DEVELOPMENT OF A MOLTEN CARBONATE PROCESS
FOR REMOVAL OF SULFUR DIOXIDE FROM
POWER PLANT STACK GASES
PROGRESS REPORT NO. 4
AUGUST 1, 1969 TO MARCH 19, 1971
Atomics International
North American Rockwell
P.O Box 309
Canoga Park. California 91304
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AI-71-37
DEVELOPMENT OF A MOLTEN CARBONATE PROCESS
FOR REMOVAL OF SULFUR DIOXIDE FROM
POWER PLANT STACK GASES
PROGRESS REPORT NO. 4
August 1, 1969 to March 19, 1971
July 28, 1971
The work upon which this publication is based was performed
by Atomics International, a Division of North American
Rockwell Corporation, Canoga Park, California, pursuant to
Contract No. CPA 70-78 with the U. S. Environmental Protec-
tion Agency. Patents covering the basic process have been
issued to North American Rockwell Corporation.
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Errata for AI-71-37
Page A-48, Figure 16
"d = 20 ft and d = 5 ft" on curves should read
D = 20 ft and D = 5 ft
"Expanded Bed Depth = 0. 4 d" in legend should read
Expanded Bed Depth = 0. 4 D.
Page A-49, Figure 17
"Expanded Bed Depth = 0. 4 d" in legend should read
Expanded Bed Depth = 0. 4 D.
Page A-54, Figure 18
"Kroft Furnace 600 psia Steam" on curve should read
Kraft Furnace 600 psia Steam
Units on ker^TTT T i** legend are
Btu/ft2-°F-hr
Page B-3, Line 16
"K2 = 1. 24 x 10"3 327' 720/RT" should read
K = 1.24xlO-3e27>720/RT
b
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Report No. AI-71-37
NOTICE TO RECIPIENTS
This document contains inventive subject matter and may not be
published, disseminated to others, or used for any purpose
without prior approval from the Patent Branch of the Environ-
mental Protection Agency and from the Patent Counsel of the
Contractor.
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TABLE OF CONTENTS
Page
Summary ..... .... . . iii
I. Introduction ... ....... 1
A. Process Description ...... .1
B. Status of Development of the Process ... 4
1. Scrubbing .......... 5
2. Reduction ....... . 5
3. Regeneration ....... .7
4. Filtration . ... . 7
5. Lithium Recovery System ....... 8
6. Materials and Components . .8
C. Development Program During Report Period . . 9
II. Progress During Report Period . . . . . . .11
A. Filtration Studies 11
1. Introduction ....... ..11
2. Experimentation ... .... 11
3. Evaluation ....... 13
B. Mist Recovery Tests 16
C. Materials Studies 20
1. Introduction .......... 20
2. Corrosion Tests ......... 21
3. Materials Test Loop 27
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TABLE OF CONTENTS (Contd)
Page
D. Reducer Engineering Analysis . . . . . . .27
E. Regenerator Engineering Analysis ...... 36
F. Heat and Mass Balances ........ 37
1. Mass Balance ......... 40
2. Heat Balance . . . ... . . . .42
G. Cooperation with Process Evaluator ..... 44
1. Introduction .......... 44
2. Conclusions and Recommendations ..... 44
3. Process Economics ... .... 45
III. Future Work 46
IV. Appendices
A. Preliminary Analysis of Two Region Molten
Carbonate Reducer ......... A-l
B. Preliminary Engineering Analysis of
Regenerator ........... B-l
11
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SUMMARY
This report presents the results of the process development work for the
•period August 1, 1969 to December 31, 1970. During this period, the follow-
ing advances were made:
(1) Filtration Studies
Fly ash filtration tests were performed with a Croll-Reynolds filter
cartridge having 25-micron spacings between the wire windings. The filter per-
formed well, with an average removal efficiency of 93%. Data from the tests
were used to obtain the relationship between throughput, pressure drop, and
ash loading over the range of ash loadings from 0 to 1. 24 Ib ash/ft filter area.
This relationship was then used in a reevaluation of the filter requirements for
an 800-Mw installation. The results showed that while Croll-Reynolds type
cartridge filters could be used in a full-scale unit, there would be an economic
advantage to using a continuous filter. However, this would probably require
building up a thicker cake, through use of a filter aid or filter precoat material;
tests of filter aids and precoats should be done.
(2) Mist Recovery Tests
Tests of wire-mesh mist eliminator pads were performed. Very high
mist recovery efficiencies were obtained, but the tests were hampered by the
lack of a source of large volumes of hot gas. Consequently, the tests were
limited to lower gas velocities than desired (6.5 ft/sec maximum) and to
short durations. These limitations made it impossible to obtain enough carry-
over to determine accurate mist eliminator efficiencies. Accurate data will
have to be obtained from longer tests with larger equipment, -such as in a pilot
plant.
111
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(3) Materials Studies
Five long-term dynamic corrosion tests and three thermal transport
tests were completed. In the dynamic tests, stressed and unstressed specimens
of Type 347 and Type 321 stainless steel were exposed at 932 F (500 C) to car-
bonate melts containing (1) 20% sulfide and 5% chloride (both steels); (2) 20% sul-
fite and 5% chloride (both steels); and (3) 40% sulfite and 5% chloride (Type 347
only). The tests were performed in rotating capsules which initially contained
atmospheres of nitrogen, oxygen, and water vapor. Four of the tests attained
the desired one-year lifetime; the fifth was terminated after 1580 hours when
the temperature controller failed, allowing the temperature to exceed 1200 F
(650°C). The results showed that
(a) Type 347 was superior to Type 321;
(b) Type 347 was satisfactory at 932°F in both the 20% sulfide and
20% sulfite melts, having a corrosion rate of about 1-3 mils/yr;
(c) There was no enhanced corrosion due to sensitization and no
chloride stress corrosion cracking; and
(d) The Type 347 specimens in the 40% sulfite-5% chloride melt were
severely corroded, probably due to overheating and thermal
cycling when the controller malfunctioned.
In the one-year thermal transport tests, capsules containing stressed
and unstressed test specimens in various melts were mounted in furnaces with
one end maintained at 932 F (500 C) and the other end maintained at 797 F
(425 C). The furnaces (and capsules) were inverted every 3-1/4 hours, causing
the melt to flow from one temperature zone to the other, the purpose being to
determine the extent to which thermal transport of alloy constituents would
occur. Specimens of Type 347 and Type 321 were tested in melts of carbon-
ate containing 20% sulfide and 10% chloride, and Type 347 specimens were
tested in a melt of carbonate containing 20% sulfite and 10% chloride. The
IV
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results showed that no thermal transport occurred, and the corrosion rates
were similar to those in the one-year dynamic tests. As a result of the dynamic
and thermal transport tests, Type 347 stainless steel remains the construction
material of choice for use in the low-temperature (below 950 F) part of the
process. However, further testing should be done to define the upper tempera-
ture limit for this alloy.
A forced-circulation materials test loop design was completed and the
procurement package for it was prepared. A revised design was also com-
pleted, in which a commercial pump is used and only Type 347 stainless steel
is tested. This latter design should be built and put into operation, to study
corrosion under flow conditions and to obtain experience with the pump, valves,
controls, and instrumentation.
(4) Reducer Engineering Analysis
An engineering analysis of the two-zone reducer using petroleum coke
was performed. The results showed that an alumina liner is needed to control
corrosion and reduce heat losses. The controlling design considerations were
found to be the maximum allowable air velocity in the oxidation region, and the
reaction rate (and required residence time) in the reduction zone. The dimen-
sions, heat losses and coke and air requirements for pilot-plant and full-sized
reducers were calculated. Experimental data are needed to verify the design;
this can be obtained by modeling the reducer hydraulics, studying the rate of
internal circulation and testing methods for introducing the melt, coke, and air.
(5) Regenerator Engineering Analysis
An engineering analysis of the regenerator was completed. The results
indicate that 9 theoretical plates will be required, and that melt cooling will be
required at a tray above the middle and at the regenerator outlet. The actual
tray location for cooling and the number of actual trays will have to be esti-
mated, and then verified when tray efficiencies are measured experimentally
in a pilot plant.
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(6) Heat and Mass Balances
Based upon a flow diagram which included a two-zone reducer using
petroleum coke, revised preliminary heat and mass balances were derived
for an 800 Mw installation.
(7) Cooperation with Process Evaluator
Information and assistance were provided Singmaster & Breyer in
their engineering evaluation of the Molten Carbonate Process. Data supplied
included process flow diagrams, material and energy balances, equipment con-
figuration, fly ash filtration data, and process economics. Singmaster &:
Breyer made a thorough evaluation of the process, pointing out potential prob-
lem areas and making valuable suggestions. They concluded that the process
was feasible and that its development should continue, with a pilot plant and
materials test loop as the next logical steps.
VI
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- 1-
I. INTRODUCTION
The Atomics International Division of the North American Rockwell Cor-
poration is developing a molten carbonate process to remove sulfur oxides from
power plant gas streams under contract with the Air Pollution Control Office of
the Environmental Protection Agency. Work was begun in June 1, 1967, under
contract PH 86-67-128 and has proceeded since then. There have been three
previous progress reports: the first covered the period from June 1, 1967, to
February 28, 1968; the second covered the period from February 29, 1968, to
October 27, 1968; and the third covered the period from October 28, 1968 to
July 31, 1969. This report covers the period from August 1, 1969, to Decem-
ber 31, 1970; the work done was under contract CPA 70-78.
A. PROCESS DESCRIPTION
In the Molten Carbonate Process, a molten eutectic mixture of lithium,
sodium, and potassium carbonates is used to scrub the power plant gas stream.
The sulfur oxides in the gas stream react with the carbonates to form sulfites
and sulfates, which remain dissolved in excess unreacted carbonate melt. The
molten carbonate-sulfite-sulfate mixture is then regenerated chemically, con-
verting the sulfite and sulfate back to carbonate and recovering the sulfur as
hydrogen sulfide. The regenerated carbonate is recirculated to the scrubber to
repeat the process cycle, and the hydrogen sulfide is converted to elemental
sulfur in a Glaus plant.
The regeneration of spent melt from the scrubber is done in two steps; first
a reduction of the sulfite and sulfate to sulfide, followed by conversion of the
sulfide to carbonate plus hydrogen sulfide. The reduction step is accomplished
by reaction with a form of carbon, such as petroleum coke. The conversion of
the sulfide to carbonate is accomplished by reacting the reduced melt with steam
and carbon dioxide, liberating hydrogen sulfide.
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FLUE GAS
RETURN
TO BOILER
SCRUBBER
850° F
PRECIPITATOR
850°F
t
FLUE GAS
FROM BOILER
71-F4-7-16A
CARBONATE
MAKEUP
FLY ASH
CLAUS
PLANT
SULFUR
REGENERATOR
850° F
CARBONATE
ANDSULFIDE
*T*
HEAT
EXCHANGER
QUENCH TANK
950°F
COKE FILTER
CAKE
ASH FILTER
850° F
T
CARBONATE,
SULFITE
ANDSULFATE
REDUCER
1500°F
CAKE
CO,
PETROLEUM
COKE
AIR
I
tv)
i
Figure 1. Molten Carbonate Process Flow Diagram
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-3-
The alkali carbonate melt used in the process has the following advan-
tageous features:
(1) It is a liquid, easy to handle, pump, and transport;
(2) It has a negligible vapor pressure, so that it is not lost by evaporation;
(3) It reacts rapidly with sulfur oxides, so that scrubbing contact time can
be short;
(4) It has a high capacity for sulfur oxides, so that the amount of melt
being circulated is relatively small; and
(5) The chemical affinity with sulfur oxides is so great that it can remove
nearly all of the sulfur oxides from even very dilute gas streams.
Furthermore, its use in the scrubber does not cool off the gas stream or satur-
ate it with water vapor.
These features make possible a process which can effectively treat large
volumes of dilute flue gas, while requiring regeneration of relatively small
volumes of material, and recovering the sulfur in elemental form.
The process flow diagram is shown in Figure 1. The process works as
follows:
1) The gas to be treated is removed from the boiler at about 850 F and
passed through a high temperature, high efficiency electrostatic precipitator
where 99% or more of the fly ash is removed. The gas then passes through
the scrubber, where the sulfur oxides are removed by contacting the gas
stream with a spray of molten carbonate at 850 F. This gas-liquid contact
removes 95% or more of the sulfur oxides and most of the remaining fly ash
from the gas stream. The cleaned gases are then returned to the boiler for
further heat recovery, and eventually pass out the stack.
2) The molten salt stream containing carbonate, sulfite, sulfate and fly
ash from the scrubber is filtered to remove, the fly ash. The fly ash filter
cake is subsequently treated to recover the lithium carbonate it contains.
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3) The filtered melt is fed into the reducer, and reacted with carbon.
The melt temperature is raised from ~ 850 F to ~ 1500 F by heat from the
indirect combustion of part of the carbon with air, and the sulfite and sulfate in
the melt are reduced to sulfide in about 20 minutes.
4) The melt stream from the reducer is passed into a quench tank where
its temperature is lowered from 1500 F to about 950 F by mixing with melt at
850°F. The effluent from the quench tank is pumped through a filter to remove
residual coke from the reduction step, cooled further to 850 F in a heat exchanger
and then passed on to the regenerator.
5) In the regenerator, the reduced melt is reacted with carbon dioxide
(produced in the reduction step) and steam in a multi-stage, countercurrent
tower. The sulfide in the melt is regenerated to carbonate, and the
sulfur is released as hydrogen sulfide. The hydrogen sulfide is sent to a Glaus
plant, where it is converted to elemental sulfur.
6) The regenerated melt is recirculated to the scrubber, with the small
losses in the purification system being made up by addition of fresh carbonate.
The status of development of the process is described briefly in the follow-
ing section.
B. STATUS OF DEVELOPMENT OF THE PROCESS
The process development work done to date has consisted primarily of
chemistry studies, bench-scale engineering tests, materials tests, test loop
design, and process analysis. The chemistry of each step of the process has
been studied separately, and kinetic and equilibrium data have been obtained.
Bench-scale engineering tests of the scrubbing step (including mist elimination)
and filtration step have also been carried out, and extensive corrosion testing
has been done. The status of each process step and of the materials testing
program are summarized below.
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-5-
1. Scrubbing
In the scrubber, it is necessary to bring about good contact between the
huge volumes of flue gas and the relatively small melt stream. Further, power
plant integration requirements make it important to impose as small a pressure
drop as possible on the gas stream. Because of this, a spray contactor has been
selected as the most promising scrubber concept. The spray contactor requires
spray nozzles, to break the melt up into small droplets for good gas-liquid con-
tact, and a very efficient mist eliminator, to prevent the gas stream from carry-
ing melt mist out of the scrubber.
In bench-scale tests, synthetic flue gas was scrubbed in a small
(4-in. ID) spray chamber equipped with a single pneumatic nozzle, and removal
efficiencies of greater than 95% were obtained at gas velocities of up to
25 ft/sec. Mist elimination tests were done with the same apparatus; they
showed that wire mesh pads were quite effective in removing melt droplets from
the gas. However, the tests were hampered by the lack of a source of large
volumes of hot gas. This made it necessary to use such a small spray contactor,
and also limited the mist eliminator tests at high velocities to very short runs.
Consequently, it was difficult to separate out the wall effects in both the scrub-
bing and mist elimination tests, and it was impossible to obtain enough mist
carryover to get accurate mist eliminator performance data. Also, the use of
synthetic gas precluded any study of effects of fly ash or other fuel-derived
impurities. Further testing of spray contactors, mist elimination and fly ash
effects requires a side stream of flue gas from an operating power plant. There-
fore, tests of the scrubber and mist eliminator must be a part of a pilot plant
program.
2. Reduction
The melt produced in the scrubber is a mixture of alkali metal carbon-
ate, sulfite and sulfate. However, when this mixture is heated to the reduction
temperature, the sulfite rapidly disproportionates to form sulfate and sulfide.
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-6-
Consequently, the reduction step actually involves the reduction of sulfate dis-
solved in a melt of carbonate and sulfide. To accomplish this effectively with
carbon, the temperature must be raised from 850 F to about 1500 F and the
endothermic heat of reaction (~ 40 kcal/g mole of sulfate) must be supplied.
Because of materials limitations, the most feasible way to supply the heat of reac-
tion is to generate the heat internally, by the indirect combustion of carbon with
air. Generating the heat directly within the melt inside the reduction vessel
eliminates the necessity for heat transfer surfaces operating at high temperatures
in a corrosive environment.
The chemistry of the reduction step with simultaneous heat generation
is quite complex. Both oxidation and reduction must occur in the same vessel,
under conditions which yield the desired sulfide product and supply the sensible
heat required to raise the incoming melt to the reduction temperature, the endo-
thermic heat of reaction, and the heat losses from the reducer vessel. Tests
done so far have demonstrated the feasibility of performing the reduction step
with carbon (petroleum coke) while simultaneously generating the required heat
internally by air oxidation. The reduction is rapid and is complete in 15 to 20
o
minutes at temperatures of 1450 to 1500 F, and yields a product containing
essentially no sulfate as long as excess coke is present.
Analytical studies have indicated that the concept of a two-zone reduc-
tion vessel (oxidation and reductions zones separated by a ported barrier) with
internal circulation shows the most promise. These same studies indicated that
an alumina-brick-lined vessel is needed to minimize heat losses. However, the
vessel configuration, dimensions, and locations of the inlet and exit ports as
well as the methods to be used for introducing the air and carbon into the reduc-
tion vessel have not been fixed. Further work, including a laboratory program
to model the reducer hydraulic processes, is needed to supply data on which to
base the pilot plant reducer design.
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-7-
3. Regeneration
In the regeneration step, the carbonate-sulfide melt from the reducer
is reacted with a mixture of carbon dioxide and steam, converting the sulfide
into carbonate plus gaseous hydrogen sulfide. The chemistry of this step has
been studied extensively; the reaction is rapid, exothermic, and equilibrium-
controlled. For complete regeneration, a multi-stage countercurrent con-
tactor such as a tray column is required, with cooling provided to remove the
heat of reaction.
The equilibrium and heat evolution data were measured experimentally,
and used to calculate the number of theoretical stages required and the
amount of heat evolved. To complete the design of a regenerator column, tray
efficiency and heat transfer coefficient values are needed. However, it will be
difficult to measure these parameters in the laboratory, and the values can be
estimated well enough to permit the design of a pilot plant regenerator. The
tray efficiency and heat transfer coefficient values for use in designing a full-
scale unit can then be obtained accurately from the pilot plant tests.
4. Filtration
In laboratory fly ash filtration tests, cartridge filters with sintered
metal, etched disc, and wire wound filter elements have all been tested suc-
cessfully. The filters remove essentially all the ash from the melt, forming a
filter cake which contains about 60% melt and 40% fly ash. Economic considera-
tions make it necessary to remove the filter cake from the filter continuously
and as "dry" as possible, for reprocessing to recover the lithium carbonate.
The most effective way to recover a dry filter cake is to use a centrifugal
basket filter, and remove the cake continuously with a plow or scraper. It is
necessary to build up a filter cake at least 1/4 in. thick for this to work. How-
ever, fly ash forms a dense filter cake, so that a thick cake layer causes a. high
pressure drop. Improving the porosity of the cake by use of filter aid or filter
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-8-
precoat material, such as fluidized petroleum coke, would ease the filtration
problem considerably. A program to study filter aids and precoat materials
should be carried out in conjunction with the pilot plant program.
5. Lithium Recovery System
The lithium recovery system must recover a large fraction of the lithiun
carbonate from the fly ash filter cake and return it to the process. An aqueous
method has been developed and tested to separate the insoluble lithium carbonate
from the soluble material (chlorides and the relatively inexpensive sodium and
potassium carbonates). .Laboratory tests have demonstrated that this method
can recover most of the lithium carbonate from the filter cake. A program to
complete the development of this process should be undertaken in conjunction
with the pilot plant program. However, the work need not start until the pilot
plant is in operation and producing filter cake.
6. Materials and Components
A test program to select materials of construction which resist cor-
rosion by the process melts has been underway for over three years. At first,
the common metals, alloys, and ceramics were given screening tests. Success-
ful candidates were then subjected to long term tests, including one-year tests
in rotating capsules. As a result, it was found that type 347 stainless steel was
the best of the standard alloys for service below 1000 F.
After the preliminary selection of type 347 stainless steel, this alloy
was subjected to further tests to study the effect of stress in the presence of
chloride, oxygen, and water vapor, the effect of sensitization, and the rate at
which the alloy constituents are leached out and transported under the influence
of a temperature gradient. The results have been satisfactory, and type 347
stainless steel is presently the planned material of construction for all process
equipment operating below 1000 F. However, further testing should be done, to
study the suitability of this alloy when exposed to flowing melt under process
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-9-
conditions. For such tests a forced-circulation test loop is needed. This loop
will include a pump, valves, and flow meters in addition to corrosion test sections.
The need for such a test loop was recognized early in the development program,
and the design of a test loop was completed and the procurement package has been
prepared. In addition, Atomics International has already purchased the pump
and a NaK-filled differential pressure cell for a flowmeter, since these are long-
lead-time items. The test loop should be built and operated as soon as possible,
both to study corrosion under melt flow conditions, and to gain experience with
the pump, valves, and instrumentation.
C. DEVELOPMENT PROGRAM DURING REPORT PERIOD
The development program for this report period concentrated on (1) obtaining
additional data on the performance of filters to remove fly ash from the molten
salt stream, (2) studying the effectiveness of mist elimination pads to remove
entrained salt from the gas stream exiting from the scrubber, (3) continuing the
materials test program, (4) completing the design of a materials test loop,
(5) cooperating and assisting Singmaster & Breyer in their evaluation of the pro-
cess, (6) completing preliminary engineering analyses of the two region molten
carbonate reducer and the regenerator, and (7) revising the material and energy
balances for 800 and 10 Mw plants. The results of the work on each of these
topics are described in the following section of this report.
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-10-
VENT
VENT
<
LU
CC
LU
LU
O
I
<
O
O
LU
OC
Q.
XD
INLET SAMPLE
CC
LU
OUTLET RECEIVER
Figure 2. Filtration Test Stand
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-11-
JI. PROGRESS DURING REPORT PERIOD
A. FILTRATION STUDIES
1. Introduction
The flue gas which enters the scrubber contains any fly ash which passes
through the electrostatic precipitator, plus other impurities such as chlorides.
These impurities will be picked up by the molten carbonate and incorporated into
the process melt. To keep them from accumulating and building up to a high level
in the process stream, it is necessary that they be removed from the melt con-
tinuously. It has been demonstrated that fly ash can be removed by filtration, and
that chloride, after reaching its solubility limit (effectively about 4 wt %), forms
crystalline particles which should also be filterable.
