)
(THERMAL DECOMPOSITION)
Plant Size (MW) 1000
Fixed Capital Investment (FCI) - $21,600,000
Stream Time (hrs/yr) 7000
DIRECT COST $/YEAR
1. Operating Labor (4 men/shift & $4.50/hr) 157,680
2. Supervision - 15% of Item 1 23,650
3. Maintenance Labor and Materials - 4.5% of FCI 972,000
'i. Plant Supplies - 15% of Item 3 145,800
5. Utilities
A. Coolinr; Water - 8,000 GPM at $.025/M gal 84,000
B. Process Water - 400 GPM at $.20/M gal 33,600
C. Electric Power - 40,218 KW at 6.75 mills/KWH 1,900,300
;). Natural Oas -• — MSCF1I at $.40/MSCF
E. Steam * - 76.9 M Ibs/hr at $.45/M Ibs 260,lCO
C>. Chemicals and Raw Materials
A. Ammonia - 300 Ibs/hr g $50/TON 52,500
B. Limestone - 168 TPD @ $2/TON 98,000
6a. Solids Disposal (ex fly ash) - 9 TPH @ $.50/TON 31,500
7. Subtotal Direct Cost (ex credits) $ 3,759,190
3. Credits
A. Sulfuric Acid (98.556) - 692 TPD @ $12/TON .(2,422,000)
9. Total Direct Cost $ 1,337,190
INDIRECT COST
.10. Payroll Overhead - 20£ of (1 + 2) 36,270
11. Plant Overhead - 50? of (1+2+3+4) 649 570
12. Total Indirect Cost $ 6«5,«40
FIXED COST
13. Capital Charges - 18.2235 of FCI $ 3,935,520
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. iiet Production Cost - Items (9+12+13) $ 5,958,550
liTJIT PRODUCTION COST
15. Gross - Items (7+12+13) $ 8,380,550
A. Mills/KWH 1.20
D. 0/Ton S not emitted 106.6
16. Net - Items (9+12+13) $ 5,958,550
A. Mills/KWH .85
B. 3/Ton S not emitted 75-5
* 51,300 Ibs/hr (150 psia) @ $.45/M.lbs
25,600 Ibs/hr (700°F superheated) @ $.55/lb
-------
TABLE 14
ANNUAL OPERATING COSTS
TYCO
Plant Size (MW) 1000
Fixed Capital Investment (PCI) - $29,800,000
Stream Time (hrs/yr) 7000
DIRECT COST $/YEAR
1. Operating Labor (5 men/shift @ $4.50/hr) 197,100
?.. Supervision - 1535 of Item 1 29,570
3. Maintenance Labor and Materials - 5% of PCI 1,490,000
4. Plant Supplies - 15% of Item 3 223,500
rj. Utilities
A. Cooling Water - 1,720 GPM at $.025/M gal 18,060
B. Process Water - 8 GPM at $.20/M gal 1,3^0
C. Electric Power - 24,730 KW at 6.75 mills/KWH 1,168,480
D. Natural Gas - — MSCFH at $.40/MSCF
E. Steam — M Ibs/hr at $.45/M Ibs
6. Chemicals and Raw Materials
A. Activated Carbon* - 68,000 ftVyear § $13.50/ft3 918,000
B. Limestone - 1.25 TPH @ $2/TON 17,500
6a. Solids Disposal (ex fly ash) - 1 TPH @ $.50/TON 3,500
7. Subtotal Direct Cost (ex credits) $ 4,067,050
8. Credits
A. Sulfuric Acid (805?) - 971 TPD @ $6/TON 100% (1,359,500)
B. Nitric Acid (6058) - 148 TPD g $25/TON 100% ( 648,000)
C. BFW Heating - 170 MM Btu/hr g $.40/MM Btu ( 474,900)
9. Total Direct Cost $ 1,584,650
INDIRECT COST
10. Payroll Overhead - 20% of (1+2) 45,330
11. Plant Overhead - 50$ of (1+2+3+4) 970,090
12. Total Indirect Cost $ 1,015,420
FIXED COST
13. Capital Charges - 18.22% of FCI $ 5,429,560
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13) $ 8,029,630
UNIT PRODUCTION COST
15. Gross - Items (7+12+13) $10,512,030
A. Mills/KWH 1.50
B. $/Ton S not emitted 137.7
16. Net - Items (9+12+13) $ 8,029,630
A. Mills/KWH ' 1.15
B. $/Ton S not emitted 104.7
* Assume one year activated carbon life In catalytic stripper.
-------
TABLE 15
ANNUAL OPERATING COSTS
MAGNESIUM OXIDE SCRUBBING
Plant Size (MW)
Fixed Capital Investment
Stream Time (hrs/yr)
(FCI)
1000
$26,800,000
7000
DIRECT COST
1. Operating Labor (4 men/shift @ $4.50/hr)
2. Supervision - 15% of Item 1
3. Maintenance Labor and Materials - 4% of FCI
4. Plant Supplies - 15% of Item 3
5. Utilities
A. Cooling Water - 4,230 GPM
B. Process Water - 616 GPM
C. Electric Power - 22,656 KW
D. Natural Gas -
E. Steam -
@ $.025/M gal
@ $.20/M gal
@ 6.75 mills/KWH
268 MSCFH @ $.40/MSCF
M Ibs/hr e $.45/M Ibs
6. Chemicals and Raw Materials
A. Magnesia - 5 TPD @
B. Petroleum Coke - 645 Ibs/hr (dry) § $8/TON
C. Limestone - 1,500 Ibs/hr @ $2/TON
7. Subtotal Direct Cost (ex credits)
8. Credits
A. Sulfuric Acid (98.5%) - 890 TPD
9. Total Direct Cost
INDIRECT COST
10. Payroll Overhead - 20% of (1+2)
11. Plant Overhead - 50% of (1+2+3+4)
12. Total Indirect Cost
FIXED COST
13. Capital Charges - 18.22% of FCI
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13)
UNIT PRODUCTION COST
15. Gross - Items (7+12+13)
A. Mills/KWH
B. $/Ton S not emitted
16. Net - Items (9+12+13)
A. Mills/KWH
B. $/Ton S not emitted
$/YEAR
157,680
23,650
1,072,000
160,800
44,420
51,740
1,070,500
750,400
58,330
18,060
10,500
$3,418,080
(3,115,000)
$ 303,080
36,270
707,070
$743,340
$4,882,960
$5,929,380
$9,044,380
1.29
111.1
$5,929,380
0.85
71.3
-------
TABLE 16
ANNUAL OPERATING COSTS
ZINC OXIDE SCRUBBING
Plant Size (MW) 1000
Fixed Capital Investment (FCI) $21,800,000
Stream Time (hrs/yr) 7000
DIRECT COST $/YEAR
1. Operating Labor (4 men/shift @ $4.50/hr) 157,680
2. Supervision - 15% of Item 1 23,650
3. Maintenance Labor and Materials - 4% of FCI 872,000
4. Plant Supplies - 15% of Item 3 130,800
5. Utilities
A. Cooling Water - 5,320 GPM @ $.025/M gal 55,860
B. Process Water - 47 GPM @ $.20/M gal 4,000
C. Electric Power - 20,130 KW 3 6.75 mills/KWH 951,140
D. Natural Gas - 145.3 MSCFH @ $.40/MSCF 406,840
E. Steam - M Ibs/hr @ $.45/M Ibs —
6. Chemicals and Raw Materials
A. Zinc Oxide - 200 Ibs/hr @ $.15/lb 210,000
B. Sodium Carbonate - 325 Ibs/hr @ $.02/lb 45,500
C. Lime - 2.1 TPH @ $15/TON 220,500
6a. Solids Disposal (ex fly ash) - 14.2 TPH i $.50/TON 49,700
7. Subtotal Direct Cost (ex credits) $3,127,670
8. Credits
A. Sulfuric Acid (98.5%) - 754 TPD @ $12/TON (2,639,000)
9. Total Direct Cost $ 488,670
INDIRECT COST
10. Payroll Overhead - 20% of (1+2) 36,270
11. Plant Overhead - 50% of (1+2+3+4) 592,070
12. Total Indirect Cost $ 628,340
FIXED COST
13. Capital Charges - 18.22% of FCI $3,971,960
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13) $5,088,970
UNIT PRODUCTION COST
15. Gross - Items (7+12+13) $7,727,970
A. Mills/KWH 1.10
B. $/Ton S not emitted 98.1
16. Net - Items (9+12+13) $5,088,970
A. Mills/KWH 0.73
B. $/Ton S not emitted 64.4
-------
TABLE 17
ANNUAL OPERATING COSTS
CITRATE
Plant Size (MW) 1000
Fixed Capital Investment (PCI) - $21,200,000
Stream Time (hrs/yr). 7000
DTKliCT COST $/YEAR
]. operating Labor (4 men/shift § $4.50/hr) 157,630
?.. Supervision - 152 of Item 1 ' 23,650
3. Maintenance Labor and Materials - 4$ of PCI 848,000
;i. Plant Supplies - 15? of Item 3 127,200
r-. Utilities
A. Cooling Water - 34,395 GPM at $.025/11 gal 361,150
B. Process Water - 537 GPM at $.20/M gal 45,100
C. Electric Power - 18,550 KW at 6.75 mills/KWH 920,000
D. Natural Gas - 171.4 MSCPH at $.40/MSCP 479,900
E. Steam •- 25 M Ibs/hr at $.45/M Ibs 78,950
6. Chemicals und Raw Materials
A. Na2C03 - 135 Ibs/hr § $.02/lb 18,900
B. Citric Acid -- 92 Ibs/hr § $.30/lb 193,200
C. Limestone - 24.5 TPD g $2/TON 14,300
6a. Solids Disposal (ex fly ash) - 1.5 TPH @ $.50/TON 5,250
7. Subtotal Direct Cost (ex credits) $ 3,273,280
3. Credits
A. Sulfur - 251 TPD @ $20/Short TON (1,464,170)
9. Total Direct Cost $ 1,809,110
INDIRECT COST
10. Payroll Overhead - 202 of (1+2) 36,270
11. Plant Overhead - 50% of (1+2+3+4) 578,280
12. Total Indirect Cost $ 615,050
FIXED COST
13. Capital Charges - l8.22# of PCI $ 3,862,640
(Includes Depreciation, Interim Replacements,
Insurance, Taxes, and Cost of Capital)
TOTAL OPERATING COST
14. Net Production Cost - Items (9+12+13) $ 6,286,800
UNIT PRODUCTION COST
15. Gross - Items (7+12+13) $ 7,750,Q70
A. Mills/KWH 1.11
B. $/Ton S not emitted 98°. 3
16. Net - Items (9+12+13) $ 6,286 800
A. Mills/KWH 0.90
B. $/Ton S not emitted 79*.4
-------
VI. IMPACT
The potential S02 emission (without abatement) from
power plants burning coal and oil has been estimated as follows
(Reference 1):
SO„ Emission, Millions of Tons/Year
1970
20.0
16.6
36.6
— ^
1975
30.2
18.0
48.2
1980
41.1
19.8
60.9
1990
62.0
24.4
86.4
2000 2010
94.5 116
31.3
125.8
2020
88
-
_
Power Plants
All Other
TOTAL
The use of coal will continue to increase and will approximately
triple by the year 2000. Peak SO,, emission from power plants will
be reached in the year 2010, after which it will decline as nuclear
power plants become the predominant energy source. The above esti-
mate, made in 1970, differs from a previous 1967 estimate (Refer-
ence 2) primarily in a shift in the peak year of coal consumption
from 1990 to 2010. This discrepancy is due to the slower utilization
of nuclear plants than originally predicted.
Power plants burning coal and oil are the largest single
source of man-made SO2 released to the atmosphere, and consequently
the development of methods for reducing the quantity of this source
of air pollution has a high priority in order to meet standards set
by the Air Quality Act of 1967, as well as other national and re-
gional standards which are currently being formulated. The current
practice in urban east coast areas has been to use low sulfur fuels
as a means of reducing the total SO2 emission. This technique has
proven successful and will continue to be used in many urban area
-------
power plants where other methods would not be either feasible
or economical. However, the present and future availability
of low sulfur fuels (both coal and oil) is limited, and would
not be adequate to provide the degree of nation-wide abatement
in S0_ emission required to meet air quality standards now
being established.
Sulfur oxide emission from power plants can be controlled
in several ways among which are: the removal of sulfur from
fuels before combustion; removal of S02 from the combustion gas;
and the use of high stacks and remote power plant sites for gas
dispersion. The first approach is being used with oil and gas
fuels but has not been successfully developed for use on high
sulfur coals. The present report is primarily concerned with
the impact of the use of processes now under development on the
reduction of SO- emission from power plant stack gas. Since
even the most efficient of these processes will only reduce the
SO_ content to the level of several hundred parts per million,
they must be used in conjunction with high stacks and judicious
site location to ensure adequate dispersion to meet S0» ground
level air quality standards.
References No.2 and 3 give estimates of potential SO?
emission levels and the amount of reduction considered feasible
using various control methods. The basis of prediction used
in this report is geared to S02 emission control by removal from
the stack gas by any of the processes covered in the report. The
predicted reduction represents the ultimate that may be achieved
by these methods if applied where feasible on existing power
plants and on all new plants which become operative after 1975.
It is assumed that the technology of the S02 removal processes
under consideration will be adequately established for commercial
use by 1973, allowing two years for installation on existing plants
-------
by 1975 and on all new plants now in the planning stage which
will become operative by 1975 and thereafter. Assuming limited
abatement in S02 emission, the level of total S02 emission from
coal and oil burning power plants will have reached 30.2 MM tons/
year by 1975. A recent survey (Reference 4) indicated that about
75% of the present U. S. generating capacity from existing power
plants over 200 MW capacity is amenable to some type of S02 re-
moval process based on area requirements only.
When consideration is given to other factors such as cost,
waste disposal capabilities, etc., an optimistic estimate would
be about 35% of the total existing generating capacity could be
adapted to some type of SO? removal installation. Assuming that
35% of the existing S02 emission from power plants in 1975 could
be placed under control removing 90% of the S02 content of the
combustion gas, and that 100% of all new plants built from 1975
on would be placed under similar control, it would be possible
to reduce the level of S02 emission from power plants to about
15 MM tons/year in 1975. The emission level would gradually in-
crease to a maximum of 25 MM tons/year in 2010, at which time
the use of fossil fuels would dPC:line .11.d nuclear power plants
would predominate. If none of the existing power plants are
placed under control, but all new plants built from 1975 on are,
the total S02 emission from power plants would be reduced to
about 20 MM tons/year in 1975, increasing to a maximum of 30 MM
tons/year in 2010. Further reduction in the S02 emitted could
be effected by increasing the control process removal efficiency
and by the use of lower sulfur fuels. The 15 to 25 MM tons/year
SO- emission range is in reasonable agreement with previous
estimates which assumed the application of maximum control
methods. The difference is largely due to the fact that peak
SO2 emission is to be reached in 2010 rather than 1990 as
originally estimated.
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The total S02 removed from power plant flue gas- would
range from 15.5 MM tons/year in 1975 to 92.8 MM tons/year in
2010. The actual net recovery of SO,, as equivalent sulfur
or sulfuric acid from the various removal processes under
consideration ranges from 0% for the wet and dry limestone
processes to 70-90% for the various non-throwaway type pro-
cesses. Therefore, if the wet limestone process were used,
its impact would not be on the existing sulfur or sulfuric
acid supply/demand balance but on its possible effects as a
ground water pollution source. The very large quantities of
CaSO, and possibly MgSO. produced as a by-product of this
removal process would present disposal problems which must
be considered in their relation to the surrounding environ-
ment.
Using an average value of 85% SO- recovery for the
£*
regenerable type processes, the equivalent sulfur recovery
would be 6.6 MM tons/year (1975) and 10.8 MM tons/year (1980)
This compares to an estimated total U. S. sulfur consumption
of 13.4 MM tons/year (1975) and 16.2 MM tons/year (1980).
Beyond 1980, the recovered sulfur from power plants alone
would exceed total U. S. consumption. The impact of such
large quantities of sulfur would seriously disrupt the
existing market requiring a complete restructuring of the
supply sources.
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VII. ADVANCED POWER CYCLES
The report entitled, "Technological and Economic
Feasibility of Advanced Power Cycles and Methods of Producing
Nonpolluting Fuels for Utility Power Stations" has been re-
viewed from a different viewpoint than were the reports covering
the various sulfur dioxide removal processes. The multiplicity
of systems and plant cycles covered in this report coupled with
the higher priority assigned to the S02~ removal processes made
it impractical in terms of time and manpower expenditures al-
located to Task #5 to examine these systems in the detail applied
to the sulfur dioxide removal processes. Also, since many of
the conclusions reached are based on predictions of future costs
and future technology, it is more appropriate to comment on the
general concepts involved than on the specific systems, power
cycles or economic evaluations.
Fuel Desulfurization Process
The main concept presented in this section is that the
ultimate goal for a desulfurized fuel for utility power plant
application lies in producing a desulfurized gaseous fuel. We
agree with this "clean fuel" goal, particularly in terms of
producing a high pressure, high heat content, pipeline gas for
the following major reasons:
1. With conventional power plants of modern design
capital costs for a gaseous fueled plant are substantially
(about 30%) lower than for a solid or liquid fueled plant.
This margin will probably prevail for plants utilizing ad-
vanced cycle concepts, assuming they could use solid or
liquid fuel. Also, many advanced cycle plants may require
gaseous fuels.
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2. Considerable effort is currently being expended
in the development of gasification processes for solid and
liquid fuels and many potentially attractive processes are
now under consideration for the production of synthetic
natural gas (SNG). The commercialization of these processes
and/or the development of new sources of natural gas will in-
crease the economic attractiveness of gaseous fueled power
plants.
3. The production of the equivalent of natural gas
by gasification of solid or liquid fuel will make power plant
siting less dependent on the location of the fuel supply since
existing or new piping networks can economically distribute the
fuel to optimized power plant locations.
The concept of a gasification plant producing low BTU
gas for use by an adjacent, close-coupled power plant may appear
to be economically attractive and may indeed be so for the "early
years" when the plant is base loaded. However, experience indicates
that as plants age their utilization factor decreases. Thus the
close-coupi.rcl qa.s i I:i cation -power plant design would have to in-
clude "intermittent power production" capabilities as well as
face the potential obsolescence of the combined facility as
opposed to only the power plant if it and the gasification plant
were truly independent.
Advanced Cycle Power Systems
Thermodynamic fundamentals indicate that the greater the
difference between the temperature at which heat is added to a
power cycle and the temperature at which heat is rejected from
the power cycle the greater is the potential thermal efficiency
of the cycle. Since the temperatures at which heat can be re-
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jacted from the cycle are set by the environment, the efforts
at improving power plant efficiencies have been largely directed
toward increasing the temperature at which heat is added to the
cycle. Along with increasing the temperature and pressure at
which the steam is generated and increasing the superheat tempera-
ture, attempts were made to develop "topping" cycles using fluids,
such as mercury, with pressure-temperature characteristics more
suitable than water to high temperature operation. The problems
associated with mercury systems were such that the mercury top-
ping cycle was abandoned, and pre-1950 efforts were directed
toward improving the steam cycle itself. With the advent of
larger, more reliable, gas turbine renewed interest developed
in the topping cycle using the "hot gas" system to "top" the
steam system.
We agree with the conclusion that future power plants
will probably employ advanced cycles using both gas turbines
and steam turbines for the following reasons:
1. Many studies have been made, here and abroad, to
evaluate such cycles and combined gas turbine-steam turbine
plants are now commercially available in modest size (such as
the 180 MW STAG power plant from General Electric Co.). There
is no question that these plants have a higher potential effi-
ciency than the straight steam power plant without regard to
whether the practical maximum steam temperature is 1000°F,
1100°F or 1200°F.
2. Since fuel costs are a major factor in the cost of
generating power, utilities can be expected to favor cycles
yielding higher efficiencies, with a pollution-free effluent
representing an additional advantage.
-------
3. Cycles of increased complexity, such as tertiary
cycles, may possibly be designed to have higher efficiencies
than the simpler gas turbine-steam turbine cycle, but it is
extremely unlikely that the efficiency advantage would justify
the added complexity. The past history of the mercury cycles
is not easily forgotten. Most utilities will be reluctant to
accept systems utilizing potentially hazardous fluids such as
ammonia. Their lack of experience in handling these fluids
plus the added equipment requirements of tertiary cycles make
them more of academic than practical interest.
Despite the long-term potential of combined gas turbine-
steam turbine power plants and the increased use of gas turbine
power for handling peak loads, the lack of experience with large
size, base loaded, gas turbines should not be ignored. Con-
siderable reluctance may be expected on the part of utilities
to being among the first to commit themselves to carrying a base
load with gas turbine driven generators.
Integrated Power Stations
It is not -'Tear, particularly at this stage of development
of gasification processes, that close integration of the gasifi-
cation plant with the power plant would result in an optimized
system based on the anticipated useful life of the plant. Better
overall on-stream reliability may be attained by producing pipeline
gas for delivery to gas distribution systems and drawing the power
plant fuel from these gas distribution systems. Utilities marketing
both gas and electric power may find this kind of system particularly
attractive.
Research and development programs related to high pressure
gasification and desulfurization of solid and liquid fuels should
be strongly supported as these will directly affect air quality
without regard to the type of power plant in which the gasified
fuel is burned.
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Research and development programs related to development
of advanced materials and cooling techniques should also be
supported since such materials and techniques could probably
be utilized in the design of the gasification plant equipment,
as well as in the design of advanced gas turbines.
Considering the current level of experience with Lurgi
coal gasifiers and the fairly extensive experience with gas
turbine systems, it is questionable if there is very much to
be gained by constructing a pilot or demonstration plant at
this time. Detailed design studies, however, to develop rela.r
tively firm economics for a full-scale integrated gasification-
power plant would be most useful in pointing up specific problem
areas and major cost items requiring more intensive development
-------
VIII. REFERENCES
(1) "Abatement of Sulfur Oxide Emissions From
Stationary Combustion Sources"
National Academy of Engineering,
National Research Council, Washington, D.C., 1970
(2) Rohrman, F.A., Steigerwald, B.J., and Ludwig, J.H.
"SO2 Pollution: The Next 30 Years",
Power, May 1967, p. 82
(3) Hangebrauck, R.P. and Spaite, P.W.,
"Controlling the Oxides of Sulfur",
JAPCA, Jan. 1968, Vol. 18, No. 1, p. 5
(4) M.W. Kellogg Co., Letter No. K-5-6,
Applicability Survey Report, March 3, 1971
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I'X. APPENDIX
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LIMESTONE - WET SCRUBBING PROCESS
A. PROCESS SECTION
1. Process Description (see MWK Dwg. No. PD-108-D)
The limestone-based wet scrubbing process involves the
removal of sulfur dioxide and fly ash from a coal-fired boiler
stack gas through the use of an aqueous slurry of raw limestone
or lime. The overall reactions involved in the S02 removal
step are as follows:
CaCO3 -> CaO + CO2
CaO + S02 + 1/2 02 + CaS04
The particular flow scheme chosen for evaluation in the study for
APCO involves the injection of raw stone into the boiler to achieve
complete calcination and partial sulfation. The remainder of the
S02 removal is accomplished via liquid phase reactions which occur
in the scrubber system (see Chemistry Section of this write-up
for a discussion of the various reactions possible).
Raw limestone is unloaded into hoppers and transported to
storage silos and feed silos by belt conveyor. Limestone is
fed to pulverizers to reduce its size to 70% minus 200 mesh,
where it is mixed with preheated air for drying and transporting
to boiler injection ports. It is assumed that two conventional
500-MW pulverized-coal fired boilers will be used in this new
installation, hence the process flowsheet is drawn to show two
parallel stack gas treating trains.
The boiler stack gas downstream of the air preheater contains
fly ash and other solid components (CaC03/ CaO, CaSO4). The main
flow stream from each 500-MW boiler is split into three parallel
scrubbing trains, the number being set by the maximum feasible
scrubber size. The stack gas is first cooled from 310°F to 177°F
in an indirect gas-liquid exchanger employing circulating water
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as the heat transfer medium. The gas will then be fed to a counter-
current scrubber for removal of inert solids and completion of the
CaO-SC>2 reaction. The particular scrubber design selected must
have a high solids removal efficiency (>99%) to eliminate the
expensive precipitator installation in a new plant. Effluent
gas from the scrubbers is heated to 250°F to achieve satisfactory
plume dispersion by the circulating hot water system, after which
it is sent to the stack via induced draft fans.
The liquid effluent from each scrubber tower is sent to hold
tanks wherein several minutes residence time is provided to achieve
controlled crystallization of CaS04 reaction product on existing
particles. This design will hopefully minimize deposition and
plugging in liquid circulation lines. A combined purge stream
from the scrubber circuit is pumped to a settling pond for removal
of solids, e.g., fly ash, CaS04 and Ca(OH)2 formed from hydration
of the unconverted lime.
The settling pond overflow is recycled to the scrubber system
on the premise that local water pollution considerations would
not tolerate the discharge of CaS04 and MgSC>4 from the limestone
streams into an adjacent watercourse. Make-up process water is
added to the system to compensate for the evaporation loss from
the settling pond plus the scrubber water vaporization resulting
from direct gas cooling.
2. Process Design Bases
The base case flowsheet developed for wet limestone scrubbing
is for a new 1,000 megawatt installation with pulverized limestone
injection into the boiler. There are other process variations
worthy of investigation, e.g., limestone or pre-calcined lime
injection into the slurry circuit, however the Kellogg base case
was selected to be identical with the base case in the TVA con-
ceptual design study (1). The TVA process design bases were
maintained wherever possible, with exceptions noted below, and
were scaled-up from 200 MW to 1,000 MW generation capacity.
* Refers to references at end of section
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Another reason for selecting the boiler injection process
scheme as the base case is that it is the most advanced .concept
having been demonstrated in two commercial plants (125-140 MW)
by Combustion Engineering. Unfortunately, no information from
these plants was made available to Kellogg during the present
study for APCO, and the commercial experience could not be
factored into this evaluation. The ability to include this
experience would shed considerable insight into the potential
problem areas discussed in a subsequent section of the write-up.
The following is a list of specific process and equipment
design bases used with amplification provided in the section
entitled design rationale:
a) Limestone addition rate ~ 110% stoichiometric
b) Limestone composition: 91.8% CaCC>3 (dry basis)
10. % moisture
c) 100% Sulfur in Coal evolves in Stack Gas (802+803)
d) 71.6% Ash in Coal evolves as Fly Ash
e) 100% Conversion of CaC03 to CaO in Boiler
25% Conversion of CaO to CaS04 in Boiler
f) Stack Gas Ductwork: Use 3 main ducts from 500 MW
Boiler (3 parallel trains). Provide separate
>. oo.lei , scrubber, reheater aud induced-draft fan
per train (6 parallel systems total - 1,000 MW)
g) Stack Gas Cooler and Reheater Design; Indirect
exchange against circulating water (can also add
glycol), single pass gas side cross-flow unit,
overall transfer coefficient 8 Btu/hr x °F x ft2,
220,000 ft2 surface/unit (6 Coolers and 6 Reheaters
required - 12 identical units), Finned tube bundles
40 ft long x 12 ft wide x 8 ft deep,
Cooler Design AP 4 inches H2O (includes gas distributor)
Reheater Design AP 4 inches H20 (includes gas distributor)
h) Scrubber Design (6 identical units req'd)
UOP TCA Contactors
Vessel size: 15 ft deep x 34 ft x 40 ft high with
center partition (2 Gas Inlets & 2 Gas Outlets)
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3 contacting stages of 1-1/2 inch plastic spheres
Design Inlet Liquid/Gas ratio = 6.12 Ibs/lb.
Design S02 Removal = 90 %
Design Particulates Removal = 99+%
Design AP (overall) = 9" H20 (including
Demister Pad)
Design Inlet Liquid cone. = 10.0% Solids
i) Recirculation System
Provide separate tank per Scrubber for optimum control
Size liquid lines for 8 ft/sec maximum velocity
Design tank residence time = 4 minutes
j) Rotating Equipment
Use 70% efficiency for centrifugal plumps and fans
(assume motor drives for estimating purposes)
Induced-draft fan head requirements (incremental only);
Gas Cooler 4 in H20
Scrubber 9 in H20
Gas Reheater 4 in H20
New Ducts 6 in H20
Incremental AP 23 in H2
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equipment design bases cited in the preceding section. The
discussion is broken down into the various process systems of
the base case flowsheet.
a) Limestone Feed System
The design bases for the raw limestone surge hoppers and
storage silos were taken directly from the TVA conceptual design
report. Similarly, the design basis for the limestone grinding
and injection system is as specified in the TVA report. There
is no need to achieve a finer degree of grinding than specified
as the S02 conversion in the boiler is not critical in the wet
limestone process. A number of suggested improvements in the
limestone handling equipment for a 1,000-MW plant is presented
in the Mechanical Review section of this report.
b) Boiler System
Essentially 100% calcination of raw limestone was assumed
followed by 25% conversion of CaO to CaSO4 in the boiler. These
assumptions are based on the most reliable pilot plant and full-
scale data from the dry limestone process test program. The
exact conversion in the boiler is not critical as the desired
overall S02 removal of 90% has been demonstrated in both ICI-
Howden pilot plant tests (1) and Combustion Engineering com-
mercial plant tests.(1) It was further assumed that the only
deleterious effect of limestone injection is the increased solids
handling, and that boiler operability(effect on firing, deposition)
would not be adversely affected.
c) Stack Gas Cooler-Reheater System
The indirect circulating water system recommended in the
TVA report based on studies of six different heat transfer schemes
was accepted without further evaluation. The temperature limits
were established .in accordance with TVA ' s recommendation of 250°F
stack gas temperature to achieve buoyancy for plume dispersion.
The overall design heat transfer coefficient of 8.0 seems rea-
sonable based on air cooler design experience. Single-pass gas
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crossflow unit having finned-tube bundles with 40 ft tube length
were used (rather than 20 ft) to minimize the number of'rows
and the gas phase pressure drop. Even so, the resulting design
yields 4 inches water AP per unit as compared with 2 inches water
AP in the TVA design. The higher pressure drop was used in
establishing the fan static pressure requirements.
d) Scrubber Design and Recirculation System
The TVA report recommends the UOP TCA Contactor (flooded-bed
design) as the optimum unit in a limestone slurry service based
on good pilot plant performance results in an actual stack gas
service. (2) However, the specific design conditions specified
by TVA were found to be inadequate when compared to the cited
pilot plant results. It was necessary to increase the liquid
circulation rate by 50% (L/G from 4.08 -> 6.12 Ibs/lb) and the
number of contacting stages from 2 to 3 to achieve 90% minimum
SC>2 removal and 99+% solids removal. The latter removal effi-
ciency is essential as no mechanical separators or precipitators
are proposed for the new plant design. The proposed new design
results in higher pressure drop (9 inches H20 vs. 6 inches H20)
including the loss through a demister pad.
The scrubber inlet liquid concentration of 10.0 wt % solids
taken from the TVA report was based on ICI-Howden pilot plant
data. The concept of a hold tank designed for 3-5 minutes resi-
dence time to achieve controlled crystal growth demonstrated at
ICI-Howden was also incorporated by TVA. The slurry solids content
should affect both reaction kinetics and crystallization rates,
however since there are insufficient data to optimize this variable,
the TVA design bases were accepted.
e) Solids Disposal System
The ability to sluice a 10% solids slurry of "reactive ash"
(containing CaO, CaS04) to the disposal pond without major
plugging problems is a critical technical consideration. Based
on successful ICI-Howden experience, the 10% solids concentration
-------
was accepted as a design value. The disposal pond water evapora-
tion loss was also taken from calculations in the TVA report and
pro-rated to 1,000 MW capacity.
Since the pond evaporation loss is 70% of the process make-up
water, this figure could be decreased by the following means;
providing deeper ponds (less area), prethickening the solids purge
stream, and use of an additive (long-chain alcohols) to decrease
the vaporization rate. However, the TVA proposal of sending the
pond overflow water to an adjacent watercourse was deemed un-
acceptable because of build-up of sulfates, and a pond recycle
stream is shown in the present design. Additional acreage for
increased solids disposal would have to be available otherwise
expensive transportation schemes would be required. The additional
disposal pond area specified was pro-rated directly from the TVA
report.
4. Process Appraisal
The limestone-wet scrubbing process is one of the most
advanced SC>2 control processes in that it has been demonstrated
in several pilot plants and commercial installations. It has
the obvious attractiveness of being a "throwaway process", and
the additional advantage of controlling particulates emission
without installation of an expensive precipitator. Based on the
TVA conceptual design study, the investment for limestone-wet
scrubbing process equipment should be significantly lower than
the regenerable S02 control processes. There is ample incentive
for continued funding of development programs, process optimiza-
tion studies and demonstration plant testing of this process.
The major disadvantage of limestone-wet scrubbing is the
problem of water pollution and solids disposal associated with
a throwaway process. There are other serious problems involving
equipment operability which have been indicated from the limited
information published from the Combustion Engineering demonstra-
tion plants. The problems have involved equipment fouling and
scaling, with the problem areas described in greater detail
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below. Suggested research and development areas to improve the
limestone-wet scrubbing process technology are also incorporated
into the following discussion.
a) Potential Problem Areas
(1) Boiler Injection
The flowsheet chosen for evaluation shows boiler injection
of the limestone additive as this concept has been demonstrated
on a commercial scale. This process scheme raises obvious concern
about the effect of the additive on boiler operability, e.g.,
effect on firing pattern, slagging, deposition, heat efficiency.
Deposition problems were experienced initially during the TVA
tests of the dry limestone process in the Shawnee unit and at
other test installations. Attempts at generalized predictions
of limestone-coal ash deposition tendencies have been unsuccessful,
so this must be considered a potential problem area in a new
installation. The operating problems may be controllable by use
of additional soot blowing capacity.
One means of avoiding the boiler injection problems is
by limestone addition directly into the scrubber circuit. This
scheme represented the alternate case evaluated in the TVA design
study. The j investment for the alternate case was higher mainly
because of the greater limestone handling requirements. That is,
higher stone injection rates were specified for the alternate
case (130% stoichiometric vs. 110% base case) because of the
lower reaction rates observed in pilot studies, possibly due to a
slower rate of solution with limestone than lime. There are ways
of improving this kinetic limitation, e.g., finer grinding of
raw stone, increased scrubber residence time. It is felt that
the concept of limestone injection into the scrubber system could
well represent the optimum process configuration and merits addi-
tional test inj.
(2) Scrubber Design Optimization
The other major incentive for use of a wet scrubbing process
is the cost reduction possible by eliminating an electrostatic
precipitator. This presumes that better than 99% (by weight) of
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the particles can be removed in a wet scrubber. Such claims
have been made by various scrubber vendors and have been- achieved
in pilot plant tests of a mobile-bed scrubber (Hydrofilter or TCA
Contactor).
There are several areas of concern relating to the scrubber
design. Firstly, 99 weight percent removal of particles with a
typical fly ash size distribution would still leave 10-30% of the
smallest particles (<2y size) in the gas which could create
problems with the visible plume appearance. Secondly, although
mobile-bed scrubbers have performed well on a small scale, there
could be flow distribution problems in scale-up to commercial
power plant units. Such problems have already been encountered
with the Hydrofilter (flooded glass marble bed) in Combustion
Engineering's commercial operation, and could also occur with
the UOP TCA Contactor specified in the TVA design. Considera-
tion should be given to testing alternate scrubber designs which
have a better established scale-up basis, e.g., venturi devices
possibly including staging, or tray-type contactors with or
without impingement baffles.
The other potential scrubber operating problem is solids
deposition in the circuit. Major plugging problems have been
encountered in the Combustion Engineering process in the region
of the scrubber gas inlet. It has been postulated that the in-
crease in pH occurring at the point of lime addition will tend
to precipitate calcium sulfite and sulfate. This would indicate
another advantage for injection into the slurry circuit, pre-
ferably between the scrubber and the hold tank, which would
ideally cause precipitation to take place on existing crystals
in the hold tank. This operating technique permitted the ICI-
Howden process to run with a minimum of plugging problems, but
the factors controlling crystallization rates need further study.
(3) Stack Gas Reheat Design Optimization
Various direct and indirect heat exchange schemes were
studied by TVA to select the optimum system for stack gas reheat.
The system selected was a cyclic liquid (hot water possibly con-
taining glycol) indirect heat exchange system to recover the
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stack gas heat. The heat exchanger costs were derived by TVA
assuming carbon steel tubes for the gas cooler and reheater.
This assumption may be too optimistic considering the presence
of free CC>2 and S02 in the gas which would result in the forma-
tion of weak acids as the dew point is approached in these units.
Upgrading the tube material to a stainless alloy could increase
the exchanger investment several-fold, and make it more attractive
to consider direct fuel combustion (oil or gas) ahead of the
stack as the source of heat. The TVA further assumed that 250°F
was the reheat temperature necessary to achieve the desired
ground-level SC>2 concentration, which is also determined by
other variables. For example, use of a taller stack and lower
reheat temperature could favor selection of a direct combustion
reheat system.
The large finned-tube exchangers specified in the proposed
process design also present several operating problems, i.e.,
gas distribution across the bundle, fouling by the fly ash and
reaction products. Some pilot plant experience has suggested
that deposits could be removed by intermittent water washing
using boiler condensate. The other alternative would be to
specify expensive soot blowing equipment. It seems obvious that
a process optimization study of means to provide gas reheat
would be justified.
b) Research and Development Areas
The preceding discussion cites a number of potential problem
areas in the limestone-wet scrubbing process which warrant further
study. Specific research and development programs to satisfy
these needs are summarized below. These suggested areas of study
are weighted to provide data on the scrubber system injection
process, which is judged to have the greatest chance of technical
success at this time.