2. Experimentation
As part of the continuing effort to develop an effective filter for the
removal of fly ash from molten salts, additional tests were conducted with a
Croll Reynolds 25 micron filter to study the effect of filter cake buildup on
removal efficiency and pressure drop across the filter. The apparatus that was
used for the tests is shown schematically in Figure 2. By applying pressure on
the melt tank, the melt-ash suspension was caused to flow from the tank through
the filter into containers. The flow rate was controlled by varying the pressure
on the melt tank. Filtration efficiency was determined by analyzing inlet and out-
let samples for ash content. t
During the test, the filter was operated with an average melt throughput
rate of ~ 150 g/min, until the pressure drop across the filter necessary to main-
tain the flow rate increased to 40 psi. A total of 26. 6 kg of melt containing 1 wt %
of Colorado Public Service fly ash was passed through the filter during the experi-
ment. This corresponds to an ash loading of 0.609 g/cm filter surface. The
data from this test were compared with data for the 25 micron Croll Reynolds
filter reported in Progress Report No. 3; the results are shown graphically in
Figure 3.
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1000
800
600
n.
ao
a-*
400
200
0
Low Loading
Equation.
0.2
High Loading
Equatiou
Q Previous Data
O This Test
0.4 0.6 0.8 1.0
o
Ash Loading, a (Ib/ft )
1.2
1.4
Figure 3. Experimental Measurement of Dependence of AP/0 on Ash Loading for
th 25-u Croll-Reyfiolds Filter
i
H->
IV
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13-
Two equations were developed to fit the data for the Croll-Reynolds filter over the
entire range of ash loadings:
AP
Low capacity range: —- = 16. 3 -f 2530 a 0 <. a £ 0.067
. . , ,. lb ash
where a = ash loading——
ft filter area
, AP _ pressure drop psi
0 unit flow rate ' , ., 2
gal. /ft mm
AP
High capacity range: — = 88 4- 610 a 0.067 £ a <; 1.24
The buildup in filter cake appeared to have little effect on filter efficiency.
This could possibly be explained by channeling. The average removal efficiency
during the experiment was 93%. Attempts to preserve the filter cake on the filter
while it was being removed from the filter housing were unsuccessful. This
supports earlier observations that the filter cake can be removed from the
filter body by simply draining the housing.
3. Evaluation
a. Fly Ash
The fly ash filter requirements were re-evaluated, based upon the
above pressure drop data. The following criteria were used in the evaluation:
(1) coal producing 10% fly ash, (2) all of the fly ash entering the absorber being
removed by the molten carbonate spray, (3) 1 hour filtration time between filter
cake removal and (4) a 50 psi pressure drop allowance across the filter. The
calculations were made for a 99.5% and a 99.75% efficient electrostatic
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-14-
precipitator and a 719 gpm and 180 gpm flow rate (the former reflecting a 3:1
recirculation rate around the scrubber and the latter reflecting a once-through
stream). One additional estimate was prepared for a 99.9% efficient electro-
static precipitator. Table I shows the results of the calculations and the budgetary
prices obtained from the Croll Reynolds Engineering Company. The prices were
based on Type 347 stainless steel and were for both 70 micron and 25 micron ele-
ments. The 25 micron element cost is 50% higher, but has been shown to be more
efficient in filtering the fly ash. The largest unit manufactured by Croll Reynolds
has 1200 ft of filter area. The quote was therefore based on multiple units where
required. In most cases 100% standby capacity is furnished to provide 60 minutes
of downtime to remove the filter cake and ready the unit for reuse. However, also
shown for comparison are two cases of lower extra capacity for less downtime.
The extra capactiy needs will depend on the cleaning time actually required and
the excess capacity desired for non-normal operation.
b> Effects of Chlorides
In the Singmaster & Breyer evaluation (see Section G), it was assumed
that a coal containing 0. 04% chloride is burned, and that all of this chloride vola-
tilizes, reacts with the carbonate melt in the scrubber, and then forms solid
potassium chloride. The potassium chloride is then filtered out of the melt along
with any fly ash which is picked up in the scrubber. These assumptions increase
the filter duty markedly, since the amount of potassium chloride formed is approxi-
mately equal to the amount of fly ash picked up when a 99% efficient electrostatic
precipitator is used. The actual fate of chloride in the process is not known.
Tests have shown that if the chloride content of the melt builds up to about 4%,
the melt is saturated and additional chloride forms a crystalline solid material!
which would be filtered out with the fly ash. However, potassium chloride will
also be sublimed out of the melt during the reduction process. This may remove
all or a large part of the chloride, so the assumption that all of it must be filtered
out is probably quite conservative. The inclusion of chlorides in the filter duty
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-15-
would change the filter requirement assumed in the process cost estimate given
in Section E of Progress Report No. 3, either by increasing the cost or by
decreasing the time of the filter cycle.
c.
Costs
Based on the latest flow diagram for an 800 Mw installation where
a 3:1 melt recirculation is used, the filter cost using 100% standby is $675, 000.
If only the stream going to the reducer is filtered or no recirculation is required,
this cost is reduced to $300, 000. However, if the fly ash filter must also remove
chlorides, the cost will increase substantially, to ~$700, 000 for the no recircu-
lation case.
TABLE I
FILTRATION PRICES USING CROLL REYNOLDS FILTER
(800 Mw Plant)
gpm
180
180
719
719
180
Fly Ash
Removal
Efficiency
99.5
99.75
99.5
99.75
99.9
Required
Filter Service
ft2
940
722
2310
1900
545
No. of
Filters
2
2
6
4
4
3
2
Model
S72-940
S60-722
S60-722
S66-1770
S72-940
S72-940
Total Filter
Surface
Provided
1880
1444
4332
3080
3760
2820
Cycle Time
(min)
60
60
56
40
60
45
60
Total Filter
Cost ($K)
70 M.
200
150
450
330
400
300
25 n
300
225
675
495
600
450
150*
* Estimated
d. Improved Filtration
On recommendation of Singmaster &. Breyer (the process evaluators)
we discussed ways to improve the filtration step with U. S. Filter Company. The
-------
-16-
experimental results showing a 40 psi pressure drop with a 1/16 in. thick filter
cake suggested that cake porosity be increased by use of a filter aid. The sug-
gested filter would have 200 micron openings and use a 1/16 in. to 1/8 in. thick
filter aid layer, operating on a one-hour cycle. Fluidized coke was suggested as
a suitable filter aid. The estimated cost of a 940 ft unit is $70, 000. This is in
line with the Croll Reynolds filter costs.
e. Conclusions
(1) Experiments should be conducted to determine the merits of
using fluidized coke as a filter aid.
(2) Experiments should be conducted to determine the effects of
chlorides on the performance requirements for the fly ash
filter.
(3) The effect of recirculating melt containing fly ash through the
scrubber should be determined.
(4) The basis for filter overdesign should be established (maximun
fly ash burst).
(5) The minimum downtime between cleaning and the penalty in
melt loss at every cleaning should be determined to find the
optimum filtering time.
(6) There is an incentive to develop a continuous filter.
B. MIST RECOVERY TESTS
In the scrubbing step of the process, the melt will be sprayed into the gas
stream being treated. It is very important that this melt be recovered, so that
it is not carried away entrained in the gas. A highly efficient mist eliminator is
I. :'
needed for this. Wire mesh pads have been used for mist recovery with aqueous
systems; their suitability for use with molten carbonate had to be studied. There
fore, a series of tests was conducted to investigate the effectiveness of woven
wire mesh pads as a means of separating entrained salt from the gas stream.
-------
-17-
Th e mist elimination test system that was utilized in these tests is shown
schematically in Figure 4. In operation, melt is pressurized from the melt tank
into a pneumatic nozzle where it is atomized by a jet of hot gas. The generated
mist is picked up by the main gas stream and carried upward through two wire-
mesh mist elimination pads. The gas then exits through a short constricting stack.
Mist which passes through the first pad is collected on the second pad. The quantity
of mist which passes through both pads will appear in an exit gas sample which is
collected continuously during a test. During the first two tests, it was found that the
quantity of material collected in the exit gas sample train was too small to be
analyzed and this technique was abandoned in favor of weighing the pads prior to and
following each test. The separation efficiency was then determined by changes in
weight and by making a material balance around the system.
As a result of operating problems (discussed below) attendant with the equip-
ment itself, only three tests were conducted. The conditions for the tests are
given in Table II.
TABLE II
MIST ELIMINATION TEST CONDITIONS
Test
No.
1
2
3
Melt Conditions
Compo sition
M CO
M CO
xVX C^O
Temp
900
1020
1100
Flow Rate
(Ib/hr)
20
35.5
36.3
Gas Conditions
Composition
Compressed
Air
Compressed
Air
Compressed
Air
Temp
950
930
850
Flow R.ate
(ACFH)
4720
4200
2410
Linear
Velocity
Through
Demister
(ft/ sec)
6.5
5.8
3.3
In the first test, a large nozzle was used, with the apparatus arranged as
shown in Figure 4. The test lasted for about 20 min, and then stopped when water
from the compressed air line got through the water trap and cooled the nozzle below
-------
TO GAS SAMPLER
MELT
TANK
3
I
1
[
3RAIN X/'
)RAIN \ *
MELT IN
/ '\
tilt
GAS
MESH
PAD
-f ' >
MESH
PAD
tt ft
GAS
ATOMIZING GAS
ir
GAS IN
ft
1
II
I)
1
1
1
1
1
1
1
u
HEATER
1
HEATER
1
t
1
FLOW
METER
t t
COMPRESSED
AIR
oo
i
MELT DRAIN
Figure 4. Mist Eliminator Test System
-------
-19-
the melt freezing point, plugging it. During the test, no melt drained from either
mist eliminator section, although faint traces of salt could be seen coming out with
gas escaping from the lower section drain. The linear gas velocity through the
mesh pads was about 6. 5 ft/sec. At this velocity the overall pressure drop through
the two mist eliminator sections was 1/8 in. of water at the start, rose to 3/8 in.
while melt was spraying, and then fell back to 1/8 in. after the nozzle plugged.
These values are quite low; the decrease after the nozzle plugged may indicate
some mist was picked up, but that it dripped back into the spray chamber instead
of draining into the circumferential weir.
During the test, the gas passing through the mist eliminator was continuously
sampled and scrubbed with water. The resulting solutions were evaporated and
analyzed for lithium, sodium and potassium using x-ray emission spectroscopy.
However, there was such a small amount of material collected in the upper pad
that the data were very inconsistent and scattered. Therefore, although the mist
collection efficiency could not be determined accurately, mist carry-over was
found to be very slight.
In the second test, the test system was modified to provide a separately-
heated source of dry nitrogen for the atomizing nozzle, a smaller nozzle was used,
and the upper pad was removed so that all of the mist passing through the first pad
would appear in the exit gas. The exit gas was sampled by sucking it through a
water bath at a controlled rate. During the test, melt did drain out of the lower
section drain, showing that the nozzle was producing mist and that the mesh pad
was removing it. However, analysis of the gas sample was inconclusive, due to
the small quantities of collected material.
In the third test a clean weighed demister pad was installed downstream of
the test demister. Based upon weight gain of the downstream demister during the
test and the total salt passed through the nozzle, the test demister was 99.92%
effective in preventing mist carry-over at a gas velocity of about 6 ft/sec. However,
there was difficulty in attaining the desired gas temperature (850 F) at the desired
gas velocities. Consequently, modifications were instituted to double the heating
capability of the gas system.
-------
-20-
The mist eliminator test stand was modified to double the gas heater
capacity, and tests were then attempted at high gas velocities (up to 25 ft/sec).
The tests failed, due to the presence of large amounts of water in the compressec
air supply. Vaporization of this water cooled the gas stream to below the melt
freezing point. It is apparent that a different gas supply will be needed if high
velocity tests are to be run for periods long enough to obtain accurate carryover
data. Hot combustion gases from a burner should be suitable. However, incor-
porating a burner system into the mist eliminator test stand was not possible
during the present report period, so the effort on mist elimination studies was
suspended. The work done to date indicates that two wire mesh pads in series
will remove all or nearly all of the fine mist produced by the aspirator nozzle
at velocities of up to 6 ft/sec and probably higher. However, caution should be
utilized in applying these results to a large scale plant due to inabilities to
separate out the wall effects. Also, the use of a synthetic gas precludes the
study of the effects of fly ash and/or other fuel derived impurities. Consequently,
pilot plant tests are really needed to determine whether the removal efficiencies
of large mesh pads will be high enough, under actual power plant conditions.
C. MATERIALS STUDIES
1. Introduction
A test program to find construction materials that are resistant to
corrosion by the process melts has been underway for over three years.
Initially, the more common metals, alloys, and ceramics were subjected to
corrosion screening tests at 932 F (500 C). The results of the screening tests
indicated that the austenitic stainless steels (preferably type 347) were suitable
for service at 900-950 F, with some other alloys (e. g. , Hastelloys G and
X and Haynes 25) also suitable as alternates. These materials were then tested
more extensively, in both static and dynamic (rotating capsule) tests. The
results of many of these tests have been presented in the first three progress
reports. During this report period, five one-year dynamic tests and three one-
year thermal transport tests were completed; they are discussed in section 2
-------
-21-
below. In addition, the design of a pumped corrosion test loop was completed.
This test loop is needed to study the corrosion effects of circulating melt
streams under process conditions. The design is presented and discussed in
section 3 below.
2. Corrosion Tests
During the report period, five dynamic corrosion tests of type 347
and type 321 stainless steels were performed in rotating test vessels with the
equipment and techniques described in Progress Report No. 1 (AI-68-104;
PB 179-908). The test specimens were stressed (clamped horseshoe) and
unstressed strips of type 347 and type 321 stainless steels exposed to carbonate
melts containing
(1) 5% chloride and 20% sulfide (both steels);
(2) 5% chloride and 20% sulfite (both steels); and
(3) 5% chloride and 40% sulfite (type 347 only).
All of the test capsules initially had atmospheres of nitrogen, oxygen, and water
vapor above the melt level; the tests were all conducted at 932 F (500 C) and
were intended to last for one year (8760 hours).
In addition to the five dynamic tests, three thermal transport tests
were conducted of both stressed and unstressed steel specimens and different
melt mixtures; in each of these tests the test capsules were mounted in furnaces
with one end of the vessel maintained'at 932°F (500°C) and the other at 797°F
(425 C). The furnaces (and vessels) were inverted every 3-1/4 hours, causing
the melt to flow from one temperature zone to the other. The purpose of these
"flip-flop" tests was to determine the extent to which thermal transport of alloy
constituents occurs. The tests involved both type 347 and 321 specimens in melts
of carbonate plus 10% chloride and 20% sulfide, and type 347 specimens in a melt
of carbonate plus 10% chloride and 20% sulfite.
-------
TABLE III
RESULTS OF ONE YEAR CORROSION TESTS
Test
No.
29
31
35
40
41
43
39
42
Metal
347
347
321
347
347
347
321
321
Test Type
Flip Flop
Flip Flop
Flip Flop
Rotating
Rotating
Rotating
Rotating
Rotating
Melt (wt %)
Cl"
5
5
5
5
5
5
5
5
so;
0
20
0
0
20
40
0
20
S=
20
0
20
20
0
0
20
0
Test
Temp
(°C)
14251
\500J
f25l
\500l
(425\
\500j
500
500
500C
500
500
Test
Duration
(hr)
8750
8850
8770
8860
8800
1580
8760
8810
Vessel Internal
Pressure (psig)
Initial
35
35
35
35
35
35
35
35
Final
-10b
19
-10b
-iob
-10
0
5
2
Internal Atmosphere (vol %)
Initial
N2
100
100
100
72
72
72
72
72
H2°
0
0
0
26
26
26
26
26
°2
0
0
0
z
2
Z
2
2
Final
N2
>99b
>99
>99b
93b
94
>99
>99
>99
°2
0
0
0
7
6
0
0.2
Ob
H2
0
0
0
0
0
0
T
0
H2°
0
0
0
Yes6
0
0
Yese
0
C°2
0
0
0
0
0
T
0
0
Corrosion
Rate
(mils/yr)
(0. 15
10. 17
/0.3
l4.5d
(0. 14
10.65
1
3.3
>29f
1
>10
Stress
Relief
(%)
43
61
54
46
88
46
50
60
58
67
g
Effect
of
Stress
None
None
None
None
None
None
None
None
None
None
Unknown
Remarks
Excellent
Excellent
SR
Heavy corrosion at
interface0' d
Excellent
Rusted-Cr depleted
SR
SR, E
Pitting factor 4f
SR
Corroded into at
interface
Based 'on weight loss over entire specimen - localized corrosion may be greater
Sample lines probably plugged with salt therefore this is probably pressure and analysis in lines and not in main vessel
°Thermal excursion to >650°C for 36 hr about 1/3 through test
Corrosion not uniform but greater at interfacial area
eWater determined qualitatively
Controller malfunction - wide thermal oscillation for several days
®Good spring back on part of stress samples left
T = Trace, SR - surface roughening or micropits, E = embittlement of wire used to tie specimens onto holders
-------
-23-
All of the tests were concluded during the report period. Seven of
the eight had attained the desired 1-year lifetime, while one of the dynamic tests
was terminated after 1580 hours due to a controller malfunction which allowed
the temperature to exceed 1200 F (650 C). The test conditions and results are
summarized in Table III, and discussed below.
a. Gas Sample Analysis
The three flip-flop test vessels were originally filled to 35 psig
with dry nitrogen. At the end of the test, one vessel contained nitrogen at 19
psig; the gas sample lines to the other two vessels were plugged so that the
internal pressure could not be measured. The loss of nitrogen was probably
via a line leak, too small for melt to escape through.
The five rotating test vessels were originally filled to 35 psig
with a 72% nitrogen-26% steam-2% oxygen mixture. At the end of the tests,
one of the gas sample lines was plugged. The other four vessels had all lost
pressure, possibly due to leaks. However, there is also evidence that the
oxygen and part of the steam were consumed during the tests, probably by
reaction with the melt on the vessel walls. The only gas samples containing
significant amounts of oxygen (tests 40 and 41} had negative pressures, and the
oxygen probably came from a small amount of air leaking into the vessel during
cooldown. Thus, in all these dynamic tests oxygen and probably steam were
lost by reaction sometime during the test. These tests will be repeated later
in a corrosion test loop with a constant composition of gas.
b. Metallurgical Examination
1) Dynamic Tests
Specimens from the one-year dynamic corrosion tests
were mounted, vibrapolished and etched with 10% ammonium persulfate to deter-
mine if carbide precipitation took place. Both intergranular and intragranular
carbide precipitation were found in the type 347 stainless steel specimens kept
at 500 C (932 F) for one year; as expected, much less carbide precipitation was
-------
-24-
^2.0
fl
o
. 5
CJ
fl
o
Ul.O
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£0-5
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Manganese
-Q-
Niobium
Titanium
Molybdenum
K Film Base | Metal |
0
Iron
A Type 347 in 20% Sulfite- 5% Chloride
° Type 347 in 20% Sulfide-5% Chloride
O Type 321 in 20% Sulfide-5% Chloride
Chromium
Nickel
:£t
2345
Material Removed (mils)
Figure 5. Concentration of Alloy Constituents in Film and Base Metal
After 1 Year in a 75°C Melt Thermal Gradient (500°C to 425°C)
-------
-25-
observed in the specimens kept at 425 C (797 F) for one year. Maximum sensi-
tization due to chromium carbide precipitation in the grain boundaries always
occurs in Type 347 stainless after about one year at 500 C. Enhanced intergranu-
lar corrosion,due to chromium depletion along the grain boundary when chromium
carbide precipitation occurs, would be a serious problem. However, in the three
samples examined little (if any) enhanced intergranular corrosion occurred. After
one year at 500 C what appears to be intergranular corrosion was found to a depth
of 1/2 to 1 mil; but this is the same depth of penetration which was found after
1500 hr at 500 C. Such intergranular corrosion thus does not appear to increase
with time or sensitization, and therefore will probably not be significant in plant
applications.
2} Flip-Flop Tests
In these tests metal specimens were exposed to carbonate
melt containing 5% chloride and 20% sulfite for SS 347 and 5% chloride and 20% sul-
fide for SS 347 and SS 321. The melt flowed from a 500°C hot zone to a 425°C cold
zone and vice versa at 200 min intervals. The purpose of these tests was to
determine if alloying constituents of the metal (particularly chromium) would be
transported from the hot zone to the cold zone. If this were to occur, enhanced
corrosion with time could be expected and extended use of such an alloy would be
questionable.
The test samplers were analyzed metallographically to see
if thermal transport occurred. Mounted specimens from the hot zone of the
three flip-flop tests were repeatedly examined by x-ray fluorescence as the sur-
faces of the samples were ground away. In this manner the metallic composition
of the specimen v£ depth was determined. The results are shown in Figure 5.
In the film on the metal surface, the chromium content was higher than in the
base metal while iron was lower; this reflects the formation of the protective
lithium chromite film. However, at one-half mil below the film-metal interface
(i.e., into the metal), the concentrations of all major (Fe, Cr, and Ni) and
minor (Mn and Nb for SS 347 and Mn and Ti for SS 321) constituents
-------
-26-
were found to be the same as those for the bulk metal, in both alloys. Since
little, if any, thermal transport of alloying constituents occurred, these austenitic
stainless steels apparently are not susceptible to thermal transport of alloying
components in chloride-sulfite (SS 347) or chloride-sulfide (SS 347 and SS 321)
melts where 75 C (i. e. , 500 C to 425 C) thermal gradients occur.
c. Corrosion Rates
Based on weight loss, the corrosion rates of type 347 stainless
steel in the 5% chloride-20% sulfide melts were 0. 15 mils/year at 425°C (797 F)
and 0. 17 to 1. 0 mils/year at 500°C (932°F); the rates in 5% chloride-20% sulfite
melt were 0. 3 mils/year at 425°C and 3. 3 to 4. 5 mils /year at 500 C. For type
321 stainless steel, the rates in the 5% chloride-20% sulfide melts were 0. 14
mils/year at 425°C and 0. 65 to 1. 0 mils/year at 500°C; in the 5% chloride-20%
sulfite melt, the rate was greater than 10 mils/year. These results show that
type 347 is more corrosion resistant than type 321 and that sulfite is more cor-
rosive than sulfide.
The test of type 347 stainless steel in a 5% chloride-40%
sulfite melt at 500 C (test 43) was terminated after 1580 hours due to a con-
troller malfunction which allowed the temperature to rise to •-'675 C for 36 hours.
The specimens from this test showed corrosion rates of greater than 29 mils/
year, with pitting also evident. The high corrosion rate and pitting were probably
caused by the overheating and thermal cycling. Further testing is needed to
verify this.
As a result of the tests, type 347 stainless steel appears to
remain acceptable as a construction material. However, care will have to be
used to prevent overheating; pitting was found when thermal excursions occurred
(tests 31 and 43). Further testing is needed to determine the maximum temper-
ature at which type 347 can be used, and also to study the effects of thermal
cycling.
-------
-27-
d) Stress Effects
No effects of stress were found. It appears as if sodium
carbonate (a known inhibitor of aqueous chloride stress cracking) prevents
stress cracks from developing. Corrosion is evidently due predominantly to
the sulfur compounds in the melt.
3. Materials Test Loop
As was pointed out above, further testing should be conducted to
verify the suitability of the chosen materials of construction when exposed to
flowing melt under process conditions. For such tests, a forced circulation
test loop is required. The need for such a loop was recognized early in the
development program and the design of the test loop using an AEG pump (from
Oak Ridge National Laboratory) with separate test legs of Type 347 and 321
stainless steel was completed and a procurement package prepared. The draw-
ings are presented as follows:
P & I Diagram Figure 6
Layout Drawing Figure 7
Immersion Tank Figure 8
Melt and Drain Tank Figure 9
A revised version of the loop, using a commercial pump and providing only a
I
single type 347 stainless steel leg, has also been designed. The revised P&I
diagram is shown as Figure 10. The procurement package for this loop was also
prepared.
D. REDUCER ENGINEERING ANALYSIS
A key component of the Molten Carbonate Process is the reducer, which
provides for the reduction of the oxidized sulfur compounds formed in the
scrubber. The melt exiting from the scrubber is a mixture of alkali metal
carbonate, sulfite and sulfate. When this mixture is heated to the reduction
temperature, the sulfite rapidly disproportionates to form sulfate and sulfide.
-------
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-------
-33-
Consequently, the reduction step involves the reduction of alkali metal sulfates,
dissolved in a melt of carbonate and sulfide. The sulfide reduction product is
then subsequently regenerated into carbonate with the release of hydrogen
sulfide. The reduction is carried out at a temperature of about 1500 F, using
carbon in the form of petroleum coke as the reducing agent,and providing the
heat required for the reduction internally through the indirect oxidation of excess
coke with air.
Analytical studies have indicated that the concept of a two-zone reduction
vessel (oxidation and reduction zones separated by a ported barrier) with internal
circulation shows the most promise. In this concept, shown schematically in
Figure 11, the reducer vessel will be divided into two zones separated by a
ceramic baffle. The melt and coke will enter at the top of the oxidation zone or
at the top of the reduction zone. Air will be blown into the bottom of the oxidation
zone. The air will oxidize the sulfide in the melt, generating heat, and the un-
reacted nitrogen (plus some carbon dioxide from reduction) will bubble up through
the melt, causing a convective circulation. The circulation will carry hot melt
over the baffle into the reduction zone, where most of the reduction will take
place, absorbing heat and releasing carbon dioxide. Part of the reduced melt will
be removed at the bottom of the reduction zone, and the rest will be recirculated
into the oxidation zone to be reoxidized, supplying more heat. An extension of
the baffle through the gas space to the top of the vessel will separate the two gas
streams, making it possible to use the relatively nitrogen-free carbon dioxide
from the reduction zone in the regenerator.
An engineering analysis of this reducer concept was performed and is
presented in detail in Appendix A. A summary and the main conclusions of this
analytical study are presented below.
-------
- 34-
N2,CO2 CO2
OUT OUT
ALUMINA-LINED
WALLS
AIR IN
MELT IN
COKE IN
REDUCED MELT
OUT
Figure 11. Two-Zone Reducer Vessel Concept
-------
-35-
1. The heat loss from the melt bed must be minimized in order to
minimize the coke and air requirements of the reducer and its physical dimensions.
2. The coke utilization in the reducer must be maximized in order to
minimize the melt recovery and make-up requirement resulting from the melt
losses associated with filtration of the unreacted coke.
3. The use of a frozen melt skull (cold wall) reducer vessel does not
appear as economical as the use of an alumina-lined, internally-insulated vessel,
because of the inherently high heat losses associated with maintaining the frozen
skull.
4. The concentration of sulfur compounds in the carbonate melt must be
as high as feasible, subject to solubility limitations, to minimize melt flow rate
and preheat requirements.
5. Increasing the rate of the reduction reaction will result in a more
significant reducer design improvement than increasing the superficial air velocity
in the oxidation zone.
6. The internal melt recirculation between the oxidation and reduction
zones of the reducer can be controlled through proper sizing of the orifices in the
baffle between the two zones.
7. For the case of the alumina-lined vessel, accurate knowledge of
the melt wall interfacial heat transfer coefficient is not essential for design of
i
the pilot plant reducer.
Typical dimensions were determined for a vertically oriented two-zone
alumina-lined, internally-insulated reducer as a function, of processing capacity
and heat loss. The controlling contraints in the design were found to be the
maximum allowable superficial air velocity in the oxidation zone and the reaction
rate (hence the required residence time of the melt) in the reduction zone.
-------
-36-
Table IV gives the dimensions, and coke and air requirements for a 5 Mw pilot
plant reducer and a 267 Mw reducer (one of the units proposed for an 800 Mw
power plant). The calculations assumed a residence time of 15 min in the reduc-
tion zone, a superficial gas velocity of 3 ft/sec in the oxidation zone, 30 % mole
fraction of sulfur compounds in the melt, and negligible heat losses.
TABLE IV
MOLTEN CARBONATE REDUCER CHARACTERISTICS
Plant
Size
(Mw)
5
267
Reducer
ID
(ft)
3.7
19.6
Expanded
Bed Height
(ft)
3
8
Coke
R equi r ement s
(Ib/hr)
130
7000
Air
Requirements
(SCFM)
120
6500
E. REGENERATOR ENGINEERING ANALYSIS
In the regeneration step, the carbonate-sulfide melt from the reducer is
reacted with a carbon dioxide-steam mixture, which converts the sulfide into car-
bonate plus gaseous hydrogen sulfide. An analytical study of the regeneration step
was conducted to determine the hydrogen sulfide concentrations achievable in the
product gas and to establish the operating conditions required to achieve these con-
centrations. Specific emphasis was given to the constraints imposed by the thermc
dynamic equilibrium and heat generation of the reaction, by sulfide solubility con-
siderations and by the composition of the available regeneration feed gas. The
results of this study are presented in Appendix B. The conclusions are summar-
ized below:
1. The operating parameters of the regenerator must be optimized to
satisfy the following constraints:
a. A thermodynamic equilibrium strongly favored by low temperatures
b. A highly exothermic heat of reaction
c. A solubility of sulfide in the carbonate melt which increases sig-
nificantly with increasing temperature.
-------
-37-
2. The regeneration equilibrium is strongly favored by a high steam con-
centration in the feed gas, and to a lesser extent by increased total pressure
and carbon dioxide concentration.
3. The maximum hydrogen sulfide concentration attainable in the regenera-
tion off-gas is practically independent of the sulfide concentration in the feed melt
stream from the reducer.
4. As a result of potential solubility limitations, the operating temperature
o£ the regenerator may be determined by the sulfide concentration in the melt feed
stream from the reducer.
5. Approximately nine theoretical plates are required to regenerate 95% of
the combined sulfide (15 mol %) at the temperature range of 950 to 1000 F with a
feed gas containing 20 mol % carbon dioxide and 40 mol % steam. The correspond-
ing hydrogen sulfide concentration is approximately 14 mol %.
6. The required number of theoretical plates, and the maximum attainable
hydrogen sulfide concentration in the off-gas, are controlled by a pinch point
between the operating line and the equilibrium line at the maximum allowable
operating temperature of the regeneration step. Intercooling reduces the pinch
point limitation and allows the reaction to proceed. A single cooling stage, cool-
ing the melt taken off a plate slightly above the middle of the regenerator from
1000 down to 900 F, appears to be adequate.
F. HEAT AND MASS BALANCES
Based upon a revised flow diagram, preliminary heat and mass balances were
derived for the Molten Carbonate Process installed on an 800 Mw power plant. The
results are given in Figures 12 and 13. (The data for a 10 Mw installation can be
obtained by scaling down the data for the 800 Mw installation.)
The following bases were used in these calculations.*
*This heat and mass balance work was done prior to the detailed reducer analysis
of Appendix A. Although the two studies are essentially in agreement, there are
some minor inconsistencies.
-------
-38-
FLUE GAS
RETURN
TO BOILER
CARBONATE
MAKEUP
SCRUBBER
850°F
©
REGENERATOR
850°F
PRECIPITATOR
850°F
FLUE GAS
FROM BOILER
71-F4-7-16A
HEAT
EXCHANGER
COKE
FILTER
CAKE
ASH FILTER
850°F
1
CLAUS
PLANT
SULFUR
CARBONATE
ANDSULFIDE
CARBONATE,
SULFITE
AND SULFATE
CAKE
©
FLY ASH
REDUCER
1500°F
I
el
H20
CO-
PETROLEUM
COKE
AIR
Component
N2
C02
H2°
°Z
^2
»°3
H2S
Ash, Ib
1
177,000
32, 800
19, 700
7,900
450
SO
0
45, 000
2
225
3
177, 000
33,300
19, 700
7,800
22.5
2.5
0
0
8
2592
69
688
12
2592
1582
69
2
13
1709
1042
46
1.3
14
517
15
883
540
540
0.7
16
883
27
27
512
Component
Ash, Ib
Carbon, Ib
Sulfur, Ib
Melt, Ib
5
225
450
9
149
19, 100
1,280
10
149
566
715
17
1200
Component
M2C03
M2S03
"2^4
M2S
4
4525
1316
6
3394
888
660
25
7
1132
296
220
8
11
1094
22
540
18
5000
888
682
52
Figure 12. Molten Carbonate Process Mass Balances
Hourly Basis, 800-Mwe Plant
Units: Ib-moles Unless Noted
-------
(875 if Q = 0)
FLUE GAS
RETURN
TO BOILER
A
QI = 37
SCRUBBER
850°F
PRECIPITATOR
850° F
t
FLUE GAS
FROM BOILER
SULFUR
CARBONATE
ANDSULFIDE
CARBONATE.
SULFITE
ANDSULFATE
H20
FLY ASH
= 14
PETROLEUM
COKE
AIR
I
OJ
NO
71-F4-7-16A
Figure 13. Molten Carbonate Process Heat Balances
800-Mwe Plant
Units: 106 Btu/hour
-------
-40-
1. Mass Balance
The mass balance was based on a sulfur input into the absorber of
500 Ib moles/hr for the 800 MW plant, exclusive of the sulfur introduced by the
fluid coke (6. 25 Ib moles/hr for the 10 MW plant). This would be the case for
a power plant operating under the following assumed conditions:
1) Coal characteristics
Analysis: C = 70.0% Heating value: 12,800Btu/lb
H = 4.8%
O = 5.8%
S = 3. 0%
N = 1.4%
Ash = 10. 0%
Moisture = 5. 0%
100. 0%
2) Plant heat rate = 9000 Btu/kwh ( ^38% plant efficiency)
3) Fraction of coal sulfur appearing as SO in flue gas = 95%
.nip
4) Combustion carried out with 20% excess air (containing 0. 013 Ib
H O/lb dry air, which corresponds to 60% relative humidity at 80 F). This
assumption, together with the previous ones, yields a flue gas containing 0. 21%
SO . For the calculation it was assumed that 10% of the SO was SO and the
x x 3
remainder SO . The 10% SO value is high, since most reported values are less
C* J
than 5%. It was, however, selected as a conservative upper bound value intended
also to take into account the possible oxidation of a small fraction of the recycled
sulfite in the scrubber.
5) Fly ash content of flue gas = 80% of ash originally in the coal.
It was assumed that 99- 5% of this fly ash carried by the flue gas is removed in a
high temperature electrostatic precipitator prior to entering the absorber.
6) Fluid Coke Composition
C = 90%
S = 6%
Ash = 0.7%
-------
-41-
7) Stoichiometry
a. Scrubber inlet M^CO^ = IQ (moie ratio)
Scrubber inlet SOX
Even with this rather high ratio, the resulting liquid-to-gas
ratio in the absorber is only 0. OZ78 on a mole basis and 0. 102 on a weight basis.
b. Carbon feed (in fluid coke) , 0
• ' — J.. O
Carbon required for M^SO^ reduction
This ratio was determined on the basis of the heat balance
around the reducer, taking into account the heat of combustion needed to provide
for the requirements of the endothermic reduction reaction, and the preheat of the
melt, the coke and the air to 1500 F. It allows for only relatively small heat
losses from the reducer.
8) Extent of Completion of Reactions
95% completion was assumed for the absorption, reduction,
regeneration, and sulfur-from-coke recovery reactions.
It was also assumed that any MS recycled to the absorber would
Lt
be completely oxidized to M SO . Recycled M SO was assumed to go through
the absorber without oxidation (assumption 4 was intended to take into account
any small amounts of M SO which might be oxidized).
10% disproportionation of the M SO was assumed to take place
£* $
downstream of the absorber prior to introduction into the reducer.
9) Minimum carbonate concentration in melt = 66 mole % which
occurs just downstream of the reducer.
10) Filter losses and make-up
The weight of melt removed from the system in the fly ash filter
downstream of the absorber amounts to twice the weight of the fly ash removed
by the filter.
The weight of melt removed from the system in the coke filter
downstream of the reducer is equal to the weight of uriburnt carbon plus ash.
The amount of unburnt carbon from the reducer was assumed to be that left over
-------
-42-
from the stoichiometric combustion with the air supplied to the reducer and the
95% efficient reduction of the M SO and the sulfur from the coke (this amounted
u TC
to about 3% unburnt carbon).
The overall material balance made no attempt to correct the
system flow rates shown for filter losses, except for indicating an equivalent
salt make-up stream upstream, of the absorber.
2. Heat Balance
a. Heats of Reaction
The heat of reaction data were obtained from available heats of
formation. While they are subject to appreciable uncertainties, they form an
internally consistent set of values. All the heats were taken at 850 F, except for
that of the M SO reduction which was taken at 1500 F.
2 4
The reaction considered for each component, and the heat of reaction
values used are presented in the following paragraphs. The percentage contri-
bution of each reaction to the total heat release of the component is also shown.
(1) Absorber
S02(g) + M2C03 (.£)— >M2S03tM2S04U)
-233. 8 Kcal/gm mole = -420, 000 Btu/lb mole 59. 3%
(2) Initial Disproportionation Downstream of Absorber
-11. 9 Kcal/gm mole = -21,400 Btu/lb mole
-------
-43-
(3) Reducer
4 -2"^' '' 4M2S°4
11. 9 Kcal/gm mole = -21, 400 Btu/lb mole 7.1%
+ 2C(s) — *M2S(.£) + 2C02(g)
+41. 9 Kcal/gm mole = +75, 500 Btu/lb mole -35. 5%
S(s)+ 0
-64. 5 Kcal/gm mole = -1 16, 100 Btu/lb mole 4.9%
S(s) + 0
-74. 1 Kcal/gm mole = -133,400 Btu/lb mole 0. 3%
C(s)+02(g)-*C02(g)
-94. 2 Kcal/gm mole = -169,500 Btu/lb mole 123. 2%
(4) Regenerator
M2S(J/) + C02(g) + H20(g) — >H2S(g) + M2C02(j2)
-25. 8 Kcal/gm mole = -46,400 Btu/lb mole
b. Sensible Heat Requirements of Reducer
The sensible heat requirements of the reducer were assumed
to be those necessary to preheat the melt from 850 to 1500 F, the coke from 600
to 1500°F and the air from 600 to 1500 F {it was assumed that both coke and air
would be preheated to 600 F outside of the reducer).
The sensible heat requirements were
Melt 43.4x650 = 28, 200 Btu/lb mole 52.4%
Coke 4. 8 x 900 = 4, 320 Btu/lb mole (C-I-S) 7. 9%
Air 7.0x900 = 6, 300 Btu/lb mole 23. 7%
84. 0%
These percentages represent fractions of the net heat generation
by chemical reaction in the reducer. They show that with a carbon stoichionietry
ratio of 1. 8 approximately 16% of the total net heat release is left over for heat
losses.
-------
-44-
G- COOPERATION WITH PROCESS EVALUATQR
1. Introduction
During the report period, the Molten Carbonate Process was
evaluated by Singmaster and Breyer, under NAPCA Contract No. CPA 70-76.
The purpose of this evaluation was to review and evaluate the Molten Carbonate
Process for removal of SO from coal-burning power plant stack gases and
from copper smelter gaseous effluents, and then to advise NAPCA on the con-
tinued course of action regarding the process. During the report period,
Atomics International worked closely with Singmaster and Breyer, supplying
assistance and data where requested. Information transmitted included process
flow diagrams, material and energy balances, equipment configurations and
selection, fly ash and chloride filtration data, and process economics.
2. Conclusions and Recommendations
Singmaster and Breyer concluded that the Molten Carbonate Process
is technically feasible, and that it should be tested in a pilot plant program.
Their report includes the following statement:
"The process appears to be feasible; however certain problem
areas have been uncovered which should be defined in a thoughtful
pilot program where corrective measures can be tested. A com-
plete technical and economic evaluation of the process can only
be performed after the pilot program is completed.
"It is therefore recommended that development and evaluation
of the process be continued. We believe that a pilot plant is the
next logical step in this development. Operation of a materials
test loop can either precede the pilot plant or be carried out as a
study parallel to the pilot plant program, depending on .time and
funds available to the program. "
Singmaster and Breyer made a very thorough evaluation of the
process, pointing out many potential problem areas and making valuable sug-
gestions about their investigation. Nearly all of these investigations will
require the pilot plant for their implementation. However, Singmaster and
Breyer also made several recommendations which can be carried out separately
-------
-45-
starting before the pilot plant program. As quoted above, the materials test
loop should be built and operated. Also, Singmaster & Breyer recommended
that additional work be done on the two-zone reducer concept, to insure its
operability, study its mechanical and hydraulic design, and determine the best
way to introduce coke into the reactor. Much of this work can be done in a
laboratory program, which can be completed in time to supply design data for
the pilot plant reducer. A third recommendation was that more filtration work
be done, directed toward increasing the filter cake thickness without excessive
pressure drop. This work also lends itself to a laboratory study, utilizing the
filtration test equipment available from the previous filtration studies.
3. Process Economics
In their study, Singmaster & Breyer made estimates of the capital
requirements and operating costs for the Molten Carbonate Process treating
flue gases from power plants and copper smelters. The base case selected was
the removal of 95% of the sulfur oxides from the flue gas emitted by an 800 Mw
power plant burning coal containing 3% sulfur, 0.04% chloride, and 10% ash,
and operating with a 70% plant factor. For this case, they estimated the capital
requirement to be $13, 448, 000 ($16. 81/kw) and the total annual operating cost
(including 14% capital charges) to be $4, 703, 000 (0.95 mills/kwh). These costs
did not include the cost of the Glaus sulfur plant, nor did they include any credit
for the by-product sulfur or for reduced costs of high-sulfur fuel.
i
It is difficult to compare the'above estimates with those for other
flue gas desulfurization processes, because most other processes have not had
the benefit of a thorough, independent evaluation. However, the costs are
believed to be lower than those for any other process which recovers elemental
sulfur. Also, the operating costs are competitive with the cost penalties for
low-sulfur or desulfurized fuels.
-------
-46-
III. FUTURE WORK
The process development work done to date has been successful, in that
it has defined a feasible process and developed each step of the process through
the laboratory bench-scale phase. The process as developed also appears
economically attractive, as was shown by the independent Singmaster & Breyer
evaluation. Thus continued development of the process is warranted.
The goal of the process development program must be to make available an
efficient, relatively inexpensive stack gas treatment system. For this system
to be accepted by the utility industry, it will have to be demonstrated on a full-
scale basis, with a unit capable of treating all the flue gas from a large
(800 Mw or more) power plant. To achieve this amount of scale-up will require
two steps: first a pilot plant, of about 10 Mw capacity, followed by a demon-
stration unit of about 200 Mw capacity. The 200-Mw demonstration unit,
when successfully developed, can then be used as a module in constructing
large units.
The key item in the development effort now is the pilot plant. The pilot
plant is needed to test the operability of the process cycle under realistic con-
ditions, to develop and test engineering solutions to process problems as they
arise, and to obtain the scale-up and cost data needed to develop the demonstra-
tion unit. The need for large amounts of real flue gas for accurate and reliable
scrubber and mist recovery tests makes it necessary to build the pilot plant
to treat a side stream of flue gas from an operating power plant. The use of
power plant flue gas is also required to study the effect of fuel-derived
impurities (ash, chlorides, and anything else) on the process. A program to
design, build, and operate a pilot plant at a power plant site should be initiated
as soon as possible, and carried out to completion.
The pilot plant program should also be supported by a parallel technology
development effort, to carry out those studies suggested by Singmaster & Breyer,
which can be performed in the laboratory. This effort should include
-------
-47-
(1) the construction and operation of the materials test loop, (2) studies to
improve the filtration characteristics of fly ash and other impurities, and
(3) analytical and experimental studies of the chemistry and hydraulics of the
two-zone reducer.
A pilot plant program and parallel technology development program will
be proposed to APCO in the near future.