(1) Kinetic data on the sulfation reaction step with
addition of either CaC03 or CaO to solutions of various total
solids content.
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(2) Kinetic studies on crystallization of CaS03
and CaS04 from supersaturated solutions in the presence of other
solids.
(3) Tower design data on particulates and SC>2
removal efficiency for various types of contactors.
(4) Studies of plume dispersion as a function of
stack gas temperature, velocity, stack height and atmospheric
conditions.
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5. Mechanical Review
a) Raw Limestone Storage and Handling
The storage capacity provided for raw limestone is 60 hours
at plant capacity. This is relatively small and it is suggested
that a minimum of one week's storage be provided. This, would
insure against equipment breakdown at the source of supply,
transportation delays, strikes, etc. It can be a supplemental
pile used in conjunction with active storage in a silo, or it
can be a working pile.
One week's storage would require approximately 2000 tons
for a 200 MW plant and 10,000 tons for a 1000 MW plant. These
large tonnages are most economically stored in open piles, stacked
by a belt conveyor. If space for piles is unavailable, silos
can be used, but they are expensive.
The added storage will increase the cost of the limestone
system considerably, and this should be provided for in any
cost estimates for the process. The major increase will be for
the more extensive conveying and elevatir.c systems, reclaim
tunnels and feeders.
When the limestone contains fines, freezing will occur
during winter months in lorthern olimatr Th i ~ >m .' reauire
car-thawing facilities and may also dictate that the stone be
stored under cover or in silos. If this is not possible, then
special precautions are needed to prevent the plug-up of reclaim-
ing feeders with large, frozen masses of limestone.
Another factor to be considered in space and cost require-
ments for the larger plants is the method of delivery. A 1000 MW
plant will use approximately 10,000 tons per week, or 2000 tons
per day, on a 5-day basis. This will probably require rail
delivery with extensive trackage and a rail-car unloading hopper.
If coal is received by rail, the coal unloading facilities should
be designed to handle the limestone also.
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On plants larger than 1000 MW, the limestone might be
delivered on a unit-train basis. A 2000 MW plant would use
2 unit train loads per week. The R.R. trackage, high-rate
unloading facilities and storage of stone then becomes a
major materials handling problem. The unit train concept may
make it possible to economically install this process even
though the limestone supply is distant from the power plant.
It also raises the interesting possibility of reloading the
train with waste for disposal at a distant point.
b) Dust Collection
Dust suppression and dust collecting equipment are not
mentioned, except for the items integral to the grinding circuit.
The limestone unloading station may require a dust suppression
system and all conveyor loading and transfer points should have
dust hoods vented through a bag filter.
c) Ground Limestone Injection
The ground limestone feed bin may not discharge in a reliable
manner due to arching and rat-holing of the material above the
discharge opening. It should be provided with an activated bin
bottom to assure a steady feed to the injection system.
The method used for dividing six streams to the injection
nozzles is questionable. A single volumetric feeder meters the
total amount into an air stream to a flow-splitting tank. This
tank theoretically divides the stream into six equal flows.
Details of the tank are lacking, but it appears to be a single
vessel with one inlet and six discharge's.
Much closer control over the solids flow rate to each
nozzle would be obtained by providng an individual feeder for
each, with individual conveying lines. This would give the
boiler operator the added advantage of individual control if
he desired to vary the rate to any boiler zone.
The above comments apply only to boilers fired by pulverized
fuels, with multiple injection points. Cyclone-fired boilers
were not studied in this report.
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^ Wet-Grinding Facilities
The limestone wet grinding facilities would be improved
if the feed bin was elevated to feed the mill directly. The
weigh feeder will then discharge to the mill, eliminating the
bucket elevator and its corresponding time lag. Also, the
elevator is a high-maintenance item and its removal eliminates
the mill's dependency upon the elevator's operation.
If possible, the elevator feeding the bin should be replaced
by a belt conveyor. A conveyor offers better reliability with
less maintenance.
An alternate method should be investigated for closing the
wet-grinding circuit for the mill. The spiral classifier used
in the process is inflexible in its relation to the mill due to
the required slopes in the feed launder and the discharge chute.
If space limitations in a retrofit installation preclude the
use of this type of unit, wet-cyclones should be investigated.
They require much less space and have greater flexibility of
location.
e) Waste Storage
The w~. sto stor.iqe requirements are vory large - approximately
765, QUO Lojio JKT yoar for a lOUU MW plant. Since these solids
are handled as a slurry and contain materials other than fly ash,
there is a risk of ground water contamination. For example, if
dolomite is used the magnesium sulfate produced in the reaction
with S02 is soluble. Also, unreacted limestone will enter the
pond as calcium hydroxide, creating a pH problem. Existing plants
which had no trouble with fly ash disposal may require new ponds
lined with membrane or an impervious blanket.
More information is needed on the settling characteristics
of the solids. Cyclones located at the disposal site might reduce
the settling area required. They are portable and easily relocated
as the solids build up. They are particularly useful where space
for thickeners is not available. Many minerals beneficiation
plants utilize them to build dikes for containment of tailings.
-------
LIMESTONE - WET SCRUBBING PROCESS
B. CHEMISTRY
*• Qualitative Process Description
The SO from stack gas is transferred to the aqueous
phase where it is "fixed" by a very rapid reaction with
basic ions like OH" and HCOZ to form MSO" and SO* ions.
^ 33
Additional S0_ nay dissolve to form H-SO,. In a prac-
fe *• O
tical scrubber situation it is not possible to supply
enough of these basic ions in solution to react with all
of the SO in the gas. This is because the amount of
liquid is limited by flooding considerations and the con-
centration of basic species is controlled by the limited
solubilities of the alkaline earth hydroxides and car-
bonates (3). To circumvent these limitations a slurry
containing basic solids which will react and replenish;
the ions consumed by S0_ is usually fed to the scrubber.
The rate of dissolution of these solids is the limiting
chemical step in determining the efficiency of the
scrubbing process (4).
Most process concepts involve injection of limestone
or dolomite into the hot flue gas at a temperature high
enough to decompose the carbonates to form oxides. Any
SO that is reacted with the solid oxides before it reaches
the scrubber is effectively removed and need not be con-
sidered in the scrubbing process. This may account for
nearly 2*0% of the SO- removal (5"). The remaining oxides
and all of the fly ash are trapped by the scrubber liquid.
Inasmuch as the mechanism for dissolution of calcium
and magnesium oxides involves an initial slow hydration
step, residence time in the process is usually provided
by installation of a scrubber effluent holding tan);. Part
of the slurry from the holding tank, which contains the
-------
alkaline earth hydroxides in suspension, is then used
for the scrubber liquid feed. It is normally assumed
that these hydroxides will dissolve rapidly enough to
be used in the scrubber, but this is open to question
(1,2).
Alternate process concepts which involve addition
of uncalcined limestone, or dolomite to the scrubber de-
pend upon the dissolution of these compounds by a liquid
saturated with CO,,. Very little data have been gathered
to dcuonstrate this concept.
Thr IISOT and SOZ ions formed from the S0_ are further
.-> .> 2
reacted nlonp a number of pathways resulting in a final
transfer of nost of the sulfur to the solid phase which
is with'lrawn frori the system. Some of these pathways
are :
1. So! + Ca**—-+• CaSO_(s)
2. IISO! * (,;a*t—». CaS()..(s) + H*
c + *
3. ^0 + M" 'i » '!<'SO (s 1
:. " •;;*••'
4. JfSO! * Mg*- *• Mf;S03(s) + H*
5. a. SO" + 1/2 0, * SO"
*J •« ^T
b. Ca**+ S0° »• CaS04(s)
6. a. HSOl •»• 1/2 09 *- SO* + H*
C8S04(s)
-------
A snail amount of the sulfur remains in solution -
principally as S(K ion - and is removed as a purge
stream. This is particularly true with dolomitic
linestoncs where roost of the sulfate made by reactions
5a. and (>a. is removed as a solution of MgSO. (3).
Fly ash, which is caught by the scrubber, nay
partially dissolve and influence which of the pathways
River above is predominent in the overall process. The
influence nay be catalytic - such as the acceleration
of sulfite oxidation in the presence of Mn ions -
or in shifting solubilities by mass action or ionic
strength effects. In addition, chemical reactions in-
volving fly ash conponents nay lead to undesired solids -
like hydrous oxides and hydroxides - which cause trouble
in subsequent filtration operations. Since alternate
process corrects nay consider removal of fly ash prior
to scrubbing it should be pointed out that there are
possible advantages to their presence. Among these is
the reduction of scale formation by providing additional
surface for crystallization and a nild abrasive action
on process cquipnrnt.
other components of the flue gas are partially ab-
sorbed by the scrubber and influence the chemistry of
the scrubbing process. The large excess of carbon d iox id e
in the r,as reacts rapidly with Oil" and C()~ ions in the
•J
scrubber feed to form i!COl ions which are the principle
J
basic species present in the scrubber for reaction with
SO . Until all of the HCOl ions have been consumed by
the fi()n si buffer syster.i between HCOZ ions and C0_ keeps
t* 3 £
the p!! between 6 and 7. Oxygen in the j;as converts
U i .••>:; D 1 \~ i».i sulfite species to r.ulfate wVich changes thp
-------
pathway by which the sulfur js rencvcd from the system.
I!" this oxidation could be prevented the scrubber night
have a j; renter capacity - since the U ions liberated
by reaction 6a . consune IIC!0" ions - but the sensitivity
of oxi.hit ion to accidental catalysts r. r r> sent in the fly
ash nal'cs t '* i . .--. difficult. ,'.' i t rogen ox id cs arc partially
renovo.l I ;• the scrubber. The anount removed is probably
related to the ;:ru.ur.t of NO oxidation that has tal:en
place before the flue nas enters the scrubber. Uhen N'O^
-.1 I s sol r >" •• it fr-.rii.--. r. it rate i «- n s r.rnJ .-. mixture of nitrite
ions nnti nitrous ociu. The chemistry of subsequent
re- ;i c t i rr s l-otwccn tlie latter 2 .species, dissolved SO
and lisol ion in tlie presence of oxyycn is not well known,
but, since nitrite and sulfitc are l:nov:n not to be stable
in the presence of each other in acid solutions it is
usually nr-sun-ed that only *.'0l ror.iains (3). It is intcr-
e.'stir.;; f'.tt; expr r inent s jit Battolle (4) nppear to show
that ','fi enhances the capacity of a ciilciuni hydrate
X
solution Tor ''.()., absorption by 70% beyond theoretical,
•'•' • i i^ . i .' ' ! .i ui i 1; i :.i i I a t i > :i s
It has 1'ccn calculated that an ciiu .i 1 ibrium partial
pressure ol." S00 above a liquor of pi! 4 to 4.5 is less
than 1500 ppn - equivalent to VQl removal (3). In tl^e
scrubbing systen described tliis situation would occur
after reaction with SO had depleted all of the basic
species ir. solution and additional SO had dissolved to
forn HnSf)_. The maxinum concentration of basic species
t)i at can be obtained in solution in tho scrubber feed
is controlled by solubility relationships between calcium
and magnesium hydroxides and carbonates in the presence
of many other solids and dissolved species. Radian
Corporation (3) has developed a computer program to
-------
handle the c.onplex equilibrium calculations involved. In
general it appears that, if the scrubber feed is equilibrated
in the absence of CO- (out of contact with flue [;as) , the
more soluble Ca(01!)0 will dissolve to produce Oil" inns as
*•
th.c principal basic species. At the hi>;h pll of this
solution Mtf(Oll),, is insoluble. If insufficient C.aO is
A.
available as in a high magnesium dolomite, for cxanple,
the pll is rmch lower; a relatively hi?.Th concentration of
Mj;SO is found in the liquid; nnd the concentration of
dissolved basic species is lower.
It. has been assumed that precipitation of calcium
r.arbonatr ilocs not take place when the 01!" ions in the
scrubber feed are converted to HCOl ions by reaction with
CO., in the scrubber feed even though the solubility product
of Cai!0» nay be exceeded. This assumption seems reasonable
»J
in the practical model of a backmixed reactor.
j , I' ate Li n i tat i ons
As stated previously, the rate limiting step in the
scrubbing reaction is the dissolution of basic solids in
the scrubber to replace the species lost by reaction with
SO,,. If too little reaction occurs in a single pass through
-A.
the scrubber .provision must be riado to recycle these solids.
A iiattelle study concerning measurement of x'arious rates
involved in the wet-scrubbing process (4) indicates that
line in ^articulate form reacts readily with various sulfur
species or carbonate in solution to yield a coating which
inhibits utilization of the bulk of the lime. The sane
study recommends that further consideration be .r.iven to
the importance of particle size in the overall process and
to the possibility of pred i ssn 1 vi rr, all of the lime, or
limestone in feed water. Unfortunately, due to solubility
-------
linitations, tho latter roconnend;:tion would appear to
involve an excessively hi<*h ratio of liquid to pas rates,
By £vr the slowest reaction in the standard process
is the hydvntion of the line tal.cn fron the flue f^as by
the "rain" of scrubber effluent, A lnr;;e holding tank
i s usually provided to r;ive sufficient residence tir'c
for tliis roii c:t ion. Af'.iin the Hattclle study indicates
that d «»p «>••-. \ l i on of su'. fnte on the liv.ic hinders the hydra-
t .i on process. In tliir- rase the surface coatinf. nny hove
i't'ei; fori:.c-:U by exvosure to S((n in'ior to contact Kith the
hydratin;1. linuo1* or by rcnction betweon tbo line and
suj<";ito oj sulfifc in tf;e lifjuor.
In snnnary it ciprears that a subr.tnjiti.nl anount of
:;li>rry r^.v-clc v. J 1 1 !.<: rerujred in tlic standard process
O'lu-ci-t b ce ;'.us o <>'' '-rr.birMs in rcnctint1 tV.c line at a
reasonable r:ite. Alternate process concepts such as pre-
d i 330 lv i rj1. ihc line or us in/7, inu:alc i urd linestor.c have
difficulties due to r.t\u il ibr i UM solubility limitations.
It i:; not clear whether tho irclus.ic.n of fly ash is ham-
'"ul .,r '^MM't'iti i] to overall scrubber nrerntion.
-------
REFERENCES
(1) "Sulfur Oxide Removal from Power Plant Stack Gas - Use
of Limestone in Wet-Scrubbing Process", TVA Conceptual
Design Study (1969), Final Report on Contract No. TV-
29233A.
(2) Pollock, W. A., Tomany, J. P., and Frieling, G.,
Mech. Eng. 89^ (8), 21 (1967).
(3) "A Theoretical Description of the Limestone Injection-Wet
Scrubbing Process", Radian Corp., Austin, Texas (June 1970)
Final Report on Contract No. CPA 22-69-138, Vol. I.
(4) "Investigation of the Limestone-SO2 Wet Scrubbing Process",
Battelle Memorial Institute, Columbus, Ohio, November 1969,
Final Report on Contract No. PH 86-68-84 Task Order No. 17.
(5) "Sulfur Oxide Removal from Power Plant Stack Gas-Sorption
by Limestone or Lime-Dry Process", TVA Conceptual Design
Study (19h8).
-------
BOILERS(2)
lAPACITY EACH
•~l
143°F,—. 2SO°F
I25i L8/HR SOLIDS
c-2-A.B c
43.15 MMBTU/H8K.
220,000 FTl ,'"'
If ISI LB/HR SOLIDS
(o.o^ GRAINS/SCF)
P
' M" 1
C
FLUE «*s
BOOSTER FANS
->-2-A,B,C
2350 BMP ")
(iWCREMENTAL)tEACl
J-fr
50 BMP
1975 6PM
2-INCCEWENTAL P(5ND
2,050 ACfJE-FEEr.
ISSUE DATE 3-Z9-7I
ADO QUANTITIES
"•I-27-7
THE M. W. KELLOGG COMPANY
STEAM POWER PLANT SULFUR
OXIDE CONTROL
WET LIMESTONE PROCESS
N EW ipg_Q_M EC AW ATT IN STALL>TI Q N
4092 -5O PD-IOS-D
-------
DRY LIMESTONE INJECTION PROCESS
A. PROCESS SECTION
1. Process Description (see MWK Dwg. No. PD-107-D)
The dry limestone injection process involves the removal of
S02 from stack gas by vapor-phase reactions which occur at high
temperatures (~2000°F) in the boiler. The following overall
reactions describe the calcination and subsequent sulfation steps:
CaCO3 -*• CaO + C02
CaO + S02 + 1/2 02 ->• CaS04
The chemistry and kinetics of the process are discussed in more
detail in the Chemistry Section of this write-up.
Limestoii'! rock from a nearby quarry is transported by truck,
unloaded into a receiving hopper and conveyed to the storage silo,
which has been sized for 60 hours limestone capacity. The raw
stone is then conveyed to a surge hopper and fed to a pulverizer
(Ring-roll mill or Ball mill), where it is mixed with 610°F air
for drying and classification. The air-limestone mixture is
transported to a feed tank where the air is separated and vented
through a bag filter. The solids are fed to a flow splitter vessel
containing six outlet nozzles provided with orifice plates. In-
jection air is supplied to each of the six outlet lines by individual
compressors. The air-limestone mixture will be injected into the
boiler through six wall-mounted nozzles. The nozzles should be
tiltable to vary the angle of injection and should be located at
several elevations. These design features are necessary as different
limestones will have different optimum injection temperatures.
-------
The boiler stack gas downstream of the air preheater will
contain fly ash, unreacted limestone, and the solid reaction
products (CaO, CaSC>4) . This gas (300°F) will be fed to a mech-
anical separator to remove the bulk of the entrained solids.
The design proposed for this separator is a multiple cyclone
collector, consisting of about 1,100 - 10 inch diameter cyclone
separators mounted in a tube sheet in a single rectangular
housing. The gas from the mechanical separator will be ducted
directly into an electrostatic precipitator wherein the design
stack gas fly ash emission limit must be achieved. This unit
will be about 3.0 times the size of a standard precipitator sized
for fly ash removal only, because of the lower particle drift
velocity and collection efficiency in the presence of limestone
additive. The precipitator outlet gas is sent to the stack via
an induced draft fan.
The fly ash-limestone reaction product mixture collected
in the mechanical and electrostatic separators is conveyed
either mechanically or pneumatically to a central sump, wherein
it is mixed with sluice water. The solids mixture is then
transported to the disposal pond as a 12 wt % solids slurry. It
is proposed to recycle the pond overflow water to avoid contamina-
tion of adjacent streams with sulfates and sulfites. Make-up
process water is added to the sump to compensate for pond evap-
oration losses and required water of hydration of CaO and CaS04.
2 . Design Bases
The Kellogg flowsheet developed for the dry limestone
injection process is for a 350 megawatt installation. This
size limitation was established mainly as a judgment factor
considering problems inherent in the process of achieving
satisfactory limestone distribution and reaction across a large
boiler section. Such problems have been quite evident in the
TVA full-scale limestone injection tests at Shawnee Unit No.
10 (1) which is rated at 175 megawatts but generally run at
lower capacity. To date, the Shawnee tests have not been able
* Refers to references at end of section
-------
to achieve the same degree of CaO utilization when operating at
conditions similar to. pilot plant reactor tests, presumably due
to problems of flow distribution and dead-burning. It does not
seem justifiable to propose use of the dry limestone injection
process in a new 1,000-MW installation considering its present
state of development.
The following tabulation gives the specific process and
equipment design bases incorporated in the Kellogg evaluation:
a. Process Design Bases
(1) 350 MW Power Plant assumed as largest feasible
size for Dry Limestone Injection process. Use
single Boiler and single Air Preheater, single
Stack.
(2) 100% Sulfur in coal evolves in Stack Gas as
S02 + 203. 71.6% Ash in coal evolves as Fly
Ash.
(3) Thermal Deficit due to Boiler reactions involving
Limestone and S02 is 26.04 MM Btu/hr.
Equivalent Coal Firing Rate = 2,174 Ibs/hr
(0.83% of power plant feed rate).
(4) Limestone composition: 91.8 wt % CaC03 (dry basis)
10.0 wt % moisture
(5) Limestone addition rate = 200% Stoichiometric
100% Conversion of CaCC>3 to CaO in Boiler
25% Conversion of CaO to CaSO4 in Boiler
Maximum S02 Removal = 50%
b. Equipment Design Bases
(1) Limestone Feed System (TVA Design Bases)
Unloading Hopper; 90 tons capacity (3 hr supply)
Storage Silo; 1800 tons capacity (60 hr supply)
Surge Bin; 240 tons capacity (8 hr supply)
Pulverizer Use Ring-roll mill or Ball Mill
Limestone Feed Rate = 30 tons/hr
Design for 99% minus 325 mesh (44y)
-------
(2) Boiler Injection System (TVA Design Bases)
Classifying Air = 0.052 Ibs/lb Limestone
(vented thru Bag Filter)
Injection Air = 0.475 Ibs/lb Limestone
Use Flow Splitter Tank with 6 outlet nozzles
equipped with orifices to distribute limestone
to 6 boiler injection points.
Injection air added downstream of Flow Splitter
by individual Compressors (6). Air injection
velocity = 300 ft/sec through til.table nozzles.
(3) Stack Gas Dust Removal System
Maintain TVA design emission level of 0.10
grains/SCF. Required overall solids removal
efficiency is 99.22 wt %.
Specify Mechanical - Electrostatic Collector
combination.
Mechanical Collector:
Specify multiple cyclone design with approx.
1,100 - 10 inch diameter tubes in single
rectangular housing.
Use cyclone efficiency of 7^% per TVA design
- assume same particle size distribution as
Fly Ash without additive.
Inlet Dust Loading = 13.2 grains/SCF
Outlet Dust Loading = 3.82 grains/SCF
Collector Design AP = 3.5 inches H2O
Electrostatic Precipitator:
Design Efficiency = 97.33% (higher than TVA
design of 96.1% because of higher Ash Coal in
current design, i.e., 15.2% vs. 12%).
Inlet Dust Loading =3.82 grains/SCF
Outlet Dust Loading = 0.10 grains/SCF
Design ESP size = 3.1 x Standard ESP size
required for Fly Ash only
Design Gas residence time = 7.2 seconds
Design Gas velocity - 6.5 ft/sec
-------
(4) Solids Disposal System (TVA Design Bases)
12 wt % solids concentration in sluice, water
lines to disposal pond.
Pond evaporation loss = 217,000 Ibs/hr water
Process make-up water = evaporation loss +
water of hydration (CaO -»• Ca(OH)2 and
CaS04 -> CaSO4'2H20)
Incremental disposal pond area of 105 Acre-ft/
year to be developed to handle the additional
solids re.-ulLing from limestone reaction products.
Pond overflow water must be recycled to Sluice
Sump to avoid water contamination from sulfates.
3. Design Rationale
In this section, a more do hailed discussion is given to support
the reasons behind the specific process and equipment design bases
incorporated into the Kellogg dry limestone process evaluation. The
discussion is sub-divided into the various systems of the process
flow sheet.
a. Limestone Feed System
The design bases for the limestone feed system were taken
directly from the TVA conceptual design study (2). Suggested
improvements on this system have already been outlined in the
Kellogg evaluation of the wet limestone scrubbing process. The
main criticism is that 60 hours limestone storage seems inade-
quate for a large installation, and an on-site storage pile with
mechanical conveying equipment would appear more desirable.
However, this alternate design basis would increase the process
investment, and consequently was riot included.
The TVA design basis for limestone particle size reduction
(70% minus 200 mesh) was increased to 99% minus 325 mesh. This
change was made to achieve 25% CaO utilization in the boiler
based on pilot plant studies which show increased conversion at
reduced particle sizes. The TVA study states that size reduction
to 99% minus 325 mesh is still feasible with a ring-roll mill
or ball mill. The TVA design bases were used for specifying
-------
drying air and injection air rates. The drying air quantity is
very low, capable of evaporating only 5% of the water in the raw
limestone feed. A much better system would appear to be to use
the proper amount of drying air in the pulverizer for complete
vaporization, and vent this larger air flow to the boiler rather
than to atmosphere.
b. Boiler Injection. System
The raw limestone injection rate was set at 200% stoichio-
metric in this process evaluation. This is consistent with the
TVA design basis, and also represents the maximum feasible addi-
tive rate consistent with the greatly increased solids handling
requirements of the dust collection system. The 200% stoichio-
metric injection rate combined with a 25% maximum CaO utilization
yields a design SC>2 removal of only 50%, which is an inherent
limitation of the dry limestone process.
To achieve a higher overall S02 removal would involve use
of still higher additive injection rates. In addition to increased
solids handling problems, this would raise questions of effects
on boiler operability; e.g., increased slagging or deposition
tendencies, increased erosion of convection-pass tubes, effects
on burner firing patterns. None of these possible adverse
effects were factored into the present evaluation. There is
however, a net thermal deficit due to boiler reactions involving
limestone. This deficit is 26.04 million Btu/hr in the 350-MW
design case, and the equivalent additional coal firing rate
(2,174 Ibs/hr) is included in the process utility requirements.
c. Dust Collection System
A mechanical-electrostatic collector combination was specified
to maintain the design stack gas solids emission level used in the
TVA study, viz, 0.10 grains/SCF. The repaired overall solids
removal efficiency to meet this level with 200% limestone injection
is 99.22 weight percent. After consultation with vendors, a
mechanical separator design involving use of multiple cyclone
-------
collector tubes was selected as the optimum unit to precede a
precipitator. Vendor designs were obtained for the specific
dust loadings involved, however the TVA mechanical separator
design efficiency of 71 wt % was used rather than higher values
(75-85%) claimed by vendors. This allows some design safety
factor deemed necessary since the actual particle size distri-
bution is not known for the fly ash-reaction products mixture.
The electrostatic precipitator (ESP) design basis was es-
sentially the same as used by the TVA which resulted in an ESP
size 2.8 x the standard ESP required for fly ash removal alone.
The current design involves higher solids loadings because of
the greater ash content of the coal; i.e., 15.2% Ash versus
12.0% Ash in TVA design. The increased loading was factored
into the ESP collection efficiency equation using the same
particle drift velocity as the TVA design case with additive.
This procedure resulted in an ESP size for the current 350-MW
design case of 3.1 x the standard ESP unit.
d. Solids Disposal System
The design basis for the solids disposal system was also
taken directly from the TVA study and a 12 wt % solids slurry
concentration in the lines to the disposal pond was specified.
There are some questions of operability and potential plugging
with this system due to hydration of reaction products (CaO,
CaSO4). The TVA design assumed^ the pond overflow water could
be sent to an adjacent watercourse; however, the current design
includes the cost of a recycle system to avoid the question of
water pollution. The best answers to questions associated with
the solids disposal system should be obtained from full-scale
tests at Shawnee.
4. Process Appraisal
The dry limestone injection process is one of the more
advanced SO2 control processes in that it has undergone extensive
laboratory and pilot plant studies over the past five years, and
is currently being subjected to full-scale boiler tests. Un-
fortunately, the results of dry limestone process tests have been
-------
discouraging, in that 50% SC>2 removal is the best attainable
under optimized conditions in a standard pulverized coal-fired
boiler. Extensive evaluation of pilot test results by TVA and
more recently by Kellogg has shown that 20-25% CaO utilization
is the maximum obtainable with the residence time restrictions
(2.0 seconds maximum) in a standard B&W or C-E boiler. Attempts
to achieve higher than 50% S02 removal with this restriction
would involve the use of unreasonable limestone addition rates
(>200% stoichiometric) . It could be possible to realize 90% SC>2
removal at higher gas residence times attainable with a fluid-bed
boiler design. This would represent a major development in boiler
design practice which is considered beyond the scope of the present
evaluation.
The applicability of the dry limestone injection process
would appear to be limited to power plant situations wherein
50% S02 removal, combined with stack height and atmospheric
conditions, would be able to meet local ground-level S02 con-
centration standards. This would appear to limit the process
attractiveness to smaller units, either new installations or
retrofits (assuming availability of space). The TVA full-scale
tests at Shawnee have uncovered a number of potentially serious
operating problems which must be overcome. The problem areas
are cited below with possible areas for further study indicated.
(a) Limestone Injection
A major limitation of the process is the ability to achieve
uniform limestone distribution across the entire cross-section of
a large boiler. This problem has been experienced at Shawnee in
the 175-MW boiler tests. As a point of reference, the TVA design
study for a 200-MW unit would require a boiler cross-section of
24 ft x 58 ft at the burner level. Performance data are being
obtained at Shawnee on the effect of injection velocity and
injection angle which should be useful. In addition, APCO is
funding simulator studies of the air-solids injection system.
It is Kellogg's judgment at the present state of development,
-------
that the dry limestone process should be restricted to a 350-MW
maximum boiler capacity. This conclusion is not applicable for
a wet limestone process installation wherein lower CaO conversions
in boiler can be compensated for by higher conversions in the
scrubbing system.
(b) Boiler Operability
There are several potential questions concerning the effect
of limestone injection into the fire box on boiler performance.
One problem area is the effect of the additional solids on the
stability of the burner firing pattern. Much difficulty has
been experienced in the control of the boiler system at Shawnee,
which may have been aggravated during periods of limestone addi-
tion.
Another major problem area experienced during initial tests
at Shawnee is the build-up of reaction products on reheater or
superheat tubes. There seems to be wide variation in the extent
of this problem experienced at various installations involved in
limestone injection tests. Babcock and Wilcox have studied the
various factors that would control deposition, e.g., ash fusion
temperature, viscosity, sintering strength. Their final report (3)
unfortunately, contains only the most general conclusions and does
not provide quantitative correlations for predicting limestone-coal
ash deposition tendencies. The implication presumably is that any
particular coal ash-additive combination being contemplated must
actually be studied in the field.
(c) Dust Collection
It is apparent that one of the major cost penalties incurred
with the dry limestone injection process will be the substantially
larger ESP unit required, i.e., 3.1 x standard ESP for present
design. There is much scatter in available information on the
effect of limestone additive on dust resistivity, particle migra-
tion velocity and collection efficiency. Early tests on an older
-------
ESP unit at Detroit Edison showed a severe decrease in efficiency
with additive (80% •*• 55%), corresponding to a factor of.3.7 x base
ESP size (2). B&W pilot plant test results (3) predict an ESP
size increase with limestone of 2.5 x base case. Recent TVA tests
carried out at Shawnee Unit #10 show an average decrease in ESP
efficiency of only ten percentage points (4) (92% •> 82%). This
would represent an increase in ESP size to 1.5 x base case; how-
ever there are considerable scatter in the Shawnee test data.
It has been reported (5) that there is a critical dust resisti-
vity of 10 ohm-cm above which the migration velocity decreases
rapidly and the ESP collection efficiency decreases. It has been
further postulated that resistivity is strongly influenced by the
free 803 concentration in stack gas, and this level is reduced to
virtually zero in the presence of excess CaO from limestone cal-
cination. One suggested means of controlling the resistivity value
is through the use of gas "conditioners"; e.g., S03, NH3. There
has been some successful experience involving this technique, and
it appears worthy of investigation.
Another approach to achieving high ESP efficiency in the
presence of additive is to take advantage of the lowered resist-
ivity at higher operating temperatures (>300°F). Con Ed's
Ravenswood plant has demonstrated that better than 99% efficiency
can be realized by running a "hot" (750°F) ESP unit (6). Such an
installation would require breaking into the boiler casing after
the superheater with new ductwork to the ESP. This alternate
deserves serious study as the cost of the additional ductwork may
well be less than a larger ESP unit.
-------
REFERENCES
(1) J. T. Reese, TVA Quarterly Progress Report, October-December
1970, Dry Limestone Injection Project 2438 (February 19, 1971).
(2) "Sulfur Oxide Removal From Power Plant Stack Gas - Sorption by
Limestone or Dry Lime Process", TVA Conceptual Design Study (1968)
(3) R. C. Attig and P. Sedor, "Additive Injection For Sulfur Dioxide
Control - A Pilot Plant Study", The Babcock and Wilcox Company,
Research Center Report No. 5460 (March 27, 1970).
(4) TVA Results Report No. 62 - "Electrostatic Fly Ash Collector
Performance With Limestone Injection, Shawnee Steam Plant Unit
No. 10 Precipitator A", (June 9-July 15, 1970).
(5) A. B. Walker and R. F. Brown, "Effects of Boiler Flue Gas
Desulfurization - Dry Additives on the Design of Particulate
Emission Control Systems", Dry Limestone Injection Process
Symposium (June 22-25, 1970).
(6) J. P. Carey and R. G. Ramsdell, Jr., "Ravenswood Conversion
to Coal", Vol. 29, Proc. Amer. Power Conference, 495 (1967).
-------
DRY LIMESTONE INJECTION
B. CHEMISTRY
1. Qualitative Process Description
As normally conceived the limestone injection process
involves addition of calcium carbonate (limestone) or a
mixture of calcium and magnesium carbonates (dolomite)
at a point where the temperature is high enough to de-
compose carbonates and form oxides. During a very few
seconds of co-current travel with the stack gas the solid
oxide particles react with S02 and oxygen to form calcium
(or magnesium) sulfate. Depending upon the point of
addition of the limestone and other.react ion conditions
there may also be formation of calcium sulfide and cal-
cium sulfite. The solid products are removed along with
fly ash and unreacted additives by an electrostatic pre-
cipitator or other dust remover.
Chemical reactions between the lime and fly ash com-
ponents - like silica - decrease the availability of the
additive. Incomplete decomposition of the carbonates or
recombination of C02 with the oxide in a lower temperature
zone also lowers the oxide availability. Probably the
most serious interference with the course of the reaction,
however, comes from the reaction products themselves.
As calcium sulfate is formed on the surface of the lime
particle it creates a layer through which SCU and oxygen
-------
must diffuse to the "decreasing core". Generally, be-
cause of these and other factors, the actual lime
utilization seldom exceeds 30%.
2. Equilibrium Considerations
A Battelle study (1) has examined the chemical thermo-
dynamics of the various species present in the dry lime-
stone process. Several conclusions can be reached on
the basis of this examination.
1. High temperatures favor the dissociation of sulfates
and sulfites such that CaO is incapable of removing
any SC>2 from a typical flue gas above 2250°F ai.d MgO
is incapable of removing S02 above 1550°F.
2. The tendency of S02 to react with CaO to form CaS03
is completely overshadowed by the much more favorable
equilibrium, in the presence of oxygen, with
The equilibrium constant for the Oxidation of
to CaSO* over the temperature range of interest ex-
ceeds 105.
3. With CaO at unit activity (pure solid) it is possible
to remove all but 1 ppm of S02 from flue ga: at
1770°F. CaO in solid solution will leave somewhat
x->re S02 but the difference would probably not be
large. Lower temperatures are, naturally, even more
favorable from an equilibrium standpoint.
-------
4. With MgO at unit activity the partial pressure of
S02 is equivalent to 1 ppm at 1200°F.
5. Limestone theoretically converts to CaO in boiler
atmospheres at about 1415°F; at lower temperatures
C02 in the flue gas will tend to reconvert CaO to
CaC03.
6. On an equilibrium basis, competition between S02
and C02 for reaction with CaO favors the S02 reaction.
Theoretically uncalcined limestone should be capable
of desulfurizing flue gas since CaC03 can be in
equilibrium with well under 1 ppm S02 at 1400°F.
7. This competition between S02 and silica from the
fly ash is somewhat less favorable in that 0.14 to
3.0 ppm S02 exist in equilibrium with a mixture of
calcium sulfate and silica or silicates at 1400°F.
At higher temperatures, say 1800°F, the equilibrium
is shifted such that no S02 removal can be expected
by calcium silicates.
In summary, there seems to be no significant equilibrium
limitations to reacting S02 with limestone in -, flue gas
atmosphere if the temperature is kept below about 2000°F.
3 . Kinetic Considerations
The rate of reaction of S02 and oxygen with solid
calcium or magnesium oxide particles under conditions
existing in the hot zone may be limited by mass transfer
-------
through the gas film or by the chemical reactivity of
the particle.
Mass transfer characteristics may be altered by using
a different size particle or changing the gas and particle
velocities but there is very little room for maneuvering
in these directions. If mass transfer limits limestone
utilization to about 30% or less (as has been estimated
in theoretical calculations) the only hope for improvement
would be to use something like a fluidized bed to in-
crease the residence time of the particle.
The chemical reactivity of the particle is affected
by a large number of parameters, of which the two most
important are temperature and sulfate loading. Kinetic
studies by Borgwardt (2) and Coutant (3) show a severe
inhibiting effect due to sulfate loading that appears to
limit the limestone utilization obtainable in a few seconds
residence time to 20% or less. Increasing temperature has
been shown to increase limestone utilization up to the
point where back reaction due to decomposition of calcium
sulfate becomes important (about 2100°F) but, even here,
a 20 to 25% utilization seems to be the most that can be
achieved. Again the only hope of achieving higher con-
versions is by greatly increasing solids residence time.
There are many ways of achieving even less utilization.
"Dead burning" of the lime by injection at a point where
-------
the temperature is too high - like the firebox - is
well known as a cause for decreased reactivity. The
high temperatures cause a sintering of the surface re-
sulting in greatly decreased porosity and surface area.
At the other extreme, incomplete calcination has been
shown to be a cause of decreased reactivity when the
temperature is below 1600°F even though the equilibrium
is very favorable. Precalcination has been shown to ex-
hibit a negative effect on reaction rate when carried
out for an extended time period (2) and when carried out
in 0.3 seconds (3). It has recently been argued, on
theoretical grounds, that precalcination at relatively
low temperatures in the presence of steam and use at low
temperatures where some Ca(OH)2 might exist should aid
the reaction kientics (4). It is hard to see how this
technique would avoid the inhibitory effect of sulfate
loading.