-------
APPENDIX A
PRELIMINARY PROCESS ANALYSIS OF TWO REGION
MOLTEN CARBONATE REDUCER
-------
CONTENTS
Page
Nomenclature v
Summary A-l
I. Introduction A-4
II. Reducer Design Concept Description A-5
III. Basis of Study A-10
1. Composition and Physical Properties of Melt A-10
2. Composition and Heat Capacity of Coke A-11
3. Composition and Heat Capacity of Air A-12
4. Heats of Reaction A-12
5. Reaction Completion Efficiencies A-13
6. Reducer Design Parameters A-13
7. Reducer Design Concepts A-14
8. Reducer Processing Capacity A-14
IV. Reducer Mass Balance A-15
V. Reducer Heat Balance A-18
VI. Physical Dimensions of Reducer Melt Bed A-24
1. Cross-Sectional Area of Oxidation Region A-25
2. Internal Melt Recirculation Requirement A-26
3. Cross-Sectional Area of Reduction Region A-27
4. Melt Bed Diameter and Heat Transfer Area -
Vertical Cylinder Reducer Geometry A-29
5. Melt Bed Diameter and Depth Optimization -
Vertical Cylinder Reducer Geometry A-30
6. Horizontal Cylinder Reducer Geometry with
Separation Baffle Perpendicular to Cylinder Axis A-31
-------
CONTENTS (Continued)
Page
7. Horizontal Cylinder Reducer Geometry with
Separation Baffle Parallel to Cylinder Axis A-34
8. Results A-36
VII. Reducer Heat Loss A-50
1. Estimate of Melt Heat Transfer Coefficient A-50
2. Reducer Wall Heat Transfer in Frozen Melt Skull Concept. . . A-52
3. Reducer Wall Heat Transfer in Alumina Liner Concept A-55
VIII. Internal Recirculation Between Oxidation and
Reduction Regions A-59
1. Natural Convection Driving Force A-60
2. Pressure Drop A-6l
3. Orifice Size A-62
IX. Discussion and Conclusions A-65
1. Basis of Analysis A-65
a. Technical Approach A-65
b. Independent Variables A-67
c. System Constraints A-67
2. Limitations of Present Study A-70
3. Results A-71
a. Coke and Air Requirements A-71
b. Physical Dimensions A-73
c. Five and 267 Mw Reducer Units A-75
d. Melt Bed Heat Loss A-82
i
e. Internal Melt Recirculation A-83
4. Conclusions A-84
References A-86
11
-------
TABLES
Page
I. Molten Carbonate Reducer Coke and Air Requirements A-71
II. Molten Carbonate Reducer Heat Loss and Dimensions A-73
FIGURES
1. Schematic Diagram of Typical Molten Carbonate Reducer A-6
2. Molten Carbonate Reducer Mass Balance A-16
3. Molten Carbonate Reducer Heat Balance A-20
4. Reducer Coke Requirement A-22
5. Reducer Air Requirement A-23
6. Reducer Diameter as a Function of Processing Capacity A-38
7. Reducer Bed Depth as a Function of Processing Capacity A-39
8. Reducer Wall Heat Flux as a Function of Processing Capacity . . A-40
9. Normalized Heat Loss as a Function of Processing Capacity . . . A-41
10. Effect of Heat Flux on Reducer Diameter (9 = 15 min,
v = 3 ft/sec) A-42
11. Reducer Bed Heat Loss (9 = 15 min, v = 3 ft/sec) A-43
12. Effect of Heat Flux on Reducer Diameter (9 = 15 min,
v = 5 ft/sec) , A-44
13. Reducer Bed Heat Loss (Q = 15 min, v = 5 ft/sec) A-45
14. Effect of Heat Flux on Reducer Diameter (9 = 30 min,
v = 3 ft/sec) A-46
15. Reducer Bed Heat Loss (9 = 30 min, v = 3 ft/sec) A-47
16. Effect of Heat Flux on Reducer Length (Horizontal Cylinder
Geometry) A-48
17. Reducer Bed Heat Loss (Horizontal Cylinder Geometry) A-49
18. Effect of Parameters h, U, and r on Frozen Skull Thickness . . . A-54
19. Frozen Skull Reducer Concept - Heat Flux and Skull
Thickness A-56
iii
-------
FIGURES (Continued)
Page
20. Alumina Liner Reducer Concept - Heat Transmission
and Rejection Capability A-57
21. Reducer Baffle Orifice Width as a Function of Processing
Capacity A-64
22. Reducer Analysis Schematic Diagram A-66
23. Effect of Reduction Time and Air Velocity on Reducer
Diameter (5 Mw Reducer) A-76
24. Effect of Reduction Time and Air Velocity on Reducer
Bed Heat Loss (5 Mw Reducer) A-77
25. Effect of Reduction Time and Air Velocity on Reducer
Coke and Air Requirements (5 Mw Reducer) A-78
26. Effect of Reduction Time and Air Velocity on Reducer
Diameter (267 Mw Reducer) A-79
27. Effect of Reduction Time and Air Velocity on Reducer
Bed Heat Loss (267 Mw Reducer) A-80
28. Effect of Reduction Time and Air Velocity on Reducer
Coke and Air Requirements (267 Mw Reducer) A-81
IV
-------
NOMENCLATURE
A Constant used in heat balance. Defined in Figure 3
o
ACS Cross -sectional area of reducer, ft
A^t Heat transfer area of reducer walls in contact with melt
bed (including side walls and bottom of reducer vessel),
plus "effective" free surface of melt, ft2
AQ Normalized cross -sectional area of oxidation region of
reducer, ft2 per Ib mole/hr SO
,54.
A Cross -sectional area of orifice in baffle between
oxidation and reduction regions of reducer, ft
A Normalized cross -sectional area of reduction region
of reducer, ft2 per Ib mole/hr
A Normalized cross -sectional area of reduction region of
Gas Velocity reducer based on superficial gas velocity limitation,
ft2 per Ib mole/hr SOX
A Normalized cross -sectional area of reducer, ft per
Ib mole/hr SOX
a Mole fraction of sulfur compounds in form of sulfite in
melt leaving scrubber. Assumed equal to 0.67 for most
of the numerical calculations of this study
1 -a Mole fraction of sulfur compounds in form of sulfate in melt
leaving scrubber. Assumed equal to 0.33 for most of the
numerical calculations of this study. Also written
S04=
L J
and . , ,. or
[M2S03]
I
at various places in the report.
a* = 1
i ,
1 +
4e
B Constant used in heat balance. Defined in Figure 3.
AQd
—=jr— for vertical cylinder reducer geometry
V j.
AQD
for horizontal cylinder reducer geometry
Normalized carbon feed rate into reducer, Ib mole/hr
per Ib mole/hr SO
-------
c.
M
D
d
E
Normalized carbon utilization rate (carbon reacted) in
reducer = (1-X) C, Ib mole/hr per Ib mole/hr SOX
Orifice discharge coefficient. Assumed equal to 0.61
Specific heat, Btu/lb mole °F
Enthalpy increase of reducer air feed from 600 to 1500 °F
= 33,900 Btu/lb mole of contained oxygen
Enthalpy increase of reducer coke feed from 60 to 1500 °F
= 6,500 Btu/lb mole of contained carbon
Enthalpy increase of reducer melt feed from 850 to 1500 °F
= 28,300 Btu/lb mole of melt
Internal diameter of reducer, ft
Depth of expanded reducer melt bed, ft
Constant used in heat balance. Defined in Figure 3.
Reaction completion efficiency. Assumed .to be 0.95.
0.4685 <
o*
LSJ
0.4685 + 05484
[SJ
/QL\
1+ °4£° +1.00641-
when AT = 300 °F
g
Fraction of scrubber outlet melt stream fed into reduction-
regeneration system. 1 - f = fraction recycled directly back
to scrubber
Mole fraction 803 in SOX in feed gas to scrubber
H
S -
Also used in Section VIII for
the conversion factor 32. 2 ft/ sec
Melt film heat transfer coefficient at surface of reducer
vessel walls, Btu/ft2 hr °F
VI
-------
K
k
alumina
melt
skull
L
L
M
m
N
n
QL
QL
o
QT
(Q/A)
Distance from axis or midplane of reducer vessel to
separation baffle, ft
Constant used in heat balance. Defined in Figure 3
Thermal conductivity, Btu/ft hr °F
Thermal conductivity of alumina liner, Btu/ft hr °F
Thermal conductivity of melt, Btu/ft hr °F
Thermal conductivity of frozen melt skull, Btu/ft hr °F
Length of reducer (horizontal cylinder geometry), ft
Length of oxidation region of reducer (horizontal cylinder
geometry), ft
Length of reduction region of reducer (horizontal cylinder
geometry), ft
Symbol used to represent the alkali metals (Li, Na, K) in
the chemical formulae of the salts contained in the melt
Sulfur to carbon atom ratio in coke. Assumed to be 0.025
Processing capacity of molten carbonate system, Ib
moles/hr SO entering the scrubber in the flue or other
waste gases being treated
Hydrogen to carbon atom ratio in coke. Assumed to be
0. 24. For purposes of the heat balance, however, n is
considered equal to 0 if it is assumed that the hydrogen
in the coke does not contribute to the reduction reactions.
Normalized oxygen feed rate into reducer, Ib moles/hr
per Ib mole/hr SOX
Fraction of sulfite in melt leaving scrubber which dis-
proportionates to suifide and sulfate prior to entering
the reducer. Assumed to be 0. 10
Normalized heat loss from reducer melt bed, Btu/hr per
Ib mole/hr SC>
Ji
Normalized heat loss from oxidation region of reducer melt
bed, Btu/hr per Ib mole/hr SOX
Normalized heat transport from oxidation into reduction
region of reducer, Btu/hr per Ib mole/hr SOX
Heat flux at reducer walls in contact with melt,
Btu/ft2 hr
vu
-------
Proportionality factor = total AP/orifice AP, for melt
recirculation. Assumed equal to 1. 20.
Ls]
SCFM
T
T
melt
T
MP
T
heat sink
T
outer alumina
"skull
t
alumina
tskull
U
V
a
ft
TMP Theat sink
Frozen melt skull temperature parameter = •=;
melt MP
Mole fraction of sulfur compounds in melt leaving scrubber.
Mole fractions of 0. 15 and 0. 30 were considered in this
study.
Standard cubic feet per minute taken at 60 °F.
Temperature, °F
Temperature of melt, °F
Melting point (freezing point) temperature of melt, °F.
Assumed to be 750 °F.
Temperature of heat sink (water, air) into which the heat
loss from the reducer is discharged, °F
Temperature at outer (cold) surface of alumina liner, °F
Temperature at outer (cold) surface of frozen melt skull, °F
Liner or skull thickness, ft
Thickness of alumina liner, ft
Thickness of frozen melt skull, ft
Overall heat transfer coefficient from outer surface of
frozen skull or alumina liner to heat sink, Btu/ft^ hr °F
Coke sulfur which is not recovered in reducer melt = mCQ
(1-e), Ib moles per Ib mole SOX (= Co/800 for the above
listed values of m and e)
Volumetric recirculation rate of melt between the oxidation
and reduction regions of the reducer, ft-Vhr
Normalized volume of melt bed in reduction region of
reducer, ft^ per Ib mole/hr SOX
Superficial air velocity in oxidation region of reducer,
ft/sec. Values of 3 and 5 ft/sec were considered in this
study
,, 2d .
= arc cos (1 — =r- )
Coefficient of volumetric expansion, ( °F)
viii
-------
Y = arc sin 28 = arc sin -=~-
g Separation baffle location parameter = j/D or j/L, as
may be appropriate
AH Enthalpy of reaction, Btu/lb mole. Subscripts and values
for the various reactions defined in Section III-4
A? Melt recirculation pressure drop, Ib/ft
AT Temperature differential between top and bottom of
melt bed, °F. Assumed to be 300 °F
At Temperature difference between bulk melt and reducer wall
surface in contact with the melt, °F
Ratio of radiation heat flux at free surface of melt to
convection heat flux at reducer vessel walls in contact
with the melt
S Melt residence time in reduction region of reducer, minutes,
Values of 15 and 30 minutes were considered in this study
* Fraction unreacted coke (or carbon). Assumed to be 0.05
p. Viscosity, Ib/hr ft
p Density, Ib/ft
p Effective average density of expanded melt bed in oxidation
region of reducer, Ib/ft^
p Effective average density of expanded melt bed in reduction
region of reducer, Ib/ft 3
i
p Density of unexpanded melt at average temperature of
™o oxidation region of reducer,
p Density of unexpanded melt at average temperature of
r reduction region of reducer,
IX
-------
SUMMARY
Regeneration of the process melt in the Molten Carbonate Process for the
removal of sulfur oxides from stack gases requires the reduction of the oxi-
dized sulfur conpounds absorbed in the carbonate melt. This reduction is
carried out at temperatures of around 1500 F using carbon in the form of
coke as both the reducing agent and the fuel to supply the heat required by the
reduction process.
As presently conceived, the design of the reducer is based on a separation
of the reduction and heat generation (oxidation) functions. The reducer is
divided into an oxidation region and a reduction region separated by a vertical
baffle with internal melt recirculation between the two regions. The design of
such a reducer requires a knowledge of reaction thermodynamics and kinetics,
and multiphase system hydraulics and heat transfer. Most of this information
must be obtained from laboratory and pilot plant tests. The purpose of this
study was to develop the analytical techniques required for the process design
of such a reducer and to provide preliminary data on the raw material require-
ments and physical dimensions of the reducer on the basis of presently avail-
able technical information.
Analytical techniques were developed for the process design of a two
region molten carbonate reducer utilizing either a vertical or a horizontal
cylinder vessel geometry. Most of'the process analysis itself was conducted
on the vertical cylinder geometry.
It was concluded that:
1) The heat loss from the melt bed must be minimized in order to
minimize the coke and air requirements of the reducer and its physical
dimensions.
2) The steady state sulfur compound concentration in the carbonate
melt must be as high as feasible subject to solubility limitations, thus minimiz-
ing melt flow rate and melt preheat requirements.
A-l
-------
3) The sulfate to sulfite ratio of the sulfur compounds in the melt
i i
leaving the scrubber must be minimized in order to minimze the coke consump-
tion of the reducer.
4) The coke utilization in the reducer must be maximized in order to
minimize the melt recovery and make-up requirement resulting from the melt
loss associated with the filtration of the unreacted coke.
The controlling constraints in the design of the reducer are the maximum
allowable superficial air velocity in the oxidation region and the minimum
required residence time of the melt in the reduction region. A decrease in
reduction region residence time from 30 to 15 minutes results in a more sig-
nificant reducer design improvement than an increase in superficial air velocity
from 3 to 5 ft/sec.
Typical reducer dimensions were determined as a function of processing
capacity and heat loss. Coke and air requirements were estimated as a func-
tion of heat loss. For a sulfur compound mole fraction in the melt of 0. 30,
one-third of which is sulfate, the coke and air requirements of a reduce*- with
a negligible heat loss amount to approximately 1.30 Ib and 73 scf, respectively,
per Ib of sulfur fed into the scrubber.
The heat loss from the melt bed in the reducer can be minimized through
use of an alumina-lined internally insulated (hot wall) reducer design. If it
is air-cooled, the outside wall of the reducer vessel in this design cannot be
insulated and operates at a temperature of around 500 F. Water cooling sur-
rounded by insulation would allow operation of the outside of the reducer at a
lower temperature.
Because of its inherently high heat loss a frozen melt skull (cold wall)
reducer design does not appear attractive when a low melting point carbonate
mixture such as the lithium-sodium-potassium carbonate eutectic is used as
the carrier melt.
A-2
-------
Internal melt recirculation between the oxidation and reduction regions
of the reducer can be controlled through sizing of the orifices in the baffle
between the two regions. These orifices must be small enough to assure
satisfaction of the minimum residence time requirement in the reduction
region and large enough to allow sufficient flow to transport the required heat
without exceeding melt bed temperature rise limitations.
The results obtained are based on the technical information presently
available and on various assumptions and approximations which had to be made
in lieu of experimental data. They are therefore preliminary and must be con-
firmed through a development program. They are, however, believed to pro-
vide a sound basis for the conceptual design of a molten carbonate reducer and
the planning of the required experimental program.
A-3
-------
I. IN TR ODUC TION
A key component of the Atomics International Molten Carbonate Process
for the removal of sulfur oxides from flue and other waste gases is the re-
ducer which provides for the reduction of the oxidized sulfur compounds
absorbed in the carbonate melt. The sulfites and sulfates are reduced to
sulfides which can then be regenerated into carbonates with release of
hydrogen sulfide for the recovery of sulfur. The reduction is carried out
at a temperature of about 1500 °F using carbon in the form of petroleum
coke as the reducing agent and providing the heat required for this reduction
through the indirect oxidation of excess coke with air.
The molten carbonate reducer technology is somewhat related to the
(9)
Kraft furnace technology in the paper industry. It is nevertheless sufficiently
different and novel to require the development of a strong analytical and
experimental technology base. It is the purpose of this study to provide a
preliminary process analysis of the molten carbonate reducer concept on
the basis of the technical information presently available and thus lay a
foundation for both the conceptual design of such a reducer and the experi-
mental program required to establish its technology.
A-4
-------
II. REDUCER DESIGN CONCEPT DESCRIPTION
The Molten Carbonate Process provides for recovery of the sulfur
from the waste gases treated in the form of hydrogen sulfide which can be
used for conversion either to elemental sulfur or to sulfuric acid. The
sulfur oxides, however, are absorbed in the carbonate melt in the form
of alkali metal sulfites and sulfates which require reduction to sulfides
prior to recovery of hydrogen sulfide and regeneration of the carbonate
melt. It is the function of the reducer to provide the environment re-
quired to perform this reduction. It is to be noted that such a reducer
may be applicable not only to the molten carbonate process, but to other
sulfur oxide recovery processes based on the use of carbonates as the
reactive species of the scrubbing medium.
The sulfites disproportionate rapidly into sulfides and sulfates at
the operating temperature of the reducer:
M2SO3 - 1/4 M2S + 3/4 M2SO4
The function of the reducer, therefore, is to provide for the reduction of the
alkali metal sulfates in the melt to alkali metal sulfides. The present con-
cept i.s based on the use of carbon as the reducing agent. This carbon is
provided in the form of a petroleum coke, such as fluid coke, though other
carbonaceous materials can be used if so desired.
The overall reduction reaction can be written as follows:
+ 2 C - MS +
The kinetics of this reaction are such that temperatures of about 1500 °F are
required to allow it to proceed at a practically acceptable rate. Moreover
the reaction is endothermic, absorbing approximately 75,000 Btu per Ib mole
of sulfate reduced.
The reducer must provide the high temperature, the heat of reaction,
and the residence time required to carry out the reduction. Provision of
excess coke and partial combustion of this coke with air are used to preheat
the reactants to the reaction temperature and generate the heat needed for
the reaction, and to compensate for whatever heat losses may take place.
The reducer thus becomes both a reduction and an oxidation unit.
A-5
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Feed Melt
from Scrubber
(0.35 gpm)
Basis: 1 Ib mole/hr SO
=•1.6 Mw
Regeneration Gas
to Regenerator
CO + HO + SO
sclm)
Waste Gas
to Stack
N + HO
(47 scfm)
/////////.
4M SO
Reduction
M S + 3M SO'
Mist Eliminator
j Li £ TC
M SO + 2C - M S + ECO
£* 4t £* La
2S + C + 2M CO - 2M S +
L* J £4
+ 3CO.
1400°F-s
1700°F
2 gpm
Oxidation
M S 4- 2O - M SO
L* £* £* T:
60°F
Coke
(48 lb/h.r)
Unreacted
Coke and Ash
(2. 7 Ib/hr)
600°F
Air
(53 scfm)
Reduced Melt
to Regenerator
(0.35 gpm)
Figure 1. Schematic Diagram of Typical Molten Carbonate Reducer
A-6
-------
The design of a reducer to allow performance of the reduction and the
oxidation functions within a single region would be expected to minimize the
overall size of the equipment required. Such a design, however, does not
appear to be practical at this time on the basis of our present knowledge
of the reactions involved, the reaction kinetics, and the required distribution
and circulation of reactants within the reacting mass. A design concept
was therefore selected which would provide for the separation of the reduction
and oxidation functions, and thus prevent the occurrence of competitive
reduction and oxidation reactions within a single highly turbulent region. Such
a concept allows for a considerable simplification of the analysis of the
process and a more reliable evaluation of the reducer design on the basis
of presently existing experimental information. This concept, however,
results in increased equipment size as compared with a hypothetical single
region reducer.
Figure 1 shows a schematic diagram of a typical two-region molten
carbonate reducer. The reducer consists of a reduction region and an oxi-
dation region, with a vertical separation barrier between them. This
barrier is provided with openings both just below the free surface of the
melt and near the bottom, of the reducer vessel to allow natural convection
circulation between the two regions. The heat required for the endothermic
reduction reaction and for preheating the molten salt melt and the coke and
compensating for heat losses is generated in the oxidation region and carried
into the reduction region by natural convection circulation of the hot melt.
Both the feed melt from the scrubber and the coke are schematically
shown as being introduced into the top of the reduction region. Some of
the melt may also be introduced at the top of the oxidation region and the
coke may actually be brought in with the feed air stream into the oxidation
region. Actual locations of the feed stream inlets and methods of distributing
the inlet flows will have to be established from information to be obtained
from the operation of an experimental test reducer. The melt and coke
come into contact with the oxidized recirculating "high" temperature melt,
are heated to reaction temperature, and react to form a reduced melt,
leaving a small amount of unreacted coke and ash. The temperature of the
melt decreases as the reaction proceeds. A product stream of reduced
melt and unreacted coke and ash is withdrawn from the bottom of the re-
duction region. The remainder of the reduced stream is recirculated to
the oxidation region.
A-7
-------
In the oxidation region preheated air comes into contact with the re-
circulated "low" temperature reduced melt stream. Part of the sulfide in this
stream, is oxidized to sulfate, generating heat and raising the melt to the "high"
temperature again. The required heat is then transferred to the reduction
region by the recirculated "hot" melt.
The reduction and oxidation functions are thus carried out completely
separately and the combustion of the excess coke is carried out through the
intermediate oxidation and reduction of recirculated excess sulfide-sulfate.
This separation of functions allows the separation of the off-gases
from the reducer, thus producing a high concentration carbon dioxide
stream for use in the regeneration reaction with only a minimum dilution
by nitrogen. Some carryover of CO£ from the reduction to the oxidation
region may, however, occur through release of CO2 from the molten
carbonates in the oxidation region and recombination of CO_ with the oxides
thus formed in the reduction region.
The flow rates shown in Figure 1 are normalized to the treatment of
gases carrying 1 Ib mole of SOX per hour (approximately equivalent to 1. 6
Mw electric power generation in a power plant buring 3 wt % sulfur coal).
They are typical but obviously vary with variations in the operating parameters
of the process, as shown in the study described in this report.
A key problem of the molten carbonate reducer concept is the selection
of structural or liner materials which come in contact with the melt at the
operating temperatures of the reducer. On the basis of presently available
materials compatibility data, two design concepts were selected for analysis
and evaluation. One of these is the so-called frozen skull concept, similar
(9)
to the designs used in the Kraft furnaces in the paper industry. This concept
relies on a film of frozen melt at the inner surface of a steel vessel to pro-
vide the necessary protection of the steel from corrosion-by the melt. The
frozen skull is formed by cooling the walls of the vessel. The temperature
at its surface in contact with the melt is the freezing temperature of the '
melt (approximately 750 °F) and its thickness is determined by the amount
of heat which must be removed in order to maintain this temperature.
A-8
-------
The second design concept makes use of a ceramic liner on the inside
of a steel vessel. The ceramic presently under consideration is alumina.
The alumina is compatible with the high temperature melt and provides both
for physical protection and thermal insulation of the steel. Since the alumina
is compatible with the melt, its surface temperature in contact with the melt
can be maintained relatively close to that of the melt (1300-1450 8F) to mini-
mize heat losses.
The two reducer design concepts selected can therefore be characterized
as follows:
1. Frozen Melt Skull: A cold wall concept with relatively high heat loss.
2. Alumina-Liner: A hot wall concept with relatively low heat loss.
The present study has been directed toward the analysis and evaluation
of a two region molten carbonate reducer using either a frozen melt skull
(cold wall) or an alumina liner (hot wall) for corrosion protection of the steel
shell of the reducer vessel.
A-9
-------
in. BASIS OF STUDY
1. Composition and Physical Properties of Melt
Molten carbonate make-up: Eutectic mixture of lithium, sodium, and
potassium carbonates of the following composition :
Li2C03 43. 5 mole % (32 wt %)
Na2CO3 31.5 mole % (33 wt %)
K2CO3 25. 0 mole % (35 wt %)
Sulfur compound concentration in melt feed to reducer: Two cases
were considered:
= 0.15 and 0.30 (mole fraction)
Extent of oxidation of sulfur compounds in melt leaving the scrubber
(prior to any disproportionation):
[M soj
*- ' J = 1 - a = 0. 33 (mole fraction)
[M.SOj +
Extent of disproportionation of sulfite into sulfide and sulfate in melt
upstream of reducer inlet:
= p = 0. 10 (mole fraction)
Physical properties of melt (assumed to be approximately equal to
those of the alkali metal carbonate eutectic melt at the same temperature):
Density: P = 147.60 - 18.87 x 10"3 T, Ib/ft3,with T in 0R^
Specific volume of expanded bed = 1.6 ft /lb mole of melt
(equivalent to approximately 75% bed expansion due to coke addition and
gas flow without attempt at differentiation between the melt in the oxidation
region and that in the reduction region).