Summarizing, the kinetics of the reaction involved
in the dry limestone process are too slow to permit
efficient utilization without means of providing extra
residence time for the solids.
-------
REFERENCES
Final Report on Fundamental Study of Sulfur Fixation
by Lime and Magnesia by Battelle Memorial Institute,
Contract No. PH86-66-108, June 30, 1966.
(2) Borgwardt, Robert H., Environmental Science and
Technology, £f 59 (1970).
(3) Coutant, R.W.; Barrett, R.E.; Simon, R; Cambell, B.E.;
and Lougher, E.H.; Summary Report on Investigation of
the Reactivity of Limestone and Dolomite for Capturing
S0'2 from Flue Gas, Contract No. PH86-67-115, November
20, 1970.
(4) Meyer, Beat and Carlson, C., The Reaction of S02 with
Metal Oxides, paper presented at the Conference cTF
the American Institute of Chemical Engineers,
Chicago, 1970.
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THE MONSANTO CAT-OX PROCESS
A. PROCESS
1. Process Description (See MWK Dwg No. PD-100-D)
The Monsanto Cat-Ox Process removes S02 from flue
gases by catalytically oxidizing the S02 to SO3. The 803 com
bines with H2O and is condensed and recovered in an absorber
using circulating sulfuric acid.
Flue gas leaving the boiler is passed to an Electro-
static Precipitator where >99.5% of the fly ash is removed.
The precipitator runs "hot" as the gas to the converter is
set at 900°F to achieve the 90% conversion to 803 at a
reasonable space velocity. The catalyst used is V^C^. Thero
is little or no temperature change across the converter.
Gases leaving the converter flow to an Economizer where they
are cooled by exchange against boiler feed water, and thence
to an Air Preheater where they are cooled further by exchange
against combustion air. The outlet flue gas temperature is
held to 450°F to avoid condensation of sulfuric acid and sub-
sequent corrosion of the Preheater elements.
Flue gas leaving the Preheater is ducted to an Absorber
where sulfuric acid vapor is condensed from the flue gas by
contact with cold sulfuric acid in a packed bed. Product
acid after being cooled to 200°F with recirculating acid is
further cooled to 100°F before being sent to storage. Flue
gas, together with entrained acid droplets or nist, leave
the Absorber at 225°F and enter a Brink Mist Eliminator where
most of the entering acid is recovered. Gases leaving the
Brink Mist Eliminator are routed to the stack via an induced
draft fan.
-------
The circulating sulfuric acid is heated from 2000F
to 270°F while contacting the hot flue gas and condensing
acid vapor. It has been found economical to recover this
low level heat by exchange against combustion air and boiler
feed water even though some thermal efficiency is lost
through the use of an intermediate recirculating heat trans-
fer fluid. The intermediate fluid system is felt to be
necessary to avoid the possibility of sulfuric acid leaking
into other systems.
Fine fly ash particles which were not caught by the
electrostatic precipitator or the converter appear either
in the product acid or in the Mist Eliminator. When the
build-up of pressure drop because of entrained fly ash be-
comes intolerable the Mist Eliminator can be back-washed
with acid while on stream.
2. Design Basis
The following conditions were established by
Monsanto and used in this study:
Converter Temperature 900°F
Flue gas from Economizer 650°F
Flue gas from Air Preheater 45QOF
Flue gas from Absorber 225°F
Acid Circulation Temperature 200/270°F
Acid Product Storage 100°F
SO2 Conversion 90%
Absorber Packing 2" Intalox Saddles
Absorber Ap 0.5" H2
-------
3. Process Design Rationale
The Electrostatic Precipitator, Converter and Brink
Mist Eliminator are proprietary equipment items and will
not be discussed. However, it should be mentioned that
the Precipitator runs about twice the absolute temperature
as is customary and will therefore have to handle twice the
gas volume. This, together with the need to design for a
higher temperature, per se, plus the higher efficiency
needed should result in a much greater than average cost
for this item.
Except for the Absorber, the remaining major equip-
ment items are heat exchangers of one type or another. Ob-
viously the temperature differences will directly affect
the size of this equipment, but in no case are all the
terminal temperatures specified. A number of judgments and
assumptions therefore had to be made in sizing this equip-
ment.
C-l Economizer
In this item, 900°F flue gas is exchanged against
boiler feed water. To establish the BFW flow and tempera-
ture level it was assumed that the plant had an overall
water rate of 5.5 f/KW, that the steam pressure was approxi-
mately 2400 PSIG and, because of the high flue gas heat
level, that the Economizer heated the BFW close to its
saturation temperature. The inlet BFW temperature was
backed out from duty, flow and outlet temperature.
-------
Based on the above the Economizer is as follows:
Duty = 606,000,000 Btu/hr
Flue Gas « 8,656,000 #/hr, 900—*650°F
BFW = 5,500,000 #/hr 650*—586°F
Transfer Rate = 9 finned
Surface = 500,000 SF finned = 40,000 SF bare
Use 4 parallel sections each 25' high by 40' wide
by 3' deep
Ducts leaving converter based on 60 FPS are
2 at 25' x 25'.
Pressure drop on flue gas side = 4" H2O
Acid Cooling System
Heat pick-up by acid is fixed by the specified tem-
peratures of flue gas in and out of the absorber. The acid
cooling range is also specified (by Monsanto) so the flow
quantity can be directly calculated.
As mentioned earlier the heat pick-up in the acid
is recovered by exchange with combustion air and BFW, both
through an intermediate fluid. There are a number of param-
eters which must be set before this system can be sized,
among them being:
1) The heat transfer fluid used.
2) The combination of flow quantity and tem-
perature range through which the circula-
ting fluid is exchanged. A large flow
quantity will lower temperature spreads
and reduce exchanger size at the expense
of piping and pumping costs.
3) The A T split between acid-fluid exchanger
and fluid BFW/Air exchanger.
4} The duty split between combustion air
preheat and BFW preheat.
5) Combustion air temperature and BFW
temperature.
It is obvious that the above parameters can be var-
ied in an almost infinite number of combinations. What
-------
follows is one scheme which might be used based on a judg-
ment of reasonable exchanger size, circulating flc, rat^s
and temperature approaches.
1) The circulating fluid used is water.
2) Combustion air and BFW were assumed to be
available at 80°F and 100°F respectively.
3) Minimum combustion air temperature to Air
Preheater to prevent condensation of acid -
350QF.
C-5 Primary Acid Cooler
Duty 176,000,000 Btu/hr
Acid 270—*>249°F
Water 210 •«—170
MTD 69°F
Rate 60 Btu/hr - SF - °F
Surface 42,500 SF
Number of Sections - 48 arranged 4 series
x 12 parallel 8" x 13'
Type Teflon
C-6 Secondary Acid Cooler
Duty 414,000,000 Btu/hr
Acid 249—>>200°F
Water 209-«—140
MTD 50
Rate 60
Surface 138,000 SF
Number of Sections - 150 8" x 13'
Type Teflon
Teflon is used here because it is known to be suit-
able for this service. The obvious drawback is in the large
number of sections required. Perhaps for an application of
this magnitude Du Pont would develop a larger bundle size
so that fewer sections would be required—the present maxi-
mum bundle size is 8" x 13' long. Aside from the high cost
-------
of manifolding lies the difficult control problem associated
with such large numbers of parallel sections. Not studied
were alternate materials which could be fabricated into a
fewer number of larger sections although this should defin-
itely be part of a complete study.
C-7 BFW Heater
The quantity of BFW flowing to the Economizer was
assumed to be available here at 100°F,
Duty 414,000,000 Btu/hr
Water 209—* 140
BFW 175*-100
MTD 37
Rate 300
Surface 37,300 SF
Sections 4- 42" x 30' Tubes
Mat'l TP 410 tube side
AP 10 to 15 psi each side
C-4 Combustion Air Primary Heater
Duty 176,000,000 Btu/hr
Water 210—»170
Air 170-<— 80
MTD 61.7
Rate 8
Surface 350,000 SF finned
Size 28' high by 40' wide by 3' deep
C-3 Combustion Air Steam Heater
Assume 300 psig steam available.
Duty 377,000,000 Btu/hr
Steam 422—»422
Air 350-*—170
MTD 144
Rate 7.2
Surface 340,000 SF + 20% for
cold air condition
Surface 408,000 SF
Size 28' high x 40' wide x 3* deep
-------
Air Preheater
Duty 485,000,000 Btu/hr
Flue Gas 650—»450°F
Air 550-*— 350
A Ljungstrora regenerative type preheater should be
used here because of relatively low cost. The air leakage
inherent in this design is not reflected by the above tem-
peratures.
Acid Product Cooler
Duty 4,100,000 Btu/hr
Acid 200 —>100
Cooling Water 100-*- 85
MTD 36
Rate 55
Surface 2100 SF
Size 3 i 700 SF
Absorber
A Monsanto study in which 2.5 MM SCFM of flue gas
were handled resulted in 9 towers of 30' 0 with 16* beds
packed with 2" Intalox saddles. Proration of the above
based on flue gas quantity for both pressure drop and heat
transfer would result in 4- 40* towers of 24' bed depth.
A check of this size using the generalized pressure drop
correlation would suggest that it is a little tight. For
this reason, as well as the unusually large diameter and
bed depth, 8 towers are suggested of 30' 0 with 20* bed
depths.
Number of towers 8
Tower size 30* 0 x 41'
Packed Bed 20' - 2" Intalox Saddles
Pressure Drop 10" H2O
Material Carbon steel lead lined
plus acid resisting brick
-------
Pumps
Acid Circulating Pump
80% Sulfuric Acid - Worthite (cast steel)
23,000 GPM - Say 6 @ 4000 6PM
50 pai differential
BHP - 23'X 5° " 670
Driver - Motor 6 @ 150 HP
Circulating Water Pumps
BFW - Acid
Water - 12,000 GPM
40 psi differential
BHP - 280
Driver - Motor - 350 HP
Combustion Air - acid
Water - 8800 GPM
40 psig differential
BHP -205
Driver - Motor - 250 HP
Fans
Combustion Air - Forced Draft
Capacity 8,150,000 f/hr - 1,680,000 SCFM
Static Pressure 10" Water ( A due to S02 removal)
Temperature 0 to 100°F
Type Centrifugal - 4 parallel
Drive - Motor 2830 BHP @ 70% Eff - 4000 HP -
4 6 1000 HP
Flue Gas - Induced Draft
Capacity 8,650,000 f/hr
1,750,000 SCFM
Static Pressure 29" Water ( A due to S(>2 removal)
Temperature - 225°F
Type Centrifugal - Minimum 2 parallel
Drive - Motor HP • 11,100 BHP £ 70% Eff - 16,000 HP
- 4 @ 4000 HP
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F-l High Laval Circulating Water Drum
8800 GPM £ 3 minute holding tin*
Us* 2-9' 0 x 28' Drums
F-2 Low Level Circulating Water Drum
12,000 GPM 9 3 minute holding tine
Use 2-10' ID x 30* Drums
F-3 CA-P) Product Acid Storage Tanks (16 Req'd)
Basin 104 GPM « 30 days
Capacity Required - 5,000,000 Gal.
4. Process Appraisal
General
As suggested in the chemistry review, this process
is severely handicapped by the need of a high reaction tem-
perature* The volume of flue gas handled by the Precipitator
and Reactor are twice the volume of the same weight flow at
more normal convection section exit temperatures. Also the
frictional pressure drop through the system will be higher
because of the higher temperature* A catalyst which would
permit operation at 450°F would avoid a redesign of the
boiler (i.e., external economizer) and probably result in
a better air preheat system because of the flexibility in
processing sequence. This should prove particularly advan-
tageous in add-on applications where a reheat system is
now required.
As pointed out in other evaluations, the market for
the product acid is quite variable. Also, a successful pro-
cess which depends on product sales to appear economically
attractive is self-defeating because the product will be in
over-supply.
-------
Flow Distribution
The large volumes of flue gas and combustion air
being handled will probably give rise to distribution prob-
lems* These problems should arise not only where one line
or duct divides into two or more but may cause trouble in
heat transfer equipment and the induced draft fans where
fairly uniform velocity profiles are essential for proper
operation. Perforated distribution plates will probably be
required in the former and straightening vanes in the
latter.
Because of the large diameter absorbers and deep beds,
attention to distribution of both gas and liquid is critical
for satisfactory performance in this equipment.
Fuel Flexibility
The sulfur content of the fuel used in this study
is approximately 3.5%. However, it should be remembered by
prospective users of this process that throughput will have
to be reduced for higher sulfur content fuel if emission
standards are to be met. For example, the combustion of
the design quantity of 5% S coal will result in about 85%
conversion as compared to 90% at the lower SO2 loading.
This will result in an SO2 stack gas concentration of about
2.2 times the design value. In addition, the Air Preheater
would probably suffer corrosion due to the higher dew point
of the acid.
Operation at lower sulfur levels than design does
not appear to be a problem except for the loss of "product".
-------
Fly Ash Removal
Problems associated with fly ash comprise plugging
of the catalyst beds, plugging of the mist eliminators, and
contamination of the product acid. The rate at which the
catalyst beds are plugged depends on the amount of dust in
the entering flue gas which in turn depends on the effic-
iency of the electrostatic precipitator and the ash content
of the coal burned in the power plant. When the catalyst
bed becomes plugged to the extent that pressure drop be-
comes excessive and/or conversion falls off, the catalyst
must be cleaned. This requires shutting down the unit,
cooling the catalyst, removing and cleaning the catalyst,
and recharging the unit. This operation likely will re-
quire several days. Catalyst loss owing to attrition and
breakage will, of course, be a direct operating expense.
The dust not removed either in the precipitator
or the absorber ultimately will appear in the product
acid—either directly or indirectly from back washing the
mist eliminators. The effect of this fly ash on product
acid marketability and selling price is unknown. No pro-
visions for filtering the product acid have been included.
Owing to the need for very high dust removal, the
precipitator must be maintained in good working order at
all times. This will probably give rise to higher than
normal maintenance costs. No provisions for recovering
from an outright precipitator failure have been included.
-------
EQUIPMENT LIST
C-l A,B,C,D Economizer
C-2 A,B,C,D Air Preheater
C-3 A,B,C,D Combustion Air-Steam Heater
C-4 A,B,C,D Combustion Air-Circulating Water Heater
C-5 Primary Acid-Circulating Water Cooler (48 sections)
C-6 Secondary Acid-Circulating Water Cooler (150 sections)
C-7 A, B,C,D Boiler Feed-Circulating Water Exchanger
C-8 A,B,C Product Acid Cooler
D-1A,B,C,D Converter
E-l A-H Absorber
F-l A, B, High Level Circulating Water Drum
F-2 A,B Low Level Circulating Water Drum
F-3 A-P Product Acid Storage Tank
j-1 A-H+2 spares Circulating Acid Pump & Drive
J-2 A,B+spare High Level Circulating Water Pump & Drive
J-3 A, B+spare Low Level Circulating Water Pump & Drive
J-4 A,B,C,D Combustion Air Fan & Drive
J-5 A,B,C,D Flue Gas Fan & Drive
J-6 Product Acid Pump
J-6A Spare
L-1A-H Brink Mist Eliminator
L-2 A,B Precipitator
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B. CHEMISTRY
In the Monsanto Cat-Ox Process, SO 2 is oxidized in the
gas phase through the catalysis by
V205
S02 + 1/2 02 •«. * S03
According to the equilibrium data, S02 conversion is
favored by low temperature. But the reaction rate is
favored by high temperature. Therefore, the Cat-Ox has
operated at about 900°F at the catalytic converter inlet,
giving about 90% conversion of SO2 in the stack gas. The
conversion is limited by the equilibrium constant which is
controlled by the partial pressure of S02 and oxygen.
Although the reaction appears simple, there has been
much controversy about the mechanism of the catalysis by
V2°5» The question of the mechanism remains unsettled be-
cause only the deductive approach of the Hougen-Watson type
has been used in the kinetics studies. Recently, some
workers have considered that changes in valence of the
vanadium is part of the mechanism. This is based partly
on the idea that potassium salt as a catalyst promoter
exists as a liquid phase that dissolves V2Os. However,
so far, no direct physical measurement such as electron
spin resonance has been made.
Vanadium pentoxide catalysts generally contain 7 to 9%
V2°5 anc* K/V ratio of from 2 to 4 on a siliceous carrier.
They require a temperature above about 750°F to be effec-
tive. The Cat-Ox process is said to use a specially de-
signed catalyst which prevents poisoning by ash components.
-------
An absorption tower is used to condense the sulfuric
acid. The very fine sulfuric acid mist particles are re-
moved from the flue gas stream with a Brink Mist Eliminator,
Flue gas which has passed through the Brink Mist Eliminator
contains less than 1.0 milligrams of 100% H2SO4/SCF, which
means that the stack gas contains less than 10 ppm of 100%
H2&O4 as a mist. The vapor pressure of sulfuric acid at
2250F will contribute another 11 ppm of 100% 1*2804 so that
the total sulfuric acid loading in the exit stack gas is
about 20 ppm. The inlet gas in the absorption tower con-
tains about 4000 ppm of 803. Therefore, the over-all
efficiency of the absorption tower and the Mist Eliminator
for the sulfuric acid removal is 99.8%.
The sulfuric acid produced has an acid concentration
of about 80%.
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MOLTEN CARBONATE PROCESS
A. PROCESS
1. Process Description (See MWK Dwg No. PD-124-D)
The Molten Carbonate Process for removal of SOX from
power plant flue gases was initially conceived and developed
by Atomics International who have done extensive research
and development studies. The firm of Singmaster and Breyer,
of New York City, have further developed the process and
have proposed a design*, still in preliminary form, for the
application of the process to an 800 megawatt power plant
burning coal with three percent sulfur content.
To bring the information in the Singmaster and
Breyer report to a common basis with studies being made on
other sulfur removal processes, their original equipment
sizing has been modified as required. Except for the cor-
rection of obvious errors, the Singmaster and Breyer design
bases have been retained in nearly every instance. The
major exceptions are: 1) the gas velocity in the absorber
towers has been reduced to avoid excessive entrainment, and
2) a quench step has been introduced ahead of the Claus
plant.
Drawing #PD-124-D shows a conceptual scheme for the
molten carbon process required for a 1000 MW power plant
burning coal containing 3.5% W sulfur. A Pfcl flow sheet was
not made for our process evaluation, but its concept can be
seen on Drawing PS-218-0002 in the Singmaster and Breyer
report.
* "An Evaluation of the Atomics International Molten
Carbonate Process", prepared for NAPCA by Singmaster and
Breyer, Contract CPA 70-76, November 30, 1970.
-------
The process consists of scrubbing of flue gas at
850°F with a molten eutectic mixture of lithium, potassium
and sodium carbonates (minimum melting point 750°F). The
melt, containing sulfite and sulfate salts, is regenerated
by heating to 1500°F, where the sulfates formed by dispro-
portionation of the sulfites can be reduced to sulfides with
carbon. The sulfides in the melt are converted back to
carbonates with the formation of hydrogen sulfide by re-
action at 850-950°F with carbon dioxide and steam. The
regenerated melt is circulated to the flue gas scrubber
and the H2S evolved is sent to a Claus plant for the re-
covery of elemental sulfur.
Fly ash entering the scrubber with the flue gas
is picked up by the circulating melt and must be removed
by filtration. The filtration process entails a loss of
melt with the filter cake. Most of the lithium salts so
lost are recovered, but the sodium and potassium salts
are discarded with the fly ash; therefore, the fly ash
quantity entering the absorber must be minimized. Con-
sequently, a minimum efficiency of 99.5% is required for
the power plant electrostatic precipitator and further,
the precipitator has to be relocated such that it processes
flue gas at 850°F rather than the usual temperature of
300°F.
-------
2. Basis for Material Balances
a. Flue gas specifications
Flue gas quantity 293,000 moles/hr
Fly ash 77,900 Ibs/hr
Flue gas composition - mole fraction
S02 0.00261
S03 0.00005
C02 0.13657
N2 0.74085
02 0.03276
H20 0.08714
HC1 0.00002
Chloride composition of flue gas must be
considered since reaction will occur with
K2C03 to form KC1 which must be rejected
from the process to minimize its build-up.
The amount of K2CO3 consumed is related to
the chloride content of coal.
b* Precipitator
99.5% W fly ash removal from 850°F flue gas
Absorber
Moles Absorbed
% Absorbed mole M2C03 melt
S02 95.0 0.1435
[ S03 94.0 0.0160
HCl 95.0 0.0018
I 02 1.0 0.0337
, Total - 0.1950
i
The absorber material balances have been pro-
rated from Singmaster and Breyer's report and as a somewhat
different flue gas analysis is now used, are not in 100%
balance.
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d. Reduction
Coke and combustion air are supplied in suffi-
cient quantity to maintain the reduction reaction at 1500°F
with complete consumption of carbon and oxygen and 95% re-
duction of M2S04 by carbon. All of the N2 and enough of
C02 to maintain an 85% atoichiometric excess of C02 in the
regenerator are vented to the stream of scrubbed flue gas
returning to the power plant. It is assumed that 95% of
the sulfur in the coke is converted to M2& with the re-
mainder retained in the solid residue* Metals, ash and un-
burned coke, totaling 4% of the dry coke feed are contained
in the reducer effluent melt and are removed by filtration*
For heat and material balance purposes, the 12%
volatile combustible matter (VCM) in the coke is assumed to
flow from the reducer with the exit CO2 gas, 86% passing to
the regenerator and 14% to the power plant. This assumption
may not be valid, since at 1500°F in the presence of oxygen,
the VCM of the coke probably will both burn and crack to
H2, CH4 and carbon. If this occurs the coke requirement
will be reduced by an amount equivalent to the heat released
by this combustion.
Although it is clearly stated by Singmaster and
Breyer that the green petroleum coke contains 5-8% moisture,
they have ignored this in the material balance, which has
been made on a moisture-free basis. Revision of this design
concept is beyond the scope of this report.
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e. General
All melt flow rates for the 1000 MW power plant
molten carbonate system were obtained by direct scale-up of
melt stream compositions from Singmaster and Breyer's flow
sheet PS-218-000. Due to lack of time no attempt was made
to calculate material balances based on theory of absorption
or reduction or regeneration reactions. Unit material bal-
ances were checked however, and all balances on the Sing-
master and Breyer flow sheet were verified except around
the reducer, where slight discrepancies were observed.
Therefore, the reducer balance for the 1000 MW plant has
been carried through on "scale-up" basis.
f• LJ2CO3 Recovery
Pilot plant investigations reported by Atomics
International have shown that 88% of the lithium salts lost
from the circulating melt with the fly ash and coke filter
cakes can be recovered as Li2C03 and re-introduced into the
circulating melt as make-up. None of the sodium and potas-
sium salts so lost are recovered in this process.
Carbonate make-up requirements of the 1000 MW
plant, on a water-free basis, are:
Lithium Carbonate, net 115 Ibs/hr
Sodium Carbonate 990 Ibs/hr
Potassium Carbonate 1684 Ibs/hr
An equivalent quantity of soluble mixed salts (carbonates,
sulfides, sulfites, sulfates and chlorides) are discarded
with the fly ash and coke slurry.
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g. Glaus Plant
Design sulfur recovery by the Claus Plant is 92%,
an economical value for a plant charging the type of gas
which is produced by the regenerator. The off -gas from the
Claus plant is recycled to absorber where the un-recovered
sulfur in the qas is picked up by the melt and recycled to
attain essentially complete recovery of the sulfur removed
from the power plant yap. Only a small quantity of sulfur,
about 320 lbs/hr., is lost with the soluble salt slurry as
sulfites, sulfides and sulfates.
3 . Process Design Pationale
a. Absorber
International ~tudies of the absorption
of SOX in molten carbonate resulted in a recommendation to
use a spray Lower for process design. Results from their
study showed that the melt circulation rate could not be below
that resulting in a minimum M2CO3 concentration of 68% at
the absorber outlet. This is the lowest practical concentra-
tion as set by freezing point consideration. As the melt
concentration is increased above this value, a high ratio of
M2c°3ss°x entering the absorber results. Recycling of melt
from the absorber outlet to the inlet decreases this ratio
(the recycle contains up to 32 mol % SOX) . Recycling was
recommended to maintain the ratio of liquid-to-gas as high
as practical, thus insuring maximum contact of the gas with
the liguid. Atomics International have set a M2CO3:SOx ratio
as 10:1 maximum, thus limiting the recycle ratio to about 3.
Singmaster & Breyer has recommended that the ab-
sorbers be designed for series flow on the liguid side but
with parallel flow on the gas side to increase the melt/gas
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mole ratio and so insure sufficient absorption of SO2* In
addition, M2CO3 recycle had been incorporated into the ab-
sorber design to obtain a high absorber flow rate with mini-
mized melt circulation through the rest of the process. In
the absorber Singmaster & Breyer's design criteria for 95%
removal of SOx from flue gas were based on a superficial
gas velocity of 25 ft/sec and a liquid spray mean particle
size of 100 microns. However, in our judgment, the super-
ficial gas velocity for the spray tower must be limited to
10 ft/sec in order to avoid excessive entrainment loss.
In this design for a 1000 MW plant a gas vel-
ocity of 10 ft/sec necessitates the use of eight spray
towers with the flue gas flow in parallel. The liquid flow
is four in series through two parallel trains, with the
streams combined in common pump tanks after each pair of
towers. The recycle flows are individual to each pair of
towers, i.e., the bottom liquid from the first towers is
returned to the inlets of these towers, etc. The spray
towers are specified as being 35 ft in diameter by 34 ft
high. The absorber pumps are designed as being capable of
a 3:1 recycle ratio. Additional design criteria for ves-
sel sizing, insulation, and materials of construction
areas are as specified by Singmaster & Breyer.
b. Fly Ash Filtration
Filtration requirements are as specified by
Singmaster and Breyer with scale-up for the 1000 MW power
plant. The filter requirements are for 2 units operating
alternately, with a 45 min. cycle time per filter. The
filter units are leaf type designed for maximum pressure
drop of 50 psi at full load, requiring 1350 ft2 filtration
area per unit. <
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c. Reducers
The main reactions which occur in the reducers
are:
Oxidation: M2S+2O2 »-M2SO4
C + 02 *"CO2
Reduction: M2S04 + 2C —*• M2S + 2C02
Disproportionation: 4M2S03 —*3M2S04+M2S
Coke S Recovery: S+!sC+M2CO3 —»-M2S+ 3/2 CO2
The endothermic heat for reduction of M2SO4 to M2S,
and the sensible heat required to raise the melt, coke, and air
to the reduction temperature of 1500°F are provided by the oxi-
dation of M2S and carbon (coke). At 1500°F, the M2S03 in the
melt disproportionates to M2SO4 and M2S.
Four reduction vessels, operating in parallel, have
been provided for a 1000 MW power plant facility. Each re-
ducer is a two compartment vessel—one an oxidation zone and
the other a reduction zone, with liquid passage between the
two compartments for internal melt recirculation. The basic
parameters for design of the reducers are:
Liquid Retention Time 20 min.
Static depth of Melt 2-3 ft.
Superficial Gas Velocity 3 FPS
Bed Expansion 100% max.
A further detailed discussion of the basis for re-
ducer vessel design including vessel configuration, selection
of insulation and materials of construction can be found in
the Singmaster & Breyer Report.
Combustion air for the reducers is preheated by
heat exchange with oxidation zone off-gas, and admitted to
the oxidation zone of the reducer through sparger pipes. The
oxidation zone off-gas is then sent to the power plant boiler
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with the absorber off-gas. The reduction zone off-gas,
consisting essentially of CC»2, is quenched by water and/or
steam to reduce the C02 gas stream temperature from 1500°F
to 850°F and sent to the regeneration. This quench provides
the 85% excess steam necessary for the anticipated operation
of the regenerator. The* quench steam is assumed to be avail-
able from the waste heat boiler in the Claus plant. The C02
production in the reducer exceeds the regeneration require-
ments at 85% excess, so the surplus C02 is vented to the
oxidation zone off-gas stream.
Melt from the reducers is cooled in the reducer
quench tank to 950°F by recirculating 850°F reduced melt
from the reducer product coolers. This recirculation con-
trols the temperature of the melt in the quench tank.
In the heat and material balances of the re-
ducer S & B assumed that the VCM content of the delayed
petroleum coke, which is 12% by weight, merely vaporizes
and passes off with the C02 gas. This is extremely improb-
able. In practice, this material, consisting of heavier
petroleum hydrocarbons, should be released from the coke
almost instantaneously when the coke is brought to 1500°F
by contact with the hot melt. There should be immediate
reaction with oxides in the melt to convert all the VCM to
CO2 and water. There may be a tendency at this temperature
for a water gas equilibrium producing some CO as well.
Singmaster & Breyer's treatment of this VCM as
inert is conservative for the reducer design, but puts an
undue burden on the regenerator and the Claus plant. If
credit were taken for combustion of this VCM, the coke re-
quirement of the unit would be reduced by about 12%.
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Moisture in the coke has also been ignored by
Singmaster & Breyer. This amounts to 5 to 8% of the dry
coke weight. Introduction of this much water vapor in the
reducer may promote a water gas reaction resulting in ex-
cessive amounts of H2 and CO in the CO2« An afterburner
in the reducer off-gas line might be required to oxidize
these undesirable materials.
d. Coke Filtration
Coke filtration requirements are similar to
those for fly ash, but with only 1000 ft2 filtration area
required. Unreacted coke is removed from the process melt
and sent to the Li2C03 recovery process. The filtered melt
is then sent to the regenerator.
e. Regeneration
The regeneration reaction is expressed as:
M2S+C02+H20 —*" M2CO3+H2S
The regenerator is a tower consisting of 15
bubble cap trays where melt is in countercurrent contact
with C02 and HpO. Regeneration is exothermic and coolers
are provided with melt recirculation through the tower as
shown on the flow sheet. Only one tower is required for
the 1000 MW plant facility.
The theoretical number of trays required is
approximately 4 based on equimolar quantities of CO2 and
H20 at an excess of 85% with little or no inert diluents
present in the gas. The equivalent tray efficiency is 27%.
Justification for this design is discussed in the Sing-
master & Breyer report. It is expected that the two zone
design of the reducer will minimize the inert diluent
gases that can affect the separation.
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The flow rate of regeneration gas feed, and
therefore the excess of C02 in the regenerator, is con-
trolled by venting surplus CO2 gas from the reducer to the
combustion gas stream being returned from the reducer to
the power plant.
The regenerator tower is designed for a 15 min-
ute minimum retention time on the trays. With melt re-
cycled for cooling, approximately a 2-inch level is required
on the tray, producing a static head loss of 30" for the
15 trays. Since the liquid has a specific gravity of 2.0,
this static head is equivalent to a pressure drop of 2.1
psi. There are other tower losses and piping losses as
well. With the reducer operating at 6.0 psig, the exit gas
from the regenerator will be at approximately atmospheric
pressure.
f• Treatment of Regenerator Off-gas
Off-gas from the regenerator, excluding the
"coke volatiles", is 34% H2S and 33% each CO2 and H20, at a
temperature of 850°F and one atmosphere pressure. For best
operation a Claus plant should have cold gas feed at a
pressure of at least 5 psig. High water vapor content greatly
increases the cost of the plant, as does high inlet tempera-
ture.
Cooling, de-watering and compression of the re-
generator off-gas is indicated in order to obtain an optim-
um design for the Claus plant. A quench tower, with heat
removal by water circulation, is a practical means of con-
densation. This tower, containing four trays and five baf-
fles, receives the 850°F regenerator off-gas which is
cooled to 120°F by countercurrent flow of condensate cooled
to 110°F and reheated to 160°F, the dew point of the gas,
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as it flows over the trays and baffles. Most of the water
in the regenerator off-gas is condensed by this quench.
The 160°F liquid is pumped through an air cooler and its
temperature reduced to 110°F and then recirculated to the
top of the tower. The net condensate is recycled to the
regenerator as quench water for the C02 fer»ti gas. Almost
half of the total 24,200,000 Btu/hr cooling duty is from
condensation at 160°F or lower, with the rest of the duty
originating as sensible heat between 850° and 110°, so that
no heat recovery iti justified.
The cooled gas is compressed to 6 psig by a cen-
trifugal blower, pensed through a water-cooled condenser
which remover the heat, of compression and condenses more
water, and flows to the Claus plant at 100°F and 5 psig,
conditions which are tr.ost economical for the sulfur plant
design.
g. Claus Plant
The Claus plant used to recover elemental sul-
fur from the treated regeneration gas is a conventional
unit which is offered competitively by several contractors.
The cost of the plant increases with the degree of sulfur
recovery. Here a recovery of 92% has been specified since
the off-gas is recycled to the absorber and an ultimate
recovery of nearly 100% of the absorbed sulfur will be
attained. The only sulfur loss will be that in the sol-
uble salts discarded with the fly ash and coke slurry.
This loss is estimated to be only 1.3% of the sulfur re-
moved from the flue gas.
It is important that the feed gas to a Claus
plant does not contain more than 2% hydrocarbons, and these
should be not heavier than propane. Excess hydrocarbons
will result in a "black" sulfur product and will upset the
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heat balance of the unit, resulting in damage from over-
heating.
It has been assumed that the hydrocarbon content
of the net feed (after water removal) will be a maximum of
20 mols/hr (1.3 vol. %). If more hydrocarbon than this is
found likely, it will brs necessary to remove the VCM from
the coke fed to the reducer or to incinerate the VCM in the
reducer off-gas.
h. Coke Handling Facilities
Green delayed coke of 2-inch maximum size is
delivered in rail cars and fed to a crusher to reduce it
to \ inch maximum size. The crushed coke is screened and
oversize returned in closed circuit to the crusher. The
crushed material is charged to 3 coke silos with a total
of 3 days live storage capacity. The unloading and crush-
ing facilities are sized for operation of one eight hour
shift per day.
From the silos the coke is fed to coke bins
associated with each reducer. Coke flows from the coke
bins via weigh feeders which measure coke into blow tanks
from where it is then charged to the reducers, using air
pressure supplied by the combustion air blower.
Detailed equipment layout for the coke mater-
ial handling equipment may be found in the Singmaster &
Breyer report for an 800 MW power plant facility. This
equipment has been scaled up directly for the 1000 MW
facility, requiring 3 storage silos instead of two to
maintain live storage capacity for increased coke consump-
tion requirements.
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No provision has been made for removing mois-
ture from the coke. As received this moisture normally
may be 5-8% by weight and in bad Breather it may be more.
Although not specified by Singmaster & Breyer, it would
seem advisable to provide a weathor roof over the unload-
ing and handlinq facilities, particularly since this
operation mu«t be done daily, handling some 8-50 ton cars
per 8-hour c5--./. Excess rain water accumulated in unload-
ing and handling will be detrimental to the process.
There is a possibility that drying of the coke before
charging to the reducer may be required.
If the plant is in a location subject to
freezing weather, special precautions must bo made to in-
sure that the coke r.u'-.ply is not interrupted by freezing
in the cars or in the unloading equipment.
i. M2CQ3 Make-up
Carbonate make-up requirements total 43 tons
per day, of which about 10 tons is Li2CO3 recovered from
discarded fly ash and coke filter cakes. Facilities are
provided to receive the three carbonates by bulk truck
shipment. The salts are dumped into a receiving hopper
from which they are transferred by a bucket elevator and
cross conveyor to their respective storage silos. The
recovered Li2C03 is discharged into the same silo as the
fresh. Approximately 4 to 5 days storaoe in provided, so
that the K2CO3 hopper is larger than the other two.
From th« silos/ weigh feeders and conveyors
deliver the proper carbonate quantities to the Carbonate
Melt Tank. As sized by Singmaster & Breyer, the electric
heater capacity of this tank is sufficient to heat and
melt the make-up on a continuous basis and to balance
radiation loss at 850°F. Only about 25% surge capacity
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was provided; for the 1000 MW plant the over-capacity has
been increased to 40%, The melted carbonate make-up is
pumped continuously into the regenerator bottoms pump tank.
This whole make-up system appears to be de-
signed on the basis that all three salts will be received
as anhydrous materials. No provision has been made to
cope with any water content of the raw materials, and no
heater capacity for vaporizing water of crystallization
is included.
These salts are deliquescent and must be pro-
tected from the atmosphere at all times. Although not
specified by Singmaster & Breyer, all carbonate handling
facilities must be housed to insure that the fresh salts
are kept dry.
For start-up purposes, this carbonate melt
tank is essential. However, during routine operation it
should be possible to introduce the cold make-up salts
directly into the regenerator bottoms pump tank, using
a screw conveyor or a standpipe to overcome the 2 psig
positive pressure in the tank. There is ample surplus
heat here not only to heat and melt the make-up carbon-
ate but also to drive off any water of crystallization
which the salts might contain.
j . LijCOj Recovery Process
Lithium carbonate is a relatively expensive
material (42-52C/lb). Appreciable quantities of melt are
discarded with the fly ash and coke filter cakes; the loss
of the lithium in these discards would impose a definite
economic disadvantage on the process.
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The recovery process is based on the insolu-
bility of lithium carbonate and its easy conversion to
soluble lithium bicarbonate. Laboratory work to establish
the feasibility of the process was done by Atomics
International. Singmaster & Breyer'a development of the
process is far from complete, with some sections appar-
ently in the conceptual stage.