Viscosity: fl = 11.253 x 10"3 e"65/T, Ib/hr ft, with T in °R(1)
A-10
-------
Specific heat: c = 28.41 + 9. 20 x 10 T, Btu/lb mole °F, with
(1) p
T in °RV ;
Enthalpy increase of melt from 850 to 1500 °F: c = 23,800 Btu/lb
mole (average specific heat = 43. 5 Btu/lb mole °F)
Thermal conductivity: k = 0. 5 Btu/ft hr °F, estimated (same
value assumed for frozen melt skull)
Nominal freezing point temperature = 750°F
2. Composition and Heat Capacity of Coke
Composition of typical high sulfur fluid coke:
C 90. 0 wt %
H 1.8 wt % -n = 0.24 atoms H/atom C
S 6. 0 wt % *-m = 0. 025 atoms S/atom C
O + N 1.5 wt %
Ash 0. 7 wt %
Present indications are that the hydrogen in the coke does not contribute
to the reduction of the sulfate in the reducer. The overall heat balance of the
reducer melt bed, therefore, did not take credit for any contribution from the
heat of oxidation of the hydrogen, and was thus based on the use of a value of
n = 0. In reality it is expected that some of the hydrogen will be oxidized
within the melt, thus contributing to some extent to the heat balance of the
bed.
The specific heat of the coke was assumed to be approximately equal to
that of graphite, with values ranging from 0. 142 Btu/lb °F at 0 °F to 0.435
Btu/lb °F at 1500 °F^ . Numerical integration between 60 and 1500 °F yields
an enthalpy increase of 487 Btu/lb over this temperature range (average
specific heat = 0. 338 Btu/lb °F).
On a mole basis the enthalpy increase of the coke from 60 to 1500 °F
amounts to:
c = 6500 Btu/lb mole of contained carbon
PC
A-ll
-------
3. Composition and Heat Capacity of Air
Composition of air (Ib moles/lb mole contained oxygen):
°2 l
N2 3.77
H2° 0. 10
Air 4.87
The assumed water content corresponds to 0. 013 Ib of water vapor per
Ib of dry air, or a relative humidity of 60% at 80 °F.
The enthalpy increase of the air from 600 to 1500 °F was estimated
from the data of reference (3), yielding a value of 6960 Btu/lb mole of air
(average specific heat = 7. 73 Btu/lb mole °F). Per mole of oxygen, the
enthalpy increase of the air from 600 to 1500 °F amounts to:
c = 33, 900 Btu/lb mole of contained oxygen
PA
4. Heats of Reaction
The following values were used for the heats of reaction of the various
reactions involved at the operating temperatures of the reducer :
(J2) - -1/4 M2S (J2) + 3/4 M2 SO4 (J?) (1)
AH~. = 21,400 Btu/lb mole
Disp '
M2S04U) + 2 C (s) - ~M2S (JL) + 2 C02 (g) (2)
AHc/~ = c= = + 75,500 Btu/lb mole
-»
M2S(i) + 2 02(g) - -M2 S04U) (3)
AHG= „_ = - 415,200 Btu/lb mole
O — * DVj .
4
C(s) + 02 (g) - -C02(g) (4)
= -169,500 Btu/lb mole
2
H2(g) + 1/2 02(g) - -H20(g) (5)
AH,, ~ = -106,900 Btu/lb mole
H2°
S(s) + 02(g) - -S02(g) (6)
= -133,400 Btu/lb mole
A-12
-------
S(s) + 1/2 C(s) + M2 C03(4) — M2S(J2) + 3/2 C02(g)
_ ^g= = + 53,400 Btu/lb mole
A recent reevaluation of the data used for the heats of reaction at
1500 °F for reactions (4) and (7) indicates that -169,900 and +43,200 Btu/lb
mole, respectively, would have been more accurate values. The effect of
this difference on the present calculations is negligible.
It is to be noted that the above listed values of the heats of reaction
for reactions (1), (2), (3), and (7) are actually the values for these reactions
with sodium salts. They were assumed to provide close enough approximations
for the heats of reaction with the lithium and potassium salts to allow their use
in this reducer study.
5. Reaction Completion Efficiencies
Assumed to be 95% for each step of the molten carbonate process, with
one exception: complete consumption of the oxygen provided was assumed in
the oxidation region of the reducer.
Coke loss as unreacted coke was assumed to be 5%, allowing usage of
95% of the coke feed material.
6. Reducer Design Parameters
Feed temperatures:
Melt 850°F
Coke 60 °F
Air 600°F
Reducer operating temperature = 1400 to 1700 °F, with a temperature
rise of 300 °F in the oxidation region and a corresponding temperature drop
in the reduction region. A smaller temperature gradient would have required
a higher melt recirculation rate and therefore resulted in a larger size
reducer. A larger gradient would have raised the oxidized melt temperature
to an undesirably high level.
Superficial gas velocity < 3 ft/sec, a controlling factor in the oxidation
region, but not in the reduction region, since the amount of gas generated in
the reduction region is much smaller than the amount of air used in the oxida-
tion region. While most of the analysis was based on a velocity of 3 ft/sec, one
set of calculations was performed for a velocity of 5 ft/sec to determine the
effect of this parameter on reducer design.
A-13
-------
Melt bed expansion due to gas flow and coke addition = 75 volume %( '.
No attempt was made to differentiate between the oxidation and reduction
regions (except in the melt recirculation calculations of Section VIII).
(7)
Melt residence time in reduction region >^ 15 minutes . The residence
time in the oxidation region is not a controlling factor since the oxidation of
sulfide to sulfate is rapid at reducer operating temperatures. While most of
the analysis was based on a residence time of 15 minutes, one set of calcu-
lations was performed for a residence time of 30 minutes to determine the
effect of this parameter on reducer design.
7. Reducer Design Concepts
Two-region reducer with separate oxidation and reduction regions.
Reducer wall corrosion protection provided by either frozen melt skull
(cold wall) or alumina liner (hot wall).
Reducer geometry: Vertical cylinder. Process design equations
were also developed for a horizontal cylinder geometry both with the
separation baffle perpendicular to the axis of the cylinder and the separation
baffle parallel to the axis of the cylinder. Within the scope of this work
only one set of calculations was performed on a horizontal cylinder geometry
for comparison purposes.
8. Reducer Processing Capacity
The mass and heat balance analysis was performed on a normalized
basis of treatment in the molten carbonate scrubber of a waste gas feed stream
carrying 1 Ib mole per hour of sulfur oxides (SO_ and SC< ).
The physical dimension and heat loss optimization was then conducted
as a function of reducer unit capacity over a range of 0 to 200 Ib moles of
sulfur oxides per hour in the waste gas feed stream to the molten carbonate
process system. Special emphasis was placed on capacities of 3. 125 and
166. 7 Ib moles of sulfur oxides per hour. For power plants burning 3 wt %
sulfur 12,800 Btu/lb coal with a heat rate of 9000 Btu/Kwh these capacities
are equivalent to those of a 5 Mw pilot plant size unit and a 267 Mw (one of
three units of an 800 Mw plant) full scale unit, respectively.
A-14
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IV. REDUCER MASS BALANCE
A mass balance was performed around the whole molten carbonate pro-
cess system on a normalized basis of 1 Ib mole per hour of sulfur oxides con-
tained in the feed waste gas stream entering the system. The reaction com-
pletion efficiencies were assumed to be as described in Section III. The
mass balance around the reducer is shown in Figure 2.
The key independent variable in this mass balance is CQ, the carbon
requirement, expressed in Ib moles of carbon which do react in the reducer
per Ib mole of sulfur oxides contained in the feed waste gas stream to the
scrubber. C is determined by a heat balance around the melt bed in the
reducer, as described in Section V. The actual carbon usage amounts to
C /1-A when one takes into account the fraction,X, of the carbon feed which
is discharged unreacted from the reducer.
All other parameters are determined by the composition of the feed
streams and the extent of the chemical reactions which take place upstream
of the reducer (partial oxidation of sulfite to sulfate, partial disproportionation
of the sulfite).
If one assumes that a mole fraction, g, of the sulfur oxides in the feed
gas to the scrubber is in the form of SO3, that no further oxidation of SO2
to SO3 or sulfite to sulfate takes place in the scrubber, and that all the sulfide
in the melt feed to the scrubber is oxidized to sulfate, the parameter a can
be expressed as follows: 1
r- -I
|M2SOjJ 1 +
+ [M2SOJ ( ~ !.£ i_g e(l-e)
L- -I ^ 1 + P — - 7^ —4—
scrubber outlet
The parameter a decreases, and therefore the sulfate fraction in the melt at
the outlet from the scrubber increases with increasing concentration of SC>3
in the feed gas, g, decreasing reaction completion efficiency, e, increasing
disproportionation upstream of the reducer (actually upstream of the branch-
ing off of the scrubber recycle melt stream from the reducer melt stream),?.
and increasing ratio of scrubber recycle to reducer melt flow, -j- .
A-15
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FEED MELT
FROM SCRUBBER
S0~ =-- a*'vl-p)a
REGENERATION GAS
TO REGENERATOR
CO., = ( 1+me) G
2 o
H,0 = ? C
2 2 o
SO_ = m(l-e) C = u
2 o
Total =
REDUCER
REDUCTION
OXIDATION
COKE
s
Ash
t
UNREACTED
COKE AND ASH
C =
S =
XCo
1-X
1-X
H = n
1-X
Ash
Basis; 1 Ib mole/hr So
^ 1. 6 Mw ^
WASTE GAS
TO STACK
=3.770
HO = 0.100
Total = 3.87 Q2
AIR
N2 =3.7702
H20 = 0. 10 02
Total = 4.87 O2
REDUCED MELT
TO REGENERATOR
S" = 1 +
C0 .
u
1-e
a*ii£a
e * [S]
Total = i±SL a* , *
1-e
'ul
e
Figure 2. Molten Carbonate Reducer Mass Balance
A-16
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The parameter u is given by the relationship:
u = m (1-e) CQ
It represents that sulfur in the reacted coke which is not reduced to sulfide
in the reducer and is assumed to become converted to SO^. It is a very
small quantity, amounting to C /800 at the selected values of e = 0. 95 and
m= 0.025.
The parameter a* is a reaction efficiency adjustment factor:
a* = ——? , which for e = 0. 95 gives a* = , and
i+rS a i+£
is therefore generally almost equal to unity.
It is to be noted that, as expected, the controlling parameter determining
the net melt flow rate into and out of the reducer is the maximum allowable
sulfur compound concentration in the melt, [_SJ . At sulfur compound concen-
trations of 15 and 30 mole %, the melt flow rates into and out of the reducer
will be approximately 7.0 and 3.5 Ib moles, respectively, per Ib mole of
SOX in the feed gas to the scrubber, corresponding to about 0. 70 and 0. 35 gpm
(at 850 °F) per Ib mole/hour of SOX in the feed gas.
A-17
-------
V. REDUCER HEAT BALANCE
The reducer must provide the heat necessary to heat the various feed
streams to the reaction temperatures, to furnish the endothermic heat of
reaction, and to compensate for the heat losses to the walls of the reducer
vessel. The heat is generated by the oxidation of the carbon in the excess
coke used for this purpose. The heat balance around the reducer, therefore,
determines the amount of coke required, and hence the parameter, Co, Ib
moles of carbon which react in the reducer per Ib mole of sulfur oxides in
the feed waste gas to the scrubber, in the mass balance described in
Section IV. The heat balance also determines the amount of air required
to oxidize the excess carbon and the amount of melt which must be recircu-
lated between the oxidation and reduction regions to transport the required
heat from the former to the latter. It therefore determines the physical
size requirements of the reducer.
It is important to note that for purposes of the present analysis, heat
balance around the reducer and heat loss from the reducer are defined as
heat balance around the bed of melt in the reducer and heat loss from this
bed, since all the chemical reactions involved, and their heat and temperature
requirements, are assumed to take place in this bed and not in the gas space
above the melt.
Knowing the specific heats and temperatures of the feed streams and the
various heats of reaction involved, as given in Section III, and the mass balance
from Figure 2 in Section IV, one can perform a heat balance around the reducer
if the heat loss can be either estimated or assumed.
An estimate of heat loss depends not only upon estimates of heat transfer
coefficients and. thermal conductances for the evaluation of the heat flux, but
also upon the heat transfer area across which this heat flows. This area can
only be determined once the physical dimensions of the reducer melt bed are
known, and these are derived from the results of the heat balance.
A-18
-------
The heat balance was therefore performed as a function of an inde-
pendent heat loss parameter, Q the heat loss from the reducer melt bed
normalized to a sulfur oxide flow of 1 Ib mole per hour in the feed gas to the
scrubber. All calculations were performed as a function of this parameter.
The heat flux across the reducer vessel walls was then obtained from the
assumed heat loss and the physical dimensions thus calculated, and matched
with anticipated heat transfer rates.
The heat balance around the reducer melt bed is presented in Figure 3.
The results are summarized in the reacted carbon and total coke require-
ments given by equations (1) through (4) of this figure. The numerical values,
of course, are based on the parameters and data given in Section HI, and
specifically feed stream temperatures of 850, 60 and 600 °F, respectively,
for the melt, the coke, and the air.
The reducer coke and air requirements are plotted in Figures 4 and 5,
respectively, as functions of the normalized heat loss parameter, QT , of the
sulfur compound concentration in the melt, Is] , and of the sulfate fraction,
pO2 , of the sulfur compounds in the melt leaving the scrubber. These
figures also provide typical coke cost and air compression horsepower data
for a power plant burning 3 wt % sulfur, 12, 800 Btu/lb coal with a heat rate
of 9, 000 Btu/Kwh. A discussion of these results is provided in Section IX. 3. a.
A-19
-------
Definition of Constants
A = Net heat available in coke, taking into account the heat required
for preheating of both the coke and the air for its combustion,
Btu/lb-mole of carbon
c
n \ PC / n\
= " ( AHro + mAHqn + 7 AHw n) " 1—T " (1 + m + l) Cr,
\ ^^-y ^^o «-*-'/ A ~ A. \ */ PA
L, C, £t -Ti
For \ = 0. 05 A = 128, 760 + 99, 500 m + 44, 980 n
Form =0.025 I A=131j20Q
n= 0 J
B = Heat required for reduction of sulfur in coke to sulfide, plus heat
made unavailable by consumption of carbon for this reduction,
Btu/lb-mole of sulfur
=) « • 2ZO' 70°
s
E = Heat required for reduction of sulfate in melt to sulfide, plus heat
made unavailable by consumption of carbon for this reduction,
Btu/lb-mole of sulfate
• - (2 iHco - *Hso= - r ) - 2 - 346- 7CO
K = Calculation constant, Btu/lb-mole
c
PT
n.
[S] e Disp
Fore =0.95 K= a* (346,700+ 29r>e.8,0° - 107,000;
p = 0.10 J X
Figure 3. Molten Carbonate Reducer Heat Balance
A-20
-------
HEAT BALANCE
CARBON REQUIREMENT, Ib-mole C/lb-xnole SO
Heat Available in Coke + Heat of Disproportionate*! =
- Heat for Coke Sulfur Reduction + Heat for Sulfate Reduction +
+ Heat for Preheat of Melt + Heat Loss
AC + - AH_.
o \ Disp/ e
= BmC
a* (1 - p) a =
--}
4J
e+c
M
+Q
[S] L
(A - Bme) C = (l + u)K+Q=K+QT+K (
o L L m
e)C
K+ Q.
C =
o
A - mfee + K(l e)]
(1)
For X
m
n
e
p
0. 05
0.025
0
0.95
0. 10
C =
o
a* (346, 700 + 29>800 107, 000 a] + Q.
\ I kj I -I
126, 000 - a* (433 + ^pST -134;
\ L^J
(2)
= a* 2. 7635+ '
- 0.8529aV 0.7971/QL
;
(3)
COKE REQUIREMENT, Ib Coke/lb Sulfur
For coke containing 90 wt % carbon, and same parameters as above
0.1042
Coke= a* Jo. 838+ ^-j + 0.374 [SO~ ]} + 0.350
(4)
with a* =
1 - [SO = ]
sr l
1 +
76
Figure 3. Molten Carbonate Reducer Heat Balance (Cont)
A-21
-------
DC
ID
LL
_J
CO
XI
LU
2
UJ
DC
a
LU
DC
HI
^
o
o
HEAT LOSS PENALTY PER 100,000 Btu/hr/lb-m S0x/hr
0.350 IbCOKE/lb SULFUR
0.0245 m/kwh d)
15 m% S IN MELT
--- — 30 m% SIN MELT
200
400
600
800
NORMALIZED HEAT LOSS (103 Btu/hr/lb-m S0x/hr>
0.66
0.55
0.44
0.33
0.22
0.11
1000
(1) AT 20^/106 Btu THIS HEAT WOULD BE WORTH 0.0125 m/kwh
(2) FOR POWER PLANT BURNING 12,8000 Btu/lb, 3 wt% S COAL WITH A HEAT RATE
OF 9000 Btu/kwh
Figure 4. Reducer Coke Requirement
O
O
UJ
^
O
O
A-22
-------
600
500
QC
Z3
U-
_l
co
z
UJ
2
ui
cc
a
LJJ
cc
400 —
300 —
HEAT LOSS PENALTY PER 100,000 Btu/hr/lb-m SO /hr
45.4 scfAIR/lb SULFUR X
0.726 hp/Mw
15 m% SIN MELT
— — 30 m% SIN MELT
NORMALIZED HEAT LOSS (I03 Btu/hr/lb-m S0x/hr)
(1) FOR POWER PLANT BURNING 12.800 Btu/lb, 3 wt% S COOL WITH A HEAT RATE OF
9000 Btu/kwh
(2) ASSUMING A PRESSURE REQUIREMENT OF 10 psig AND A COMPRESSOR
EFFICIENCY OF 75%
Figure 5. Reducer Air Requirement
A-23
-------
VI. PHYSICAL DIMENSIONS OF REDUCER MELT BED
The physical size of the reducer melt bed was determined by calculating
the cross-sectional areas of the oxidation and reduction regions as a function
of bed depth at various values of the normalized heat loss parameter, QT ,
assuming maximum allowable superficial gas velocity and minimum reduction
region residence time to be the controlling parameters (the melt recirculation
rate between the oxidation and reduction regions does not appear to be con-
trolling, as will be shown in Section VIII). The depth of the reducer melt
bed was then optimized to achieve a minimum heat transfer area of reducer
wall surfaces in contact with the melt. This optimization was only approxi-
mately correct as it was performed at constant value of the normalized heat
loss parameter, QT , rather than at constant value of the heat flux across the
L
wall surface, Q/A, which is the truly independent variable in the design of
the reducer, as shown in Section VII. It is to be noted that such an optimization
based exclusively on heat transfer considerations may be of only limited value
in a low heat loss reducer in which heat loss considerations may prove to be
secondary to structural considerations.
The general relationships developed in this section express the various
parameters as functions of both the normalized heat loss parameter, QT , and
L
the normalized amount, O~, of oxygen consumed in the reducer. O-, is actually
a function of the amount of carbon consumed, C , per the mass balance of
Figure 2, and C in turn if a. function of the sulfur compound concentration in
the melt, JSJ , and of the normalized heat loss parameter, Q, , per the
heat balance of Figure 3. In the general case the equations are simpler when
written with O? left in as an explicit parameter rather than expressed as a
function of fsl and QT which it actually is. In the specific case of the
present study, however, the sulfate mole fraction in the sulfur compounds
in the melt leaving the scrubber was not varied and was assumed to be equa'l
to 0.33. Under these conditions, the approximate equation (3) of Figure 3
can be used to express the value of C and therefore that of O? as a linear
function of JSJ and Q,. and thus to provide an explicit relationship between
the various parameters and IS I and Q, .
A-24
-------
For 1-a = [M2S04|/[lM?SOj + [M?SOj j = 0.33, equation (3)
of Figure 3 becomes:
n
C = 2.1758 + re-i + 0.7971
o
'L
,5
Ib moles/lb mole SO (5)
10-
From Figure 2, using the value of C given by equation (5):
x
o
O2 = 0.4960 Fp lb moles/lb mole SO (6)
.., ~ i ,
withFj = 1+
x
I r\ \
0.4685
Equation (6) provides the relationship used to express all the other
parameters as explicit functions of [Sj and Q? for the specific sulfate-sulfite
concentration ratio assumed in this study.
1. Cross Sectional Area of Oxidation Region
The oxidation of sulfide to sulfate with air at around 1500 °F is rapid.
The residence time of the melt in the oxidation region is therefore not con-
sidered to be a controlling factor in determining the volume of the melt bed
in this region. The upper limit on the maximum allowable superficial gas
velocity thus becomes the controlling parameter determining the cross-
sectional area of the oxidation region.
The mass balance of Figure 2 shows the amount of air required as
equal to 4. 87 O9 lb moles/hr per lb mole/hr of SO in the feed gases to the
Lt X
scrubber. The present calculations were based on the superficial velocity
of the air at 1500 °F. This is conservative since the oxygen from the air is
removed by reaction with the sulfide as it progresses through the oxidation
region.
.'. Air flow at 1500 °F = 1. 935 O7, acfs per lb mole/hr SO
^ x
If the maximum allowable superficial gas velocity is v ft/sec, the
minimum area. A , of the oxidation region is given by the equation:
° °2 2
A = 1. 935 —- , ft per lb mole/hr SO (7)
O V X
Hence, per equation (6):
Fl 2
AQ = 0. 9597 -^- , ft per lb mole/hr SOx (8)
A-25
-------
2. Internal Melt Recirculation Requirement
The amount of melt recirculation required between the oxidation and
reduction regions of the reducer is determined by the amount of heat which
must be transported from the oxidation to the reduction region, and the
allowed temperature rise of the melt. The temperature rise was selected
as equal to 300 °F for purposes of the present study (Section III). The heat
conduction through the separation baffle was assumed to be negligible.
The amount of heat to be transported can be calculated most easily
by determining what happens in the oxidation region. In this region heat
is generated by only one chemical reaction, the oxidation of sulfide to sulfate,
which consumes all the oxygen fed to the reducer. Part of this heat release
is utilized in the oxidation region as preheat of the air feed and as heat loss,
QT . The remainder is transported to the reduction region in the form of
j-»
increased sensible heat of the recirculated melt.