Filter cakes from the fly ash and coke filters
at 850°F are dumped periodically and are quenched with
water, flowing to the dissolving sump. Here all salts
except lithium carbonate are dissolved. Fly ash, coke
residue, and lithium carbonate are slurried in the re-
sulting salt solution, then filtered. The filtrate, con-
taining only a small quantity of dissolved lithium salt,
is discarded. The filter cake is re-slurried with a cir-
culating stream of water, which contains 0.7% of Li2C03
in solution. This slurry is charged in batches from a
receiver to a reactor when it is agitated with inert gas
containing 12% C02« The carbonate is converted to solu-
ble bicarbonate, leaving the ash and coke residue as
slurry which is dumped into a surge tank. Slurry from
this surge tank is filtered to remove the fly ash and
coke residue, which is discarded. The filtrate is
collected and charged batchwise to a reactor where it is
blown with hot air to strip out the CO2 and re-convert
the bicarbonate to carbonate. This insoluble carbonate
is discharged as a slurry to a surge tank from which it
is fed to a filter. The filter cake is passed through a
drier and sent to the Li2C(>3 storage silo. The filtrate
containing some dissolved Li2C03, is recycled to slurry
the filter cake discharged by the fly ash filter.
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Atomics International report successful re-
covery of 88% of the lithium entering the recovery system.
This efficiency has been used as the basis for the pres-
ent design.
fc« Changes to Existing Power Plant
If this process is to be applied to an existing
1000 MW power plant some revisions to the present boilers
will be required. It may be assumed that the plant .is
equipped with an electrical precipitator following the
air heater but ahead of the induced draft fans. In all
probability the efficiency of this precipitator is not
over 99%. In any case, it will not be suitable for opera-
tion at 850°r with ?9.5% dust collection efficiency so it
must be replaced. It is assumed that its dust handling
and disposal equipment is in good condition and can be
salvaged for use with the new precipitator.
The boiler casing must be modified to allow
all the flue gas to exit at a temperature of 850°F, pass
through the new precipitators and absorbers and return,
still at 850°F, to the economizer section of the boiler.
The present precipitator would be removed and the induced
draft fans would be replaced to give added capacity equal
to the increase in draft loss imposed by the new equipment.
This added power requirement resulting from the replace-
ment is 10,000 BIIP.
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4. Process Appraisal
A brief summary of engineering development efforts
required for commercializing the process are summarized
below for each step in the process including absorption,
filtration, reduction, regeneration, and lithium recovery.
a. Absorption
The main difficulty inherent in commercial
design of absorbers which needs to be further demonstrated
is the ability to eliminate melt carryover in the absorber
off-gas. Melt loss in the absorber off-gas is undesirable
from an economic standpoint and intolerable from a corros-
ion standpoint since melt carried by the scrubbed gases
may condense in the return vapor line or in the boiler
system. Condensation of melt on the cold surfaces of the
boiler would cause rapid decrease in heat transfer and pre-
mature shutdown for cleaning.
No investigation has been made on the loss of
melt to flue gas by vaporization. The carbonate salts
have a low but definite vapor pressure. The flue gas
leaving the absorber will be saturated with melt vapor at
850°F. When the gas is cooled to 300° in the economizer
and air heater of the power plant, most of this vapor will
condense and drop out as dust. Some plating of cold sur-
faces may occur.
The proposal to operate the absorber at a
superficial velocity of 25 ft/sec is outside the scope of
known commercial practice. In other processes, velocities
well below this are used for particle transport, i.e., 100%
entrainment. A maximum velocity of 10 ft/sec is suggested,
and even at this lower rate, it is doubted that entrainment
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can be low enough to permit the use of a demister pad for
effective entrainment separation.
In towers of this size (35' diameter), distri-
bution of the inlet gas becomes an important factor. Ho
inlet distributors are indicated by Singmaster & Breyer.
It is felt that elaborate turning vanes with a perforated
plate distributor above them will be required in the bot-
tom of the absorber tower.
The flue gas duct sizes indicated result in
velocities of 106 ft/sec from the precipitator and 156
ft/sec to and from the absorber tower. It is felt that
these velocities should be limited to 60 ft/sec to avoid
uneconomic pressure drops and difficult flow distribution
problems.
There is no assurance whatsoever that the 95%
absorption of SO* obtained by Atomics International in a
4" tower with a single spray can be duplicated in a 35'
diameter vessel. The possible corrections for wall effects
and other factors are mentioned by Singmaster & Breyer
who advise that extensive plant scale studies will be re-
quired.
A recycle range of 1:1 to 3sl on melt circula-
tion through the absorbers has been recommended. However,
no provision for adequate distribution of this stream by
sprays in the tower has been shown. The operating range of
a spray nozzle is extremely limited and best results are
obtained only when the flow is held very close to design
rate and pressure. It would appear advisable to design
and operate the spray system at the maximum recycle pump
capacity, particularly since no disadvantage to a high
recycle rate has been shown.
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The ability to remove fly ash from flue gas by
electrostatic precipitation at 99.5% efficiency over ex-
tended operating periods is highly questionable and must be
demonstrated. The possibility of plugging spray nozzles
by ash in the recycle melt increases greatly if actual fly
ash removal efficiency is below 99.5%, causing an increase
in fly ash concentration in the melt. Correspondingly/
continuous spray nozzle operation must be demonstrated
for increased ash contents of the melt.
Demister design requires further consideration
and demonstration that it will not plug with melt or fly
ash, can be operated with acceptable pressure drop and
will effectively eliminate carryover.
The fourth and final absorber pump tank is
oversize so that it may hold the surges in the melt circu-
lation. Therefore, its instrumentation should be only a
Level Recorder, with a Flow Controller on the filter out-
let measuring the entire melt circulation and resetting
four Flow Ratio Controllers on the feeds to the four re-
ducers .
Flow indicators are needed on the recycle melt
streams to each of the absorbers.
b. Fly Ash and Coke Filtration
A complete discussion of commercial design prob-
lems associated with filtration may be found in the Sing-
master & Breyer report. Suitability of pressure leaf fil-
ters as selected by Singnaster & Breyer for cost estimation
must be verified with prototype tests. The calculation of
filter area requirements for the 1000 MW plant facility uses
empirical relationships developed by Atomics International.
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These correlations must be verified with prototype test
data for pressure leaf filters for both fly ash and coke
filtrations.
No means of preventing excessive heat loss
and possible freezing of the flow in the filters has been
suggested. This should be investigated. The alternately
off-stream filters must be maintained at 850°F as well.
c. Reduction
Continuing development efforts are required
for satisfactory solution and demonstration of the follow-
ing:
(1) Materials of Construction
The design of a reduction vessel must be
verified regarding its structural integrity and ability
to withstand corrosive attack.
(2) Reducer Configuration
The two-zone reducer concept has not been
physically demonstrated to date and thus requires additional
effort to insure operability. The object of the two-zone
design is to separate oxidation gases from reduction gases
so that reduction gas alone may be used for regeneration.
Equipment sizing and configuration for an 800 MW plant fac-
ility was investigated by Singmaster & Breyer on an economic
basis alone. On a technical basis, it will be important to
better define the location of a zone separation baffle and
the size of baffle plate openings required to insure proper
melt circulation. The dynamics and reliability of melt
circulation between the two zones must also be verified.
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(3) Reducer Off-Gas Composition
The reducing zone off-gases should contain
a minimum of N2, CO, and hydrocarbons from coke volatiles
since the presence of any of these gases in large amounts
will affect the operability of the Claus plant. After-
burning of reducing zone off-gases may be required to con-
vert unburned hydrocarbons and CO to CO2. If excessive
unburned hydrocarbons reach the Claus plant, black sulfur
will be produced.
(4) Reducer Off-Gas Cooling
Corrosion problems inherent in the design
of reducer off-gas cooling relate to the chloride and melt
composition of the gas. Melt carryover in the reducer off-
gas must be minimized and at 1500°F, chlorides may vaporize
from the melt.
(5) Coke Addition
As reported by Singmaster & Breyer, it may
be necessary to introduce coke in the form of a slurry with
the melt in order to insure coke contact with the melt in
the reducer. Also, sufficient agitation with air must be
maintained, but is limited by the degree of melt expansion
and frothing that can be tolerated.
(6) Coke Distribution
It is obvious that the consumption of coke
in the reducing and oxidizing zones must be closely con-
trolled to obtain the desired off-gases and overall reaction,
The suggested design does nothing to achieve this distribu-
tion. Back circulation of reduced melt is required to bring
the required quantity of carbon into the oxidizing zone and
to maintain nearly equal temperatures in the two zones. A
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means of forcing this circulation under proper control
should be developed.
Temperature records of both sections of
the reactor are necessary so that the coke and combustion
air quantities can be adjusted as required. Continuous
analysis of the off-gas is needed to keep the coke and
air flows in the proper proportions.
The instrumentation shown by S & B is
not practical in that the valve shown to release excess
CC>2 from the reducing section to the oxidation section
exit gas has no pressure drop and if closed will merely
result in internal short-circuiting of gas through the
circulation ports in the internal baffle. This line
should be left open and its flow controlled by regula-
tion of the off-gas rate of the reducer which in turn is
regulated by the charge rate of the Claus plant.
As mentioned elsewhere/ studies should
be made to determine the effect of water introduced with
the coke, and if necessary provision for its elimination
should be included. Similarly, the effect of the VCM in
the coke should be determined and the current design
basis modified to reflect the reactions which occur due
to its presence. Hydrocarbon material in the C02 gas
must be eliminated as it cannot be tolerated in the Claus
plant in concentration above 2.0%.
d. Regeneration
Regeneration studies would appear to be more
thorough than those for some other parts of the process.
The proposed design seems feasible/ using bubble cap trays
to obtain the desired liquid retention while allowing for
significant changes in the gas rate as may be expected from
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the normal load variations of the power plant. These var-
iations, estimated to give a turndown of 50%, would be
greater than are practical for the use of perforated trays.
Pressure drop in this tower will be greater
than anticipated by the Singmaster & Breyer estimate.
e. Glaus Plant
It was assumed by S & B that feed gas at 850°F,
and 3" H20 pressure, containing 32% water vapor would be
suitable for feed to a sulfur recovery plant of the Claus
type. The conventional feed to such a unit is supplied at
100-120°F under 5 psig pressure, which results in a much
more economical installation and more satisfactory opera-
tion. Added facilities to cool, de-water and compress the
feed gas are specified for the 1000 MTV system.
Hydrocarbons in the feed to a sulfur plant are
highly undesirable. As previously discussed, it is impor-
tant that the reducer system be designed in such a manner
that hydrocarbon generation is minimal.
Cost of a sulfur plant is influenced by the
completeness of the sulfur recovery. In this case the
off-gas is to be recycled to attain 100% sulfur yield.
A once-through recovery of 92% has been chosen as a suit-
able design basis. For final design a more complete econ-
omic study should be made, considering the incremental cost
imposed on the other parts of the system by this recycle
and balancing this against the cost of the Claus plant.
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f. Raw Material Handling
For a final design the problems which may arise
in the handling of the coke and carbonates should receive
more attention. Delivery facilities for these materials
will be especially designed for each plant and will vary
with the source of the various materials. The railroad
facilities for eight cars of coke per day will be extensive,
requiring trackage for storage of both full and empty cars.
The size and arrangement of these facilities will be gov-
erned by the dependability of the rail service.
Availability of the carbonates in the required
quantities must be known. It is possible that the potas-
sium and sodium carbonates will be available only as
hydrates and that the storage system and melt facilities
must be designed with this in mind.
g. Lithium Carbonate Recovery
In the Singmaster & Breyer design of the car-
bonate recovery section there are several areas of uncer-
tainty. These are not insurmountable problems and should
be resolved easily for final design.
It is probable that the filter cakes from the
fly ash filter and the coke filter will be discharged al-
ternately, and quickly. It would seem advisable for these
hot slugs to drop into the dissolving sump where an apprec-
iable mass of water has been accumulated. This will keep
a reasonably steady temperature on the feed to the Li2CO3-
fly ash filter. However, this alternating slug flow
should be held in the sump for sufficient time to produce
a constant quality feed to the filter. The Singmaster &
Breyer sump is not large enough for this. Four to six
-------
cycles/ or three to four hours holding time would be re-
quired. The approximately 54 minutes provided is insuffi-
cient.
As specified by S & B the Lithium Carbonate
Filtrate Receiver (T-108) and the LiHCO3 Reactor Feed Pump
(P-109) are included with the Fly Ash Filter and are dupli-
cates of items associated with the Li2C03~Fly Ash Filter.
This cannot be, as the receiver and pump must be designed
for batch loading of the L±2C^3 reactor (T-108 must be
larger, P-109 must be for vacuum suction). This has been
corrected in the 1000 MW specifications.
No specifications at all aro given for the Hot
Air Blower which supplies air to the Li2CO3 Reactor. On
the basis of chemical equilibrium the reactor should
operate as close to 212°F as is practical. Working away
from the 25 HP motor size specified by S & B, the air de-
livered might be 2000 Ib/hr. No source is shown for the
hot air; it might be assumed that it could be taken from
the discharge of one of the air cooled exchangers at per-
haps 300°F. If this air were at 300°F, its heat on a 100%
efficiency basis would be enough to give only an 8° rise
in the water in the reactor. It is clear that this step
of the process has not received proper attention.
An inert gas generator is provided to supply
gas of about 12% CO2 content at 75 psig for the LiHCOa
Reactor. Singmaster & Breyer specification for this unit
seems to be in error as to the cooling water required by
this equipment. For this specified capacity the water
use might be 545 GPM rather than 75 GPM stated. The
fuel requirement for the 800 MW plant is 8000 SCF/hr (not
8000 SCFM). The space provided on the plot plan for this
generator, including a blower driven by a 30 HP motor and
-------
a compressor driven by a 200 HP motor, all said to be skid
mounted, is but 6'4" x ll'O" (scaled). On the basis of the
size of similar units, in other plants, this space is in-
adequate and the unit is too large for mounting on a single
skid.
This generator is specified to deliver gas of
12% CC>2 content. The necessity of the generator is ques-
tionable when a practically limitless supply of desulfurized
flue gas of higher CO2 content is available from the stack
of the power plant at 300°F. Only the compressor would be
required, and this would be located at the stack, deliver-
ing the gas at 75 psig through connecting piping.
A dryer is provided for the Li2CO3 filter cake
recovered by the process. S & B have specified a moisture
content of the feed which is eight times the quantity of
dry carbonate, but the stated fuel consumption of the drier
will evaporate only a fraction of this quantity of water.
This amount of water does not appear elsewhere in the
material balances, so it may be in error. The design basis
for this dryer should be reviewed.
h. Waste Disposal
This process requires the disposal of an apprec-
iable quantity of slurried waste containing dissolved salts.
For the 1000 MW plant studied here the slurry discard
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composition is estimated to be as follows!
Water 15,000 Ibs/hr
Air 390
Coke Residue* 1285
Total Solids 1675 Ibs/hr
KC1 685
Mixed Salts** 2157
Total Dissolved Salts 2842 Ibs/hr
Total 19,517 Ibs/hr
* Includes nickel, iron & vanadium oxide.
** Sulfate, sulfite, sulfide, and carbonate
salts of sodium, potassium & lithium
In many locations there will be a disposal prob-
lem. Consideration should be given to development of a
means of salvaging the salt, particularly the potassium and
residual lithium salts which now incur appreciable make-up
costs.
The valuable metals in the coke ash may total
as much as 0.1 wt. % of the coke, equivalent to 500 to 700
Ibs/day of vanadium and nickel. Recovery of these metals
might prove profitable.
i. Temperature Maintenance
High quality insulation has been specified for
most of the equipment in order to minimize heat loss and
avoid "cold spots" which could cause freezing of the melt.
Electric heating elements sized to overcome heat losses
are included in the pump tanks. It has been assumed that
in the case of shutdown all of the melt will be drained
into these tanks and maintained at 850°F to prevent freez-
ing. There is no overall pumpout system to rid all of the
-------
Such an arrangement is advisable. A single large tank
would be provided for total melt storage. During any
prolonged shutdown the melt could be allowed to cool and
freeze and would be re-melted for start-up. Either elec-
trical heating or direct fired flues through the tank
could be used.
In the Ringmaster & Ureyer report all of the
melt piping is specified as "electrically traced". This
is necessary not only to prevent freezing during opera-
tion but to heat the system prior to start-up so that the
melt can bo circulated. Such electrical heating, whether
by resistance or by induction is rather expensive and
much electrical gear is required.
No suggestion has been made?an to how the re-
qenerator, the coolers, the reducers, and the filters are
to be heated prior to start-up to prevent freezing of the
melt at the start of circulation. This will require fur-
ther study and development of operating techniques.
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B. CHEMISTRY
1. Literature Reviewed
The following list of reports from Atomics Interna-
tional have been reviewed:
1. Development of a Molten Carbonate Process for Removal
of Sulfur Dioxide from Power Plant Stack Gases, 6/1/67 -
2/28/68.
2. Progress Report No. 2, Part 1, Process Chemistry, Re-
duction, 4/1/68 - 10/27/68.
3. Progress Report No. 2, Part 2, Process Chemistry,
Regeneration, 4/1/68 - 10/27/68.
4. Progress Report No. 2, Part 5, Fly Ash Studies, 4/1/68
10/27/68.
5. Progress Report No. 3, 10/28/68 to 7/31/69.
In reviewing the chemistry, the process flow will be
followed from the absorber to the filter, reducer>regen-
erator and back to the absorber.
2. Absorber
Although a reasonable number of basic melt properties
for the pure eutectic of lithium, sodium and potassium
carbonates has been mentioned, additional data should be
obtained at or near the actual composition of the steady
state absorber fluid. First of all, the pure melt has a
melting point of 747°F but with 20% sulfite-sulfate the
melting point appears to be about 78C°F, thus to operate
-------
at 800°F without knowledge of viscosity and the effect
of fly ash on this melt may be dangerous. Although, the
last report recommended 800 to 850°F, without further
knowledge 850 to 880°F looks like a somewhat safer range
for operation at this time. It is recommended from the
literature that melt viscosities should not exceed 200
centipoises for reasonable flows and we concur with this.
Atomics International indicates that S02 absorption is
not affected by the level of sulfite-sulfate in the melt
up to reasonable quantities, by the presence of 10% fly
ash and by the presence of water vapor and/or oxygen in
the flue gas. Sulfur dioxide absorption is very rapid,
irreversible and apparently gas-film controlled. The
reactions of interest in the absorber are:
S02Cg) + M2C03(1) » M2S03(1) * C02Cg) • CD
S03Cg) + M2C03(1) » M2S04(1) + C02(g) C2)
1/2 02(g) + M2S03(1) > M2S04(1) (3)
There is no reason to believe that S03 will not be taken
up by melt although no experimental evidence was noticed.
Some study of reaction (3) indicated it was not as fast
as reaction (1), but that a substantial amount of the sul-
fite formed in reaction (1) could be oxidized to sulfate
Cflue gas contains 3.3% 02). This is a function of the
time of contact of melt and gas, thus dependent on the
type of contactor.
-------
Although results showed nitrogen oxide absorption
at a slower rate than S02 absorption, no identification
of the products was mentioned. One might expect that
decomposition or reduction to nitrogen gas occurs. How-
ever, nitrite-nitrate buildup in the melt should be checked,
Experiments were performed which showed that ^S was not
emitted into the flue gas from the melt, at least at the
level of 0.4 wt.% sulfide in the melt, despite the pre-
sence of water vapor and carbon dioxide in the flue gas.
Therefore sulfide level must be low in the absorber to
present no problems in this direction.
Some of the experiments performed by AI were done in
a bomb type container with gas being bubbled through a .
few inches of melt. The rate of flow of gas was low,
usually well under 0.1 ft/sec superficial gas velocity,
whereas good mixing and more conventional flows would
require greater than 0.25 ft/sec superficial gas velocity.
Naturally at these flows baffling must be done to prevent
excessive entrainment or carryover out of the reactor.
In other words, the experiments were performed at flow
velocities which did not approach practical values and
this must be kept in mind. Larger contactor tests were
made - wetted wall, baffled tube, spray tower - at super-
ficial gas velocities up to 30 ft/sec. Removal of S02 was
easily obtained. The results indicated the wetted wall
-------
had to be about 9 times as long as the baffled tube to
get similar S02 removal efficiency. About 5 ft/sec linear
velocity in the baffled tube showed 4" of H20 pressure drop
which was considered excessive. Spray tower tests have
not been completed. Contactor design does appear to be
the major problem for the absorber.
3. Fly Ash Removal
AI has shown that it is preferable to remove the fly
ash from the melt coming from the absorber and before going
to the reducer. From a temperature standpoint, it is also
the lowest melt temperature in the process and may allow the
use of a metal filter. The sulfide content is also low
which is a desirable condition from corrosion and solubility
limitations. At higher tmeperatures fly ash can be signi-
ficantly attacked by a molten carbonate system and show
high solubility of silica.
Enough work appears to be done to fix the composition
of the filter cake at roughly 2/3 melt and 1/3 fly ash.
However, it was not established whether this tie up of melt
was due to just physical absorption or partially to chemical
reaction.
The solubility of iron and calcium oxides to the ex-
tent of about 0.2 gram metal/lOOg melt at 840°F is of in-
terest. The formation of a red color due to iron oxide
-------
solubility in a 62% M2C03, 37% M2S04 melt may be indic-
ative of ferrite or ferrate formation. For instance,
I^FeO* is a deep red in color.
The fact that 25 to 65 micron opening filters can be
used to filter out fly ash which contains a large amount
of particles less than 25 microns is indicative of melt
addition to fly ash to form particles larger in size
which do not pass even 65 micron openings.
It is suggested that future filter development in-
clude life tests on the material to be used especially
under the cycling conditions envisaged for this operation.
Naturally, design of the filter requires maintaining equip-
ment in contact with melt above the freezing point of the
system especially during potential upset conditions.
4. Reducer
The advantages of using fluidized coke instead of the
gaseous reductants are well documented by A.I. It appears
that the filtered melt at approximately 850°F will be pumped
to the reducer which should operate at about 1500°F in
order to get a rapid rate of reduction of all sulfur to
the sulfide form in about 15 minutes. Heat of reaction
must be supplied as well as heat required to go from 850
to 1500°F. Excess carbon must be present and sufficient
oxygen for the heat generating reaction to maintain tempera-
ture .
-------
AI has run tests which indicate that out of a large
number of possible reactions, only three overall reactions
predominate, namely, 4 N^SO^ » 3 N^SO, + f^S (1)
M2S + 2 02 > M2S04 + heat (major) (2)
M2S04 + 2 C > M2S + C02 - heat (3)
Therefore, it is necessary to continually supply coke
and melt with sufficient oxygen to maintain the melt tem-
perature at about 1500°F.
It appears that the disproportionation reaction (1)
will be very rapid and so will the oxidation of sulfide
to sulfate, reaction (2), which liberates about 230 Kcal
per mole Na2S (1340°F). Heat liberated from reaction (1)
is very minor. Thus heating up of the melt appears to be
no problem. However, carbon must be present to reduce the
sulfate to sulfide which will consume about 43 Kcal/mole
Na2S (1340°F). Although this last reaction may not be as
rapid as the others, it is still a fast reaction if good
contact and sufficient coke (excess) is available.
Since all the reactions are very rapid, the major con-
cern should be with the mechanical end of getting good
mixing in order to take advantage of these rapid reactions
If the melt is quiescent, most of the coke will rise to
the surface; only when a gas, i.e., air and/or recycle
gas, is bubbled through the melt may sufficient agitation
be available for good mixing. Recycle gas is mentioned
-------
because the air must be controlled to amounts which may
be less than sufficient for good mixing, above 0.25 ft/sec
superficial gas velocity. As the fluid coke is consumed,
the remaining particles will tend to flow with the melt
as well as some of the initial fine particles, thus a
coke separation problem exists and should be studied. A
settling zone (upward for coke) may be a solution.
Another point of importance is the concentration of
carbon dioxide in the off-*gas from the reducer. If air
and/or recycled gas are used in the reducer, will the
carbon dioxide content in the off-gas be enough to use in
the regenerator? Perhaps a particular design for the re-
ducer can yield the proper C02 content stream.
A,I. have considered the effect of bed expansion and
did show advantageously that coke in the melt at the level
of 6.25% did significantly lower bed expansion. One
questions the superficial gas velocities on the figure
reporting this work, it was presumed that 20 was 0.2 and
40 was 0.4 ft/sec. More investigation of bed expansion,
foaming or frothing, and salt entrainment should be done
with melts simulating reducer conditions. This will be
important for design.
The C02/CO ratio of the off gas from the reducer has
been looked at and reported to be about 5-6/1. One must
be cautious in analyzing gases of this type coming out of
-------
a reactor at high temperature because of possible gas
phase reactions occurring after leaving the hot melt.
Normal techniques use probes which take samples directly
above the melt and attempt to freeze the gas composition
by sudden cooling. If all the oxygen gets consumed in
the melt a probe is unnecessary, but if oxygen exits
then any CO can be oxidized to C02 at these temperatures
and the results do not reflect the true conditions.
After—burning may be necessary to convert any CO to C02
before using this gas in the regenerator. Possibility
of COS formation exists.
Since the rates of reaction are quite fast and
materials handling is probably more of a problem, then
it hardly seems necessary that catalysis by added iron,
except for that which is present in the melt, need be
considered at this time. It may also cause frothing
and require separation before the melt can be regenerated
Specifications on coke may be required. Volatile
matter can vary widely and AI has shown that calcined
(temperature unknown) delayed coke was poor. Since the
reducer operates at about 1500°F any coke could be pre-
heated to about this temperature to lower volatile matter
and not change the reactivity of the carbon appreciably.
5. Regenerator
The regeneration reaction is,
M2S(1) + C02(g) + H20(g) y==^ M2C03(1) + H2S(g) + heat.
-------
From literature data on sodium sulfide and potassium sul-
fide, the reaction as written is favored at low tempera-
tures and the reverse reaction is favored at high tempera-
tures. This was substantiated by experimental data at
800°, 930° and 1110°F. Choice of about 900'F for regenera-
tion appears to give rapid equilibrium and a suitable
equilibrium constant of about 10. Equilibrium runs made
by Atomics International established that the reaction is
exothermic,AH is -23 Kcal/mole M2S in the range 850-1125°F
using mole fractions instead of activities for the liquids
and ignoring .side reactions. Estimate of the effect of
the side reaction, C02 + M2S « M2C02S, monothiocarbonate
formation, indicated that A H could be -28 as an upper
limit. Non ideal liquid behavior, i.e. effect of activities,
would tend to lower the AH value. Therefore the result
appears satisfactory.
Considerable chemistry is involved in passing C&2
into the J^S • M2C03 melt. AI have performed a number
of experiments to try to unravel this system. The results
of these experiments and the negative effect of increasing .
C02 pressure on I^S removal indicated that thiocarbonate
formation appears to be the best plausible answer. No
direct analysis for thiocarbonate was made however. The
effective Ke for regeneration which incorporates the pres-
sure effect of carbon dioxide appears to give reasonable
-------
agreement between calculated and experimental values,
K.
1.90 x 10~fe exp (28,000/RT)
[1 + Pco x 1.24 x 10-3 exp (-15,40011
i RT
Therefore less than 0.5 atm of C02 in regenerator gases
is recommended by AI to keep thiocarbonate formation
minimal.
Investigation of the ratio of C02 to h^O in the gas
to the regenerator indicated 1 to 1 gave maximum con-
version providing the melt did not absorb large amounts
of C02.
The desired concentration of sulfide in and out of
the regenerator was not established. However, it was
shown that when the mole fraction of N^S reached less
than 0.02 mole % (1.4 wt.%), there was a rapid drop in
H2S content in the effluent gas. If the absorber can
tolerate 1.4 wt.% N^S without loss of sulfur as f^S into
the flue gas, then this looks like a satisfactory working
melt.
Analysis of the rate of reaction data by AI indicated
equilibrium was attained very rapidly and that there was
a dependence on the incoming gas rate. It appeared that
when gas rates were pushed up to what would be considered
for engineering design, the equipment plugged up due to
-------
salt entrainment, etc. Better designed experimental equip-
ment is needed which will allow proper gas flows. Design
must consider melt entrainment in the off-gases.
A disadvantage of having iron present in the melt while
regenerating was the enhanced foaming and insoluble potassium
iron III sulfide, KFeS2, formation. Thus, it does appear
that added iron or other catalysts should be avoided in
the melt to the regenerator.
Regeneration time of 20 - 30 minutes appears feasible
with properly designed equipment and with good gas-liquid
contact. The 1500 ° F melt out of the reducer will have
to be cooled down to 850-900°F for the regenerator and
heat will have to be removed from the regenerator due to
the exothermic reaction.
6. Mi s ce11aneous
From the very complete corrosion investigation of metals
only 347 stainless steel appears satisfactory at temperatures
below 950°F, therefore this may be the metal for the ab-
sorber, pumps, filter, piping and possibly the regenerator.
For the reducer and melt piping which operates at about
1500°F, AI suggests Haynes LT-1 (cermet), tf-A^Oj, or a
cooled steel surface with frozen melt on the surface to
protect it. Additional testing and evaluation of these
high temperature materials at process conditions appears
des i rab1e.
-------
The chemistry of the lithium recovery appears sat-
isfactory; One wonders whether potassium and sodium re-
covery should also be practiced since consideration must
be given to possible pollution from discarded streams.
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C. MECHANICAL
1. Coke Handling
This process uses green delayed high-sulfur petroleum coke
as a reducing agent. Although 2" coke might be used for reduc-
tion, a crushing and screening circuit has been included to
reduce the coke size from 2" to minus 1/4". If the need for
crushing is uncertain, this indicates that the reducing equip-
ment might easily accept a nominal amount of oversize. If so,
the crushing and screening circuit could be simplified by
eliminating the closed circuit screen entirely and mak.inq ,H
single pass through the crusher. This would not only remove
the cost and operation of a piece of equipment, but it would
eliminate the circulating load on the bucket elevator and crusher,
reducing their size and cost.
The crusher choice for this service is unusual. It is a
two-rotor cage mill and it will produce a product with a large
percentage of fines. If the process doesn't require this, a
simple impactor could be substituted to save cost and power.
A few errors and ommissions were noted in the estimate of
the coke handling equipment.
a. The bucket elevator which feeds the coke crusher
was not rated to include the circulating load ,.
b. A belt conveyor is required to convey the cj-.u:.jhe>i,
sized coke from the screen to the coke silo feed
elevator. This conveyor is shown on the arrange-
ment drawings but is omitted from the flow sheet
and cost tabulation.
c. Two feeders should be provided at the coke silos
to feed the elevating belt conveyor.
d. Dust collecting equipment should be added to
provide pick-up points at the following loca-
tions:
1) The feed chute to the reversing conveyoi. "var
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2) The feed chutes from the coke silos to the
elevating belt conveyor.
3) The tranfer chute from the elevating conveyor
to the tripper belt over the coke bins.
4) The four coke bins.
2. Li2 C03 Recovery Area
The flow diagram and equipment cost estimate have omitted
a feeder under the surge drum between the filter and dryer.
The drum should be equipped with a type of feeder which will
insure positive discharge from the drum and a steady, controll-
able feed to the dryer.
t,
3. M2C03 Make-up Area
The carbonate bucket elevator has not been provided with
a truck hopper, or feeder. A truck dump hopper with capacity
for one and a half to two truckloads should be added to the
estimate.
A dust collecting connection should be provided at the
loading point for the silo feed conveyor.
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LOW 5HOWM IS F~ -
IMftECYCLE RATIO. P"OR MAX. DES "»
3:i RECYCLE RATIO, FACM J-3 FLOW
IS 1136 6PM lOfc BHP
RECYCLE IS 5t<»,300
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POTASSIUM FORMATE PROCESS FOR REMOVING SOa FROM STACK GAS
A. PROCESS
1. Process Description (See MWK Dwg No. PD-111-*D)
In this process, stack gas is scrubbed at 200°F with an
excess of 85% aqueous solution of potassium formate to form
potassium thiosulfate. The solution is then heated to 540°F at
520 psig and the excess formate reduces the thiosulfate to potas-
sium hydrosulfide. The hydrosulfide is then stripped with carbon
dioxide and steam thereby producing hydrogen sulfide from which
elemental sulfur is recovered in a three-stage Claus plant. The
stripped solution is predominantly a mixture of potassium bicar-
bonate and carbonate. This mixture is regenerated to potassium
formate by reaction with carbon monoxide at a total pressure of
2000 psig and a temperature of 540°F. The formate solution is
then recycled to the stack gas scrubber.
The main chemical reactions involved in the potassium
formate process are summarized below in the same sequence they
occur in the process.
Scrubbing Reaction
2 KOOCH + 2 S02 « K2S2°3 + 2 CO2 + H2°
Reduction Reactions
4 KOOCH + K2S203 = 2 K2CO3 + 2 KHS + 2 C02 + H2O
K2C03 + C02 + H20 - 2 KHCO3
-------
Stripping Reactions
2 KHS -I- CO2 + H20 = K2C03 + 2 H2S
KHCO3 + KHS = K2C03 + H2S
Regeneration Reaction
K2C03 + 2 CO + H20 = 2 KOOCH + C02
The overall reaction of the formate process may be written
as:
S02 + 3 CO + H2O = H2S + 3 C02
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2» Process Design Basis
This process has been evaluated for operation with a new
1000 MW power plant. The design criteria used in evaluation of
the formate process are given below:
Plant Size
Heat Rate
SO2 Removal
Excess Air
Carbon Burned in Coal
Heating Value in Coal
Wet
Dry
Sulfur to S02
Sulfur to SO3
Ash in Coal to Fly Ash
Temperature out of Air
Preheater (nominal)
Minimum Gas Temperature
to Stack After Treatment
Preferable Gas Temperature
to Stack After Treatment
Gas Velocities in Ducts
Flue Gas Quantity
Flue Gas Molecular Weight
Fly Ash
1000 MW
9000 Btu/kwhr
90%
20%
100%
11,980 Btu/lb
12,580 Btu/lb
98%
2%
71.7%
300°F
200°F
250°F or higher
60 ft/sec
293,000 moles/hr
29.54
77900 Ibs/hr
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Flue Gas Composition
Component
S02
SO 3
C02
N2
02
H20
Mole Fraction
0.00261
0.00005
0.13657
0.74087
0.03276
0.08714
Coal Composition
Component
Ash
S
H2
C
N
0
Water in Coal
Dry Wt. %
15.2
3.5
5.0
67.2
1.6
7.5
4.8 lbs/100 Ibs wet
Design criteria related specifically to the potassium formate
process are discussed in the section Process Design Rationale.
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3. Process Design Rationale
The lack of information and data on this process made
it necessary to make many assumptions before a process flow
sheet could be defined and evaluated. The subsequent discussion
will give the assumptions made and design criteria used in eval-
uating this process.
Stack gas flows through cyclone separators where 80% of
the fly ash is removed. It is assumed that since the Formate
process is a wet scrubbing process, the fly ash remaining in the
flue gas effluent from the cyclones will be removed in the stack
gas scrubbers.
Flue gas flows from the cyclones into the stack gas
scrubbers. From the article describing the Formate process,
there does not appear to be any appreciable back pressure of
BC>2 in the aqueous formate solution. The absorbers were there-
fore sized assuming no SC>2 back pressure. A three stage spray
column would give 90% removal of SO2 from the stack gas assuming
tower efficiency to be 60%. A linear gas velocity of 500 ft/min
was used to determine the tower area. Based on the limited
amount of available data, it appears that a liquid recirculation
rate lower than that used in the laboratory would be adequate.
For this evaluation a liquid rate of 230 Ibs/hr ft2 was used
but it should be noted that if rates 5-fold higher were used,
the effect on the overall plant economics would be slight. The
tower operates at 200°F with an 85% solution of potassium form-
ate. The experimental work done at these conditions indicates
that water is not evaporated and the scrubber stays in water
balance. Since enthalpy data were not available on the absorp-
tion system it was assumed that the absorbers are in heat bal-
ance. It was further assumed that the £03 in the stack gas is
converted to K2SO4 and the sulfate precipitates out of solution
but all other potassium salts would stay in solution. The
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formate solution was assumed to remove the fly ash remaining in
the stack gas downstream of the cyclones and the fly ash could
be slurried in the formate solution.
It was assumed that centrifuges could handle the formate-
fly ash slurry effluent from the scrubber to separate out the fly
ash. If filters were used instead of centrifuges an electrostatic
precipitator would be required instead of cyclones since filters
can handle a solution with only about 0.1% fly ash; thus fly ash
removed in the precipitators would have to be 99.5%. The solids
stream discharged by the centrifuge is assumed to contain 20
weight percent liquid. To keep potassium losses low, three cen-
trifuges in series are used with a water wash (reslurry) between
each centrifuge.
The ash free solution is sent to the reducing reactor
which is designed as a continuous stirred tank with a 20 minute
holding time. This is the same holdup as that reported in the
ACS paper. Temperature and pressure are also the same as that
demonstrated in the laboratory (540°F, 520 psig). In the reactor
all the potassium thiosulfate is assumed to be converted to the
hydrosulfide. While passing through the centrifuges, the solu-
tion was diluted from 85% to 58% potassium salts but it is
assumed that this dilution does not affect the reaction rate in
the reducer.
The reduced potassium salts from the reducing reactor
are sent to the E-2 stripper. The concentration of H2S in this
stripper off-gas was assumed to be 25% based on the batch strip-
ping data presented in the ACS report where about 30% I^S was
initially observed. Stripping is done with CO2 and steam. En-
thalpy data are again lacking and a heat balance could not be
made on this tower.
-------
The H2S-rich gas ia sent to a Glaus reactor where ale-
mental aulfur is recovered. Recovery of the available sulfur
in the Glaus plant is 97.5%; oxygen rather than air is used in
this plant for the combustion of the H2&. Since carbon dioxide
leaving the Claus plant is recovered but must be compressed be-
fore absorption/ use of air instead of oxygen would greatly in-
crease the compressor size due to the nitrogen in the air.