°2
/. Heat Release = - (AH0= __= ) -r* = 207,600 O0, Btu/hr
'
per Ib mole/hr SO
Air Preheat = c O0 = 33,900 O0, Btu/hr per Ib mole/hr SO^
PA 2 2
Heat Loss = Q , Btu/hr per Ib mole/hr SO
J—t X.
o
Heat Transport = QT = -(l/2AHg=_^so = + c ) °2 ~ QL
= 173,700 O9 - QT , Btu/hr per Ib mole/hr SO
£ J—» X.
o
A simplifying approximation was made at this point by assuming the
heat loss from the reducer to be equally distributed between the oxidation
and the reduction regions: Q = -y Q... Then:
o
Q™ = 173,700 O- - QT /2, Btu/hr per Ib mole/hr SO (9)
J. Lf J_l X
Using equation (6):
CU = 86,160 F_, Btu/hr per Ib mole/hr SO (10
X £* X
with
A At-OC
- + 1.0064
2
A-26
-------
For a melt temperature rise of 300 °F with a specific heat at these
temperatures of about 45 Btu/lb mole °F, this heat transport requires:
/QT\
Melt Recirculation = 12.87 O? - 3.7041—=H, Ib moles/hr per Ib
\ioV
mole/hr SO (11)
X
Using equation (6):
Melt Recirculation = 6. 3819 F-, Ib moles/hr per Ib mole/hr SO (12)
^ x
Should it prove desirable or necessary to use a melt temperature rise,
AT, different from the 300 °F assumed in this analysis, equation (12) can
be rewritten as follows:
Melt Recirculation = — F> Ib moles/hr per Ib mole/hr SO (12a)
3. Cross-Sectional Area of Reduction Region
The volume and cross-sectional area of the reduction region can be
determined from the melt flow through this region and the residence time
required to allow the reduction reaction to proceed to the desired degree
of completion. The melt flow is calculated as follows:
Reduction Region Melt Flow = Melt Recirculation + Melt Feed
/•Q \ 1 1
= 12. 87 O, - 3. 704 (—W + —— a* -^
2 \1QS/ e [S]
= 12.87 O0 - 3.704(—^7 + rcf5 » H> moles/hr per Ib mole/hr SO (13)
2 MO^ LbJ x
Using equation (6): '
Reduction Region Melt Flow = 6. 3819 F,, Ib moles/hr per Ib mole/hr SO (14)
1.0064
If one assumes a specific volume of 1.6 ft /Ib mole for the expanded
melt bed (equivalent to a 75% bed expansion due to gas flow and suspension
of coke), a minimum reduction region residence time of B minutes requires
a minimum expanded bed volume, V in this region such that:
A-27
-------
Vr = 1>6 TT (ReductionRe«ion Melt Flow)
= JO. 3431 O, - 0. 09877 (-— ~) + °h3f810f ft3 per Ib mole/hr SOv (15)
[ £. \ l?J J x
= 0. 1702 F, 8, ft3 per Ib mole/hr SO (H)
•
X
For a vertical cylinder reducer geometry, the cross-sectional area
of the reduction region, A , can be determined as a function of expanded
melt bed depth, d:
( SQT\ 0 0281 o 2
A = JO. 3431 O, - 0. 09877 ('—•*-) + V"H9 [ -f, ft per lb mole/hr SO (1?)
r [ £ \10y IrJ J a x
= 0. 1702 F3 -|, ft2 per lb mole/hr SOx (18)
Should it prove desirable or necessary to use a melt temperature rise,
A T, different from the 300 °F assumed in this analysis, equations (14), (16),
and (18) can be rewritten as follows:
1915
Reduction Region Melt Flow = • Arp F^, lb moles/hr per lb mole/hr SO (1-
1.0064
O
Vr ="f F3 6> ft per lb mole/hr SOX (16a|
Ar =^T^ F3 -f, ft2 per lb mole/hr SOx (18a|
In general it is found that the cross-sectional area of the reduction
region thus determined on the basis of minimum melt residence time is
greater than that obtained on the basis of maximum superficial gas velocity:
C _
A =0. 3973 (1 + m + -|) —-, ft per lb mole/hr SO (19)
rGas Velocity v x
Using equation (5):
A =— |o.9898 + O'r£i'x + 0.3626
rGas Velocity v
ft2 per lb mole/hr SO (20)
Jfi
A-28
-------
Under the conditions of this study the reduction region residence time is
therefore the controlling parameter in determining the cross -sectional area
of this region.
4. Melt Bed Diameter and Heat Transfer Area - Vertical Cylinder
Reducer Geometry
The total cross -sectional area of the reducer, A is the sum of the
cross -sectional areas of the oxidation and reduction regions:
NVJ
d
* The value of e is difficult to determine as it depends upon the emissivity
of the surface and the amount of radiation shielding which may take place.
In all the numerical calculations in this report it was assumed to be
negligible (e= 0)
(21)
The actual cross -sectional area, A , of a reducer with a capacity
OS
to handle N Ib moles/hr of SO in the feed gas to the scrubber is:
X.
A = NA,. (22)
cs t
For a vertical cylinder reducer geometry the diameter of the reducer, D,
is: /A
D =
The heat loss from the melt bed takes place across the area of the
walls in contact with the melt (side walls and bottom, assumed to be flat,
of the reducer vessel), and, by thermal radiation, across the free surface
of the melt. If one assumes that the radiation heat flux at the free surface
of the melt amounts to a fraction, e , of the convection heat flux at the walls,
the effective heat transfer area, A , in a vertical cylinder reducer
geometry, can be expressed as follows:
Aht " Acs (! + e ) + 7rDd ACs ( + e ) + 2dV"" cs
•} (!+€)+ 2\/7TNd(Aod + Vr) (25)
A-29
-------
5. Melt Bed Diameter and Depth Optimization - Vertical Cylinder
Reducer Geometry
The optimization of the diameter and depth of the reducer melt bed was
directed toward achieving a minimum heat loss from the bed, neglecting any
structural considerations which may actually prove to be the controlling
factors. At a given rate of heat transfer across the walls of the reducer in
contact with the melt, such an optimization is equivalent to minimizing the
area, A, ,, available for this heat transfer. To be completely rigorous the
optimization must be done at constant heat flux. This, however, would have
become excessively complex for the present scope of the analysis. A simpli-
fying approximation was used by performing the necessary differentiation of
Equation (25) at constant normalized heat loss, QT , instead of constant heat
flux, (-/r\ such that:
NQ
-7-^ (26)
Aht
The optimization was therefore performed by minimizing the value
of the effective heat transfer area, A, , in equation (25), with respect to
bed depth, d, at constant normalized heat loss, Q . The heat transfer area,
A, ,, is minimum when the following relationships prevail:
N _ TTd3 (l+2b)2
- vr i + b
D = 2d (l+2b) YT7 (281
bfl+2b) _I_
(3+2b) 1+e
JL 1~b
8 Hb
_
sin/
_7T 1-b 1
8 1+b S2 S4 6
, L. O C. O 4O
A-30
-------
with b = o = ratio of volume of oxidation region to volume of
reduction region
Fl d
= 5. 639 —p— —r for the case of the present analysis with
3
a 0. 33 mole fraction sulfate in the sulfur compounds in the
melt leaving the scrubber*.
j = distance from axis of reducer vessel to separation baffle,
ft, positive or negative depending upon whether the volume
of the oxidation region is smaller or larger than that of
the reduction region, respectively.
7 = arc sin rr*pr = arc sin 28.
It is to be noted that the value of the ratio b increases with increasing
reducer processing capacity, becoming equal to unity when
9J
2
,
A
o
6, Horizontal Cylinder Reducer Geometry with Separation Baffle
Perpendicular to Cylinder Axis
For the horizontal cylinder reducer geometry with the separation
baffle perpendicular to the axis of the cylinder, the lengths, LQ and L^, of
the oxidation and reduction regions are given by the following equations:
NA
T _ _ 2-
o ~ D sina (31)
4NVr
o j ~r\
with a = arc cos (1 - --) and therefore d = -- (1 cosa).
* It is to be noted that there is an inconsistency in the time units used in
the numerical equations as N is expressed in Ib moles/hr, v in ft/sec,
and Q in minutes.
A-31
-------
The total length, L, of the reducer is therefore:
NA
bf4—co.aV
\ sma y
AoD
with b = j== (note the difference between this definition and that used in
•* vr
Section VI-5).
The effective heat transfer area, Aht, across which the heat loss from
the melt bed takes place is the area of the walls in contact with the melt (lower
part of the cylindrical section plus end plates which for purposes of this
calculation are assumed to be flat) and the partially effective free surface of
the melt:
(34)
An optimization of reducer dimensions similar to the one performed
for the vertical cylinder reducer geometry yields a minimum value of the
effective heat transfer area, Aht, at constant values of the normalized heat
loss parameter, Q^, when the following relationships prevail:
_3 4 NVr
J /!
d = "Sl = cos a)
T T^ /1 sin 2 a . , , , .
L=D(1 .____) 1 + b(____ cos a) +^ina (37)
2b
cosa)
1 - b( ' - cosa )
ri
sina
(39)
1 + b ( — r—i ' - cos a \
'
A-3 2
-------
a = arc cos (1 - -
j = distance from vertical midplane of reducer vessel to separation
baffle, ft, positive or negative depending upon whether the volume of the
oxidation region is smaller or larger than that of the reduction region,
respectively.
As in the case of the vertical cylinder reducer geometry, the ratio
of the volume of the oxidation region to that of the reduction region increases
with increasing reducer processing capacity, becoming equal to unity when
'
a
—: cos a
sma
and hence
3 —,SL- (. . n - cosa ) , , , sin a
A sina sma 1 + € ~~Z
o u
The free surface of the melt should preferably be somewhat below
the horizontal midplane of the reducer vessel, with, for instance, a value
of expanded melt bed depth to diameter ratio d/D = 0.40. Assuming e = 0,
equations (35) through (39) can then be written:
D3 = 3.9779 N V (40)
r
d = 0.4D (41)
L = 0.8569 D (1 + 1. 1977 b) (42)
_ .QL 1.7138 b (43)
" AQ 3 + 2.3954 b v '
- 1.1977b {44)
l '
_
- 2 1 + 1.1977b
A D F, D
with b = — Q — = 1 410--rr" ~~^ f°r the case of the present analysis with
a 0. 33 mole fraction sulfate in the sulfur compounds in the melt leaving the scrubber.*
* See footnote on page 31
A-33
-------
7. Horizontal Cylinder Reducer Geometry with Separation
Baffle Parallel to Cylinder Axis
For the horizontal cylinder reducer geometry with the separation baffle
parallel to the axis of the cylinder, the basic equations can be written:
NAQ = - (sina sin/) (45)
NVr =
+ 2. siny(cosy-cosa)] (46)
S) (47)
with: a= arc cos (1 -~r^~)
Y = arc sin 28
j = distance from axis of reducer vessel to separation baffle, ft,
positive or negative depending upon whether the volume of the
oxidation region is smaller or larger than that of the reduction
region, respectively.
An optimization of reducer dimensions similar to the one performed
for both the vertical cylinder reducer geometry and the horizontal cylinder
reducer geometry with the separation baffle perpendicular to the axis of the
cylinder yields a minimum value of the effective heat transfer area, A, , at
constant values of the normalized heat loss parameter, QT , when the follow-
JLj
ing relationships prevail:
3 _ 8 NVr (1 +
/, sin_2a > ( , , sin 2QU „ n sin 27 . _ . , .1 (48)
(1--^ - )M1 — 2a — ' ^ ' — 2y ' 2 sma(cos r-cos v
d = -- (1 cos a) (49)
2
2 b (cosy - cosq)| ^ ^.^ (50)
2Siaq(cosy- cosa)}
(cosy- cosa)} (i +« Jj
I A- 34
-------
(52)
cosy ~ 2 cos a )
A D
with b = -jZ—
r
As in the previous cases, the ratio of the volume of the oxidation region
to that of the reduction region increases with increasing reducer processing
capacity, becoming equal to unity when
b =
and hence
N =
•SETT -cosa
8 Vr2 sinotjl + S1ga (2-3 cos a)}
sma ' a
If one assumes, as before, a typical case with a value of expanded
melt bed depth to diameter ratio d/D = 0.40, and a value of €= 0 , equations
(48) through (52) become:
3.4058 NV-
(53)
1 + 0.36487(1 S1"y) 0.7149(1 cos7)
d = 0.4 D (54)
L = 0. 8569 DJ1 + 2 b (cosr- 0.2)1 (55)
3£\ j^L 2. 0017 (l + 0.36487 (1 si^y2y) - 0.7149 (1 cos7)} (56)
A^ 3 + 4b (cos7- 0.2)
8=0.4899 - _ - (57)
1 1.1977b
= 0.4899 2 - 4 - 6 -
i i / i- / i 58 8 58 \ /eo\
1 + 1. 6 b ( 1 -- ---- . . . ) (58)
6 27
A-35
-------
AD Fl D
with b = ^r^— = 1.410 -=r^- -=^r- for the case of the present analysis
with a 0. 33 mole fraction sulfate in the sulfur compounds in the melt leaving
the scrubber.*
8. Results
The results of the dimensional optimization of the reducer melt bed
for a vertical cylinder reducer geometry are plotted as a function of reducer
processing capacity in Figures 6 through 9.
Equations (27) through (29) were used to calculate the melt bed diameter,
expanded bed depth, and heat flux for various values of the normalized heat
loss parameter (Figures 6, 7, and 8). Since the heat flux is the real independent
variable rather than the heat loss parameter, this parameter was plotted in
Figure 9 (a cross-plot of Figure 8) as a function of reducer processing capacity
at various values of the heat flux.
Special emphasis was given to reducer processing capacities correspond-
ing to 3. 125 and 166.7 Ib moles/hr of SO in the feed gas entering the scrubber,
X,
meeting respective requirements of electric power generating capacities of 5
and 267 Mw (pilot plant size, and full scale size of one of three reducers of an
800 Mw plant, respectively).
Equations (23), (24), and (26) were used to calculate the melt bed diameter
and heat loss as a function of heat flux for the 5 Mw reducer with expanded melt
bed depths of 2 to 4 ft, and the 267 Mw reducer with expanded melt bed depths
of 6 to 10 ft. The results are presented in Figures 10 through 15 for the
reference case of a maximum superficial gas velocity of 3 ft/sec and a minimum
reduction region residence time of 15 minutes, and the two additional cases of
a superficial gas velocity of 5 ft/sec with a residence time of 15 minutes, and
a superficial gas velocity of 3 ft/sec with a residence time of 30 minutes.
The scope of the present analysis did not allow calculation of optimised
reducer dimensions as a function of reducer capacity for the horizontal? cylinder
reducer geometries. On the basis of Equations (26), (33), and (34), Figures 16
and 17 were plotted to show the reducer length and heat loss as a function of heat
flux for the special cases of the 5 and 267 Mw reducers with diameters ranging
* See footnote on page 31
A-36
-------
from 5 to 10 ft, and from 20 to 30 ft, respectively, with separation baffle
perpendicular to the axis of the cylinder.
A discussion of these results is provided in Sections IX. 3.b and
IX.3.c.
A-37
-------
EQUIVALENT Mw
I
W
00
267
300
15 m% SIN MELT
30 m% SIN MELT
NORMALIZED HEAT LOSS <103 Btu/hr PER Ib-m S0x/hr)
20
40
60
80
100
120
140
160
180
200
REDUCER CAPACITY (Ib-m S0x/hrl
220
Figure 6. Reducer Diameter as a Function of Prpcessing Capacity
-------
14
EQUIVALENT Mw
100
200
267
300
UO
sO
12 -
10
o
UJ
m
O
UJ
0
— _ —
15 m% SIN MELT
30 m% SIN MELT
NORMALIZED HEAT LOSS (103 Btu/hr PER Ib-m SOx/hr)
. ^__ _ . 2,000
__ 1.000
I
I
20
40
60
80 100 120
REDUCER CAPACITY (Ib-m 50,,/hr)
140
160
180
200
220
Figure 7. Reducer Bed Depth as a. Function of Processing Capacity
-------
EQUIVALENT Mw
200
267 300
15m%SIN MELT
30 m% SIN MELT
NORMALIZED HEAT LOSS (103 Btu/hr PER Ib-m 50,,/hr)
I I I I I I
200 220
REDUCER CAPACITY (Ib-m S0x/hr)
Figure 8, Reducer Wall Heat Flux as a Function of Processing Capacity
-------
1400
EQUIVALENT Mw
200
267
300
15 m% SIN MELT
30 m% SIN MELT
REDUCER WALL HEAT FLUX (103 Btu/ft2-hr)
20
80 100 120 140
REDUCER CAPACITY (Ib-m SOx/hr)
200
220
Figure 9. Normalized Heat Loss as a Function of Processing Capacity
(Cross-Plot of Figure 8)
-------
70
60
50
40
_ 30
CC
s
o
Ul
tc.
20
15
[S0|] + [SOJ]
= 0.33 AT SCRUBBER OUTLET
REDUCTION TIME - 15 min
SUPERFICIAL AIR VELOCITY = 3 ft/sec
GEOMETRY - VERTICAL CYLINDER
d = EXPANDED BED DEPTH
15 m% S IN MELT
30 m% SIN MELT
10
I
I
10
20 30 40
HEAT FLUX (103 Btu/ft2-hr)
50
60
70
Figure 10. Effect of Heat Flux on Reducer Diameter
A-42
-------
2000
1800 -
1600 -
1400 -
1200 -
1000 -
i
Cl
Ul
N
OC
O
0.33 AT SCRUBBER OUTLET
REDUCTION TIME = 15 min
SUPERFICIAL AIR VELOCITY = 3 ft/sec
GEOMETRY = VERTICAL CYLINDER
EXPANDED BED DEPTH = 3 AND 8 ft
15 m% SIN MELT
— — —30 m% S IN MELT
20 30 40
HEAT FLUX (103 Btu/ft2 hr)
Figure 11. Reducer Bed Heat Loss
A-43
-------
60
50
40
15
10
ISOj]
- = 0.33 AT SCRUBBER OUTLET
[S0§] + [SOp
REDUCTION TIME = 15 min
SUPERFICIAL AIR VELOCITY - 5 ft/sec
GEOMETRY - VERTICAL CYLINDER
d - EXPANDED BED DEPTH
15m%SINMELT
30 m% S IN MELT
267 Mw
5Mw
10
20 30 40
HEAT FLUX (103 Btu/ft2-hr)
50
60
70
Figure 12. Effect of Heat Flux on Reducer Diameter
A-44
-------
2000
1800
5-
x
E
1400
1200
§ 1000
UJ
X
a
UI
NJ
O
2
= 0.33 AT SCRUBBER OUTLET
REDUCTION TIME = 15 min
SUPERFICIAL AIR VELOCITY = 5 ft/sec
GEOMETRY = VERTICAL CYLINDER
EXPANDED BED DEPTH = 3 AND 8 ft
15 m% S IN MELT
30 m% S IN MELT
30 40 50
HEAT FLUX (103 Btu/ft2-hr)
Figure 13. Reducer Bed Heat Loss
A-45
-------
tr
LU
u
20 -
15
10
5Mw
[so;
[S0]
• 0.33 AT SCRUBBER OUTLET
REDUCTION TIME = 30 min
SUPERFICIAL AIR VELOCITY = 3 ft/sec
GEOMETRY = VERTICAL CYLINDER
d = EXPANDED BED DEPTH
15 m% SIN MELT
30 m% S IN MELT
10
20 30 40
HEAT FLUX (103 Btu/ft2-hr)
50
60
70
Figure 14. Effect of Heat Flux on Reducer Diameter
A-46
-------
2000
1800 -
= 0.33 AT SCRUBBER OUTLET
REDUCTION TIME = 30 min
SUPERFICIAL AIR VELOCITY = 3 ft/sec
GEOMETRY = VERTICAL CYLINDER
EXPANDED BED DEPTH = 3 AND 8 ft
15 m% S IN MELT
— 30 m% SIN MELT
30 40
HEAT FLUX (103 Btu/ft2-hr)
Figure 15. Reducer Bed Heat Loss
A-47
-------
= 0.33 AT SCRUBBER OUTLET
REDUCTION TIME = 15 min
SUPERFICIAL AIR VELOCITY = 3 ft/sec
GEOMETRY = HORIZONTAL CYLINDER
(BAFFLE 1 AXIS)
EXPANDED BED DEPTH = 0.4 d
15 m% SIN MELT
— — —30 m% S IN MELT
30 40
HEAT FLUX (103 Btu/ft2-hr)
60
70
Figure 16. Effect of Heat Flux on Reducer Length
A-48
-------
2000
1800 -
1600 -
1400 -
1200 -
1000 -
o
ui
N
-J 800 -
(E
O
BOO -
400 -
200 -
= 0.33 AT SCRUBBER OUTLET
[S0|] + (S0|)
REDUCTION TIME = 15 min
SUPERFICIAL AIR VELOCITY = 3 ft/sec
GEOMETRY = HORIZONTAL CYLINDER
(BAFFLE I AXIS)
EXPANDED BED HEIGHT = 0.4d
REDUCER DIAMETER = 7.5 AND 25 ft
15m%SIN MELT
—30 m% S IN MELT
HEAT FLUX (103 Btu/ft2-hr)
Figure 17. Reducer Bed Heat Loss
A-49
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VII. REDUCER HEAT LOSS
The previous sections of this report have shown that the coke and air
consumption of the reducer and its physical dimensions must be determined
on the basis of a heat balance around the reducer melt bed. This heat balance
was performed and the key parameters were defined as a function of the inde-
pendent variable parameter, QT , the normalized heat loss from the melt bed,
or of the heat flux, (Q/A), at the walls of the reducer vessel in contact with
the melt, which is the truly independent variable determining the magnitude
of QT . The question which must therefore be answered is "What is the
magnitude of the heat flux, (Q/A), and how can it be minimized? "
The heat flux at the walls of the reducer vessel is determined by the
temperature driving force between the melt and whatever heat sink is pro-
vided, and by the thermal resistance between the melt and this heat sink.
Both of these can be controlled by the selection of the design concept. As
stated in Section IV, the two concepts evaluated in this study are the "frozen
melt skull (cold wall)" and the "alumina-liner (hot wall)" concepts. This
section of the report provides an estimate of the heat fluxes which one may
expect to encounter at the reducer vessel walls in contact with the melt in
these two concepts.
1. Estimate of Melt Heat Transfer Coefficient
The heat transfer coefficient across the melt film in contact with the
surface of the walls of the reducer vessel controls the heat transfer rate
from the melt to the outside of the vessel in the frozen skull design concept.
In this concept the temperature of the skull in contact with the melt is fixed
(at the freezing point temperature of the melt or approximately 750 °F) and
therefore the temperature driving force is fixed. The heat flux at the reducer
wall in the frozen skull concept is:
(Q/A) = h (Tmelt TMp ) = h (1550-750) = 800 h (59)
An accurate knowledge of the expected, magnitude of the melt film heat
transfer coefficient is therefore essential to the evaluation of the frozen
skull reducer design concept.
A-50
-------
In the alumina-liner reducer design concept on the other hand, the
resistance of the film of melt represents only a small part of the overall
resistance to the heat transfer across the walls of the reducer vessel. There
is therefore little need for an accurate knowledge of the film heat transfer
coefficient in the evaluation of this concept.
The liquid film at the walls of the reducer vessel consists of the
carbonate-sulfate-sulfide process melt, containing suspended particles of
partially reacted coke. This melt is kept in a highly turbulent state by the
flow of air in the oxidation region and by the flow of gases generated in the
reduction region. The melt thus represents a multiphase environment with
complex circulation patterns for which reliable heat transfer coefficient
values can only be obtained experimentally.