Gases leaving the Claus plant are rich in carbon dioxide.
As suggested by the ACS paper, a hot potassium carbonate
absorption-desorption system is used to recover 98% of the avail*
able C02» Since additional C(>2 is required to keep the plant in
carbon balance (CO2 is lost in the stack gas scrubber), makeup
C02 from the reformer is used. Sulfur dioxide leaving the Claus
plant is assumed to be absorbed by the K2C03 to give potassium
sulfite.
This K2SC>3 represents an expensive potassium loss but
potassium values probably could be recovered by additional pro-
cessing in equipment (L-5) which has yet to be defined.
In the formate regenerator (D-2), experimentally it re-
quired two hours to convert the carbonate to formate. The
reason for the long residence time was attributed to the experi-
mental equipment. It was assumed with proper agitation only a
25 minute residence time would be needed. Experimentally pure
CO was used with an inlet partial pressure of 700 psi and the
water vapor pressure of the 60% carbonate solution at reaction
conditions was 500 psi. A commercial reformer will produce only
about 50% CO maximum. Consequently, the commercial reactor must
operate at 2000 psig to keep the inlet partial pressure of CO
at 700 psi. Thermodynamic data were not available for this re-
action. Estimates of the heat of reaction were made based on
using sodium formate and the reaction was found to be exothermic.
It was assumed that the excess water would be evaporated from
-------
the reactor to bring the exit product formate concentration to
the desired 85% level. The regenerator exit gas, which is rich
in CC>2 and CO, is recycled to the reformer.
Natural gas reforming of C02 is used to generate the CO
needed for the regenerator. The main reaction in the reformer
was assumed to be
CH4 + CO2 « 2 CO + 2H2
The heat required for the endothermic reaction is provided by
combustion of natural gas with air. Part of the combustion gas
is used to supply the makeup CC>2 needed in the plant.
-------
4* Process Appraisal
The basic concept of the potassium formate process is defined
in the ACS paper. The experimentation done on this process
demonstrates only the basic concept and this limited data pan be
used only for a very preliminary process design and economic eval-
uation. Additional data are required in every facet of this process.
Before discussing those areas which need further clarification, the
advantages and disadvantages of this process will be given.
Advantages
Wet scrubbing of flue gas at 200°F eliminates the need
for reheat to get plume rise.
Tii': app.-.ir'.'.nt e I. .1. j ciency of absorption indicates that the
S02 absorbed hns no back pressure in the formate solution.
This great affinity keeps the liquid recirculation rate
down to relatively low levels and provides nearly complete
removal of SC>2 •
The use of liquid scrubbing may make the use of electro-
static precipitators unnecessary if centrifuges can be
designed to handle the fly ash slurry formed with the
potassium formate solution.
The SC>2 is recovered as elemental sulfur rather than HaSOi*
thur providing greater flexibility in handling, storage,
-------
Pisadvantages
• Being a liquid scrubbing process, some spray will pass
through even the most efficient demisters. Potassium
salts are expensive and losses represent large operating
expense. It would also be undesirable to have appreciable
quantities of formate or carbonate contained in the exit
flue gases since there could be as bad or worse pollutants
than the sulfur oxides.
c The operating pressure of the formate regenerator is high
(2000 psig) because the reducing gas in a commercial unit
contains a maximum of 50% CO. The slow reaction rate and
high pressure of regeneration impose severe design problems
on the reactor.
0 The need for recovering the C02 after the Glaus plant may
pose problems in the hot potassium carbonate recovery
system as suggested in the ACS paper. Any unconverted
sulfur leaving the Glaus system as S02 will be absorbed
and reacted with the K2CC>3. This represents intolerably
high potassium loses and consequently requires additional
processing which has yet to be defined.
• As in the stack gas scrubber, the hot carbonate recovery
system is a source of potassium loss due to mist formation.
The data required to give a complete definition of the process
for detailed design must be obtained experimentally. Literature
sources are expected to give only a minimal amount of information
pertinent to this process. Experimental data required are proces-
sing and thermodynamic. The thermodynamic data requirements are as
-------
follows:
• Vapor pressure of water over the potassium salt solutions.
• Solubility data of the different potassium salts in water
and each other, i.e. phase diagrams.
• Heat of reaction data — probably available in the
literature.
• Heat of mixing and solution data.
The different chemical processing steps of the formate process
must each be experimentally studied. Those areas where processing
studies are required are discussed below:
• Scrubbing
Experimental tests must be made to obtain data on the
effect of the variables such as temperature (concentration),
liquid to gas rate and conversion levels of potassium
formate on mass transfer rate. The effect of nitrogen
oxides on the absorption must also be determined. The
rheological properties of the resultant fly ash slurries
produced with the aqueous potassium solutions should also
be determined and proposed methods of separating the fly
ash should be confirmed.
• Reduction
This reaction must be studied to determine the effects
of temperature, pressure, residence time, fly ash content,
water concentrations and stability of the potassium
-------
formate. The results given in the ACS paper shows.
potassium sulfate being formed and eliminated. Side
reactions producing I^SO^ are a source of potassium
loss.
• Stripping
Variables affecting the stripping of the sulfur values
in the reduced potassium salt solution are CC>2/steam ratio,
temperature, pressure and solution concentration. This
reaction step does not seem to present any major problems.
• Regeneration
This reaction step appears to be the most difficult one.
The variables to be studied which affect reaction rate
are temperature, pressure, CO partial pressure in the
reducing gas, H2/CO ratio, concentration of the reacting
solution and mechanical mixing effects on the mass trans-
fer rate. Effects of foaming and expansion may pose
serious problems (increased reactor size) and must be
determined.
In addition to the above basic area of study required in the
formate process, some of the associated facilities used in this
process also must be examined.
• Recovery of C02 in Hot Potassium Carbonate
The use of potassium carbonate absorption of carbon
dioxide may be undesirable owing to the losses of potas-
sium incurred by reaction with sulfur dioxide leaving
the Glaus plant. Other absorption systems should be
evaluated for applicability to the formate process.
If hot potassium carbonate sorption is used then the
additional processing to recover the potassium values
from the sulfite formed must be defined.
-------
• Reforming
It was proposed in the ACS paper that natural gas
reforming with C02 be used to produce the required
reducing gas. Although this reaction is known it has
not been used on a commercial scale and may require
additional experimentation.
In addition to the studies made on the individual steps of
this process an integrated pilot, plant must be operated to
demonstrate to total system. This is required because recycle
streams may build up undesirable compounds and bleed streams
would have to be used to eliminate these compounds.
Source of Data
The process design and evaluation is based entirely on a paper
presented at the ACS National Meeting in Houston on February 22-27,
1970 . No other sources of data on this process were available
to Kellogg for this evaluation.
(1) Yavorsky, P. M., Mazzocco, N. J., Rutledge, G. D., and Gorin,
E., Potassium Formate Process for Removing S02 from Stack Gas,
Paper Presented at the ACS National Meeting in Houston,
February 22-27, 1970.
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B. CHEMISTRY
1. Scrubbing
The chemistry is likely to be feasible. As pointed out
in the process discussion, water scrubbing upstream of the
absorber eliminates the advantage of the hot scrubbing. However,
eliminating the water scrubbing introduces contamination by
soluble ions in the ash of the solution. This is mentioned in
the previously cited ACS paper, but is not discussed in detail.
No mention is made of any other ion than 804.
On this point the flow sheet is not worked out to a sufficient
degree.
It is noted also that foaming is mentioned in a later step,
where good gas-liquid contact is desired. .Foaming tendencies,
if they exist, are a danger to the process in that they would
demand extreme care in handling. It would not be desirable or
permissable to spray the countryside with mixtures of potassium
thiosulfate and formate.
2. Reduction
There is uncertainty as to the water vapor pressure over the
solution before and especially after reduction. It is quite likely
that even with a low vapor pressure, say at 200°F and a, high .heat
of vaporization, a relatively high water vapor pressure at 540°F
will result.
In the actual experiment a reflux on the cool autoclave head
can easily have hidden the evaporation by causing a sizeable reflux.
This may lead to expensive changes in the flow sheet.
Table III of the ACS paper indicates feed and product analysis.
from this a decrease in both potassium and sulfur content follows,
which is impossible. Clarification is desirable.
-------
Further, K2SO4 appears and disappears without explanation.
This could possibly lead to expensive bleed stream treatments.
3. Stripping
Stripping has been carried out only after dilution of a typical
reduced product. It is implied that one could also do this step in
concentrated solution. This raises several questions.
First, is it possible to have solutions before and after
stripping? Second, what is the stripping factor if stripped at
pressure and at temperature?
4. Conversion
Conversion seems to be limited by mass transfer, it is stated.
A constant rate from zero to 80% conversion seems to bear this out.
If, however, the break occurs at 80%, or in other words, if at 80%
the process becomes reaction rate controlled, the overall improvement
by better stirring can never be five times greater reaction rates as
stated on page 21 of the ACS paper.
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AMMONIA SCRUBBING ( (NH.) SO, - FERTILIZER )
4 2 4
A. PROCESS
Sulfur Dioxide removal from power plant stack gas by ammonia
scrubbing followed by production of ammonium sulfate for use in
fertilizer manufacture.
The process under evaluation is designated as Process "A"
in the TVA report entitled "Sulfur Oxide Removal From Power Plant
Stack Gas, Conceptual Design and Cost Study Series, Study No. 3
Contract No. TV-29233A".
1. Process Description (Refer to MWK Flow Sheet Dwg. No.
PD-112-D.
a. Sulfur Dioxide Removal and Production of Ammonium
Sulfate
Flue .gas exits the air preheater at 300°F and is split into
.six parallel streams which flow thru identical scrubbing systems
before being exhausted into a common stack. The gas is first
cooled to 172°F in gas cooler C-l. Ammonia is injected into the
gas stream upstream of the cooler to neutralize S03 for corrosion
prevention in the lower part of the scrubber. The cooled gas
containing 4.9 gr/scf of solid particulates is scrubbed with water
on the bottom tray of a four stage scrubbing tower E-l. 99.5 wt %
of the solids are removed and discharged to settling tank F-4
common to the six scrubbers. The solids are settled out and
diluted with water to a 12 wt % ash slurry which is pumped to a
settling pond. The solids free gas flows countercurrent to the
wash solution containing NH3, SO2, (NH4)2SO3, NH4HS03 and (NH4)2S04.
Three stages are required to remove 90 wt % of the SO2 by reaction
with NH3 and (NH4)2SO3. The proposed scrubbers are countercurrent
seive tray type Jmpinjet gas scrubbers manufactured by the W. W. Sly
Mfg. Co.
Six scrubbers are required to handle the flue gas from a
1000 MW plant since the largest scrubber available is 20' wide
x 40' long. The total sidestream drawoff from each of the three
-------
wash trays is segregated flowing respectively to the recirculation
surge tanks F-l, 2, 3. Each tank is common to three scrubbers
resulting in two recirculation systems each of which feeds three
scrubbers via pumps J-l, 2, 3. Fresh NH3 makeup is added to each
recirculation stream and the liquid composition on the trays is
regulated by drawing off a slipstream from each drum flowing
stepwise from F-3 to F-2 to F-l. The scrubber effluent product is
withdrawn from drums F-1A, B and pumped to a common surge tank F-6.
The scrubber effluent consists of a 50 wt % salt solution containing
a mixture of SC>2; NH3 and ammonium sulfite, bisulfite and sulfate.
The flue gas is cooled adiabatically in the scrubber by the evapora-
tion of water and leaves the scrubber saturated at 116°F. Makeup
water is injected into the scrubber above the top tray and liquid
entrainment is removed by a fixed blade impingement baffle located
in the top of the scrubber. The gas is reheated to 247°F in gas
reheater C-2 to ensure adequate dispersion of the gas leaving the
stack. Flow through scrubbers is maintained by exhaust fan J-9
located downstream of gas reheater C-2. The six fans exhaust to a
common stack and maintain the entire system under slight negative
pressure. A pumparound system utilizing surge tank F-5 and pump
J-8 pumps tempered water thru cooler C-l and gas reheater C-2 in
series transferring heat from C-l to C-2.
The combined effluent product from the six scrubbers is pumped
from surge tank F-6 to neutralizer D-l. The neutralizer is sparged
with NH3 to convert the remaining bisulfite to sulfite in order to
minimize SO2 loss during the subsequent oxidation reaction. Water
is added to the neutralizer to lower the salt concentration to 40
wt % to prevent crystallization of ammonium sulfate in the oxidizer.
Off-gas from the neutralizer containing ammonia and water is vented
back to the scrubbers. The neutralized solution is pumped to
oxidizers D-2A, B (2 in parallel) where the ammonium sulfite is
reacted with air at 100 psig and 185°F to produce ammonium sulfate.
Heat of reaction is removed by recirculating thru solution cooler
C-4 where the salt solution is cooled from 185°F to 135°F. About
-------
8 vol % of the cooled salt solution is pumped to surge tank F-7
from which it is fed to the fertilizer plant precipitator. . Off-
gas from the oxidizers containing nitrogen, excess oxygen and
water is vented back to the scrubbers.
b. Production of Ammonium Nitrate-Phosphate Fertilizer
Phosphate rock is reacted with 60% nitric acid to form soluble
calcium nitrate and phosphoric acid. The resulting extract slurry
is heated to 165°F and charged to the precipitator where it reacts
with ammonium sulfate to form insoluble CaS04, H3P04 and NH4NC>3.
The CaS04 is removed by filtration and discharged to a sluice tank
where it is diluted with water to a 20 wt % slurry and pumped to
the settling pond. The filtrate consisting mainly of ammonium
nitrate and phosphoric acid is neutralized with ammonia and then
concentrated to 99.7 wt % in a steam heated falling film type
evaporator. The concentrated ammonium nitrate-phosphate product
is spray dried with air in a prilling tower, cooled, screened
for particle size control and sent to bulk product storage. Under-
size and oversize product is dissolved and slurried and recycled
to the neutralizer feed tanks. Total production is 2370 tons/day
of fertilizer ranging in content from 25-15-0 to 28-14-0 (N2, ?20s,
Potash) depending on the type of phosphate rock used as feed and
the plant operating conditions. 60% nitric acid used in the ex-
traction reaction is produced in a 1275 T/D (100% HN03 basis) acid
plant. Ammonia storage, unloading and pumping facilities are
provided to supply the acid plant, fertilizer plant, S02 scrubber
and bisulfite neutralizer requirements.
2. Process Design Basis
Basic process design criteria common to all of the S02 removal
processes being evaluated are not included in the following list.
Overall Plant Capacity Factor, % of Nameplate Rating
1st to 10th year 80%
llth to 15th " 57%
16th to 20th " 40%
21st to 35th " 17%
Average Capacity Factor 43%
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a. Ammonia Scrubbers E-1A, B, C, D, E, F
The following data pertaining to the scrubbers selected as
being most suitable for this service was obtained from the W. Sly
Mfg. Co. bulletin and is applicable to the cross flow sieve tray
Impinjet type gas scrubber.
Quantity (6) Six
Dimensions 20' x 40' (rect.) x 27' high
Design Gas Velocity, Nor. 420 fpm, max. 500 fpm, actual 461 fpm
Fly Ash Dust Removal 99.5% (total fly ash from boiler)
Water recirculation to bottom scrubber stage for dust removal,
combined plate and spray water, 3 gpm/1000 ACFM gas
Ammonia injection upstream of gas cooler C-l for corrosion control
in bottom of scrubber is based on the stoichiometric quantity
required to react with the 2% of sulfur in coal which forms 863.
b. Scrubber Operating Conditions
Gas inlet temp., 172°F
Gas outlet temp.,116°F
Effluent temp., 120°F
Effluent salt concentration, 50 wt %
Effluent NH3:S mole ratio, 1.49
Salt solution recirculation rate, 1.14 gpm/1000 ACFM
Fly Ash Slurry concentration to settling pond, 12 wt % solids
c. Bisulfite Neutralizer D-l
Salt inlet temp., 120°F
Salt outlet temp., 188°F
Effluent salt concentration, 40 wt %
Effluent NH3:S mole ratio, 2.0
Approx. retention time, 10 minutes
d. Sulfite Oxidizer D-2A, B
Salt inlet temp, 188°F
Salt outlet temp.,185°F
Recirculated salt solution temp., 135°F
Operating temp., 185°F
Air to liquid ratio, 7.6 SCF/gal
Operating pres., 100 psig
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PH, 6.5
Approx. retention time, 1 hr
Recirculated salt solution to product ratio, 11:1
e. Ammonium Nitrate - Phosphate Fertilizer Plant, L-2
Plant Capacity (25-15-0 to 28-14-0 Range), 2370 Tons/CD
Precipitator Effluent Temp., 150°F
Filter Effluent Temp., 100°F
Neutralizer Effluent Temp., 150°F
Evaporator Salt Concentration (28-14-0), 99.7 wt %
Prilling Tower Recycle Ratio, lb Recycle/lb Product, 1:4
Neutralizer NH3:H3PC>4 mole ratio, 1.37
f. Nitric Acid Plant, L-l
Capacity (100% HNO3 Basis), 1276 T/CD
Capacity ( 60% " ), 2126 T/CD
g. Miscellaneous Storage Facilities
Liquid Ammonia Storage (10 days consumption), 2,765,000 gal
(insulated atmos. pres. storage @ -28°F)
Phosphate rock storage (2 days consumption), 2500 Tons
60% Nitric Acid Storage (2 days production), 750,000 gal
Fertilizer Product bulk storage (90 days production), 215,000 Tons
h. Settling Pond
Basis: 10 years combined production of gypsum and recovered fly ash
Fly Ash 3096 Acre-Feet
Gypsum 4644 " "
3. Process Design Rationale
The TVA report discussed three alternate methods for the
production of the intermediate product, ammonium sulfate. These
are complete oxidation of the sulfite and bisulfite products to
sulfate in the scrubber proper, external oxidation of the scrubber
effluent in a separate vessel and acidification of the scrubber
effluent with H2SO4. Some production of sulfate occurs in the
scrubber and the various factors which inhibit or promote the
production of sulfate have not been fully evaluated. This reaction
does tie up ammonia in the system thereby reducing the efficient
removal of SC-2 .
-------
In view of the many unanswered questions and lack of experi-
mental data this procedure was not considered to be a workable
system at this time, however it is planned to evaluate partial
oxidation in the scrubber up to an optimum point to minimize the
size of the external oxidation vessel.
Production of ammonium sulfate by acidification with H2S04
was discussed from the standpoint of the pilot plant work done by
TVA and the commercial scale operation of Cominco. however no
definite reason was given for not using this method. Presumably
oxidation in a separate vessel was chosen as the preferred method
for producing (NH4)2SC>4 primarily because this method has been
proven commercially in Japan and complete details for carrying out
this reaction are available under license from the Japanese Engine-
ering Consulting Co. (JECCO) . A fourth method which has potential
and which will be evaluated is the crystallization .by cooling, of
from a scrubber sidestream.
The technology involved in the production of ammonium nitrate-
phosphate fertilizer from (1^4)2304, phosphate rock and HNO3 has
been well established both in the U.S. and Europe therefore further
evaluation was not considered to be necessary to carry out this
study. Various types of scrubbers were looked at including packed,
venturi, spray, mobile bed, orifice and sieve tray towers. The
selection was made on the basis of the following criteria;
(a) an overall pressure drop less than 12" H2O at 90% S02
absorption.
(b) High efficiency particulate removal, minimum 99.5%.
(c) resistant to plugging by solids in the gas stream.
(d) a high turndown ratio.
(e) adaptable to stagewise contacting of gas and liquid
with complete segreation of each recirculation stream
to permit controlling NH3 concentration on each tray
for the purpose of minimizing NH3 loss. This require-
ment was considered to be the single most important
criterion .
-------
The crossflow sieve tray (impingement type) appeared to meet
these basic requirements better than the other tower types.investi-
gated. Further this type of scrubber is available in the large
size required for 500 and 1000 MW power plants which is not the
case for a number of the tower types investigated. The plugging
problems experienced with this type of scrubber in the wet limestone SC>2
removal process which involves injecting a limestone slurry on the
top tray apparently do not arise in the NH3 system since essentially
all solids in the gas stream are removed on the bottom tray with
water sprays. This type of scrubber is normally supplied with a
fixed vane type mist remover. The entrainment removal efficiency
of this type of demister is low compared to wire mesh or fiber
bed demisters. Therefore if provisions are made to prevent plugg-
ing the latter type of demister may be more effective in removing
entrained salts which could possibly cause trouble in the gas
reheater or exhaust fan.
A cost comparison was made of five methods for cooling and
reheating the flue gas. The indirect liquid to gas system proved
to be the lowest priced in terms of $/ton of coal burned. The TVA
pilot plant study of NH3 scrubbing involved use of a 2' x 10'
high packed tower. Variables studied were circulation rate, pH,
solution concentration, packing depth and gas velocity. The
results of these tests plus data obtained from articles pub-
lished by Johnstone, Chertkov and others provided the design
basis for the scrubber used in this study. Further pilot plant
evaluation is required to firm up the optimum scrubbing variables
including scrubber type, gas velocity, mist remover type, inlet
gas temperature and the point of NH3 addition.
4. Process Design Appraisal
This process which combines NH3 scrubbing with fertilizer
production was designed to determine the economic potential of a
recovery type process as applied to S02 removal processes. Process
"A" which shows the best economics of the three processes evaluated
does not offer a very attractive investment potential. Since it is
known that plant cost estimates increase substantially as engine-
ering becomes more definitive it is reasonable to predict that the
-------
rate of return or payout for this process based on a more definitive
engineering study would be even less attractive than that shown in
'the TVA report. This approach does point up the relative diffi-
culty in attempting to fit any type of recovery scheme whether
the end product is sulfur, sulfuric acid or fertilizer into the
general economy without disrupting the existing supply demand
pattern. In the process under evaluation the fertilizer plants
would have to be located in the midwest area of the U.S. where
marketing conditions are most favorable. The number of such plants
would need to be limited to 2 or 3, 500 MW plants in order not to
over saturate the existing feetilizer markot so that even under
the most favorable conditions this process has very limited appli-
cability.
The ammonia scrubbing process is a relatively cheap and
effective method .for removing S02 from flue gas. It has a high
absorption efficiency and requires fairly ;•; implo handling tech-
niques in the scrubbing and regeneration steps. A disadvantage
of this process which is common to all of the wet scrubbing
processes is that the flue gas must be cooled and reheated.
As compared to limestone/ammonia is expensive and therefore
must either be regenerated for recycle or become a constituent
of a saleable product. In the process under evaluation all of
the nitrogen remains in the product resulting in a fertilizer
with a high nitrogen to phosphate ratio. This type of fertilizer
has a lower market demand than fertilizers with lower nitrogen
to phosphate ratios.
The ammonia scrubbing process is well developed and has
been tested extensively in Russia, Japan and France. However it
is still felt that substantial additional pilot plant testing is
justified to optimize the operating variables of the process.
TVA currently plans to evaluate, (a) the degree of oxidation
taking place in the scrubber and the variables which affect
oxidation, (b) optimum bisulfite/sulfite ratio for best scrubber
efficiency, (c) determination of how much dust can be tolerated
-------
in the subsequent oxidation and fertilizer production steps,
(d) steady state composition of the scrubber liquor and the
effect of dissolved solids buildup on scrubber operation, (e)
pH control for corrosion prevention.
The oxidation reaction and the fertilizer plant and auxiliary
plants represent established technology proven on a commercial
basis and therefore do not require further evaluation.
The nitric acid plant required for a 1000 MW power plant
has a capacity of 1276 T/CD (100% HN03 Basis). The ammonia
required for the combined scrubbing operation and nitric acid
plant is 716 T/CD. This amount of NH3 would normally be produced
on site rather than being shipped in by rail. This consideration
may place some further restriction on the location of a plant
of this type.
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B. CHEMISTRY
The chemistry associated with the ammonia scrubbing
process, for sulfur dioxide removal from stack gas, falls
into three separate categories, namely:
(a) absorption
(.b) recovery
(c) production of finished product.
(a) In the first of these, sulfur oxides are removed
from the stack gas by absorption into a concentrated aqueous
solution of ammonium sulfite. The chemical reaction shown
in equation 1 describes the principle chemistry of the
scrubbing process.
(NH4)2S03 + S02 + H20 = 2 NH4HSO_ (equation 1)
Detailed and reliable thermodynamic and kinetic data are
available for this system. These are sufficient for ade-
quate definition of the process parameters as well as for
selection of optimum conditions for most efficient sulfur
dioxide removal. Equilibrium data indicate that C0~, pre-
sent as a major component in stack gas, can not effectively
interfere with the efficiency of the SC^ removal. Varia-
tions of SC>2 partial pressure with temperature, solution
pH and solution concentration have been measured and ac-
curate data are available.
-------
(b) The recovery of sulfur values from the spent
scrubber solution can be accomplished in several ways.
The method of choice depends almost entirely on the type
of end product desired. For example, when sulfuric acid
is the end product, a regenerative process can be used
that essentially comprises the reversal of the scrubbing
reaction (equation 1) by means of air stripping at elevated
temperature. Regenerated solution is recycled to the
scrubber and the sulfur dioxide is used as feedstock to
;i sulfuric acid plant. In situ oxidation (equation 2)
must be held to a minimum for regenerative operation as
2(NH4)2S03 + 02 = 2(NH4')2S04 (equation 2)
this represents a net loss in scrubbing capacity. Chem-
ical factors required to control the extent of this
reaction are known.
When ammonium sulfate is the desired recovery product,
the oxidation reaction (equation 2) must be carried out in
a vessel separate from that employed in the scrubbing
operation since conditions that favor equations 1 and 2
are mutually incompatible.
(c) The chemistry involved in the production of the
finished product depends entirely on the nature of the end
product. However, for those cases where fertilizer materials
are to be produced, all of the required chemistry is well
known in the art.
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AMMONIA SCRUBBING (ALTERNATE A)
(STEAM STRIPPING REGENERATION)
A. PROCESS
Sulfur dioxide removal from power plant stack gas by scrubbing
. ith ammonia followed by regeneration by stripping and production
' sulfur dioxide for use in sulfuric acid manufacture.
1. Process Description (Refer to MWK Flow Sheet Dwg. No.
PD-113-D)
a* Sulfur Dioxide Removal from Flue Gas
Flue gas exits the air preheater at 300°F and is split into
t>
six parallel streams which flow through identical scrubbing systems
before being exhausted into a common stack. The gas is first
cooled to 172°F in gas cooler C-l. The cooled gas containing 4.9
gr/scf of solid particulates is scrubbed with water on the bottom
tray of a four-stage scrubbing tower E-l. 99.5 wt % of the solids
are removed and discharged to settling tank F-4 common to the six
scrubbers. The -solids are settled out and diluted with water to
a 12 wt % ash slurry which is pumped to a settling pond. The solids-
free gas flows countercurrent to the wash solution containing NH3,
S02, (NH4)2S03, NH4HS03 and (NH4)2SO4. Three stages are required
to remove 90 wt % of the SO2 by reaction with NH3 and (NH4)2S03.
The proposed scrubbers are countercurrent sieve tray type Impinjet
gas scrubbers manufactured by the W. W. Sly Mfg. Co.
Six scrubbers are required to handle the flue gas from a
1000 MW plant since the largest scrubber available is 20* wide
x 40' long. The total sidestream drawoff from each of the three
wash trays is segregated,flowing respectively to the recirculation
surge tanks F-l, 2, 3. Each tank is common to three scrubbers
resulting in two recirculation systems each of which feeds three
scrubbers via pumps J-l, 2, 3. Fresh NH3 makeup is added to each
recirculation stream and the liquid composition on the trays is
regulated by drawing off a slipstream from each drum flowing step-
wise from F-3 to F-2 to F-l. The scrubber effluent product is
-------
withdrawn from drums F-1A, B and pumped to regenerator E-<2. The
scrubber effluent consists of 48 wt % salt solution containing a
mixture of SC>2, NH3 and ammonium sulfite, bisulfite and sulfate.
The flue gas is cooled adiabatically in the scrubber by the evap-
oration of water and leaves the scrubber saturated at 116°F. Makeup
water is injected into the scrubber above the .top tray and liquid
entrainment is removed by a fixed blade impingement baffle located
in the top of the scrubber. The gas is reheated to 247°F in gas
reheater C-2 to ensure adequate dispersic-rt of the gas leaving the
stack. Flow through the scrubbers is maintained by exhaust fan J-9
located downstream of gas reheater C-2. The six fans exhaust to a
common stack and maintain the entire system under slight negative
pressure. A pumparound system utilizing surge tank F-5 and pump
J-8 circulates tempered water through cooler C-l and gas reheater C-2
in series transferring heat from C-l to C-2.
Regeneration of Scrubber Effliiuut and Production of
b. Sulfur Dioxide
The combined scrubber effluent from ,the six scrubbers is pumped
from tanks F-1A&B through preheaters C-4 and C-l, where it is heated
to- 185°F and discharged into the top of Regenerator E-2. The re-
generator is a packed column operated under vacuum at 185°F and
348 MMHg (absolute pressure) with an external steam heated thermo-
siphon reboiler providing the vapor boilup. The feed contains a
mixture of (^4)2203, NH4HSO3 and (NH4> 2S04 in aqueous solution
with an S/C* ratio of 0.85. The unstable NH4HS03 breaks down when
heated releasing NH3 and S02- These gases along with water vapor
are carried overhead. The regenerated stripper bottoms leave at 187°F
with an S/C ratio of 0.76. A slipstream comprising 2.8% of the regen-
erator effluent is pumped to the sulfate reactor D-l where it is reacted
wi£h CaCC>3 to form NH3, CC>2 and CaSC>4. This removes the net make
S MOLES SULFUR
C MOLES AMMONIA
(contained in (NH4> 2SO3 anc* NH4HSO3 only)
-------
of (NH4)2SC>4 produced in the scrubber. The reactor off gas con-
sisting of NH3, C02 and H20 vapor is cooled to 120°F in C~6 where
part of the water condenses and returns to the reactor. The remain-
ing vapors are recycled back to the scrubber via the sidestream
recirculation circuits. The insoluble CaS04 settles out and is
pumped to sluice tank F-8 where it is diluted with water to a 20
wt % slurry. It is then pumped to the settling pond. The balance
of the regenerator effluent is cooled to 120°F in C-8 and recycled
back to the scrubber recirculation drums F-l, 2, 3.
The regenerator overhead vapors are partially condensed in C-4
which exchanges heat with the cold scrubber effluent. Essentially
all of the NH3 vapors condense in C-4 along with S02 to form a weak
aqueous solution of (NH4)2SC>3. The condensate discharges to seal
drum F-6 from which it is pumped into the regenerator effluent
stream and then recycled back to the scrubber. The remaining vapors
are cooled to 122°F in C-5 Lo condense the bulk of the water vapors.
Approximately 8.5% of the S02 condenses along with the water and is
collect^*] in seal -iinm v -7. The remaining v.ipor which contains 90
wt % S02 is pumped via vacuum pump J-15 to a sulfuric acid production
unit. The condensato collected in drum F-7 is pumped to stripper
E-3 where the S02 is stripped out with compressed air and discharged
to the sulfuric acid production unit. Total S02 recovery is 15, 190
Ib/hr in terms of sulfur equivalent representing 68.8% of the sulfur
removed from the flue gas.
2. Process Design Basis
Basic process design criteria common to all of the SC>2 removal
processes being evaluated are not included in the following list.
Overall Plant Capacity Factor, % of Nameplate Rating
1st to 10th year 80%
llth to 15th year 57%
16th to 20th year 40%
21st to 35th year 17%
Average Capacity Factor 43%
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a. Ammonia Scrubbers E-1A, B, C, D, E, F
The following data pertaining to the scrubbers selected as
being most suitable for this service was obtained from .the W. Sly
Mfg. Co. bulletin and is applicable to the cross flow sieve tray
impinjet type gas scrubber.
Quantity (6) Six
Dimensions 20' x 40' (rect.) x 27' high
Design Gas Velocity, Nor. 420 fpm, max. 500 fpm, actual 461
Fly Ash Dust Removal 99.5% (total fly ash from boiler)
Water recirculation to bottom scrubber stage for dust removal
Combined plate and spray water, 3 gpm/1000 ACFM gas
b. Scrubber Operating Conditions
Gas inlet temp., 172°F
Gas outlet temp., 116°F
F.ff.lii-nt temp. , .120°F
Affluent salt concentration, 48 wt %
Effluent S:NH3 mole ratio, S/C = 0.85
"Salt solution recirculation rate, 1.06 gal/min-ft2
Fly ash slurry concentration to settling pond, 12 wt % solids
c. Regenerator E-2
Liquid inlet temp., 185°F
Liquid outlet temp., 187°F
Tower top abs. pres. 348 MMHg
Tower Bottom abs. pres. 363 MMHg
Feed composition, (S/C)^ = 0.85
Effluent composition, (S/C)F = 0.76
Theoretical stages, 5
Liquid flow rate L, 1.15 lb/sec-ft2
Gas flow rate G, 0.26 lb/sec-ft2
Packing type, 3" Intalox saddles
d. Sulfate Reactor, D-l
Operating temp., 187°F
Operating Pres., 18 psia
Retention Time, 30 minutes
-------
e* SOg Stripper, E-3
Oper. Temp., 120°F
Oper. Pres., 5 psig
Pres. drop thru internals, 0.25 psi
Stripping medium, air (8 moles air/mole S02)
Packing type, 2" Intalox saddles
Liquid flow rate L, 7.1 lb/sec-ft2
Gas flow rate G, 0.315 lb/sec-ft2
f. Settling Pond L-l
Basis: 10 years combined production of gypsum and recovered fl,y
ash.
Fly ash, 3096 acre-ft
Gypsum, 854 acre-ft
3. Process Design Rationale
The scrubbing operation, with proper modification of operating
conditions, Is assumed to be capable of producing an effluent with
an S/C ratio of 0.85 using the three stage scrubber selected for the
TVA base case. Literature references indicate that a ratio of 0.92
or higher is attainable. If true this would enhance the regeneration
aspects of the process and show a more profitable overall operation.
Scrubbing liquid sidestream recirculation rates are based on the test
data of Chertkov reproduced in the TVA report (1). This data was
obtained from a 6 plate sieve tray absorber with an effluent S/C
ratio of 0.90. A sulfate make of 15% of the total sulfur absorbed
was assumed to be constant for all sulfate concentrations in the
recirculation liquid. Since the production of sulfate involves a
simple oxidation reaction this should be a reasonably accurate
assumption.
The stripper design is based on the Johnstone vapor pressure
formulas for NH3, SC>2 and I^O over an aqueous sulfite, bisulfite
solution. The Russian references claim that the Johnstone formulas
predict higher than actual steam consumption rates in the stripper.
(1) "Sulfur Dioxide Removal From Power Plant Stack Ga.s, Ammonia
Scrubbing", Conceptual Design and Cost Series Study No. 3, 1970,
Contract No. TV 2$233A.
-------
Since we cannot reconcile these two methods of predicting vapor
pressure we have retained the Johnstone method as given in. the
TVA base case report. The theoretically possible separation value
(limiting S/C ratio) attainable in the stripper and the practically
attainable separation based on temperature and number of transfer
units were calculated by the method of Gordeev and Chertkov (2).
Calculations made to find rough optimum operating conditions indi-
cated that 85°C was better than 95°C (stripper top temperature) for
the assumed feed composition and that five stages gives maximum
separation efficiency. The final design is based on tower top
conditions of 85°C at 348 MMHg absolute pressure using a packed
column with 5 theoretical stages.
h
The net sulfate make is removed by reacting with CaC03 to
form CaS04 which is discarded and NH3 which is recycled to the
scrubber. Since the stripper effluent contains a mix of sulfite,
bisulfite and sulfate some quantity ol: sulfite and bisulfite will
be reacted along with the sulfate thereby adding to the total
sulfur Lhrowaway. The quanLity oi sulfur lost in this manner can
be arbitrarily set by fixing the sulfate/sulfite-bisulfite ratio
and consequently the circulation rate. Taking the relative value
of the sulfate/sulfite-bisulfite ratio as 1/1 at a zero recircula-
tion rate gives a net sulfur recovery of 10% of the total absorbed.
A ratio of 5/1 results in a recovery of about 69% and a 10/1 ratio
in 76% recovered. The 5/1 ratio is about optimum since further
increases in sulfur recovery result in a disproportionate increase
in the liquid recirculation rate. The final design shows a total
equivalent sulfur recovery of 68.8% of the sulfur absorbed in the
scrubber. To accomplish this it is necessary to air strip the weak
S02 water solution which condenses out of the SO2-rich stripper
overhead vapors. Approximately 6% of the total sulfur equivalent
is recovered in the air stripper.
(2) "Investigation of the Steady State Characteristics of a Process
for Obtaining 100% S02 from Ammonium Bisulfite", Gordeev, L. S.
and Chertkov, B.A. (Vol. 7, No. 4) Inter. Chem. Eng.
-------
The reaction of the sulfate, suifite, bisulfite solution
with CaCC>3 is assumed essentially 100% complete thus eliminating
the need for a filtration step to recover ammonia values in the
filtrate.
4. Process Design Appraisal
The ammonia scrubbing process is a relatively cheap and
effective method for removing S02 from flue gas. It has a high
absorption efficiency and requires fairly simple handling techniques
in the scrubbing and regeneration steps. A disadvantage of this
process which is common to all of the wet scrubbing processes is
that the flue gas must be cooled and reheated.