Using many simplifying assumptions an attempt was made to obtain
an estimate of the magnitude of the melt heat transfer coefficient which one
might expect to prevail at the walls of the reducer vessel. For this purpose
the film was assumed to consist of a single liquid phase with the physical
properties of the carbonate melt at the prevailing temperatures. Preliminary
calculations showed that at the low net flow velocities and the relatively high
temperature differences encountered in the reducer, natural convection
would provide the main contribution to the heat transfer at the walls. The
estimate of the heat transfer coefficient was therefore based on McAdams'
recommended correlation for natural convection heat transfer for vertical
IQ \
planes and vertical cylinders in the turbulent range:
it OA, 1/3
(60)
h = 0. 13 k (g —^Tk/SAt) ' = 0.13k
1/3 / u± y_t\ 1/3
The physical properties used were those of the carbonate melt at the
average temperature of the film (approximately 1150°F in the case of the
frozen skull reducer design and 1400 °F in the case of the alumina-liner
reducer design).
With these assumptions, the correlation yielded heat transfer coefficient
i of 317 and 303 Btu/ft2 h
design concepts, respectively.
2
values of 317 and 303 Btu/ft hr °F for the frozen skull and the alumina-liner
A-51
-------
Obviously many factors which may have significant effects upon the
magnitude of these heat transfer coefficients were neglected in such a
calculation. These include the increased turbulence attributable to the flow
of gases and to the suspended particulates, the increased effective viscosity
due to the presence of the particulates, and the possible effect of gas phase
heat transfer. In addition, there are uncertainties in our knowledge of the
properties of the melt at the prevailing temperatures and in the correctness
of the use of these properties in the correlation at an arithmetic average
temperature between that of the bulk melt and that of the walls. It must be
noted, however, that even a 10-fold increase in the viscosity values used in
the calculations would still yield the relatively high heat transfer coefficient
values of 117 and 112 Btu/ft2 hr °F, respectively.
From the practical point of view, a melt heat transfer coefficient in
excess of 50 Btu/ft hr °F yields a reducer wall heat flux in excess of
40,000 Btu/ft hr in the frozen skull design concept (per Equation (59)).
Figures 4 through 17 clearly show that a heat flux of this magnitude results
in an unacceptably high heat loss from the reducer, with consequent excessive
coke and. air consumption and large vessel dimensions. The present analysis of
the frozen skull design concept was therefore based on a range of melt heat
transfer coefficients from 10 to 70 Btu/ft2 hr °F, which is realistic only if
actual melt heat transfer coefficients are found to be lower than predicted
on the basis of the simplifying assumptions made above.
In the case of the alumina -liner design concept, a value of 50 Btu/ft
hr °F was assumed for the melt heat transfer coefficient. As stated before,
this represents only a small fraction of the total heat transfer resistance in
this concept and therefore has little effect on its analysis and evaluation.
2. Reducer Wall Heat Transfer in Frozen Melt Skull Concept
The heat transfer rate across the walls of the reducer vessel in the
frozen skull design concept can be expressed as follows:
(Q/A) = h (T -T ) = _ heat sink (61)
IU/A) - n Umelt -L; V '
where t is the thickness and k , ,, the thermal conductivity of the skull,
S kUli S KLUll
and U is the overall thermal conductance from the outer (cold) surface of the
skull to'the cooling medium of the heat sink.
A-52
-------
Equation (61) gives:
. . f TMP Theat sink 1 1
skull skull 1~hTf~T . „)U
1 melt skull '
- ".tail (f - ¥> <62)
with „ _
MP " heat sink ,, ,,
r = rp ' T (63)
melt " MP
Figure 18 presents a plot of frozen skull thickness as a function of melt
film resistance for various values of the temperature parameter, r, and of the
thermal conductance to the heat sink, U (the appropriate equivalent skull thick-
ness, k , ,,/U, must be subtracted from the skull thickness shown on the
' skull
ordinate for U—»-oo). The thermal conductivity of the skull was assumed to
be 0. 5 Btu/ft hr °F.
With a skull freezing temperature of approximately 750 °F and cooling
with boiling water at 50 and 600 psia, the parameter r takes values of 0. 59
and 0. 33, respectively. If one assumes a melt heat transfer coefficient of
50 Btu/ft2 hr °F and a thermal conductance from the skull to the heat sink
2
of 200 Btu/ft hr °F, the thickness of the skull becomes 40 and 10 mils,
respectively. No skull could be maintained if the melt heat transfer co-
2
efficient were to be greater than 118 Btu/ft hr °F with 50 psia cooling water
and 67 Btu/ft hr °F with 600 psia cooling water.
The temperature parameter, r, is relatively small in the molten
carbonate process which uses the low melting eutectic as the molten salt
mixture. Other processes may be able to use a higher melting point melt,
such as sodium carbonate, with resulting higher values of r. These higher
values of r would make it considerably easier to maintain a protective frozen
skull of appreciable thickness on the walls of the reducer vessel. A typical
(9)
value of r in a Kraft furnace, for instance, might be about 1. 5, which for the
above-considered conditions of h = 50 and U = 200 Btu/ft2 hr °F would yield
a skull thickness of 150 mils with 600 psia cooling water.
A-53
-------
200
100
67
50
40
h (Btu/ft2-°F-hr)
33 25
20
200
180 —
0.59,
«• U =50 Btu/ft2-°F-hr
I
u/ft2-°D-hr
HEAT TRANSFER COEFFICIENTS
h = FROM MELT TO SKULL
U = FROM SKULL TO WATER
I
0.01 0.02 0.03 0.04 0.05
MELT FILM RESISTANCE (ft2-°F-hr/Btu)
0.06
0.07
Figure 18. Effect of Parameters h, U, and r
on Frozen Skull Thickness
A-54
-------
Figure 19 provides a plot of heat flux and frozen skull thickness as a
function of melt heat transfer coefficient in the range of 0 to 70 Btu/ft2 hr °F.
It clearly shows how rapidly the heat flux and therefore heat loss increases
and the skull thickness decreases with increasing magnitude of the melt heat
transfer coefficient.
3. Reducer Wall Heat Transfer in Alumina Liner Concept
In the alumina liner reducer design concept the temperature of the
walls in contact with the melt is no longer fixed as it was in the frozen skull
concept. The only fixed temperature is that of whatever heat sink is used to
absorb the heat loss from the reducer. The heat transfer rate equation across
the reducer walls can therefore be written:
(Q/A) = ^elt Theatsinki (64)
h"+ (~) . . + ~U~
alumina
where t , . is the thickness and k , . the thermal conductivity of the
alumina alumina
alumina lining, and U is the overall thermal conductance from the outer (cold)
surface of the alumina to the cooling medium of the heat sink.
In this concept it is highly desirable and possibly essential that liquid
melt not be allowed to penetrate all the way through the alumina liner to the
structural steel of the reducer vessel. The outer (cold) surface of the alumina
liner must therefore be kept below the 750 °F freezing point temperature of
the melt.
To give visibility to this requirement, the overall thermal resistance
across the walls of the reducer was divided into two series components, the
thermal resistance from the melt to the outer surface of the alumina, and
the thermal resistance from the outer surface of the alumina to the heat sink.
For this purpose Equation (64) can be rewritten in the form:
,„,., Tmelt outer alumina _ T , T > „_,
(Q/A) = —j- -j- - U
-------
400
BASIS: AVERAGE MELT TEMPERATURE = 1550°F
SKULL SURFACE TEMPERATURE = 750°F
THERMAL CONDUCTIVITY OF SKULL = 0.5 Btu/ft-hr-"F
HEAT CONDUCTANCE FROM SKULL TO WATER = U
SKULL THICKNESS CORRECTION
U = 400 8tu/ft2-hr-°F ADD 15 mils
U = 200 Btu/ft2-hr-°F ADD 0 mils
U - 100 Btu/ft2-hr-°F SUBTRACT 30 mils
U = 50 Btu/ft2-hr-°F SUBTRACT 90 mils
281"FH,O(50psia)
486"F H20 (600 psia)
20 30 40 50
MELT HE AT TRANSFER COEFFICIENT (Btu/ft2-°F-hr)
60
70
Figure 19. Frozen Skull Reducer Concept — Heat Flux
and Skull Thickness
A-56
-------
16
14
12
m
"o
10
BASIS: AVERAGE MELT TEMPERATURE = 1560°F
MELT HEAT TRANSFER COEFFICIENT - 50 Btu/ft2-°F-hr
HEAT CONDUCTANCE FROM ALUMINA TO WATER = U
•• K = 5 Btu/ft-°F-hr
— — — K = 2 Btu/ft-°F-hr
NATURAL
CONVECTION
U, Btu/ft2-°F-hr
200 100 200
100
281°F H20
(50 psia)
200
400 600 800 1000
ALUMINA OUTER SURFACE TEMPERATURE (°F)
1200
1400
Figure 20. Alumina-Liner Reducer Concept — Heat Transmission
and Rejection Capability
A-57
-------
transmitted through various alumina linings. The straight lines with the
positive slope and the curve show the heat flux which can be rejected from
the outside of the lining to various cooling systems. The intersections of
the negative slope lines with the positive slope lines and curve provide
solutions to the overall heat transfer rate Equation (65) from the melt through
the walls of the reducer vessel to an external heat sink.
Figure 20 shows that with natural convection air cooling, approximately
12 inches of high thermal conductivity alumina, such as Monofrax A, are re-
quired to keep the temperature of the outer surface of the alumina below 750 °F
(a major part of the alumina not in direct contact with the melt could be re-
placed with a thermally equivalent amount of higher porosity but lower thermal
conductivity alumina such as alundum). The corresponding heat flux at the
walls of the reducer vessel in contact with the melt is less than 4000 Btu/ft^ hr.
The temperature at the outside surface of the steel reducer vessel is below
750 °F and can be lowered to below 650 °F by a slight increase in thickness of
the lining, thus possibly allowing the use of carbon steel for the reducer vessel.
Natural convection air cooling requires a high temperature at the outside
surface of the steel reducer vessel and thus does not allow installation of
thermal insulation on this surface. If it should prove desirable that the outside
temperature of the reducer be below 200 °F, water cooling of the outside sur-
face of the alumina liner, or of the steel surrounding it, is required. Thermal
insulation can then be installed around the outside of the cooling pipes to
lower the outside temperature of the reducer to whatever temperature is
desired. The heat flux at the walls of the reducer in contact with the melt
will be around 5000 Btu/ft hr if one uses the equivalent of a 12 inch high
thermal conductivity alumina lining. The thickness of lining required can
be decreased in this case, but at the penalty of increasing the heat flux and
thus the heat loss from the reducer.
In view of the major incentive to minimize heat loss from the melt, the
actual design of the reducer will involve the use of a combination of liner
materials between the melt and the steel shell of the vessel. Low porosity
high thermal conductivity alumina will provide the inner lining in contact with
the melt. It will in turn be surrounded by one or several layers of more
porous alumina with better thermal insulation properties. Figure 20 shows
that a reduction of the heat flux at the walls to below 2000 Btu/ft hr is thus
readily achievable with either natural convection air cooling or water cooling
of the outside walls of the reducer vessel.
A-58
-------
VIII. INTERNAL RECIRCULATION BETWEEN OXIDATION AND
REDUCTION REGIONS
The two-region reducer concept is based on the use of melt recirculation
between the oxidation and reduction regions to transport the heat generated in
the oxidation region into the reduction region. This recirculation must be
fast enough on the one hand to transport the required amount of heat without
exceeding the specified temperature differential between the top and the
bottom of the melt bed. On the other hand it must be slow enough to allow
for the required minimum residence time of the melt in the reduction region
and thus permit the reduction reaction to proceed to the required degree of
completion.
Internal melt recirculation takes place as a result of the difference in
effective average densities of the melt in the oxidation and reduction regions.
The melt density in the oxidation region is lower than that in the reduction
region because:
1) The greater superficial gas velocity in the oxidation region
results in a greater bed expansion in that region than in the reduction region.
2) The anticipated non-linearity with distance of heat generation
in the oxidation region and heat consumption in the reduction region results
in the major part of the melt temperature rise taking place in the lower section
of the oxidation region and the major part of the melt temperature drop occurring
in the upper section of the reduction region, thus yielding an effective difference
in average melt temperature between the two regions and a consequent difference
in average melt density.
The density difference of the melt provides the driving force necessary
to establish a natural circulation flow pattern between the oxidation and
reduction regions of the reducer. Opposing this driving force Is the pressure
drop resulting from the circulation thus established. The actual internal
circulation rate is determined by the equalization of these two forces. Since
the main resistance to the flow is provided by the baffle openings connecting
the oxidation and reduction regions at the top and bottom of the melt bed, these
openings will control the melt recirculation rate and must therefore be sized
to provide the rate required to allow transport of the heat and adequate
residence time for completion of the reduction reaction.
A-59
-------
1. Natural Convection Driving Force
The natural convection driving force is generated by the differences in
effective average melt densities between the oxidation and reduction regions
of the reducer. Experimental information is presently not available on either
the difference in bed expansion which one might expect to prevail, or the non-
linearity with distance of the chemical reactions taking place which would
establish the average temperature differential between the two regions. The
analysis was therefore based on the following somewhat arbitrary assumptions:
1) The net effect of the difference in superficial gas velocity on
melt bed expansion and therefore melt density was assumed to be approximately
5 percent. A melt bed expansion by a factor of 1. 75 was assumed for the
oxidation region and a factor of 1. 75/1. 05 = 1. 67 for the reduction region.
2) The average temperature differential between the oxidation and
reduction regions was assumed to be approximately one -third of the temperature
rise from the bottom to the top of the melt bed.
The natural convection driving force, A P, can be expressed as:
AP = dAp (66)
with:
. p
r T
where p and p are the effective average densities of the melt in the reduction
and oxidation regions, respectively, and P^ and P are the unexpanded melt
ir -"-o
densities at the average temperatures prevailing in these regions. It is to be
noted that of the two terms on the right hand side of Equation (67) the first
represents the effect attributable to different gas flow rates and the second the
effect attributable to the prevailing temperature differential.
Using the melt density data of Section III, and assuming an average
melt temperature of 1500 °F in the reduction region and an overall temperature
rise AT from the bottom to the top of the melt bed:
A-60
-------
A 0.05 no ,, , 18.87 x 10"3 AT
&p- li?5 110.61 + p-yg -r
= 3. 160 (1 + 1. 137 x 10"3 AT), lb/ft3 (68)
for AT = 300 °F:
= 3. 160 + 1. 078 = 4. 238 lb/ft3 (69)
Equations (68) and (69) show that, under the conditions of the specific
assumptions made above, the effect of differences in gas flow rates is con-
siderably more important than the effect of temperature differences between
the oxidation and the reduction regions.
The natural convection driving force is therefore:
AP = 3. 160 d (1 + 1. 137 x 10"3 AT), lb/ft2 (70)
2. Pressure Drop
Internal recirculation of the melt between the oxidation and reduction
regions of the reducer results in a pressure drop across the openings
(orifices) in the baffle between these two regions at the top and bottom of the
melt bexi, plus an additional pressure drop for the flow through each region.
The pressure drop across the orifices can be expected to form the major
part of the total pressure drop. For simplifying purposes one can write:
AP = qAPor (71)
where APor is the pressure drop across both orifices and q is a proportionality
factor somewhat larger than unity-.
The orifice equation can be written:
Lt „_ f i vr. £•
(72)
where V is the volumetric flow rate, AQr the orifice area, c the orifice
discharge coefficient, and the subscripts 1 and 2 refer to each of the two
orifices.
A-61
-------
Assuming both orifices to have the same area and approximately the
same volumetric flow, and using an average value of the melt density,
Equation (72) becomes:
V 2
Therefore:
AP =i£. _v (?4)
g \ or;
V can be expressed as a function of the heat transported from the
oxidation to the reduction region and the temperature difference between the
top and bottom of the,-melt:
NQT
V = —
where Q_, is the normalized heat transport, as given by Equations (9) or
(10).
q NQT j 2
p or
Arbitrarily assuming that the unaccounted-for pressure drop may
represent about 20% of the pressure drop across the orifices results in an
approximate value of q = 1. 20. Substitution in Equation (76) of the appropriate
average values of p and c and a typical orifice discharge coefficient of 0.61
yields:
1Q NQ 2
AP = 6. 127 x 10 U (-T 4=-) , Ib/ft ('
-Ti tJ X
or
This is the pressure drop resulting from the required melt re-
circulation between the oxidation and reduction regions of the reducer.
The value of Q_ can be obtained from equations (9) or (10).
3. Orifice Size
From equations (70) and (77):
NQiT
c
A = 2.475x10
or
= 1.392* ID
c; NQ_
"5 - T
AT-/ d (1 + 1. 137 x 10'^ AT)
A- 62
-------
Substitution of the value of Q from equation (10) yields:
1.2 F7
or AT/I + 1. 137 x 10'3 AT
For AT = 300 °F :
A = 0.003453 F, Jk ft2 (80)
or
V a
Typically the openings in the baffle at the top and bottom of the melt
between the oxidation and reduction regions may be rectangular, with a length
possibly equal to one -quarter of the internal diameter, D, of the reducer.
Under these conditions the width, w, of these openings would be:
w
= 0. 1657 F, -~= , inches <81)
,
D/d
Figure 21 presents a plot of orifice width as a function of reducer
capacity, based on the use of Equation (81) with the values of D and d obtained
in Section VI (Figures 6 and 7). In view of the several rather arbitrary
assumptions made it cannot be considered to represent anything more than
trends and a very rough approximation of the actual size of the orifices
required. Nevertheless it does show that no difficulty should be expected
in obtaining adequate recirculation of the melt and that the orifices must be
rather small even in large reducer units to control the recirculation rate
to provide the minimum melt residence time in the reduction region required
for completion of the reduction reaction.
A-63
-------
EQUIVALENT Mw
100
300
1000
NORMALIZED HEAT LOSS (103 Btu/hr PER Ib-m SO IM
BASIS: (OXIDATION fl EG ION/R EDUCTION REGION)
BED EXPANSION = 1.05
AVERAGE TEMPERATURE DIFFERENCE = 100°F
TOTAL AP/AP ACROSS TWO ORIFICES - 1.20
ORIFICE LENGTH/REDUCER ID =0.25
15 m% SIN MELT
____30M%SIN MELT
80 100 120
REDUCER CAPACITY (Ib-m S0x/hr|
140
160
180
200
Figure 21. Reducer Baffle Orifice Width as a Function of Processing Capacity
-------
IX. DISCUSSION AND CONCLUSIONS
1. Basis of Analysis
A meaningful discussion of the results obtained in this study of the
molten carbonate reducer requires a review of the technical approach and
parameters on which it has been based. A diagrammatic flow-sheet of the
steps involved in the procedure used to size the reducer and determine its
coke and air requirements is shown in Figure 22. This figure also lists
the independent variables considered, the system constraints involved, and
the results obtained.
a. Technical Approach
Starting from a mass balance around the reducer, a heat balance (1)
yields the coke (2) and air (3) requirements. The air requirement combined
with a superficial air velocity limitation determines the cross-sectional area
of the oxidation region of the reducer (4). The air requirement also determines
the amount of heat which must be transported from the oxidation region to the
reduction region of the reducer and therefore, in combination with the
temperature rise from the bottom to the top of the melt bed, determines
the melt recirculation rate required to transport this heat (5). The melt
recirculation rate, and the required minimum melt residence time in the
reduction region determine the required volume of the melt bed in this
region (6). This volume in conjunction with an assumed bed depth yields
the cross-sectional area of the reduction region (7). Adding the cross-
sectional areas of both regions gives the normalized total cross-sectional
area of the reducer per unit SOx input into the molten carbonate system (8).
The cross-sectional area for the actual SO input is then calculated, yielding
X,
the reducer diameter (9). This in turn determines the area for heat transfer
from the melt bed to the walls of the reducer in contact with the melt (10).
This area is minimized, thus yielding an optimized relationship between the
melt bed diameter and depth. Finally the value of the heat flux at the re-
ducer walls in contact with the melt is obtained and matched with the actual
design heat transfer rate for a given reducer design concept (11). The
A-65
-------
[S]l
(1) ( Heat
Requirement
[S0=]
V
(2) .-, (3) .^
o 2'
AT]
(4)
r *\
O
(5) Melt 1
— -*• Recirculation
Rate
9
•
(6) 1
r
d
(7)
r
12J-A1
t
1M
(9)
-D
(10) A ,
Q
L (") , /^Q
/
\A
N
Independent Variables
System Constraints
N
QL
[S]
V
9
AT
= 0.15 + 0.30
= 0.33
= 3 ft/sec
= 15 min
= 300°F
Results
C (coke) Figure 4
O (air) Figure 5
D Figure 6
d Figure 7
8 + 9
Figure 22. Reducer Analysis Schematic Diagram
-------
optimization shown by a set of double arrows in Figure 22 is a heat loss
minimization and therefore does not take into account any other considerations
which may influence the overall optimization of the reducer design.
b. Independent Variables
The key independent variable, aside from the obvious and trivial one
of processing throughput, N, is the heat loss from the melt bed to the walls
of the reducer vessel in contact with the melt, expressed in this analysis
by its normalized value, QT . This is as might be expected since the molten
JL<
carbonate reducer is basically a furnace designed to provide the heat and
space requirements for the reduction of the alkali metal sulfates dissolved in
the carbonate melt. The net heat generated by the oxidation of the excess coke
fed into the reducer provides the heat required for the reduction of the sulfates
in the process melt, the heat required for preheating the feed streams to the
operating temperature of the reducer, and the heat losses to the walls of the
vessel in contact with the melt. Once the processing capacity of the reducer
has been specified, the reduction and preheat requirements are established,
and the only variable which can be controlled by the design of the unit is the
heat loss to the walls. This heat loss is determined by the heat flux at the
walls of the vessel in contact with the melt. The heat flux is therefore the
real independent variable. It is controlled by the nature of the heat sink and
the thermal resistances at and through the walls, and therefore by the design
concept and physical dimensions involved. In this analysis the normalized
heat loss was varied from 0 to 2 x 106 Btu/hr per Ib mole/hr SOX, with lower
values generally achievable in large capacity reducers and higher values
prevailing in small capacity units.
c. System Constraints
As shown in Figure 22 the major constraints considered in the present
study were the concentration of sulfur compounds in the molten carbonate melt,
[s] , the fraction of these sulfur compounds present in the sulfate state in the
melt leaving the scrubber, [so|]= 1-a, the superficial air velocity in the
oxidation region region, v, the melt residence time in the reduction region,9,
and the temperature differential between the top and the bottom of the melt
bed, AT. Some of the other constraints listed in Section III include the
composition of the coke, and the extent of initial preheat of the air feed into
the reducer, the extent of sulfite disproportionate upstream of the reducer,
the extent of completion of the chemical reactions involved, and the geometry
of the reducer vessel. A-67
-------
The concentration of sulfur compounds in the carbonate melt,[Sj, should
be as high as possible to minimize the melt flow requirement, with consequent
minimization of equipment and piping sizes throughout the system. With respect
to the reducer, the main effect of melt flow rate is upon the melt preheating
requirement and therefore the overall heat requirement. Another effect is on
the contribution of the net melt flow through the reducer to the total flow rate
of melt through the reduction region and therefore upon the volume of the bed
required to meet the reduction reaction residence time requirement.