As compared to limestone, ammonia is expensive and therefore
must either be regenerated for recycle or become a constituent of
a saleable product. In the process under evaluation all of the
nitrogen in the scrubber effluent is regenerated either as (NH4>2SQ3
or NH3 and recycled to the scrubber. Within the time and scope
limitations of this study an attempt was made to select realistic
operating conditions which would approach the optimum. Further
study is indicated in the following areas to define the optimum
process both with respect to cost and efficiency:
a. Establish the maximum S/C ratio attainable in the
scrubber compatible with scrubber cost, ammonia
losses, etc.
b. Establish the temperature-pressure conditions for
carrying out the regeneration operation most effi-
ciently.
c. Establish the optimum sulfate/sulfite-bisulfite
ratio in the regenerator feed. This involves a
comparison of the value of the recovered sulfur
with the cost of higher recirculation rates.
d. Investigate the merits of direct steam injection
versus indirect heating in the regeneration step.
-------
e. Reconcile, by laboratory teats, the vapor pressure
calculations of Chertkov and Johnstone for the
HN3-S02-H20 system.
f. Consider alternate methods of recovering the
sulfur value from the sulfate rather than the
proposed throwaway as CaS04.
, Note: In addition to the references specifically cited in
the text, the following two rofer-ence.n also were used
in this ovr'liMi inn.
i — .
"Recovery of S02 from Waste Gases", Johnstone, H. F.,
IEC, Vol. 29 (1937). ~
"Recovery of S02 from Waste Gases", Johnstone, H. F.,
i and Keyes, D. B., IEC, Vol. 27 (1935).
i
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AMMONIA SCRUBBING (ALTERNATE B)
(THERMAL DECOMPOSITION REGENERATION)
A. PROCESS
Sulfur dioxide removal from power plant stack gas by scrubbing
with ammonia followed by acidification with NH4HS04 and thermal
decomposition of (NH4)2S04 for recovery of NU.3. Sulfur dioxide
produced in the acidification step is used in the production of
suifuric acid.
1. ProcessDescription (Refer to MWK Flow Sheet Dwg. No.
PD-114-D)
a. Sulfur Dioxide Removal From Flue 'Gas
Flue gas exits the air preheater at 300°F and is split into
six parallel streams which flow thru identical scrubbing systems
before being exhausted into a common stack. .The. gas is first
cooled to 172nF in gas cooler C-l. The cooled gas containing 4.9
gr of solid particu.lat.es/SCF is scrubbed w: I h w*.ter on the bottom
tray of a four-stage scrubbing tower E-l. 99.5 wt % of the solids-
are removed and discharged to settling tank F-4 common to the six
scrubbers. The solids are settled out and diluted with water to
a 12 wt % ash slurry which is pumped to a settling pond. The solids
free gas flows countercurrent to the wash solution containing NH3,
SO2, (NH4)2S03, NH4HS03 and (NH4)2SO4. Three stages are required
to remove 90 wt % of the SO2 by reaction with NH^ and (NH4)2SO3.
The proposed scrubbers are countercurrent seive tray type Impinjet
gas scrubbers manufactured by the W. W. Sly Mfg. Co.
Six scrubbers are required to handle the flue gas from a
1000 MW plant since the largest scrubber available is 20! wide
x-401 long. The total sidestream drawoff from each of the three
wash trays is segregated flowing respectively to the recirculation
surge tanks F-l, 2, 3. Each tank is common to three scrubbers
resulting in two recirculation systems each of which feeds three
scrubbers via pumps J-l, 2, 3. Regenerated NH3 is added to each
recirculation stream and the liquid composition on the trays is
-------
regulated by drawing off a slipstream from each drum flowing step-
wise from F-3 to F-2 to F-l. The scrubber effluent product is
withdrawn from drums F-1A,.B and pumped to acidifier/stripper E-2.
The scrubber effluent consists of 48.5 wt % salt solution containing
a mixture of SC>2, NH3 and ammonium sulfite, bisulfite and sulfate.
The flue gas is cooled adiabatically in the scrubber by the evapora-
tion of water and leaves the scrubber saturated at 116°F. Makeup
water is injected into the scrubber above the top tray and liquid
entrainment is removed by a fixed blade impingement baffle located
in the top of the scrubber. The gas is reheated to 247°F in gas
reheater C-2 to ensure adequate dispersion of the gas leaving the
stack. Flow thru the scrubbers is maintained by exhaust fan J-9
located downstream of gas reheater C-2. The six fans exhaust to a
common stack and maintain the entire system under slight negative
pressure. A pumparound system utilizing surge tank F-5 and pump
J-8 circulates tempered water thru gas cooler C-l and gas reheater
C-2 in series transferring heat from C-l to C-2.
b. Regeneration of Scrubber Effluent for Recovery of
Ammonia and Sulfur Dioxide
The combined effluent from the six scrubbers is pumped from
tanks F-1A, B and heated to 158°F in preheater E-3. The scrubber
effluent is then combined, by line mixing, with the molten salt
slurry from the electric furnaces B-lA, B, C, D, E and discharged
into the top section of the acidifier/stripper column E-2. The
acidifier/stripper is a packed column with top outlet operating
conditions of 158°F and 378 mm Hg absolute pressure. Thermosiphon
reboiler E-4 heated by the overhead vapors from evaporator/crystal-
lizer E-3 provides vapor boilup. The acidifier/stripper column is
primarily a mixing and heating device to enable the following reactions
to rapidly proceed to completion;
1. NH4HS03(AQ SOLN)+NH4HS04(AQ SOLN) HE-AT (NH4 ) 2S04 (AQ SOLN) +SO2 t+H20 (1)
HFAT
2. (NH4)2S03 (AQ) +2NH4HSO4 (AQ) 5 2 (NH4)2SO4 (AQ) +S02++H2O (1)
-------
The acidif ier/stripper effluent consisting of (NH4)2S04 and excess
NH4HS04 in aqueous solution containing a small amount of dissolved
S02 is pumped thru preheater C-7 where it is first heated to 266°F
and then discharged into evaporator /crystallizer E-3. The overhead
vapor from E-2 is cooled with water to 122°F in C-6 where most of
the water condenses out and drains to condensate drum F-6. The
remaining S02-rich vapor is pumped via vacuum pump J-17 to a sulfuric
acid plant L-2. Total sulfur recovery is 18,473 Ib/hr representing
83.7% of the sulfur removed from the flue gas. The acid plant
produces 692 tons/CD of 98% acid. The evaporator/crystallizer E-3
is heated by forced circulation reboiler C-10 which vaporizes 96%
of the water of solution along with the dissolved S02« The vapor
from E-2 flows in series thru exchangers C-3, "C-4 and C-5 where it
is condensed and cooled. The combined condensate from exchangers
C-3, C-4, C-5 and C-6 is collected in condensate drum F-*6 from
which it is pumped to the sulfate reactor D-l where the 579 Ib/hr
of dissolved S02 is neutralized with CaC03 .
The wet ammonium salts which crystallize out in E-2 are con-
veyed to the electric furnaces B-lA, B, C, D, E. The following
decomposition reaction takos place in the furnaces;
(NH/|)2S04(G) juoj,, NH4HS04(.l)+NH3t
Higher temperature will promote undesirable secondary decomposition
reactions producing unusable ammonium salts. By maintaining the
temperature at 700°F and by injecting 700°F superheated steam through-
out the furnace melt these secondary reactions are minimized.
Approximately 80% of the (NH4)2SC>4 in the feed is converted to
NH4HSO4 . The furnace effluent stream containing 83 mole % NH4HSO4
and 17 mole % (NH4)2SC>4 is recycled to the acidif ier stripper after
first being cooled to 350°F in exchanger C-7. A slipstream compris-
ing 9 wt % of the total furnace effluent is pumped to the sulfate
reactor D-l after first being mixed with the S02/H20 solution from
condensate drum F-6 and cooled to 175°F in exchanger C-9. The
ammonium salts contained in this stream are reacted with CaCOs for
-------
removal of the net sulfate make in the scrubbers. The following
reactions take place in the sulfate reactor;
4-HKAT
1. (NH4)2S04(AQ SOLN)+CaC03(s) + CaS044- + 2NH3t+C02++H20
—HP AT
2. NH4HS04(AQ SOLN)+CaC03(s) + CaS044-+NH3++C02t+H20 (1)
3.
The net overall heat effect in the sulfate reactor is exothermic
requiring the use of water cooled internal coils for heat removal.
The resulting insoluble CaSO4 is settled out and pumped to the
settling pond. The vapor from the reactor containing NH3, CO2
and H20 is combined with the vapor from the electric furnaces con-
taining NH.3 and H20 and is then cooled to 120°F in exchanger C-8
where most of the water condenses out and is sewered. The non-
condensible NH3 and CO2 vapor is recycled back to the scrubbers.
2. Process it(v,i'in iVi."vi. :•;
Basic procosy design criteria common to all of the S02 removal
processes being evaluated are not included in the following list.
Overall Plant Capacity Factor, % of Nameplate Rating
1st to 10th year 80%
llth to 15th year 57%
16th to 20th year 40%
21st to 35th year 17%
Average Capacity Factor 43%
a. Ammonia Scrubbers E-1A, B, C, D,E, F
The following data pertaining to the scrubbers selected as
being most suitable for this service was obtained from the W. Sly
Mfg. Co. bulletin and is applicable to the cross flow sieve tray
impinjet type gas scrubber.
Quantity (6) Six
Dimensions 20' x 40' (rect.) x 27' high
Design Gas Velocity, Nor. 420 fpm, max. 500 fpm, actual 461 fpm
-------
Fly Ash Dust Removal 99.5% (total fly ash from boiler)
Water recirculation to bottom scrubber stage for dust removal
Combined plate and spray water, 3 gpm/1000 ACFM gas
b. Scrubber Operating Conditions
Gas inlet temp., 172°F
Gas outlet temp., 116°F
Effluent temp., 120°F
Effluent salt concentration, 48.5 wt %
Effluent S»NH3 mole ratio, S/C =0.85
Salt solution recirculation rate, 1.06 gal/min-ft2
Fly ash slurry concentration to settling pond, 12 wt % solids
c. Acidifier/Stripper E-2
Liquid inlet temp., 215°F
Liquid outlet temp., 160°F
Tower top abs. pres., 378 mm Hg
Tower bottom abs. pres., 398 mm Hg
Feed composition, (S/C)j = 0.85
NH4HS04 in feed * 1.25 x stoichiometric requirements
Theoretical stages, 4
Liquid flow rate L, 1.17 lb/sec-ft2
Gas flow rate G, 0.423 lb/sec-ft2
Packing type, 3" Intalox saddles
d. Sulfate Reactor, D-l
Operating temp., 185°F
Operating Pres., 18 psia
Retention time, 30 minutes
CaCO3 in feed, 1.25 x stoichiometric requirements
e. Evaporator/Crystallizer, E-3
Oper. Temp., 212°F
Oper. Pres., 930 mm Hg
Retention time, 22 minutes
Reboiler recirculation rate, 415 gpm
% of water vaporized, 96
-------
f. Electric Furnace, B-lA, B, C, E, E
Operating temp., 700°F
Operating Pres., 3 psig
Feed temp., 212°F
(NH4)2S04 conversion, 80%
Electrical Power Consumption, 350 KWH/ton of (NH4)2S04 in feed
Steam injection rate, 0.2 Ib stream/lb (NH4)2S04 in feed (70Q°F
superheated)
Design capacity of unit, 5000 KW
Number units required, (5) five
Total working capacity, 22400 KW
Power supply characteristics, 3 phase, variable voltage 120 to 164 V
Furnace dimensions, 20' Dia. x 10' Deep
g. Settling Pond L-l
Basis: 10 years combined production of gypsum and recovered
fly ash.
Fly ash, 3096 acre-ft
Gypsum, 576 acre-ft
3. Process Design Rationale
A scrubber effluent with a high S/C ratio is beneficial to
the acidification/stripping process since a high Ni^HSO^ content
would require fewer moles of NH4HSO4 for acidification thereby
reducing the duty of the thermal decomposition furnaces. Therefore
the scrubbing operation as used for alternate A case (regeneration
by steam stripping) which had a scrubber effluent with an S/C
ratio of 0.85 is used without change for the alternate B design.
The operating conditions for the acidification/stripping column
were chosen to give a favorable S02/H2O ratio in the overhead vapor
stream and to permit utilization of a part of the heat available in
the evaporator/crystallizer overhead vapor stream for preheating
the feed stream. The regenerated molten NH4HS04 stream and the
scrubber effluent are line mixed prior to injection into the
acidifier/stripper to facilitate heat transfer and promote the
acidification reactions. This will more closely simulate the
basic design assumption that the mixture of SO2 and H20 vapor is
-------
in equilibrium with an aqueous solution of S02 and sulfates. The
NH4HSC>4 feed rate to the acidifier/stripper is set at 1.25 times
the stoichiometric requirements to ensure complete reaction of
the sulfites.
The performance data pertaining to the electric furnaces used
for decomposition of the (NH4) 2804 were obtained from the report()
covering an experimental plant for producing alumina. The perform-
ance of these furnaces which is critical to this process seems to
be well enough established to remove any doubts as to the feasibil-
ity of the process. The important operating parameters covered in
the report are 80% conversion of (NH4)2S04 to NH4HS04 at 700°F with
minumum secondary reactions provided superheated steam is injected
into the melt to retard such reactions and an actual power con-
sumption of 350 KWH per ton of (NH4)2S04 charged to the furnace.
The combined condensate from the acidifier/stripper overhead and
evaporator/crystallizar overhead streams contains such a relatively
small amount of S02 (579 Ib/hr, 0.8 wt %) that it does not warrant
further treatment for recovery of the S02 value. Therefore this
stream is combined with the slipstream from the electric furnace
effluent and reacted with CaCO3 in the sulfate reactor. CaCC>3
feed to the sulfate reactor is 1.25 x stoichiometric to ensure
complete recovery of the NH3 from the ammonium salts.
4. Process Design Appraisal
The ammonia scrubbing process is a relatively cheap and
effective method for removing SC>2 from flue gas. It has a high
absorption efficiency and requires fairly simple handling tech-
niques in the scrubbing and regeneration steps. A disadvantage
of this process which is common to all of the wet scrubbing processes
i§ that the flue gas must be cooled and reheated.
(1) Engineer-Contractor's Report on Alumina-From-Clay, Experi-
mental Plant (Plancor 1865) at Salem, Oregon. Chemical
Construction Corp., N.Y., N.Y.
-------
As compared to limestone, ammonia is expensive and therefore
must either be regenerated for recycle or become a constituent of
a saleable product. In the process under evaluation all of the
nitrogen in the scrubber effluent is regenerated as NH3 and recycled
to the scrubber. Within the time and scope limitations of this study
an attempt was made to select realistic operating conditions which
would approach the optimum. Further study is indicated in the
following areas to define the optimum process both with respect to
cost and efficiency!
a. NH3 Scrubber, establish the maximum S/C ratio attain-
able in the scrubber compatible with scrubber cost, ammonia loses,
etc.
b. Acidjfier/Stripper, establish the optimum operating
temperature and pressure. Determine the best method of handling
the molten effluent from the electric furnaces particularly with
respect to premixing, cooling and injection into the acidifier/
stripper. Determine the most suitable vessel internal configura-
tion to obtain adequate mixing for rapid and complete sulfite
reaction. Determine actual composition of overhead vapor in
equilibrium with non-idealized liquid in vessel. Determine the
minimum excess NH4HS04 required for complete reaction of sulfites.
c. Electric Furnace, establish the optimum operating
temperature, and steam injection rate compatible with percent
conversion, extent of secondary reactions and power input. De-
termine required residence time for equipment sizing.
d. Sulfate Reactor/Settler, establish optimum reaction
temperature, residence time and excess CaCOs requirements for
reaction to go to completion. Determine the most satisfactory
slurry concentration with respect to handling characteristics.
e. NH3 Losses; determine overall system NH3 losses.
-------
Additional References
1. U.S. Patent 2,405,717, Recovery of Acidic Gases,
A. W. Hixson and R. Miller (1946).
2. "Sulfur Dioxide Removal Prom Power Plant Stack Gas,
Ammonia Scrubbing", Conceptual Design and Cost Series
Study No. 3, 1970, Contract No. TV 29233A.
-------
NOTE-EQUIPMENT INSIDE THIS ABO 15
PART or POWER PLANT & NOT TO
BE INCLUDED IW ESTIMATE.
i T/HR CAPACITY ' /
K'WCLUDCS ALL EQuiPMCNT i/ (
LwiTtmi_pASMEO_ARCA)_K —I
226PM ' _SfCl..UJ_SI,300t.»t
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TYCO PROCESS
A. PROCESS
1. Process Description (See MWK Dwg No. PD-116-D)
The process described below is based on the TYCO Modified
Chamber Process for the removal of S02 and NQX from electric power
plant flue gases (1). The sulfur and nitrogen oxide pollutants are
removed as sulfuric and nitric acids. The flue gas to be treated
is based on the burning of 3.5 wt % sulfur coal in a 1000 megawatt
coal-fired power plant. Sulfur levels in the flue gas are 0.261
mole % S02 and 0.005 mole % 803; NOX is assumed to be 500 pprn, by
volume. Removals of 90% SOX and 80% NOX are assumed.
*
Flue gas (95% fly ash-free after electrostatic precipitation)
is split into four parallel trains, and the pressure increased via
flue gas fans, J-la-d, from 14.05 psia to 16..05 psia to supply the
necessary head to move the gas through the downstream equipment.
The flue gas is then injected with an N02-rich (10 mole %) gas
recycled from the downstream catalytic stripper/ E-2. Enough N0£
is injected into the flue gas such that .about ,90% of the S02 is
oxidized to 803 in S02 oxidizer reactor, D-l., with a resulting N02
to NO ratio of 1 to 1. The reaction is:
SO2 + N02 = "S03 * NO (1)
The 803 combines with water vapor to form H2S04. Any oxidation of
NO to N02 is assumed negligible - a reasonable assumption consider-
ing the 300°F temperature and low oxygen partial pressure.
The effluent from D-l, at 328°F, is sent to N203 scrubber, E-l,
where NOX is absorbed as N203 and the H2S04 recovered via condensa-
tion of 503 and H20. Over 98% absorption of ®2°3 *~s rec3uired in
oirder to achieve the overall 80% NOX recovery; this is owing to the
(1) Gruber, A., and Walitt, A., "Oxidation of SO2" - Final Report,
Contract CPA 70-59 for NAPCA, 1970.
-------
fact that 11-fold as much NOX is recycled as is present in the
original flue gas. The NO2-NO is absorbed by contacting with
267°F, 80% H2S04 in the packed scrubber, E-l. The absorption-
reaction produces 0.3 wt % nitrosylsulfuric acid dissolved.in
80 wt % H2S04 according to the following reaction:
NO + N02 + 2H2S04 = 2HNSOs + H20 (2)
The scrubbed gas passes through a high velocity mist eliminator,
L-2, to recover entrained sulfuric acid mist before the flue gas
enters the stack. The scrubber overhead temperature is set such
that the partial pressure of water vapor over 80% H2S04 will result
in a scrubbing operation in which there will be no unnecessary
absorption or stripping of water. The partial pressure of H2S04
over 80% acid (1) at 267°F yields a 66 ppm H2S04 concentration in
the flue gas - corresponding to a loss of about one ton per hour.
The scrubber, E-l, will also pick up the fly ash which has not
been removed in the precipitator. According to experimental work
performed by TVA (2) about 20% of the fly ash constituents are soluble
i.n hot, 80% sulfuric acid. (The saturation point is not reported.)
The undissolved asli hat; Lo bo removed irom the circulating acid before
it is sent downstream to the catalytic stripper. The process flow-
sheet shows 272°F bottoms liquid from E-l pumped by J-2 to an undefined
"black-box", L-l., which contains the ash removal facilities. A 50%
by weight ash/acid mixture is shown leaving L-l and is neutralized
by addition of sluice water and limestone. About 3100 Ib/hr of acid
is lost to neutralization with the ash. The undissolved ash amounts
to less than .005 wt % in the liquid acid. A more detailed study of
solids/liquid separation' at such low solids levels is needed before
this section can be specified. The dissolved ash will build up until
the purge rate of dissolved ash in the sulfuric acid product equals
the input ash rate with the flue gas. The build-up is to about 0.9
wt-% dissolved ash in the circulating acid.
The solid-free acid leaves the ash removal facilities and
flows to F-4 surge tank. This sulfuric-nitrosylsulfuric mixture
is then pumped via J-3 to the top of the oxidizing stripper, E-2.
(1) Gmitro, J.I., and Vermeulen, T., "Vapor-Liquid Equilibria for
Aqueous Sulfuric Acid", UCRL-10886, Univ. Calif., Berkeley,
June 24, 1963.
(2) TVA, "Sulfur Oxide Removal From Power Plant Stack Gases - TYCO
Process", Preliminary Conceptual Design and Cost Study (Draft)
Contract TV-29233A, 1/18/71.
-------
The liquid flows over a bed of granular activated carbon counterr-
current to a stream of air. The TYCO concept indicates £hat this
activated carbon will act as a catalyst to promote the oxidation
of the dissolved HNSOs to gaseous N02 and H2S04:
4HNS05 + 2H20 + 02 = 4H2S04 + 4NQ2 (3)
CAT
An extremely high stripping efficiency of 99% and an off-gas from
the stripper of about 10% N02 are required. The high efficiency
gives the required NO -free scrubbing acid to be recycled to E-l,
and an NO -free H2S04 product. The 10% NO2 level allows for the
economic production of HN03 by cold absorption of N02 into water.
The necessary air for the reacting/stripping operation is
supplied by air blower J-5, operating at a AP of 30" w.g. A
slip-stream of the 80% H2S04 from E-2 is cooled to 100°F in product
cooler C-2 and sent to storage. Net product acid amounts to 80,920
Ib/hr of 80% acid. The remaining hot acid from E-2 is pumped by
J-4 to recycle acid cooler, C-l, and then scrubber E-l to complete
the acid circulating loop. Cooler C-l cools the acid from 272°
to 267°F, by exchange with boiler feed water.
About 90% of the N02-rich effluent gas from E-2 is recycled
to SO2 Qxidizer D-l to provide the necessary stoichiometric amount
of N02 for the SO2 oxidation and for the equimolar N02/NO absorption
in E-l. The remaining 10% of the stripper overhead is sent to a
single train nitric acid plant where the NOX values from the flue
gas are recovered as 60% nitric acid.
The purge of gas from the stripper at 272°F is cooled in C-3
to 105°F. A vapor-liquid separation is made in drum F-l, from which
some 40 wt % (approx) HNO3 condenses out and is pumped up via J-8
to 100 psia and fed to the HNO3 absorption tower, E-3. The gas off
F-l is compressed to 105 psia in two-stage compressor J~6. Inter-
stage cooling, vapor/liquid separation, and condensed HNO3 recovery
are carried out in C-4, F-2, and J-9 between the compressor stages.
The gas exiting from the second compressor stage is at 351°F.
-------
This gas is cooled to 166°F by exchanging heat with the tail gas
from the HN03 Absorber. Final gas cooling to 105°F occurs, in
Cooler C-6 and condensed acid removed in vapor/liquid separator
F-3. Most of the HN03 is produced in absorption tower E-3 at 105°F
and 100 psia. Sixty percent nitric acid can easily be produced in a
standard nitric acid absorption tower containing bubble cap trays
and cooling coils on each tray (C-9). The exothermic reaction is:
3N02 + H20 • 2HN03 + NO (4)
The presence of oxygen at a reasonable partial pressure, and the
relatively low temperature, 105°F, affords substantial re-oxidation
of NO to N02 in the absorber. Tail-gas from the absorption tower
at 105°F and 90 psia is heated, by exchange with tower feed gas,
to 301°F in C-5, then expanded from 90 to 16 psia through a power
recovery turbine to deliver almost half of the horsepower needed
to drive the feed gas compressor, J-6.
Nitric acid from E-3 is treated in HN03 bleacher E-4 at 15
psia by countercurrent contact with hot air. The latter is supplied
by bleacher air compressor J-12, and steam heater C-8 which heats
the bleacher air to 275°F. Spent bleacher air and the expanded tail
gas are combined and recycled to the S02 oxidizer reactors.
The clean nitric acid from the bleacher is pumped to storage
by J-ll, first being cooled to 100°F in product acid cooler C-7.
Sixty % nitric acid product amounts to 12,300 Ibs/hr.
The process start-up involves the use of an initial inventory
of stored liquid NO2, which is bled into the S02 oxidizer as a
flashed gas when the recovery plant is put on stream. The start-up
inventory is replaced, whenever necessary, by purging a slip-stream
of" the HNOj absorption tower feed gas, and cooling to condense N02
via refrigeration unit L-3, and condenser C-10. A start-up steam
heater, C-ll, is used for heating up the large acid inventory as
each train is brought on stream, initially, or after a shutdown.
-------
2. Design Bases
The design bases for the process flowsheet are obtained
primarily from information presented in TYCO Lab's final report,
"Oxidation of S02M.(1) Secondary references include the draft
of TVA's report on the TYCO process (2), and., use of certain
M. W. Kellogg standards for process specifications. The important
design bases aret
S02 Level in Flue Gas
NOX Level in Flue Gas
as NO
as N02
SOX Removal (as H2S04)
NOX Removal (as HN03)
Flue Gas Temperature
S02 Oxidizer-Reactor Residence Time
Scrubber Packed Height:
Scrubber Overhead Temperature
Scrubber Packing
L/G in Scrubber
Catalytic Stripper Space Velocity
NO2 Concentration off Stripper
NO "
Nitric Acid Plant Operating Pressure
Nitric Acid Absorption Temperature
Nitric Acid Strength
Product Acids Storage
Sulfuric Acid Strength
0.261 mole %
500 ppm (vol)
95%
5%
90% .
80%
300°F
15 sec.
30 ft
267°F
3" Intalox Saddles
7 Ib liq/lb gas
10 liq volumes
Hr-Vol Activated Carbon
10%
0.1%
100 psia
105°F
60 wt %
30 days
80 wt %
(1) Gruber, A., and Walitt, A. "Oxidation of SO2" - Final Report,
Contract CPA 70-59 for NAPCA, 1970
(2) TVA, " Sulfur Oxide Removal From Power Plant Stack Gases -
TYCO Process", Preliminary Conceptual Design and Cost Study
(Draft) Contract TV-29233A, 1/18/71.
-------
3. Design Rationale
The rationale behind the design of the major equipment items
and process sections is cited below:
D-l, S02 Oxidizer-Reactor
TYCO's data indicate that at least 15 seconds residence time
is needed for the required oxidation of SO2 to S03 at 300°F. The
commercial reactor is assumed to operate under plug flow conditions,
and that no gas mixing problems are encountered upon introduction
of recycle N02 into the Jj'Lue gas stream. It is assumed that a
direct scale-up can be made from the TYCO bench-scale experimental
results, and the 15 seconds residence time can be used to calculate
the oxidizer volume.
E-l, N203 Scrubber
TYCO experimental work with NO-N02 absorption into hot 80%
su.lfaric acid was carried out in beds packed with 1/4 and 3/8"
saddles at L/G's of 4 to 18 Ib liquid/lb gas. The commercial
scale flow sheet indicates scrubber operation in a tower packed
with 3" saddles at an L/G of 7.7. The latter figure is approxi-
mately that L/G ratio suggested by TVA as the optimum for the
overall process. For our case time did not permit optimization
calculations to be made, so the value of 7.7 was selected as
representative; it is within the range of the experimentation.
TYCO's miniplant scrubbing results indicate that the pre-
sently specified 98.2% ^03 scrubbing efficiency would require
about 30 feet of packing based on their 3/8" packing column •
scaling up to 3—inch packing at the same liquid loading (in Ib/hr-
ft2) would require more packed height to effect the same absorption
efficiency. At the same L/G (7.7 Ib/lb) the miniplant scrubber
operated at L and G's of about 2100 and 270 Ib/hr-ft2 respectively,
and would require about 30 ft of the 3/8" packing to achieve the
desired absorption efficiency. The commercial scale scrubber is
designed for liquid and gas loadings of 13,500 and 1750 lb/hr-ft2,
respectively. Very rough examination of general packed tower data
-------
indicate that, even with the less efficient larger packing (in
terms of transfer unit heights), the large increase in liquid
loading (2100 to 13,700) should cause the commercial scale tower
to give the same scrubbing efficiency with 30 ft of packed height.
The scrubber, E-l, consists of four, 40 ft diameter units.
It was assumed that a 40 ft diameter would be the largest vessel
size practicable for a packed tower, and.the number of vessels
determined using a superficial gas velocity of about 500 ft/min.
E-2, Oxidizing Stripper
TYCO lab data showed that the best performance of the cata-
lytic oxidizing stripper occurred at about 290°F with an L/G
of about 16 and a space velocity of about 1. vol liquid/hr-vol cata-
lyst. Close to 99% denitration stripping .efficiency of NOX was
achieved. TYCO used a bed of granular activated carbon, 4 x 10
mesh, the largest commercially available ,acti,vated carbon. This
process evaluation also specifies the same grade (Witco 256).
Very optimistic conclusions had to be l^ade about the perform-
ance of this catalytic stripper in order to obtain an overall
reasonable process design. These are as follows:
1) A liquid-hourly-space velocity of 10 can be used t9
determine catalyst volume. TYCO attempted minipIant operation at
an LHSV of 69, but failed. Using the demonstrated bench-scale
LHSV of about 1 seems impractical for a commercial plant. It is
assumed for the present design that this catalytic bed can be
made to perform at a space velocity of 10.
2) Stripping efficiency will be 99% at 272°F with 99% of
the NOX stripped as NO2, and 1% as NO with no N20 formed.
3) Stripping air rate is set so as to give an overhead vapor
of 10% N02« Such an N02 level is deemed necessary for reasonably
economical nitric acid production downstream. Consequently, the
resulting L/G for the stripper is fixed at .about 165, 10-fold
higher than the laboratory operation. It is assumed that the
-------
stripping factor is sufficient to remove the NC>2 from the .liquid
acid. More air could be used to increase the stripping factor,
if necessary, but this would decrease the overhead NO2 concentra-
tion, which at 1% NC>2 seems to make the process unattractive.
4) The stripping reactor cross-sectional area was determined
on the basis of roughly calculated flooding characteristics for a
granular packed bed. The 4 x 10 mesh carbon has an average par-
ticle size range of about 1/4 to 1/16 inch. Crushed stone has a
wetted packing factor of about 2190 for 1/4-inch size, and 1/4-
inch carbon Raschig rings, about 1900. Optimistically, flooding
characteristics for the granular carbon were assumed to be the
h
same as for the 1/4-inch rings. The bed would flood at a Q of
about 80 Ib/hr-ft2 whereas TYCO operated at a G of around 6 Ib/hr-
ft^. The commercial stripper was sized for a gas loading of
58 lb/hr-ft2, or 70% of flooding. In order to have four trains, as
in the scrubber section of the flow sheet, and to keep a maximum
vessel diameter of about 40 feet, the strippers are specified as
horizontal vessels 24 feet in diameter by 80 feet long. For com-
parison, four vertical vessels would have to be 48 feet in diameter.
HNO-3 Plant
A somewhat standard nitric acid absorption plant at 100 psia
and 105° is specified for the recovery of NOX as aqueous HN03.
The purge rate of the 10% NO2 from the catalytic stripper can be
varied, as well as the nitric acid strength obtained by the absorp-
tion of NO2 into water. The purge to the nitric acid plant was
set on the basis that this stream should contain the stoichiometric
amount of N02 to give the required HN03 production according to the
reaction:
3N02 + H20 = 2HNO3 + NO
A purge with excess NO2 would aid the absorption reaction,, but at
the expense of the HN03 plant handling a larger total volume ot gas.
-------
For the given purge rate, calculations were made to see what
HNC>3 acid strength might reasonably be obtained. At 65%, the ab-
sorption approaches a pinch point, at which the number of tower
trays is extremely large. Sixty percent acid, though, can be
easily produced in a small tower.
L-l, Ash Removal Facilities
Fly ash which is not removed from the flue gas by the power
plant electrostatic precipitator is picked up by the hot, sulfuric
acid in scrubber, E-l. Approximately 20% of the ash is soluble
in the acid; the remaining undissolved ash has to be removed from
the circulating acid before being sent to the ^catalytic stripper
for denitration. The undissolved ash in the 80% sulfuric acid
will be about 0.005 wt % at 272°F. At this time, no definitive
process design can be made in specifying the required ash removal
equipment. A detailed study of solids/liquid separations at such
low solids levels in the vqry corrosive acid medium is needed
before any specifications can be made.
4. Process Appraisal
The TYCO process for SOX and NOX removal presents several
inadequacies in terms of the soundness of the process definition.
The major concern is the catalytic stripping operation. It is
felt that this important denitration step has not been fully demon-
strated at conditions suitable for scale-up to a commercial level.
The experimental catalytic stripper gave the most favorable de-
nitration results only at very low liquid space velocities (^1)
at 290°F, and at a liquid to gas loading which is incompatible
with the extremely high liquid circulation rate required for the
upstream scrubbing operation. The commercial scale flow sheet has,
o£ necessity, assumed operating conditions for the oxidizing
stripper which have not been demonstrated in the laboratory. A
continuous, integrated scrubbing/stripping operation has not been
demonstrated; it was attempted unsuccessfully in the TYCO mini-
plant runs.
-------
Another problem with the catalytic stripper operation concerns
the catalyst, the activated carbon. Experimentation has not verified
catalyst life over 200 hours, and indications are that the carbon
activity to catalyze the HNSOs oxidation is reduced by impurities
in the acid. The commercial plant will have acid circulating with
dissolved fly ash components, and some corrosion byproducts; the
effect on the activated carbon performance must be investigated.
There are some indications that the activated carbon manufacturers
might be introducing larger sized carbon granules than the current
maximum of 4 mesh. Experimentation should be carried out with the
largest size activated carbon, since operating a packed bed of 1/8"
granules under countercurrent liquid/gas flow presents flooding
and pressure drop problems, with ensuing costly designs,
Some comments are made about the S02~oxidizer chemistry in
section B of this report. Each stage of TYCO's experimentation
indicates an upward trend in the required residence time for the
flue gas in order to have 90% SO2 oxidation. Early work indicated
several seconds residence time; the latest TYCO data indicates
about 20 seconds is required. Also, in the laboratory equipment
the surface to volume ratio in the S02 reactor is very high compared
to the duct type commercial reactor. The effects on the S02 oxidar
tion caused by the possibility of small scale reactor behaving as a
wetted-wall type device, owing to sulfuric acid condensing on the
walls, would not be the same as in the low surface/volume commercial
reactors. The residence time calculated in the small scale equipment
might not be applicable on a larger scale.
The hot scrubbing operation seems to have been demonstrated
experimentally at very high liquid scrubbing rates in the order of
800 Ibs circulated per Ib of sulfuric acid product. Further experi-
mental work should be directed toward decreasing the very high
scrubbing liquid rates as much as possible.
A major problem area in process definition concerns the ash/acid
separation section of the flow sheet. As mentioned earlier, removal of
this undissolved ash at levels of 0.005 wt % must be demonstrated,
otherwise the overall process will be inoperable. Also, the
disposition and effect on the process chemistry of the dissolved
-------
ash components in the system presents a problem which must be
clarified experimentally.
The operability of this TYCO process presents unusual
problems of process control. The NO/NOj ratio of 1 in the
scrubber must be controlled very closely, since a high scrubbr
ing efficiency (over 98%) is required. For example a 5%
increase in the oxide ratio to 1.05 would decrease the overall
NOx removal from the flue gas from 80% to 25%. Of necessity,
a power plant operates at fluctuating loads/ depending on power
demand. The process must be attuned to frequent turndowns and
variations in NOX and SOx levels in the flue gas.
The TYCO Modified Chamber Process depends on the presence
of both NOX and S02 in the flue gas in order to effect removal
of both, if upstream changes in boiler operation or fuel changes
are made which tend to increase or decrease one of the oxide
pollutants, there arises the possibility that the new system (at
steady state conditions) may be one which does not recover the
necessary S02 or NOX to meet the pollution control requirements.
Process control of the TYCO process is probably more crucial than
for most other air pollution removal processes.
A final major item of concern in this process involves the
partial pressure of sulfuric acid over the 80% acid solution. The
equilibrium partial pressure of H2S04 in the scrubber outlet gas
is 66 ppm by volume, excluding any physically entrained acid,
based on Gmitro and Vermeulen's data. This corresponds to 8.3
mg/SCF in the flue gas. Chemico, in its 1970 report for NAPCA,
"Engineering Analysis of Emissions Control Technology for Sulfuric
Acid Manufacturing Processes," states that a probable future
maximum on H2SO4 emission will be about 3 mg/SCF. The gas effluent
from the scrubber would have to be 225°F over 80% acid to have 3
mg H2SO4 vapor/SCF in the stack gas. However, at this low temp-
erature, substantial water would be absorbed from the flue gas by
the colder acid. It seems unlikely that future pollution control
codes will allow one ton per hour H2S04 emission into the atmo-
sphere. This means that the TYCO process must be subjected to
-------
extensive modifications to reduce the sulfuric acid emitted
with the stack gas.
-------
REFERENCES
Graitro, J. I., and Vermeulen, T., "Vapor-Liquid Equilibria for
Aqueous Sulfuric Acid"/ UCRL-10886, Univ. Calif., Berkeley,
June 24, 1963.
Oliver, E., "Sulfur Dioxide Removal From Stack Gases," SRI
Report No. 63, July"1970.
Walitt, A., "A Process for the Manufacture of Sulfuric and Nitric
Acids From Waste Flue Gases," Sec. Int'l Clean Air Congress
of IUAPPA, Washington, D.C., Dec. 6-11, 1970.
"TYCO Process For Sulfuric Acid Production From Power Plant
..*•
Flue Gases - Heat and Material Balance Analysis," Contract
No. PH-86-68-92 for NAPCA by A. M. Kinney, Inc., March 31,
1969.
"Sulfur Oxide Removal From Power Plant Stack Gases - TYCO
"Process," Preliminary Conceptual Design and Cost Study
(Draft) NAPCA Contract TV-29233A by TVA, Nov. 18, 1970.
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B. CHEMISTRY
The following review of the chemistry of the process
is based on Final Contract No. CPA 70-59 by A. Gruber and
A. Walitt. Basically the Tyco process removes S02 and NO
from stack gas by reacting N02 with S02 and absorbing the
resulting S03 and NO (with more N02) into hot (250-300°F)
sulfuric acid. The advantage of the hot acid (approximately
80% sulfuric) is that there is no net absorption of steam
from the flue gas and the gases do not have to be reheated
for entrance to the stack.
Absorption of S07 and NO (For NO = 1/2 S02 in stack gas)
2H2S04 + S02 + 1/2 NO + 2 1/2 N02 (added via recycle)
> 3IINSOs + l/2 H20
Ox i (I i zer
According to Tyco this is accomplished via the homo-
geneous gas phase oxidation of S02 with N02. They state
that approximately 15-20 seconds are required in order
to convert 90% of the S02. Therefore for a 1000 megawatt
power station, an oxidizer volume of 750,000 to 1,000,000
ft3 would be required. N02 and S02 concentrations in the
gas are, as stated, about 6000 and 3000 ppm respectively,
at about 300°F. Even this huge oxidizer volume would be
optimistic if indeed the homogeneous gas phase reaction
(1) Gruber, A., and Walitt-, A., "Oxidation of SO2" - Final
Report, Contract CPA 70-5A For NAPCA, 1970.
-------
did control in view of the kinetic work of Boreskov and
Illarionov (Zh. Fiz. Khim. Vol. 14, page 1428 (1940))
which indicates that at the 3000 to 6000 ppm concentra-
tion region the reaction rate would be infinitesimally
s low.
N^OT; Absorber
The chemistry of this reaction is reasonably clear
cut. Tyco data look reasonable. The reaction
NO + N02 + 2H2S04 = 2HNS05 + H20
is well known. The absorber should be sized on gas film
controlling mass transfer at appropriate liquid and gas
loading rates. The only questionable item is the back
pressure to be used (i.e. the proper equilibrium ex-
pression). The equilibrium expression used by Tyco is
based on their own work and that of Berl and Saenger who
both used vapor pressure measurements assuming that in
the gas phase NO = N02 . Their NO + N02 partial pressures
are probably considerably higher than thermodynamic equi-
librium due to the fact that they neglected any small a-
mount of HNOj in the sulfuric acid solution (in fact,
throughout their work they neglect it except for a few
brief remarks). Even the presence of small amounts of
HNOj in the acid solution along with HNSO^ can elevate
the partial pressure of N's above the solution by very
-------
large amounts. This is probably the reason why in the
absorption experiments Tyco reports HQQ'S of about 8
feet which seems high (they probably had much higher
back pressure in their absorption experiments than they
thought they had). In general, Tyco's estimation of the
size of the ^2^3 absorber is felt to be conservative.
"Catalytic" Stripper
This catalytic stripper is an extremely important
k
part of the process scheme. Nitrogen values absorbed in-
to 250-300°F 80% sulfuric acid must now be stripped back
into the gas phase at a high enough partial pressure to be
able to make nitric acid product. The stripping efficiency
must be high (approximately 99% N removal), carbon cannot
be consumed, ^0 and N2 cannot be produced and the "catalyst"
must have a significant life.
In general, on the basis of Tyco's report it is almost
impossible to design a commercial "catalytic stripper"
since most of the experimentation was of an exploratory
nature and did not provide careful material accounting.
Comments
1) Granular carbon (4 x 10 mesh) was used. This is
certainly not a feasible commercial size for a reactor
handling millions of pounds per hour of liquid and large
volumes of gas.
-------
2) "Catalytic" behavior has not been demonstrated.
The highest HNS05 loading to the carbon bed Cstudied by
passing 42,000 ml. of HNS05 solution to the bed) was
approximately 4 g.moles/8 g.moles of carbon (30 in-*) . The
actual nitrogen oxides stripped amounted to about 2 g.moles/
8 g.moles of carbon.
3) The carbon bed is apparently sensitive to im-
purities in the acid.
4) Only during two time periods (105 and 120 minutes.
Table XXXII, page 131) was the nitrogen material balance
better than 90%. The overall balance in this table is
about 81%.
5) Only NO and N02 figures are given for the catalytic
stripper exit gas. Since carbon is a reducing agent
there is a definite possibility that the exit gas contains
N2, N20, CO, C02 and less possibly S02. It must be
realized that in the overall scheme of stripping, if as
little as 17% of the N's stripped appear as N2 or N20
there will be no net production of nitric acid in this
proposed clean-up scheme. (The recycle ratio of N02
added per NO in the flue gas is high - 5/1). If larger
than 17% then there will be a net purchase of N02 to re-
move the S02. To compound the problem, if CO or C02 is
formed then carbon is being consumed.
-------
Again, extrapolation of Sanfourche et al. data in-
dicate that at 300°F and 80% acid with HN03 and HNSOg both
in solution at a concentration of 0.023 g.moles/liter*)
the highest pressure of N02 obtainable, at a total pres-
sure of 1.0 ata., is equal to 0.02 ata or 20,000 ppm ,
hardly a high enough pressure to make significant amounts
of nitric acid in the nitric acid section of the plant.
Nitric Acid Production
k
This is a conventional nitric acid tower and can be
sized in the conventional manner provided the "catalytic"
stripper produces a high enough partial pressure of N02
and a high enough ratio of N02/N0. Since this unit (and
also the catalytic stripper) does not handle the huge
volumes of stack gas, increasing pressure may be advan-
tageous .
a) Tyco, Table XXXII, used 0.047 g.moles/liter of HNS05
in feed acid to stripper which we are going to assume is
0.023 g.moles/liter of HNSOs and 0.023 g.moles/liter of
HN03 to see what the absolutely highest thermodynamic
equilibrium pressure of N02 is above this solution.
-------
An ultra-optimistic way to size the "catalytic strip-
per" would be to assume that true catalytic behavior is
realized (i.e. there is no carbon consumption (no CO and
C02), there is 100% selectivity in oxidizing HNS05 to N02
(no N2 or N20 formation), a long enough life to make
carbon cost insignificant, and that large size activated
carbon packing (1 or 2 inch rings, saddles, aggregate, etc.)
can be effectively used instead of the 4 x 10 mesh that
Tyco used. (Further, we make the assumption that oxidation
is instantaneous - no kinetic or mass transfer limitation -
which is an extreme simplification to say the least). In
this case, sizing would then be based on the ability (strip-
ping factor, VK/L) to strip N02 from an equimolar mixture
of HNS05 and HN03, the most favorable stripping factor of
all nitrogen components in the sulfuric acid.
UNO3 (soln) + HNS05 (soln) = 2N02(g) + H2S04
Extrapolation of Literature data (Sanfourche and Rondier:
Bull. Soc. Chim. France 4_3_ (1928) 815) indicate that at
300°F and 80% H2S04 a VK/L = 2.0 can be obtained at a
liquid loading of L = 1000 Ib acid/hr/ft2 and a total gas
loading of 1.5 Ib moles/hr/ft2 (Tyco's loading Table XXXII,
page 131, was L = 145 Ib/hr/ft2 and GT = 0.29 Ib moles/hr/
ft2} for a VK/L = 2.7). Depending on size of carbon used
the stripping tower can be ul tra-optimistically sized on
flooding velocities using a L/G of about 700 Ibs/lb mole of gas
-------
J
3S4.1 NO, 211.
CO, 10,004-
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MAGNESIUM OXIDE SCRUBBING PROCESS
A. PROCESS
1. Process Description (See MWK Dwg, No. PD-118-D)
Flue gas from the power plant at 300°F, together with the
gases vented from the dryer of the regeneration system and the
sulfuric acid plant, are cooled to 170°F in an indirect heat
exchanger. The heat is recovered and used to reheat the absorber
effluent gas from 120°F to 250°F before it is sent to the stack.
The 170°F gases enter the venturi scrubber where a minimum of 98%
of the ash is removed by a circulating stream of ash slurry in
water.
In the scrubber enough water is evaporated to result in a
saturated exit gas at 120°F. The ash is removed from the system
in a slip stream which is discharged to a settling pond. Overflow
water is pumped from the pond back into the circulating slurry to
minimize the need for fresh water make-up.
As 803 dissolves in water forming H2S04, limestone (CaC03)
slurry is added to the discard stream to neutralize acid so formed.
The effluent gas from the venturi scrubber flows to a three
stage absorber wherein the gas is contacted with a circulating
stream of magnesium oxide/magnesium hydroxide slurry in water
which is saturated with MgS03 (1.24%) and contains an appreciable
quantity of dissolved MgSO4. In the absorber a small amount of
the SO2 in the gas oxidizes to SO3. All of the 803 and 95% of the
SO2 react with the magnesium oxide to form magnesium sulfite,
which hydrates insoluble MgSO3 • 6H2O, and magnesium sulfate,
MgS04, which remains in solution.
The exit gas from the absorber is reheated to 250°F by the
indirect heat exchange system previously described. From the ex-
changer the gas is discharged by the induced draft fan into the
stack.
-------
The concentration of solids in the absorber bottoms is held
at 10% by the withdrawal of a slip stream to the magnesia regen-
erating system. The build-up of ash, which enters the absorber
with the flue gas from the scrubber, is limited to 1% of the
absorber bottoms by removal of the net ash from a portion of the
stream of regenerated magnesia.
The slip stream to be regenerated is fed to a thickener where
a clear solution is produced from the launder, and a bottoms of
30% solids content is withdrawn and fed to a centrifuge. The
centrifuge reduces the solids to a cake containing 5% water, and
the soluble salts associated with water. The clear liquor from
the centrifuge, together with the thickener overflow, collects
in the solution tank, from which part is returned directly to the
absorber and the rest is used to slurry the reclaimed magnesium
oxide/hydroxide.
The cake from the centrifuge is discharged into the cake
hopper. The cake hopper also receives the dried magnesium sulfate
recovered from the ash disposal system. The mixture of sulfate
and wet centrifuge cake is then fed into the dryer. This dryer
is of the direct fired type, preferably natural gas fuel with about
200% excess air to give an entering gas temperature of 1525°F.
Flow from the dryer is countercurrent with the dried solids being
heated to 500°F, a temperature sufficient not only to evaporate
the free water but to drive off the water of hydration as well.
The gases leave the dryer at 300°F, and after passing through a
cyclone separator to recover any entrained dust, are discharged
into the stream of flue gas entering the scrubber. Thus any dust
that escapes the cyclone will be removed and any sulfur that may
be in the dryer fuel will be recovered. The dryer effluent solids
and the dust recovered by the cyclone are discharged into the
calciner charge hopper.
Crushed green petroleum coke used for the reduction of MgS04
to MgSC>3 is received by truck and stored in the coke silo from
where it is fed into the calciner charge hopper. As previously
mentioned, this hopper also receives the discharge of dried solids
from the dryer. The mixed solids from the hopper are discharged
into the calciner.
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The direct fired calciner uses natural gas fuel with about
100% excess air to obtain hot gases at 2240°F entering the' kiln.
In the kiln the flow is countercurrent, with the solids heated
from 500° to 1600°F while the exit gas, including the SC>2 and
C02 generated by the calcination, are at 600°F.
The gas from the calciner is vented to a spray tower for
cooling, while the calcined solids, MgO plus ash, are discharged
into the quench tank. In the quench tank, which is held at 180°F,
sufficient fresh water is added to slake the oxide to hydroxide.
This slaking is a slow reaction and is only partially completed
in the course of the recirculation of the slurry to the absorber.
The quench tank is provided with a vent stack in case a local
high temperature zone at the 1600°F solids inlet into the liquid
produces steam. The fresh water addition is via sprays in this
stack which will condense and knock back any steam so formed.
Liquid from the solution tank is added to the quench tank to
dilute the slurry to the desired solids content. Temperature
in the quench tank is maintained by circulation of a portion of
the slurry through the external quench cooler. The ash discard
stream also flows through the cooler and thence to the acid treat-
ing step at 120°F.
The net bottoms from the quench tank is transferred at 180°F
to the surge tank where it is diluted to a 10% slurry. Make-up
MgO and fresh water to balance the system losses are both added
in this tank. The equilibrium temperature is 139°F but the slurry
is cooled to 120°F in the solution cooler before it is returned to
the absorber.
To prevent build-up of ash and other inert materials in the
slurry, a continuous purge is required. This purge stream is
taken from the quench tank and also contains Mg(OH)2 in suspension
and MgSC>3 and MgSC>4 in solution. These salts must be recovered to
avoid excessive replacement requirement. After the addition of
sufficient water from the base of the spray tower to balance
evaporation loss, the discard stream is mixed with sulfuric acid
-------
(recycled from the acid plant) in stoichiometric quantity. The
insoluble Mg(OH)2 is converted to soluble MgS04 • 7H20 which then
dissolves leaving the ash in suspension. The slurry from the acidi-
fying tank is pumped to the ash filter which separates the ash (and
other solid inerts, if any) as a wet cake that is discharged into
the ash slurry tank. The filtrate, a saturated (or nearly so)
solution of MgSC>4 • 7H20 with some MgSOs and MgS04 dissolved as well,
is collected in the sulfate filtrate receiver from which it is
pumped to a spray dryer.
The spray dryer is heated with 1200°F hot air produced by
burning gas fuel at 300% excess air in the air heater. Here the
solution is evaporated at 300°F to produce dry sulfate crystals
which are returned to the scrubbing circuit via the sulfate
reduction system. The effluent gases from the spray dryer flow
through cyclones for dust recovery and then into the spray tower.
In the spray tower the gas from the calciner at 600°F, the
gas from the spray drum at 300°F and the steam from the acidify-
ing tank at 220°F are quenched with 110°F water thus cooling the
gases to 120°F and condensing most of the water contained in the
gases. The condensate is withdrawn at 180°F, pumped through an
external cooler where it is cooled to 110°F, and split into three
streams. The net recovered water is discharged to the surge tank,
a stream is recycled to the acidifying tank to provide cooling,
and the remainder is recirculated to the spray tower as quench
water.
The cooled gases from the spray tower, containing 11.1% SO-
are vented to the acid plant, which converts the SC>2 to 98% sul-
furic acid. Vent gases from the acid plant containing unreacted
S02 are recycled to the flue gas entering the scrubbing system.
Storage tanks for 30 days net production of sulfuric acid are
provided along with an unlading pump. A separate pump is
provided to supply acid from storage to the ash acidifying tank.
-------
Sulfur trioxide in the flue gas will be absorbed by the
slurry circulating in the venturi scrubber case producing a
weak solution of H2S04 in the ash slurry stream. To neutralize
this acid and thus prevent an excessive build-up of acid in the
water circulated from the ash pond, facilities are provided for
adding ground limestone (CaC03) to the ash discard stream. This
will react with the acid to form CaSC>4 which is deposited in the
pond along with the ash.
2• Design Bases
The present flow sheet developed for magnesium oxide scrubb-
ing is for a new 1000 megawatt installation. General design
criteria used in evaluating this process are the same as the
basic criteria used for all the present evaluations and are given
elsewhere in this report. Major design criteria specifically
pertinent to the MgO process are listed below.
a) Six parallel scrubbing trains are provided including
separate cooler, scrubber, reheater and induced-draft
fan per train.
b) Shut-off gates are provided such that each train can
be isolated from the other five.
c) Stack gas cooler and reheater system:indirect exchange
against circulating water (with glycol added), single
pass gas side cross-flow unit (12 identical units
required -- 6 coolers and 6 reheaters).
d) Venturi scrubbers - 6 identical units (one per train),
98% ash removal, circulating and purge slurry @ 10%
ash, design pressure drop @ 10" H2O, vessel size:
34'-6" diameter by 66 ft high.
e) Scrubber (absorber) design (6 identical units required),
UOP TCA contactors. Vessel size: 15 ft deep x 34 ft x
40 ft high with center partition (2 gas inlets and 2
gas outlets),3 contacting stages of 1-1/2 inch plastic
spheres; design pressure drop @ 9" H2O; 95% SC>2 removal
-------
(B&W data); bisulfite concentration of 0.762 wt %
(as SO2) in outlet slurry (B&W data).
f) Thickener - to concentrate outlet slurry from absorber
from 10% solids to 30% solids; vessel sizei 113 ft
diameter and 10' high, 3 identical units.
g) Centrifuges - to separate solids from thickened slurry.
h) Dryer System - rotary dryers 16 ft dia x 200 ft, 2
identical units; to dry solids from centrifuges and
drive off water of hydration; natural gas fuel, 200%
excess air.
i) Calciner - rotary kilns 12.5 ft dia x 20.0 ft, 2
identical units, natural gas fuel, inlet temperature
of 2240°F. To regenerate MgO for recycle by reduction
with carbon and heat.
j) Acid Plant - to convert SO2 liberated in calciner to
concentrated H2SO4 which can be sold as a by-product.
k) Acid storage - storage capacity equivalent to 30 days'
production is provided; unloading pumps are sized to
pump out one day's production in 4 hours.
1) Equipment is provided to reduce MgO losses in the
fly ash purge stream by converting insoluble Mg(OH)2
to the soluble sulfate, filtering, and spray drying.
3. Design Rationale
The process design information specifically related to the MgO
system was obtained from reports issued by The Babcock & Wilcox
Company(1)(2). The general basic design criteria established for
the present evaluations were used and the available B&W values pro-
rated as needed. Where data were missing or not applicable to the
present design, literature and/or vendor data were used if available,
and when not available, assumptions were made as needed to complete
the design. The judgment factors behind the major design bases are
presented below.
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a) Venturi Scrubber - used by B&W in their conceptual
design. Most (98%) of the fly ash is removed in
this vessel thereby reducing the possibility of
plugging in the absorber and also providing an
easier separation of the fly ash from the scrubbing
slurry owing to the reduction in quantity of fly
ash in the slurry.
b) Stack Gas Cooler-Reheater System - the indirect
circulating water system was recommended by TVA
based on a study of six different heat transfer
schemes. The gas reheat temperature of 250°F
also is based on TVA's recommendation.
c) Scrubber (Absorber) Design - based on B&W's ex-
perimental data for a floating bed absorber; UOP' s
TCA contactor was selected for the commercial
plant. Flow rates, absorption efficiency, and
pressure drop are based on B&W data (1).
d) Dryer - designed to dry the solids and drive off
the water of hydration but the temperature is kept
below that needed to decompose the magnesium sulfite,
Natural gas fuel is used as the heat source and the
gas temperature is regulated by the amount of excess
air supplied with the combustion air.
e) Calciner - designed to regenerate the magnesium
salts by thermally driving off S02 from the MgSC>3
while the MgSC>4 is reduced to MgO by reaction with
carbon (coke) which is fed into the calciner along
with the magnesium salts. The calciner temperature
is a key parameter in that if it is too high the
MgO produced will be unreactive (dead-burned), while
the reduction reaction will not occur if the tempera-
ture is too low. A temperature in the range of 1100-
1600°F appears necessary but the optimum has not yet
been defined. For the present evaluation a tempera-
ture of 1100°F was assumed to insure good kinetics
and reduce the size of the calciner.
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f) Acid Plant - the S02 liberated during the regenera-
tion of the MgO is used as the feed to a sulfuric
acid plant for the recovery of sulfur as concentrated
(-98%) acid. Unreacted SC>2 from the acid plant is
recycled to the absorber. Storage capacity equiva-
lent to 30 days' production of acid is provided to
insure continuity of operation in case of short term
interruptions of acid disposal (e.g., rail strikes,
market fluctuations, etc..).
g) MgO Recovery System - the fly ash not removed in the
venturi scrubber will be scrubbed out in the absorber,
thus requiring a purge stream to avoid build-up of
fly ash in the scrubbing slurry. To minimize MgO
losses a recovery system is specified wherein the
insoluble Mg(OH)2 is converted to soluble MgSO4 and
the fly ash is separated by filtering. Sulfuric
acid is recycled from the acid plant to react with
the Mg (OH)2 and the resulting MgS04 is recovered by
spray drying and returned to the dryer-calciner
circuit.
4 . Process Appraisal
There is no complete process of this type presently in com-
mercial operation. The Chemical Construction Corp's. process for
acid plant gas, without particulate matter, apparently has been
developed in detail. However no design operating conditions, or
actual operating data were available for this evaluation. Therefore,
a number of assumptions were required to develop the present flow
sheet which should be considered as preliminary only.
The Babcock & Wilcox report (1) is concerned entirely with
pilot plant experiments on particulate removal and the removal of
S02 from the gas by MgO reaction. No experiments were reported
on the regeneration of the sulfite.
-------
The primary difficulty encountered by B&W in this laboratory
work was formation of crystalline deposits, primarily MgS03.• 6H20,
with traces of MgSC>4 • H20. If the bisulfite, MgH2(S03)2 exceeded
about 1 gm/100 ml as S02 the deposition of crystals was extremely
rapid. However the presence of fly ash in the solution apparently
prevents deposition. The design under consideration assumes 0.9%
fly ash in the absorber bottoms liquid (1.1% by weight of water
alone). By B&W findings, this should be sufficient to eliminate
deposits.
If this ash concentration is not high enough to prevent
deposition, it may easily be increased by reducing the flow of
the ash discard stream.
Another possible difficulty would be the accumulation of
chlorides in the solution. These would be formed by reaction of
HCl in the flue gas with Mg(OH)2 • MgCl2 is quite soluble and does
not decompose under heat. Since there is essentially no water lost
from the regeneration system there is no way of purging the solution,
so the build up could be to a point where precipitation would occur.
This would be in the base of the absorber, the thickener, or the
centrifuge. Solids formed there would be carried to the dryer and
calciner, where at 712°C (1313°F) the crystals would melt and possibly
cement the MgO at the quench point. The tubes of the quench cooler
would collect deposits and plug. Resolution of this problem remains
to be done.
Deactivation of the MgO by calcination may be a problem. B&W
report that a temperature below 1100°F is advisable, and that at
2200°F the MgO is essentially unreactive. A more complete study of
the mechanism of deactivation is requied, together with details of
the exact temperature where deactivation starts. It is possible
that the 1100°F quoted is a mass bed temperature resulting from
severe local overheating, and that with careful application of
heat appreciably higher temperatures can be withstood.
-------
Also lacking is detailed information as to when the reduction
of MgSC>4 to MgS03 by reaction with carbon will occur. It is pre-
sently estimated that 1600°F may be required, but it is possible
that the reaction may proceed at a much lower temperature so that
such severe calcining may not be necessary.
The use of a once-through direct fired calciner may not be
advantageous as there is appreciable oxygen in the hot flue gas
(100% excess air at 2240°F). There is a possibility that some or
most of the coke added to reduce the sulfate may preferentially
burn to CO2 under these conditions, leaving the sulfate unchanged.
A recirculation type calciner operated at very low excess air
would eliminate this possible difficulty, and could provide a
lower initial gas temperature. This would be advantageous in
avoiding overburning of the magnesium oxide.
No data have been developed for the direct loss in the dryer
and the calciner. The design presented herein assumes that dust
entrainment with the exit gases will be minimal. A simple cyclone
separator has been used to recover dust from the dryer off-gas
while the dust from the calciner is to be recovered in the spray
tower and returned to the system with the reclaimed water. If
severe dust entrainment occurs in the calciner, a cyclone separator
will be needed in the off-gas. However no information is readily
available as to the type of crystals which will be formed in the
absorber and the subsequent drying and calcining. If these mater-
ials are prone to entrainment, it may not be possible to employ
direct fired apparatus for drying and calcining.
Erosion could be a major problem in the various circulating
slurry streams. The use of rubber lined pumps and piping might
be considered as a means of combating the combined corrosion-
erosion problems.
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REFERENCES
(1) Babcock & Wilcox "Magnesia Base Wet Scrubbing of Pulverized
Coal Generated Flue Gas - Pilot Demonstration", NAPCA order
4152-01, September 28, 1970.
(2) Babcock & Wilcox "Conceptualized Fly-Ash and Sulfur Dioxide
Scrubbing System with By-Product Recovery", Contract No.
CPA 22-69-169.
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MAGNESIUM OXIDE SCRUBBING
B. CHEMISTRY
System Description
Two reports have been submitted for comment: The
Babcock and Wilcox Company "Magnesia Base Wet Scrubbing
of Pulverized Coal Generated Flue Gas - Pilot Demonstra-
tion", dated Sept. 28, 1970; and a flow sheet and three
pages of an untitled report from Chemical Construction
Corporation.
The Babcock and Wilcox report describes the pilot
plant removal of 95.5% to 99.2% of the sulfur dioxide
and removal of 97.7% to 99.5% of the particulate matter.
The Babcock and Wilcox system used a venturi type fly
ash scrubber (water washing) and sulfur dioxide removal
by venturi scrubbers and/or tray scrubbers using magnesium
oxide slurry as absorbent for S02- The Babcock and Wilcox
study did not include regeneration of the magnesium oxide.
The Chemical Construction Corporation pilot plant
work was done on tail gas from a sulfuric acid plant
(no particulate matter) and presents a schematic for the
regeneration of the magnesium oxide. The Chemical Con-
struction Corporation recovery system consists of cen-
trifuging a portion of the Mg(OH)2-MgS03 slurry, drying
the MgSOj • 6 H2<) to remove water of hydration, calcination
of the MgS03 to make MgO and S02 (concurrent calcination
-------
of any MgSC>4 with coke to convert sulfate to sulfite),
recovery of S02 for production of sulfuric acid and re-
covery of MgO for recycle.
Comments - MgO Recovery System
The MgO recovery scheme, as per the Chemical Construction
Corporation's schematic, is based on a SC^-SOg feed that
does not contain particulate matter (tail gas from a
sulfuric acid plant) .
Particulate matter and other inert materials may con-
centrate in the MgO recovery system and decrease the
efficiency of the MgO system by adding to the burden of
centrifuging, drying, calcining, etc.
Inert material may build up in the recovery system
from several sources:
(1) A particulate scrubber working at 99.5% removal of
particulate matter that processes 77,900 Ibs fly ash/
hour would leave 390 Ibs particulate matter/hour
available for possible collection in the Mg(OH)2
slurry.
(2) The reduction of sulfate to sulfite with coke
(assuming an ashless coke) will probably mean that
excess coke will be used in the calcination and will
be recycled with the MgO. The coke concentration
in the slurry will probably stay at some steady
-------
state level but will add to the inert content of
the MgCOH)2 slurry.
(3) The regeneration of MgO does not mean the MgO will
be reactive. Care must be taken during the re-
generation to insure that inert MgO is not made.
Babcock and Wilcox did not go into this point
other than to comment in section 7.3-C, "The re-
generated magnesia must be reactive td be reusable.
Therefore, the thermal decomposition temperature
should be minimized". And in section 5.4, "- - - -
the optimum temperature for calcining is between
750 and 1100°F. Above calcining temperatures of
about 2200°F, the resulting MgO is essentially un-
reactive and is referred to as Dead Burned - - - -".
Comments - Particulates in Mg(OH)2 Slurry
Large amounts of particulate material in the Mg(OH)2
slurry are undesirable; however, small amounts of partic-
ulate matter in the Mj>(01l)2 slurry are advantageous in
that they have been found to decrease the rate of de-
position of MgS03-6H20 on the scrubber surfaces due to
a slight abrasive action of the particles and large par-
ticulate surface areas for crystal growth.
Comments - Venturi Particulate Scrubber
The venturi particulate scrubber removed 97.7% to
99.5% of the particulate matter and none of the sulfur
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dioxide. The aqueous slurry of particulate matter was re-
circulated through the venturi, a portion of the slurry
was discarded, and de-ionized water was used for make-up.
The pH of the slurry was 2 to 3.5.
The sulfur content of the slurry is. not reported
but it appears that the slurry is absorbing sulfuric acid
and/or bisulfate.
The used particulate slurry is being dumped in the
Babcock-Wilcox scheme and is not useable in the Chemical
Construction Corporation regeneration scheme.
An advantage may be gained by a more efficient
particulate scrubber in that it would permit less particu-
lates to nccumtilntc i" Hm MgO recovery system and would
decrease the nmount. of MI If ate that would need to be re-
duced with coke in the calcination step. The calcination
might then be accomplished at lower temperature and shorter
times and yield a more reactive magnesium oxide.
Comments - Gas Scrubbers
The venturi gas scrubber and tray gas scrubber were
designed to remove sulfur dioxide but both are very effi-
cient ( "> 98% removal) in removing particulate matter
from the gas stream. A small amount of particular^ maLl or
will not be trapped by the particulate scrubber during
normal operation and may find its way into the MgO recovery
-------
system. A failure in the venturi particulate scrubber
(i.e.} failure of water-slurry circulating pump) would
cause a rapid buildup of ash particles in the MgO re-
covery system.
The MgO recovery system is a closed system and does
not make provision for discharge of inert materials that
may build up in the Mg(OH)2 slurry.
Reaction - Slaking
Fresh or recovered magnesium oxide is slaked with
water to produce magnesium hydroxide
MgO t- H20 > Mg** + 20H~ (Eq. 1)
Mg + + + 20H- — NMg(OH)2CS) (Eq. 2)
Two to three hours at 180°F were required to con-
vert 94.9% of a reactive MgO to Mg(OH)2.
Babcock and Wilcox determined that MgO and Mg(OH)2
had the same ability to absorb sulfur dioxide.
Reaction - Formation of Sulfate
The presence of sulfate in the recovery system is
undesirable since an additional treatment is required
to reduce the sulfate to sulfite.
Sulfite may be oxidized to sulfate by several mech-
anisms, either in gas phase or in the slurry; by reaction
with excess air in the flue gas or with oxygen dissolved
in make-up water, i '
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S03 + 1/2 02 -- > S04= (Eq. 5)
or by reaction with nitrogen oxides from the flue gas.
N02 + S03= - > NO + S04= (Eq. 4)
The sulfate present will react with magnesium in the. slurry
to form magnesium sulfate
HOH
Mg + + + S04 - * MgS04-7H20 (Eq. 5)
The reactions in Eq . 3, 4 and 5 are rapid and may
occur anywhere from in the furnace to in the Mg(OH)2
absorber .
Reaction - Absorbing Sulfur Dioxide
The overall reaction for the removal ,of sulfur
dioxide with MgO is
HO
MgO + S02 - - — > MgS03-6H20(S) (Eq. 6)
in which sulfur dioxide is absorbed in aqueous MgO slurry
to make bisulfite and sulfite.
S02 + OH' - ? HS03" (Eq. 7)
HS03- + OH~ -— S03= + HOH (Eq. 8)
The sulfite will react with the magnesium in the slurry
to precipitate MgS03-6H20.
S03~ + Mg++ - * MgS03-6H20 (Eq. 9)
The dissolution of MgO with water (Eq. 1) is the
slowest and would be the rate setting step. The rate of
dissolution of MgO in the slurry is enhanced by the bufferin;
-------
effect of the sulfite-bisulfite system (Eq- 7 and 8)•
The enhancement of the rate of solubility of MgO is such
that Babcock and Wilcox report MgO and Mg(OH)2 had the
same ability to absorb sulfur dioxide.
Reaction - Absorption Stoichioaetry
Babcock and Wilcox report that using a reactive
MgO, that had not been recycled, one mole of MgO was re-
quired to absorb one mole of sulfur dioxide.
Reaction - Absorption pH
The pH of the system was controlled by regulation
of the MgO makeup.
At pH more acid than 7 the sulfur dioxide absorption
decreased and high sulfate levels had an adverse effect
on the sulfur dioxide absorption. At pH more basic than
7 the rate of sulfur dioxide absorption was constant and
the sulfate level had no effect on sulfur dioxide ab-
sorption .
Reaction - Absorption-MgO Coating
The MgO base system (as opposed to the CaO base
system) is not dependent on the available surface area of
the MgO. The solubility of magnesium sulfite is high as
compared to the solubility of calcium sulfite (485 mole/
mole @ 140°F) and does not tend to coat the MgO.
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Reaction - Absorption-Nitrogen Oxides
Nitrogen oxides are not absorbed by the MgO slurry
system as described. Babcock and Wilcox suggest an addi-
tional MgO absorber and injection of recycled NC>2 to re-
move NO from the flue gas.
Reaction - Calcination and Sulfate Reduction
The calcination process is to convert dehydrated
magnesium sulfite to sulfur dioxide gas and make active
magnesium oxide for recycle.
MgS03 >MgO + S02 I (Eq. 10)
The dehydrated solids fed to the calcination contain
magnesium sulfate which is reduced with coke to sulfite.
2MgS04 + C ^2MgSOs + 002^ (Eq. 11)
The activity of the MgO produced by the calcination
(Eq. 10) will be critical.
Chemical Construction Corporation make no mention
of the activity of the produced MgO, calcination tempera-
ture, calcination time, quantity of carbon used in cal-
cination, disposition of any excess of carbon. The only
hint in regard to the calcination is that it is done at
less than atmospheric pressure.
SUMMARY
The chemistry of the removal of sulfur dioxide with
MgO slurry appears to be quite workable and, in fact, has
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been done on pilot plant scale.
The undefined areas appear to be:
., (1) Does the recovery system produce an active MgO
for recycle?
(2) How much inert material can the recovery system
tolerate ?
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WOT£_-EQUIPMENT INSIDE DASHED LIKE IS
PART OF POWER PUNT * NOT TO BC
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ZINC OXIDE PROCESS FOR REMOVING S02 FROM STACK GASES
A. PROCESS
1. Process Description (See MWK Dwg No. PD-110-D)
The major reactions involved in the zinc oxide process are
summarized below.
Scrubbing Reaction
Na2SC>3 + H2° + S02 =
Na2S03 + 1/2 02 =
Liming Tanks
a V-
2NaHS03 + CaO = Na2SOa + CaSO3 + H2O
Gasif ier
CaSO3 + SO2 + H2O = Ca(HS03)2
Ca(HS03)2 + Na2S04 = CaSO4 + 2NaHS03
Mixer
2NaHSO3 + ZnO = ZnSO3 + Na2SO3 + H2O
Calciner
= ZnO + SO2
In the zinc oxide process, stack gas is scrubbed with an
aqueous solution of sodium sulf ite-bisulfite. The spent solution
is mixed in a clarifier with a slurry of calcium sulfite in an
aqueous solution of sodium bisulfite. The thickened calcium
sulfite-fly ash slurry from the clarifier is next reacted in a
gasifier with sulfur dioxide whereby the calcium sulfite is
solutized as calcium hydrogen sulfite. Sulf ate ions that are
present in the gasifier precipitate out as calcium sulf ate. Waste
solids of fly ash and calcium sulfate from the gasifier are dis-
carded and the soluble portion is reacted with lime (CaO) to form
the insoluble calcium sulfite needed for mixing with the spent
absorber solution.
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The supernatant liquid from the clarifier is reacted in a
mixer with zinc oxide to give insoluble zinc sulfite. The re-
sulting lean solution from the mixer is refortified with soda
ash to make up for losses and this solution is sent back to the
stack gas scrubber. Zinc sulfite is processed by drying off
the excess water and then indirectly calcining to give zinc oxide
and sulfur dioxide. Zinc oxide is recycled back to the mixing
vessel while part of the sulfur dioxide is used in the gasifier
with the remainder sent to an acid plant for recovery of the
sulfur values as 98.5% sulfuric acid.
2. Process Design Basis
This process has been evaluated for operation with a new
1000 MW power plant. General design criteria used in evaluating
this process are those given elsewhere in this report as the basic
criteria for all the present evaluations. Design criteria related
specifically to the zinc oxide process are based on the flow sheet
presented in a report by Envirogenics (1) to the National Air
Pollution Control Administration (NAPCA). Process flows and equip-
ment sizes have been scaled to the specified plant size (1000 MW)
and adjusted for the sulfur level used in this evaluation.
Since the zinc oxide process uses aqueous scrubbers, it was
assumed that mechanical separators would remove 80% of the fly ash
and the scrubbers would remove the remainder, thereby eliminating
the need for electrostatic precipitators. A stack gas cooler-
reheater system has been added to the process so that the treated
stack gas is reheated to get plume rise.
Sulfur dioxide removed in the zinc oxide process is recovered
as 98.5 percent sulfuric acid. The equipment and processing cost
for converting the sulfur dioxide to sulfuric acid has been
included in this evaluation.
Other design criteria related to the zinc oxide process are
discussed in more detail in the following section, viz., Process
Design Rationale.
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3. Process Design Rationale
The evaluation of the zinc oxide process is based on the
process flow sheet given in a 1970 report submitted to the
National Air Pollution Control Administration by Envirogenics (1).
A 1400 MW new power plant was the basis of the material balance
prepared and the equipment sized by Envirogenics; the material
balance was adjusted to the process design basis used by Kellogg.
The assumptions made and the design criteria used for evaluation
of the zinc oxide process are discussed in detail below.
Stack gas flows through cyclone separators where 80% of the
fly ash is removed. The remaining fly ash in the flue gas will
be removed in the stack gas scrubbers and is slurried with the
scrubbing solution. Therefore, this solution will contain about
0.5 weight percent fly ash which should not present any handling
problems.
Prior to entering the absorber system, the effluent stack
gases from the cyclones are cooled from 300°F to 167°F by a
circulating water system and the heat removed is then used to
reheat the gases downstream of the scrubbers to obtain the buoyancy
required for plume rise in the stack. In addition, less water is
evaporated in this system as opposed to allowing the solution to
directly cool the 300°F stack gas.
Sulfur dioxide is removed in the absorbers by the aqueous
sodium sulfite-bisulfite solution. The liquid recirculation rate
of the scrubbing solution is scaled down from that used in the
1400 MW plant and adjusted for the sulfur values in the Kellogg
design case. This "scaling" down factor is used throughout the
zinc oxide process to determine liquid and gas rates and the heat
duties. The absorbers used for removing the SO2 are four stage
T M
Impinjet ' " wet scrubbers (manufactured by the W. W. Sly Mfg. Co.)
which are assumed to give the desired 90% removal. Based on the
vendor's recommendations, an allowable superficial gas velocity
of 500 ft/min was used in sizing the tower area. The tower operates
at 122°F with the 14% solution of sodium sulfite/bisulfite. Since
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enthalpy data were not available on the absorption system, it
was assumed that the absorbers are in heat balance. Sodium
sulfate is formed by oxidation of sulfite in the absorbers. The
molar ratio of sulfite to sulfate is assumed to be 3/1 which is
the same as that used by Envirogenics (1)*
The spent solution leaving the absorbers is mixed with a
ca-lcium sulfite solution in the clarifier L-l. This unit size
was scaled down from the one given by Envirogenics since no
settling data are given in their report (1). Clear liquid leav-
ing the clarifier is sent to the ZnO tanks where ZnSC>3 is formed.
Again, the processing equipment needed for formation, crystalliza-
tion and thickening of ZnSC>3 is scaled down from the Envirogenics
report because of the absence of reported data on the reaction of
ZnO to form ZnSC>3 and crystal growth rate. Superdecanting centri-
fuges are assumed to be capable of separating the ZnS03 crystals
from the lean solution of sulfite-bisulfite. This lean solution
is refortified with soda ash to make up losses in a lean, solution
surge tank which was scaled down directly from that used in the
larger plant. The solution in the lean solution surge tank is
pumped to the stack gas absorbers, thus completing the cycle.
Wet crystals discharged from the centrifuges contain 17%
free water which is evaporated in a direct dryer to produce the
dry crystals of ZnSC>3 ' 2-1/2 H2O. Drying is carried out concur-
rently with combustion gases entering at 700°F and the dried solids
leaving the dryer at 200°F. No liberation of SC>2 would take place
in the dryer since only free water is evaporated. The dryer duty
used for design was scaled down from the larger plant owing to the
lack of specific heat data on ZnSC>3 • 2-1/2 H2O. However, as an
approximate check, the dryer duty was estimated independently and
the estimated and scaled-down heat requirements were in close
agreement.
The Envirogenics process flow sheet shows a cyclone to
recover the fines in the effluent flue gas from the dryer and a
hammer mill on the solids discharge stream. Since particle size
distribution data are not available for the dryer effluent solids
and, since "clinkers" may be formed in the dryer, use of the
* Refer to references at end of section.
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cyclone and grinding system was assumed necessary. The cyclone
was scaled down from the Envirogenics case. However, a grinding
mill system was substituted for the hammer mill shown by Enviro-
genics in order to produce a finer particle size, viz., 70% minus
200 mesh.
The dried zinc sulfite (hydrated) is feed to a tripping belt
feeder so that the ZnSC>3 can be distributed to a 5 bin hopper
which in turn is used to feed the 5 indirect kilns. Thus each
kiln has an independently controlled solids feed.
The ZnS03 • 2-1/2 H2O is indirectly calcined at the same
temperature used by Envirogenics (850°F). Five of these calciners
are required to handle the total solids flow since maximum kiln size
available is limited by the need for tight seals to prevent leakage
of air into the system which would produce sulfate. The estimated
heat requirement of the calciner was lower than that obtained by
scaling down the larger plant. Consequently, since enthalpy data
were not available, the scaled-down value (larger) was used to
insure an adequate design.
A cyclone is needed on the calciner effluent gas to recover
the entrained zinc oxide since 70% of the precalcined particles
were assumed to be less than 200 mesh. It was further assumed,
since no specific data were available, that essentially all the
calcined particles were greater than 40 microns, thus resulting
in a collection efficiency of about 100% in the cyclone. However,
if significant quantities of particles less than 40 microns are
produced, other means of dust collection probably would be required
to keep zinc oxide losses low.
An absorber tower is used to cool the gases leaving the
calciner cyclone. This absorber is simply a packed tower sized
to give the gas cooling and water absorption required. Recircu-
lated water is used as the absorbent and the heat gained in the
water is removed by an external exchanger.
The SC>2 leaving the water absorber system is in part sent to
the acid plant for conversion to concentrated (98.5%) sulfuric
acid. Product acid storage equivalent to 30 days' production was
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assumed necessary to allow for continuous plant operation even
if modest delays in shipping (e.g., R.R. strike) were encountered.
The unloading system is designed such that a day's production of
sulfuric acid can be pumped out of storage in six hours.
The remaining SC>2 is recycled and used to dissolve the
insoluble CaSC>3 in the gasifier. This is accomplished in a
continuous stirred tank reactor equipped with a turbine-type
agitator. This reactor holding time is based on that used by
Envirogenics.
Liquid leaving the gasifier contains insoluble CaSC>4 and
fly ash. This slurry is thickened in a waste thickener which
has been sized by scaling.
Clean liquor from the waste thickener is sent to the liming
tanks where insoluble CaSC>3 is formed as shown by the process
reactions given earlier. These liming vessels are agitated and
have simple covers. Two reactors in series are used to minimize
by-passing of the CaO and to insure formation of the CaSO3.
A tilting pan filter with three water wash stages is used
to recover the occluded liquor from the thickened waste slurry.
Sizing of this filter is based on filtration data for gypsum.
Waste solids are sluiced to a slurry tank to give a 10 wt % slurry
and the slurried waste is pumped to a settling pond.
Solid hopper feeder systems for make-up ZnO and soda ash are
sized to hold a week's supply of feed. To insure proper control,
weigh feeders are used for adding the make-up ZnO and soda ash
to the system. The lime storage is for one week but the lime
feeder-hopper is only one day. Lime is fed into the liming tank
using a weigh feeder to insure proper control. To insure that the
dried ZnSC>3 • 2-1/2 H2O is properly distributed to each of the five
calcining kilns, weigh feeders are again used. Delivery of make-up
solids is assumed to be by truck with pneumatic unloading and all
solids hoppers are vented to the atmosphere through bag filters for
dust control.
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4. Process Appraisal
The wet scrubbing zinc oxide process has been shown to' be
technically feasible (3). Although a process flow sheet and
material balance were prepared and major equipment were sized by
Envirogenics (1), additional data are required in every facet of
this process before a definitive design can be made. The major
advantages and disadvantages of this process will be given before
discussing those areas which need additional information.
a. Advantages
• Wet scrubbing of the flue gas eliminates the need for high
efficiency precipitators. Fly ash not removed in the pre-
cipitator is removed with the scrubbing solution.
• Calcination of a sulfite (ZnSC>3 • 2-1/2 H2O) to produce by-
product SC>2 and regenerated ZnO is carried out at a rela-
tively low temperature.
• High degrees of SC>2 removal from the flue gas should be
readily obtained.
• Concentrated H2S04 is produced rather than dilute acid
as in some processes.
b. Disadvantages
• Presence of nitrogen oxides in the flue gas causes the
loss of sodium values in the scrubbing liquor.
• Oxidation of sulfite to sulfate results in a loss of
calcium thus increasing both operating costs and the
quantity of waste.
• Wet scrubbing cools and saturates the flue gas and
requires reheat to get plume rise.
• Disproportionation of zinc sulfite in the kilns forms
zinc sulfide and represents a potential loss of zinc
or will require expensive additional processing.
• By-product sulfuric acid is produced instead of elemental
sulfur, thus complicating storage and handling problems.
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c. General Considerations
The various individual steps in the zinc oxide process have
been experimentally demonstrated (3). The literature contains
some laboratory data which have been obtained on this process.
Although the process concept has been demonstrated, additional
processing and thermodynamic data are required in all areas of
this process before a prototype demonstration unit can be designed
in detail.
The thermodynamic requirements are as follows.
• Vapor pressure of sulfur dioxide over aqueous solutions
of sodium sulfite-bisulfite. Some data are available in
the literature but should be checked for thermodynamic
consistency to determine if additional experimentation
is required.
* Solubility data on the different precipitated salts, i.e.,
zinc sulfite, calcium sulfite, calcium sulfate.
• Heat of reaction data may be available in the literature.
• Specific heats of zinc sulfite and zinc oxide
• Heat of hydration of the zinc sulfite.
The different chemical processing steps of the process also
must be studied experimentally. Processing studies are required
in all the seeps of the process and are as follows.
* Scrubbing
As pointed out above, vapor pressure data of sulfur
dioxide over the sodium sulfite-bisulfite system are
required. In addition, processing studies must be
made to obtain data on the effects of variables such
as temperature and liquid—to-gas—rate on mass transfer
rate. The effect of nitrogen oxides on the absorption
step must be determined. Nitrogen oxides, oxygen and
possibly fly ash can affect the amount of sulfate formed
and the effect of these process variables must be more
clearly defined.
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Liming Tank
Particulate lime in suspension in the presence of
sulfate ion can form a shell of calcium sulfate
which hinders the hydration and dissolution reactions.
The same phenomenon also can occur in the precipita-
tion of zinc sulfite from zinc oxide. Unless particle
size is properly controlled, full utilization of the
zinc and calcium oxides will not be realized. Con-
sequently quantities larger than stoichiometric must
be used thereby increasing the solids load through
the plant.
This reaction must be studied to determine the reaction
rate of the calcium oxide with the sodium bisulfite.
Since the sulfate in solution can coat the calcium
oxide, the effect of particle size on conversion (i.e.,
calcium oxide utilization) should be determined.
Gasifier
Temperature, pressure and sulfate concentration effects
on dissolution of sulfite and precipitation of sulfate
need to be determined.
Mixer
Rate data need to be determined experimentally on the
formation of zinc sulfite, including the effects of
temperature and sodium bisulfite concentration. Parti-
cle size of the zinc oxide may affect the conversion
level owing to a coating of zinc sulfite on the particle
surfaces. This effect must be understood to determine
if greater than stoichiometric quantities are needed.
Calciner
Additional experimental data on the calcination of
zinc sulfite are needed to enable proper design of
the reactor. The flash-type calciner recommended
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by Field et al (3) may pose severe design problems.
Solids transport data are required to determine if
this reactor is desirable or if an indirect rotary
kiln is preferable. In calcination, the degree of
agglomeration and attrition of solids need to be
understood to determine if grinders and cyclones
also are required. The problem of disproportiona-
tion has in part been studied (2) but more detailed
experimental data are required before this problem
can be completely understood.
Dryer
Drying data on the wet zinc sulfite are needed since
sizing the drying kiln depends on the heat transfer
rate. Drying rate will depend on the hot gas cir-
culation rate and the particle size distribution of
the wet crystals. The quantity and size distribution
of the fines produced by attrition and the degree of
agglomeration of particles is needed for sizing the
cyclone and grinding mill downstream of the dryer.
In addition to the above basic areas of study, studies
should also be made in the following.
Miscellaneous
Settling data are required for proper sizing of the
clarifier, zinc sulfite thickener and waste thickener.
The angle of repose and flow characteristics of the
solids used in the process should be determined to
provide design data for storage, feeding, and convey-
ing equipment.
The rheological properties of the thickened solids
slurries should be determined to insure that pumps
are properly designed.
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In.addition to the above studies made on the individual
steps of this process, an integrated pilot plant should
be operated to demonstrate the total system. Recycle
streams may build up undesirable by—product which may
necessitate bleed streams or additional processing.
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REFERENCES
The process design and evaluation are based on the references
stated below. No other sources of data were made available to
Kellogg for this evaluation.
(1) Envirogenics Company, A Division of Aerojet-General
Corp., Applicability of Aqueous Solutions to the
Removal of S02 From Flue Gases, Final Report, Volume
1, Contract #PA 86-68-77, Submitted to National Air
Pollution Control Administration (Oct. 1970).
(2) Ibid, Volume 2.
(3) Field, J. H., Brunn, L. W., Haynes, W. P. and Benson, H. B.,
Cost Estimate of Liquid Scrubbing Process for Removing
Sulfur Dioxide From Flue Gases, Bureau of Mines, Report
of Investigations 5469 (1959).
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B. CHEMISTRY
In the zinc oxide process the stack gas is scrubbed counter-
currently with a sodium sulfite solution. The sulfur dioxide in
the stack gas reacts with the sodium sulfite to produce sodium
bisulfite.
Na2S03 (1) + H20(l) + S02 (g) ^ 2NaHS03 (1)
The effluent from the scrubber is mixed with zinc oxide to form
insoluble zinc sulfite and to regenerate sodium sulfite.
ZnO + 2NaHS03 + 1-1/2 H20 ->- ZnS03• 2-1/2 H2O + Na2SO3 (2)
The solid zinc sulfite is separated and calcined to regen-
erate zinc oxide and produce sulfur dioxide in a concentrated
form for recovery as such or conversion to elemental sulfur or
sulfuric acid.
ZnSOa-2-1/2 H2O ->• ZnO + SO2 + 2-1/2 H20 (3)
Perhaps the most vexing problem of all bulk water scrubbing
systems is the oxidation of the sulfur dioxide-sulfite system to
a sulfur trioxide-sulfate system due to the presence of oxygen,
nitrogen oxides and/or fly ash. This problem is of importance
in an oxidation product in that sulfate ion, formed as an oxida-
tion product in the scrubber, does not produce an insoluble zinc
salt which can be regenerated to zinc oxide and, therefore, builds
up in the scrubber by consuming effective Na20 content. Some
sulfate is also formed in the calciner which also finds its way
into the scrubber. Data of Johnstone and Singh (2) show that the
overall production of sulfate amounts to 10% whereas the Bureau
of Mines reports a 14% overall production of sulfate.
A separate system is incorporated in this process to remove
the sulfate by converting it to calcium sulfate.
2NaHS03(l) + CaO(s) -* Na2S03 (1) + CaS03 + H20 (4)
CaS03(s) -I- H20(l) + S02(g) -> Ca(HSO3)2(l) (5)
Ca(HS03)2(l) + Na2S04(l) •*• 2NaHS03 (1) + CaS04 (6)
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A side stream of the sulfate-containing scrubber solution is
reacted with a less than stoichiometric quantity of calcium oxide
to produce sodium sulfite and insoluble calcium sulfite. The
solid calcium sulfite is further reacted with product sulfur
dioxide to produce soluble calcium bisulfite. The above step
provides the calcium ion concentration necessary to precipitate
only calcium sulfate.
The solid calcium sulfate is separated to waste and the liquid
bisulfite recycled to the liming step. Since only a portion of
the sulfate ion is being removed there will be a steady state
concentration of sulfate ion in the entire system.
Aside from the problem of oxidation of sulfite to sulfate
in the calciner is the problem of the disproportionation reaction.
4ZnSO3-2-l/2 H20 + 3ZnSC>4 + ZnS + 10H2O (7)
The presence of zinc sulfite in this system would be detrimental
to its operation since it will not decompose to zinc oxide, is
difficult to remove and will not react with bisulfite solution.
Therefore, any appreciable production of ZnS per pass will eventu-
ally use up all of the zinc oxide.
A systematic study of the disproportionation reaction shows
that this reaction begins at 300°C, increase in rate with increas-
ing temperature to 375°C and then decreases in rate with a further
increase in temperature. According to Johnstone (1) at 350°C
disproportionation occurs to the extent of 4%. Subsequent work (1)
indicates that if the calcining is done in the presence of steam,
no disproportionation occurs since the decomposition of ZnSO3•2-1/2
H2O can be carried out at 290°C.
It has also been reported (1) that the maximum concentration
of sodium ions in the scrubbing solution should be limited to 3.5
moles/100 moles of water to prevent the formation of the double
salt of sodium and zinc sulfites.
6NaHSC>3 + 3ZnO + H20 -»• Na2SO3 • 3ZnSC>3 • 4H20 + 2Na2S03 (8)
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It has been shown from other investigations that particulate
lime in suspension in the presence of sulfate ion will form a
shell or skin of calcium sulfate which hinders the hydration and
dissolution reactions. As a consequence the full utilization of
the lime is not realized. It is conceivable that this same
phenomenon will occur during the precipitation of ZnS03 using
particulate zinc oxide. The result of this condition is that
the amount of solid zinc oxide which must be cycled will be much
larger than the stoichiometric quantities required by the chemistry.
Others
Since the stack gas contains nitrogen oxides and fly ash as
well as other gases, consideration should be given to the effect
of these components on the overall process. Unfortunately, the
chemistry of the interactions of NO + N02 with sulfite and bisulfite
solutions is not clearly understood. There are data however that
indicate that the presence of N02 in the gas stream greatly in-
creases the rate of oxidation of sulfite to sulfate. There is no
apparent effect of NO and NO2 on the efficiency of absorption of
sulfur dioxide.
Similarly the presence of fly ash has been shown to increase
the rate of oxidation. The direct causes of the increase in the
rate of oxidation have not been established but the end result is
that the increase in the amount of sulfate produced obviously will
require a larger clean-up process.
Summary
In summary, the zinc oxide process has the advantage of using
a homogeneous scrubbing system as compared to a solids injection
system and scrubbing with slurries. The major difficulties appear
to be in the oxidation of sulfite to sulfate, the hydration of
lime and tne presence of nitrogen oxides and fly ash. A major
process improvement would be to suppress or eliminate the oxida-
tion reaction with a subsequent saving in process steps.
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REFERENCES
(1) Envirogenics Company, A Division of Aerojet-General
Corp., Applicability of Aqueous Solutions to the
Removal of SO? From Flue Gases, Final Report, Volume
1, Contract #PA 86-68-77, Submitted to National Air
Pollution Control Administration (October 1970).
(2) Ibid, Volume 2.
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CITRATE PROCESS
A. PROCESS
1. Process Description (See MWK Dwg. No. PD-121-D)
In the citrate process, S02 is removed from the flue gas by
contacting the gas stream with an aqueous solution of citric acid
and sodium carbonate. The S02, which complexes with the citrate
ion, is then reacted with H2S to precipitate sulfur. After sepa-
ration of the sulfur from the citrate solution, about one-third
of the sulfur, equivalent to the amount absorbed less losses, is
removed as product while two-thirds are converted to H2S for
recycle.
Flue gas from the air preheater is split into 6 parallel
trains and cooled from 300°F to 157°F in coolers C-l A-F. The
gas is then scrubbed with an ash slurry in E-l A-F to remove 98%
of the fly ash and all of the 303, while cooling the gas to 105°F.
A low gas temperature and minimal 303 content are important for
efficient S02 removal in the subsequent absorption step. Part of
the slurry from E-l A-F is purged to D-5 where it is neutralized
with limestone prior to disposal as waste. The bulk of the slurry
is recycled through coolers C-3 A-F to the scrubbers.
Cooled flue gas from E-l A-F is sent to absorbers E-2 A-F
where 90% of the S02 is removed from the gas by contact with a
citric acid/sodium carbonate solution at 115°F. The cleaned
flue gas leaves the absorbers saturated with water at 107°F, is
reheated in C-3 A-F to 250°F (using the heat recovered from the
flue gas in C-l A-F) and compressed in J-l A-F to stack pressure
before being exhausted to the atmosphere. Water,which is produced
in the H2S-S02 reaction step, is removed from the citrate solution
in the absorbers by saturating the flue gas through contact with
the warmer liquid.
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The S02-rich solution from E-2 A-F is reacted with H2S in
D-l A-F to precipitate sulfur according to the reaction:
S02 + 2H2 S £ 3S + 2H20
Part of the solution bypasses D-l A-F and is combined with the
bottoms to react with any dissolved H2S which leaves the reactor.
The resultant sulfur slurry is thickened from 0.5% solids to 10%
solids in thickener L-3 and centrifuged in L-4A-G to remove all
solids. The filtrate from L-4A-G is combined with the thickener
overflow and sent to citrate feed tank F-2 for recycle to the
absorbers. Residual solids in the solution are removed in line
filters L-5 A,B.
Wet cake from centrifuges L-4A-G is conveyed to tank F-4
where it is re-slurried with part of the thickener underflow.
This slurry, containing 33.3% solids, is heated in C-6 from 117°F
to 280°F, thereby melting the sulfur. Phase separation of the
liquid sulfur and the residual citrate solution occurs in F-5.
The citrate solution is removed, cooled to 120°F in C-4, and
pumped back to the feed tank F-2. Any SO2 dissolved in the
citrate solution from F-4, and released by heating, is vented
from F-5 back to the absorbers E-2.
Liquid sulfur from F-5 contains the fly ash not removed in
E-lA-F. This stream is centrifuged in L-6A,B. The ash is slurried
in water in tank F-13 and pumped to waste disposal. Purified
sulfur from L-6A,B is combined with recovered sulfur from sulfur
separator F-l. About one-third of the total sulfur is pumped to
storage as product while the remainder is recycled for conversion
to H2S. The net sulfur product equals the total sulfur removed
from the flue gas less losses.
Recycle sulfur, at 276°F, is heated to 550°F in sulfur pre-
heater F-6, using a non-flow system in order to circumvent the
peculiar viscosity characteristics of sulfur in the intervening
temperature range. Sulfur from F-6 is vaporized and superheated
to 1200°F in B-l, which also preheats a methane feed stream to
759°F. The two streams are combined and fed to carbon disulfide
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reactor at 1165°F and 40 psia where the following reaction occurs
over a silica gel catalyst:
CH4 + 4S J CS2 + 2H2S
At the process conditions indicated, 90% of the methane is converted
and the reaction is slightly exothermic.
The reactor effluent, at 1200°F, is used to preheat sulfur in
F-6, after which it is cooled in a series of heat exchangers (C-7,
C-8, C-9) to condense unreacted sulfur for recycle. Condensed
sulfur is removed from the gas in sulfur separator F-l and the
effluent gas, after heating to 600°F, is combined with superheated
steam and fed to the hydrolyzer D-4 at 493°F and 32.5 psia.
In D-4, CS2 is hydrolyzed over an activated alumina catalyst
according to the following reaction:
CS2 + 2H2O t 2H2S + C02
About 98% of the CS2 is converted, and the temperature of the
reaction is controlled at 600°F by generating steam. The effluent
is cooled to 200°F in a series of heat exchangers (C-10, C-ll, C-12).
In addition to that produced in D-4, steam is also generated
in exchangers C-8 and C-ll. The combined steam from these vessels
is used in C-6 to melt sulfur. A balance is maintained between
the steam generated and the steam required, resulting in a closed-
loop system.
The cooled gas from C-12 is fed to reactors D-l A-F where the
H2S-SO2 reaction occurs. The effluent gas from D-l containing both
CS2 and H2S, is sent to absorber D-2A or D-2B which removes CS2 by
adsorption on activated carbon. The resulting gas stream is vented
to stack through E-2A absorber which removes H2S by reaction with
the SO2-rich citrate solution. This results in precipitation of
some sulfur in the absorber. Adsorbed CS2 in D-2A, B is removed
with superheated steam and recycled to hydrolyzer D-4.
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To remove sulfate formed either in the process or via 803
carry-over to the absorber, a purge stream is taken from the
recovered citrate solution and pumped by J-ll to neutralizer D-5
before disposal. Make-up solution from F-3 is added to feed tank
F-2 as required.
2. Process Design Bases
General design criteria used in evaluating this process, which
is designed for installation in a new 1000 MW power plant, are those
given elsewhere in this report as the basic criteria for all the
present evaluations. The process design and material balance shown
on Dwg. PD-121-D for the citrate process are based primarily on
brief status reports and summaries of laboratory and pilot plant
tests by the U.S. Bureau of Mines, including equilibrium data for
the SO2~absorption step. Basic design data for the sulfur recycl-
ing step, i.e., the conversion of sulfur to H2S for recycle to the
sulfur reactor, were lacking however, and these were obtained from
the literature (1,2,3). Major design bases are listed below.
a. Absorbers E-2 A-F
gas inlet temperature 105°F
citrate solution inlet temperature 115°F
S02 removal from gas 90%
SO2 loading on solution 4.3 gms S02/1 solu (approx)
citric acid concentration 0.87 gm moles/1 (approx)
Na/citric acid ratio 5.3 atoms/mole (approx)
b. Sulfur Reactors D-l A-F
SC>2 conversion 100%
excess H2S 5%
inlet superficial gas velocity (nominal) 0.1 fps
liquid residence time 10 minutes
c. Carbon Disulfide Reactor D-3
temperature 1200°F
space velocity (SCFH gas/CF catalyst) 560
catalyst silica gel
methane purity 100%
methane conversion 90%
excess sulfur 10%
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d. Hydrolyzer D-4
temperature 600°F
contact time 0.4 seconds
catalyst activated alumina
carbon disulfide conversion 98%
excess steam 100%
3. Process Design Rationale
The citrate process has been developed through initial pilot
plant stages by the U.S. Bureau of Mines. However, studies have
been done primarily on smelter gases containing 1-3 mole percent
S02 rather than on flue gases having a low SC>2 content. The in-
formation received, which contains equilibrium data on SO2 absorp-
tion by solutions of citric acid/sodium citrate, also suggests
that the major experimental effort has been directed towards the
absorption step. Although the pilot plant has demonstrated sulfur
recovery and recycle of the citrate solution, no attempt was made
to recycle sulfur via conversion to H2S.
The present process design is based, as far as possible, on
the results of laboratory and pilot plant studies by the Bureau
of Mines. Where necessary, process conditions were adjusted from
those indicated by the U.S.B.M. to reflect the lower S02 gas con-
centration and other pertinent process factors that apply when
treating power plant stack gases.
Prior to absorption of SO2 by the citrate solution, 663 must
be removed from the flue gas to minimize sulfate build-up in the
recirculating solutions. This is done by scrubbing the gas with
water, which also serves to remove fly ash while cooling the gas
to 105°F.
For efficient absorption, temperatures in the absorber should
be kept as low as possible. However, other factors such as gas
cooling and reheat requirements resulted in the selection of a gas
temperature of 105°F as a practical minimum. The solution tempera-
ture was chosen not only to allow absorption to occur, but also to
provide a slight thermal driving force which serves to heat and
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saturate the gas, thereby removing water which is formed elsewhere
in the process. At the temperatures selected, equilibrium curves
show that citrate solution loadings of greater than 5 gms SC>2/liter
solution can be achieved. The design loading of 4.3 gms SC>2/liter
solution represents roughly an 80% approach to equilibrium.
Sulfate which is formed in the process must be continually
purged from the system. In order to reduce the accompanying loss
of citric acid to an acceptable level, it was assumed that sulfate
could be allowed to build up to a high level in the solution before
being purged. The required pH could be maintained by addition of
soda ash. This scheme results in approximately the same citric
acid concentration as that used by the Bureau of Mines (0.87 molar
vs. 1.0 molar), but a higher sodium to citric acid ratio (5.3 vs.
1.5). If this does not prove to be feasible, the purge stream
would have to be treated to remove sulfate before being returned
to the process. It is not expected that this would have any
significant effect on either process operation or economics.
Very little information was available for the design of the
sulfur reactor, where sulfur is precipitated by reaction of SC>2
and H2S. The Bureau of Mines states that the required residence
time of the loaded citrate solution is about 10 minutes and that
reaction occurs in a stirred, closed vessel. This was used for
design, and it was assumed that 5% excess K^S would keep a suffi-
cient high back pressure of P^S over the solution to drive the
reaction to completion.
Some H2S may dissolve in the citrate solution in the sulfur
reactor. To eliminate its presence (and the possibility of its
release) downstream in the process, a by-pass stream containing
dissolved SC>2 is mixed with the reactor bottoms prior to thicken-
ing, thereby converting any t^S present to sulfur.
will also leave the reactor in the gas phase. It is
expected that this stream could be vented through the absorber,
thus reacting the H2S to extinction. This precipitates sulfur
in the absorber, and its effect on tower operation is not known.
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The H2S required in the sulfur reactor is produced by the
hydrolysis of carbon disulfide, CS2. The behavior of residual
CS2 in the sulfur feed gas is not known. Presumably it will
not react in the sulfur reactors, but will be carried out in the
effluent gas. The CS2 must be removed from the gas prior to vent-
in^, and one method of achieving this, viz., adsorption on activated
^hircodl, is shown on the flow sheet. The effects and behavior of
oUier sulfur compounds, such as COS, which may be present in the
reactor feed gas also are uncertain.
Since the Bureau of Mines has not recycled sulfur (as H2S)
in the pilot plant, no information was obtained from them for the
design of this system. Several approaches could be taken, includ-
ing the following possible routes:
1) natural gas reforming followed by sulfiding
2) single-step reaction of methane, steam, and sulfur
3) two-step reaction system consisting of reaction between
methane and sulfur to form CS2 and P^S, followed by
hydrolysis of the CS2 to H2S.
Some information for the design of the latter two was available
in the literature, and the two-step system was selected for the
present flow sheet design. This system minimizes formation of
by-products such as COS and S02 by sulfiding methane in the
absence of oxygen. The subsequent hydrolysis of CS2 proceeds
easily at moderate temperatures with high conversions. The
system chosen represents a workable scheme, but no attempt at
optimization was made.
Recovery of sulfur from the citrate solution and recycling
of the solution has been demonstrated by the Bureau of Mines in
the pilot plant. The present flow sheet basically follows the
pilot plant scheme, although no design data were available for
sizing purposes.
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4. Process Appraisal
Based on the available information, the citrate process is
a workable system for the efficient removal of S02 from power
plant stack gases. Data do not appear to be missing for any
critical areas which would cause the feasibility of the process
to be questioned from a technical point of view.
Design data, in most instances, are lacking or incomplete
and more development work is needed to adequately define process
variables. Equilibrium data for SC>2 absorption are available, but
these should be confirmed in the low range of SO2 gas concentra-
tions encountered in flue gas streams. Solubility and kinetic
data are needed to establish the design of the sulfur reactor,
and an investigation to determine the best commercial design
(sprayed vessel, staged reactor, etc) should be made.
The H2S generating system, although not piloted by the
Bureau of Mines, is reasonably firm. According to the literature (3),
the production of CS2 from methane and sulfur has been commercial-
ized. The subsequent hydrolysis of CS2 is easily performed and
represents no new technology.
Pilot plant testing on simulated flue gases containing fly
ash is needed. Recycle of both citrate solution and sulfur should
be demonstrated to determine the effects and levels of impurities
in the system. It is also necessary to establish the permissible
sulfate level in the citrate solution.
The preceeding points were mentioned to point out some of the
deficiencies in existing design data. A confident commercial design
could only be made after resolution of these deficiencies through
further experimental work.
One advantage of the citrate process is that the sulfur is
recovered in elemental form. It is more easily handled, stored,
and marketed compared to other forms of sulfur (e.g. sulfuric acid)
which are produced by some of the other SO2~removal processes.
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The citrate process should be easily adaptable to varying
sulfur content in the power plant fuel as well as varying turn-
down ratios (corresponding to the cycling of power demand). Six
separate trains on the front end of the process operate indepen-
dently, and since the entire process with the exception of the
H2S generating section operates at about ambient temperature,
start-up of the single train can be done easily by establishing
circulation of the citrate solution. Shutdown is just as readily
achieved. No heating or cooling periods are necessary in either
case.
A major disadvantage of the process is that since absorption
must be done cold (about 110-115°F) and in an SO3~free environment,
the gas must be cooled, scrubbed, and reheated before being dis-
charged to stack. This requires extensive equipment and represents
a significant part of the overall cost of the process.
Since fly ash is removed from the flue gas during the water
scrubbing operation .to remove SG>3, electrostatic precipitators
are not required. For a new power plant, therefore, elimination
of the precipitators represents a reduction in investment charged
to the citrate process. The cost of the wet scrubbers would, of
course, comprise a portion of the citrate investment. Retrofit
to an existing power plant would pose no unusual problems, provided
sufficient area were available, since the flue gas is withdrawn
just upstream of the stack. Existing precipitators could be left
in service or by-passed without adversely affecting operability
of the citrate process.
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REFERENCES
1. Bacon, R. F. , and Boe, E. S., Ind. Eng. Chem., 37, 469-474
(1945).
2. Fisher, R. A., and Smith, J. M., Ind. Eng. Chem., 42, 704-
709 (1950).
3. Folkins, H. 0., Miller, E., and Hennig, H. , Ind. Eng. Chem.,
42, 2202-2207 (1950).
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B . CHEMISTRY
The Bureau of Mines citrate process has as its core,
the facile near quantitative reaction of sulfur dioxide
with hydrogen sulfide at low temperature (50°C).
S02 + 2H2S * 3S -i- 2H20
Since the hydrogen sulfide cannot be present in the flue
gas absorption tower, a 'conveyor1 is used to move the
sulfur dioxide to a place where hydrogen sulfide can be
present - the stirred closed reactor. The 'conveyor' is
the citrate anion, produced from citric acid and sodium
carbonate. This complexes with bisulfite (sulfur dioxide
in water).
Other compounds will be formed in the hydrogen sul-
fide reactor, and if the pH is not carefully controlled,
in the SC>2 absorber solution as well. Thiosulfate,
po lythionates and sulfate are likely the only ones of
consequence and these products and their equilibria as
outlined by the Bureau of Mines are consistent with the
experimental work of other groups. For more efficient
operation of the citrate process, they should be kept
to as low a value as possible. It is indicated that
they do not interfere in the process if their concentra-
tions are 1ow.
The pH plays a most important role in the citrate
process and allied closely to this is the problem of
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sulfur trioxide carry-over into the SOo absorber due to
incomplete gas washing. SOg carry-over will eventually
lower the pH of the S02 absorbing solution and reduce
its effectiveness as an SC>2 absorber (.502 is not ab-
sorbed in strong acid solutions). Strong acid tends to
decompose thiosulfate and polythionates according to the
reactions
8 SS03 + 16 H+ ) 8 S02 + 8 H20 + Sg
S40^ + 2H+ + H20 ^ H2S04 + x/4 S8 + S02 + H2°
thus possibly precipitating sulfur in the S02 absorber.
Sodium carbonate could be added to the SO^ absorber to
neutralize 503, but this action would be limited since
it would lead to increased sulfate concentration in the
absorbing liquor.
Sulfate can form via 803 carry over into the S02
absorber and also in the hydrogen sulfide reactor via
the reaction
3HSO~ » 250^ + S + H+ + H20
and even if the carry-over problem is solved, it appears
that there will be slight continuous sulfate buildup
in the citrate liquor, eventually necessitating some
treatment to remove sulfate.
Carbon dioxide is not expected to be absorbed in the
citrate solution. The solution already has a high salt
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content (190 g/1 citric acid, 80 g/1 sodium carbonate,
20 g/1 thiosulfate + smaller amounts of polythionate and
sulfate), and a pH range 3.2 - 3.7. Water saturated with
carbon dioxide has a pH of 3.7. Possible nitrogen oxide
absorption in the citrate solution has not been studied
but Bureau of Mines indicates it has begun investigating
NOX absorption in the gas washing and S02 absorption
steps .
In conclusion, the SO^ removal from the flue gas is
an important procedure in this process. A scheme to re-
move sulfate also appears necessary. If these steps can
be accomplished, the process (on a chemical basis) appears
suitable for smelter gas clean-up.
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BIBLIOGRAPHIC DATA 1- Report No. 2.
SHEET ' APTD-0807
4. Title and Subtitle
EVALUATION OF S02 - CONTROL PROCESSES
7. Author(s)
9. Performing Organization Name and Address
The M. W. Kellogg Company
Research & Engineering Development
Piscataway, N. J.
12. Sponsoring Organization Name and Address
Office of Air Programs
The Environmental Protection Agency
Research Triangle Park, N. C. 27711
3. Recipient's Accession No.
5- Report Date
October 15, 1971
6.
8- Performing Organization Rept. !
No. ''
10. Project/Task/Work Unit No.
11. Contract/Grant No.
CPA 70-68 '
13. Type of Report & Period i
Covered 1
Final
14.
is. supplementary Notes DISCLAIMER - This report was furnished to the Office of Air Programs
by The M. W. Kellogg Company , Research & Engineering Development, Piscataway, N. J,
_Jn fnlfi'llmpnt nf r.nnt.rart No. P.PA 70-fift
16. Abstracts
The results are reported of technical and economic evaluations of the feasibility
of different SO^-control processes. In addition to process evaluations, studies
were made of; the applicability of these processes to both power plants and
S02~control processes, and the impact S02~control technology would have on emission
reduction. Preliminary process designs and economic evaluations have been prepared
for twelve different S0.2~control processes. Investments are reported in millions
of dollars for a 1,000 MW plant so the investment cost in dollars per Kilowatt has
the same value. The operating costs are reported as dollars per ton of sulfur not
emitted, as mills per Kilowatt-hour, and as a percent increase of power cost. Of
the twelve processes evaluated, none can be ranked as commercially available at the
present time.
17. Key Words and Document Analysis. 17a. Descriptors
Air pol1ut ion
Desulfuri zati on
Sulfur di oxi de
Evaluat ion
Economic analysis
Feas i b i1i ty
Investments
Costs comparison
17b. Identifiers/Open-landed Terms
Air pollution control
17c. COSATl Field/Group
18. Availability Statement
Unlimi ted
19. Security Class (This
Report)
UNCLASSIFIED
20. Security Class (This
Page
UNCLASSIFIED
21. No. of Pages
271
22. Pric
FORM NTIS-35 UO-70)
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