The maximum allowable concentration of sulfur compounds in the melt
is determined by the solubility of the least soluble of these compounds at the
operating temperatures of the various steps of the process. The controlling
solubility appears to be that of the alkali metal sulfides in the melt to be
cooled in the quench tank upon discharge from the reducer (the possibility of
the presence of these sulfides in the actual form of thiocarbonates may provide
some relief from this constraint). It is expected that a 30 mole per cent
concentration will be acceptable, but the possibility exists that the maximum
allowable concentration may be as low as 15 mole per cent. The reducer
analysis, therefore, was performed for both cases of 30 and 15 mole per cent
sulfur compounds in the melt.
The oxidation state of the sulfur compounds in the melt leaving the scrubber
is determined by the ratio of 803 to SO2 in. the feed gas to the scrubber, the
amount of melt recycle directly back to the scrubber, the amount of sulfate
returned to the scrubber from the reduction-regeneration system, and the
extent of sulfur compound oxidation taking place in the scrubber itself. The
greater the extent of oxidation of the sulfur compounds, the greater will be
the amount of coke required for their reduction as well as the amount of coke
required to supply the heat for this reduction. ' For purposes of the present
study, the effect of the oxidation state of the sulfur compounds in the melt
leaving the scrubber upon the coke and air requirements of the reducer has
been shown as a function of the variable parameter [§04 J. In the remainder
of the analysis, the scope of the work limited the effort to the case of a single
value of the sulfate fraction, 33 mole per cent, as a reasonable approximation
to most of the conditions encountered in actual practice.
A-68
-------
For a given air feed rate requirement the maximum allowable superficial
gas velocity determines the minimum required cross-sectional area of the
oxidation region of the reducer. As the gas velocity is increased the expansion
of the melt bed increases. On the basis of the experimental data presented in
Reference (6) on the expansion of beds with a quiescent height in excess of one
foot, a value of 3 ft/sec was selected for the maximum allowable superficial
air velocity, v, in the oxidation region. The corresponding total expansion
of the melt bed was assumed to be approximately 75 per cent. Most of the
analysis was done at this velocity. The possible improvement to be achieved
if the superficial air velocity could be increased to 5 ft/sec was evaluated for
the special cases of a 5 Mw and a 267 Mw reducer capacity.
For a given melt throughput through the reduction region of the reducer
the volume of this region is determined by the minimum residence time re-
quired to allow the reduction reaction to proceed to the desired degree of
completion. The data of Reference (7) indicate that at temperatures above
1472 °F (800 °C) the reaction time required to achieve 95 per cent reduction
is less than half an hour. The present study has been based on a minimum
reduction region residence time, Q , of 15 minutes since the reaction tempera-
ture ranges from 1700 to 1400 °F. The possible penalty which might result if
the residence time had to be increased to 30 minutes was evaluated for the
special cases of a 5 Mw and a 267 Mw reducer capacity.
Another parameter which may be considered either as a constraint or
an independent variable is the temperature differential, AT, between the top
and bottom of the reducer melt bed. The larger the temperature differential,
the smaller will be the internal melt recirculation requirement, and therefore
the volume of the reduction region melt bed required to satisfy the minimum
residence time requirement. The temperature at the top of the bed, however,
must be kept as low as possible to minimize the potential problems associated
with the structural materials of the reducer and the decomposition of the
carbonate salts. The temperature at the bottom of the bed must be high enough
to allow the reduction reaction to proceed at a reasonable rate. In this analysis
the top and bottom temperatures of the melt bed were selected as 1700 and
1400 °F, respectively, with a consequent temperature differential of 300 °F,
considered as a fixed constraint throughout the study.
A-69
-------
2. Limitations of Present Study
The reducer analysis performed in this study is believed to provide a
reasonable approximation to the actual process performance of a two-region
molten carbonate reducer, and to the effect of key parameters upon this per-
formance. It also provides a starting baseline for the more detailed process
analysis to be performed in conjunction with the actual design of such a
reducer and the evaluation of performance data to be developed from future
experimental work.
It must be emphasized, however, that this analysis has of necessity had
to be based on the limited experimental information available at this time.
The following qualifications must be noted:
a) The present study is concerned only with the process engineering
aspects of the reducer design. It does not take into account mechanical and
structural problems, such as method and location of process stream intro-
duction, distribution, and discharge, minimization of melt carry-over,
installation of vessel lining and baffle.
b) The analysis assumes perfect separation of the oxidation and re-
duction functions. It assumes that all the direct air oxidation takes place
through the oxidation of sulfide to sulfate and that the oxidation of the coke
occurs only as a result of reduction of sulfate to sulfide. This entails trans-
port of a larger amount of heat between the two regions of the reducer than
would otherwise be necessary. The analysis also ignores any potential contri-
bution of oxidation of the hydrogen in the coke to the heat balance around the
melt bed, and any lack of contribution or other effects associated with the
volatile matter in the coke.
c) The analysis assumes idealized component and flow distributions in
the melt bed and thus ignores any potential problems associated with poor flow
i
distribution, differences in the densities of the various components of the melt
(feed melt, reduced melt, coke, partially oxidized coke, ash), and differences
in particulate sizes.
d) The analysis is based on the limited information presently available
on the physical property and thermodynamic data of the process materials in-
volved, and the heat transfer and hydraulic behavior of the three phase system
existing in the molten carbonate reducer.
A-70
-------
3,
Results
a> Coke and Air Requirements
Figures 4 and 5*show the normalized coke and air requirements of the
two region molten carbonate reducer as a function of normalized heat loss,
at sulfur compound mole fractions in tne melt of 0. 15 and 0. 30, and sulfate
mole fractions of the sulfur compounds in the melt leaving the scrubber of
0, 0.33, and 1.
Table I presents typical values of the normalized coke and air require-
ments at the reference sulfate mole fraction of 0. 33:
TABLE I
MOLTEN CARBONATE REDUCER COKE AND AIR REQUIREMENTS
Sulfur
Compound
Concentra-
tion, [s]
0. 15
ii
11
it
ii
0.30
ii
ii
ii
ii
Normalized
Heat Loss,
QL,
103 Btu/hr
per Ib mole/
hrSOx
0
50
100
250
1,000
0
50
100
250
1,000
Normalized
Coke Reqmt,
Ib per Ib Sulfur
1.64
1.82
1.99
2. 52
5. 14
1.30
1.48
1.65
2. 17
4.79
Normalized
Air Reqmt,
scf per Ib Sulfur
118
141
163
232
572
73
96
119
187
527
Coke
Cost,
mills / kwhr
0.180
0.200
0.218
0.276
0.566
0.143
0.164
0.182
0.239
0.526
Air
Compressor
Power Reqmt
HP/Mw
1.9
2.3
2.6
3.7
9.2
1.2
1.5
1.9
3.0
8.4
It is to be noted that the conversion factors used to obtain the right hand
ordinates of Figures 4 and 5 and therefore the last two columns of Table I are
based on the use of a 12, 800 Btu/lb, 3 wt % sulfur coal in a power plant with a
heat rate of 9000 Btu/Kwh.
The coke and air requirements increase linearly with increasing heat loss.
An increase in heat loss of 100, 000 Btu/hr per Ib mole SOx/hr results in coke
air requirement increases of 0. 35 Ib and 45 scf per Ib of sulfur, respectively.
*Pages 22 and 23
A-71
-------
Similarly decreasing the sulfur compound concentration in the melt from
a mole fraction of 0. 30 to a mole fraction of 0. 15 increases the coke and air "
requirements by 0. 34 Ib and 45 scf per Ib of sulfur, respectively.
The extent of oxidation of the sulfur compounds in the melt leaving the
scrubber can be seen from Figures 4 and 5 to result in an increase in coke and
air requirements of 0.39 Ib and 21 scf per Ib of sulfur, respectively, when the
sulfate fraction increases from 0 to 1.
If, as is most probable, the oxidation of the coke in the reducer is less
than 100 per cent complete, some unreacted coke will be discharged to the
quench tank with the reduced melt and ash. Filtration of the melt will then
entail the loss of salt associated with a filter cake consisting not only of ash
but also of coke, thus requiring increased recovery plus make-up of molten
carbonate. For a constant fractional coke usage an increase in coke require-
ment therefore also results in an increase in carbonate recovery and make-up,
with an economic penalty directly proportional to the increase in coke requirement,
If one assumes 95 per cent utilization of the coke, 0. 7 per cent ash in the
coke, and a 1:1 ratio of melt to coke plus ash in the filter cake, the resulting
carbonate recovery and make-up requirement amounts to 113 Ib per ton of
coke fed into the reducer. At an average carbonate recovery plus make-up
cost of 8 cents per Ib this would add $9 per ton of coke fed into the reducer
to the cost of the coke itself, thereby effectively almost doubling the cost
directly associated with the coke requirement (assuming a coke cost of $11
per ton). Major emphasis must therefore be placed on achieving the greatest
possible utilization of the coke, possibly through recycle of part of the filter
cake back into the reducer (this may be limited by the resulting ash build-up
in the melt).
It is therefore essential that:
(1) The heat loss from the reducer melt bed be minimized
(2) The sulfur compound concentration in the melt be as high as
possible without exceeding their solubility limits.
(3) The oxidation of the sulfur compounds going through the scrubber
be minimized. The two parameters which may be most controllable in this
respect are the direct recycle of melt to the scrubber and the return to the
scrubber of unreduced sulfate and unregenerated sulfide (which becomes
oxidized to sulfate) from the reduction-regeneration system.
A-72
-------
(4) The coke utilization in the reducer be maximized. The fraction
of unreacted coke can be minimized through good distribution of the coke
in the melt and possibly partial recycle of the coke-ash-salt filter cake
back to the reducer.
b. Physical Dimensions
Figures 6 through 8*show the melt bed diameter and expanded depth, and
the heat flux at the walls of the reducer vessel in contact with the melt, as a
function of processing capacity and normalized heat loss. Since the real inde-
pendent variable is the heat flux rather than the normalized heat loss, Figure 9*
presents a cross-plot of Figure 8 showing the normalized heat loss as a function
of processing capacity and heat flux. Figures 6, 7 and 9 therefore provide the
means for obtaining the physical dimensions of the reducer melt bed as a function
of processing capacity and heat flux.
Table II presents typical values of the normalized heat loss, melt bed
diameter and expanded depth, at reducer processing capacities of 5, 40, 60,
and 200 Ib moles/hr SO , corresponding to equivalent nominal power plant
capacities of 8, 64, 96 and 320 Mw, respectively.
TABLE II
MOLTEN CARBONATE REDUCER HEAT LOSS AND DIMENSIONS
Reducer Capacity,
Ib m/hr SOX
Reducer Capacity
Equiv. Mw
i
0.15
n
11
ii
n
0.30
it
H
it
n
, Q/A
103 Btu/ftz
hr
0
10
20
30
40
0
10
20
30
40
1
5 40
8 - 64
&L
103 Btu/hr
per Ib m/hr
SOX
0
240-120
640-290
>1000-550
> 1000-980
0
170-90
540-210
>1000-410
>1000-770
D,
ft
6.5-15
8-17
10-19
>ll-23
>ll-28
5.5-12
7-14
9-16
>11-19
>ll-23
d,
ft
2.2-4.2
2.5-4.3
2.7-4.5
>2. 9-4. 7
>2. 9-5.0
2.0-3.6
2. 1-3.7
2.4-3.9
>2. 8-4. 2
>2.8-4. 5
60 - 200
96 - 320
QL,
103 Btu/hr
per Ib m/hr
SOX
0
110-70
250-170
470-310
830-500
0
80-50
180-120
350-220
620-380
D,
ft
18-29
20-32
23-35
26-39
30-44
15-24
17-26
19-29
22-33
26-37
d,
ft
4.7-6.5
4.8-6.6
4.9-6.7
5.2-6.9
5.4-7. 1
4.0-5.5
4.1-5.6
4.3-5.8
4.6-6.0
4.9-6.3
*Pages 38 through 41
A-73
-------
It is to be noted that the plots of Figures 6 through 9 and the data of Table u
are based on a melt bed diameter to depth relationship optimized for minimum
heat loss since this is considered the most important independent process variable
It does not take into account any other factor, such as structural requirements
which may affect the detailed design of the reducer. These, however, are not
expected to introduce any major changes in either the trends or the actual values
presented in the figures and table.
Both Figure 9 and Table II show that at a given reducer processing
capacity the normalized heat loss increases exponentially with increasing heat
flux. This is as would be expected since an increase in heat flux increases
not only the heat loss per unit wall area but also the size of the reducer vessel
required to accomodate the resulting higher heat requirements.
At a given heat flux, the normalized heat loss increases with decreasing
reducer processing capacity. For pilot plant size reducer units, the normalized
heat loss becomes large even at relatively low heat fluxes. At high heat fluxes
it becomes so great that a small pilot plant unit cannot provide an adequate
technical representation of a high heat flux reducer design (a design which would
be of interest only if there were a demand for the production of large amounts
of steam).
Table II shows that the actual magnitude of the heat loss from the reducer
melt bed will be quite large unless the heat flux at the walls is kept low. At
2
10,000 Btu/ft hr the heat loss from 60 and 200 Ib mole/hr SOX capacity reducers
would amount to 4. 8 to 6. 6 x 106 and 10. 0 to 14. 0 x 106 Btu/hr, respectively, at
sulfur compound concentrations in the melt of 0. 30 to 0. 15.
The following conclusions can be drawn from these results:
(1) The heat flux at the walls of the reducer vessel in contact with
the melt must be minimized not only to minimize the heat loss but also(to
minimize the physical size of the reducer vessel. As discussed in Section IX-3.^
this entails the use of the alumina-liner reducer design concept.
(2) The sulfur compound concentration in the melt must be as high
as possible without exceeding their solubility limits.
A-74
-------
(3) The data developed in this work do not allow any conclusion to
be drawn with respect to the effect of the state of oxidation of the sulfur com-
pounds in the melt leaving the scrubber upon the physical dimensions and heat
loss of the reducer. It is evident, however, that here again the sulfate content
of these compounds should be minimized as much as possible.
(4) If for some reason there were an interest in developing a frozen
skull or other high heat flux reducer design concept, such a reducer could not
be mocked up adequately in a small scale pilot plant unit as the excessively high
normalized heat loss in such a unit would entail the consumption of disproportionately
high amounts of coke and air.
c. Five and 267 Mw Reducer Units
Figures 10 through 17^'show the physical dimensions? and heat loss for the
special cases of 3. 125 and 167 Ib mole/hr SOX capacity reducer units, equivalent
to reducer units for nominal power plant capacities of 5 and 267 Mw , respectively
(pilot plant size, and one of three units of a fullscale 800 Mw plant). The
reduction region residence time was varied from 15 to 30 minutes and the super-
ficial air velocity from 3 to 5 ft/sec.
For comparison purposes the results are summarized in Figures 23 through
28 which show the effects of varying reduction region residence time and super-
ficial air velocity upon the reducer dimensions, normalized heat loss, and
normalized coke and air requirements. It is to be noted that the abscissa
scale on these figures is purely arbitrary and that the shape of the curves
drawn represents assumed trend curves since actual data were available only
at 15 minutes and 5 ft/sec, 15 minutes and 3 ft/sec, and 30 minutes and 3 ft/sec.
Figures 23 through 28 (and 10 through 17) show that:
(1) There is a greater incentive to decrease the minimum required
reduction region residence time from 30 to 15 minutes than to increase the
maximum allowable superficial air velocity from 3 to 5 ft/sec. The reduction
region residence time could be decreased by increasing the operating temperature
of the reducer, at a penalty, however, of increased preheat requirements and
potentially greater materials and carbonate salt decomposition problems.
(2) The incentive to achieve these improvements increases with
increasing heat flux from the melt bed to the walls of the reducer vessel.
*Pages 47 through 49
A-75
-------
24
15
REDUCTION TIME (min)
IS
30
20
16
14
H12
UJ
O
O
UJ
<£ 10
[S05
[so]
-• 0.33 AT SCRUBBER OUTLET
GEOMETRY = VERTICAL CYLINDER
EXPANDED BED DEPTH = 3 ft
15 m% S IN MELT
30m%SINMELT
Q/A = 20,000 Btu/ftz-hr
Q/A = 10,000 Btu/ft2-hr
Q/A
2 —
SUPERFICIAL AIR VELOCITY (ft-sec)
Figure 23. Effect of Reduction Time and Air Velocity
on Reducer Diameter
A-76
-------
1800
1600
1400
_ 1200
h
1000
< 800
ui
I
O 600
40C
20C
15
REDUCTION TIME (mini
15
[SO?]
[SO§1 + [S0|]
" 0.33 AT SCRUBBER OUTLET
GEOMETRY = VERTICAL CYLINDER
EXCEPT O = HORIZONTAL CYLINDER
-^^— 15 m% SIN MELT
30 m% S IN MELT
O
30
Q/A = 20,000 Btu/ft2-hr
x10°
5x 10°
:106
3
m
3
3x1062
g
2x10b
Q/A - 10,000 Btu/ft2-hr
1x106
SUPERFICIAL AlR VELOCITY (ft/sec)
Figure 24. Effect of Reduction Time and Air Velocity on
Reducer Bed Heat Loss
A-77
-------
15
REDUCTION TIME (mini
15
30
LLJ
5
ui
cc
a
LU
cr
ui
900
800
700
600
500
I
IS07]
[S0]
= 0.33 AT SCRUBBER OUTLET
GEOMETRY = VERTICAL CYLINDER
EXCEPT(D= HORIZONTAL CYLINDER
- 15 m% S IN MELT
30 m% S IN MELT
1800
Q/A, Btu/ft2-hr
400
300
1600
- 1400
1200
- 1000
- 800
- 600
400
- 200
- 100
ui
-------
15
REDUCTION TIME (min)
15
30
55
50
[SO;
ISO)
= 0.33 AT SCRUBBER OUTLET
45
40
£
Q
tr
LU
o
35
30
25
20
16
GEOMETRY = VERTICAL CYLINDER
EXPANDED BED DEPTH = 8 ft
15 m% SIN MELT
— 30 cn% SIN MELT
Q/A « 40,000 Btu/ft2-hr
Q/A - 20,000 Btu/ft -hr
Q/A = 10,000 Btu/ft2-hr
t Q/A = 20,000 Btu/ft^-hr
Q/A = 0
Q/A = 10,000 Btu/ft2-hr
Q/A = 0
3 3
SUPERFICIAL AIR VELOCITY (ft/sec)
Figure 26 Effect of Reduction Time and Air Velocity on
Reducer Diameter, 267-Mw Reducer
A-79
-------
1600
1400
1200
8
1000
m
"o
X
Q
UJ
N
O
800
600
400
200
15
REDUCTION TIME (min)
15
30
T
iscm
(305) + isoy
-- 0.33 AT SCRUBBER OUTLET
GEOMETRY = VERTICAL CYLINDER
EXCEPTO HORIZONTAL CYLINDER
15 m% SIN MELT
— — — 30 m% S IN MELT
Q/A = 40,000 Btu/hr-ft2
Q/A = 20,000 Btu/ft2-hr
-__ ^ Q/A = 10,000 Btu/ft2-hr
I
250x106
200 x106
150x10'
,6
100 x 10'
,6
50 x 10°
SUPERFICIAL AIR VELOCITY (ft/sec)
Figure 27. Effect of Reduction Time and Air Velocity on Reducer
Bed Heat Loss, 267-Mw Reducer '
A-80
-------
35,000
15
REDUCTION TIME (min)
15
30,000
25,000
CO
I-
z
HI
UJ
SE 20,000
2
oc
UJ
15,000
10,000
5,000
30
ISOJ
-= 0.33 AT SCRUBBER OUTLET
[S0§] + [SOJ]
GEOMETRY = VERTICAL CYLINDER
EXCEPTO HORIZONTAL CYLINDER
15 m% SIN MELT
— — —30m%SIN MELT
Q/A, Btu/fT-hr
_ 20,000
55,000
50,000
45,000
40,000 >
m
D
35,000 5
a
m
5
m
30,000 ^
25,000
20.000
15,000
10,000
5,000
SUPERFICIAL AIR VELOCITY (ft/sec)
Figure 28. Effect of Reduction Time and
Air Velocity on Reducer Coke and
Air Requirements, 267-Mw Reducer
A-81
-------
(3) On the basis of the very superficial look at a single case of a
horizontal cylinder reducer geometry with separation baffle perpendicular to
the axis of the cylinder, and within the limitations of the assumptions used in
the analysis, there appears to be no significant difference from the process
engineering point of view between this geometry and that of a vertical cylinder.
(4) As stated in Section IX-3.b. a small pilot plant unit cannot be
used to provide adequately representative technical information on a high heat
flux reducer design concept (a concept which would be of interest only if there
were a demand for the production of large amounts of steam).
d. Melt Bed Heat Loss
The conclusions reached thus far have all emphasized the essential
need for minimizing the heat loss from a molten carbonate reducer and there-
fore minimizing the heat flux at the walls of the reducer vessel in contact with
the melt.
The results of Section VII have shown that the frozen skull (cold wall)
reducer design concept does not lend itself to the use of a relatively low melting
point melt such as proposed in this process. The high temperature driving
force (~800 °F) between the melt and the walls of the vessel combined with a
relatively high heat transfer coefficient (> 50 Btu/ft hr) from the melt to the
walls yields a heat flux of over 40,000 Btu/ft^ hr considerably in excess of
^f
what would be acceptable in a molten carbonate reducer (Figure 19). Such a
high heat flux also imposes very stringent cooling requirements to allow the
maintenance of the frozen melt skull (Figures 18 and 19).* The frozen skull
reducer design concept is therefore not suitable for the molten carbonate
process. This conclusion, however, must be limited exclusively to the process
based on the use of the lithium, sodium and potassium carbonate eutectic as
the carrier melt. It does not necessarily apply to a process using a higher
melting point salt mixture, such as one based on sodium carbonate as the
carrier melt.
The alumina liner (hot wall) reducer design concept on the other hand in-
volves the use of internal insulation which, in the form of a combination of low
porosity, high thermal conductivity and high porosity, low thermal conductivity
liner materials, allows reduction of the heat flux at the walls of the reducer
vessel in contact with the melt to less than 2,000 Btu/ft^ hr and thus minimizes
&
the heat loss from the reducer vessel (Figure 20). Though this concept may
#Pages 54, 56 and 57
A-82
-------
present some materials problems it appears to be well suited for application to
the reduction step. It is to be noted that the minimum heat loss will be
achieved with natural convection air cooling of the outside surface of the reducer
vessel. Such cooling precludes the installation of thermal insulation at this
surface and therefore requires the outside wall of the vessel to run hot (~500 °F).
If a cold outside wall temperature should be required, the vessel walls would
have to be water-cooled, with thermal insulation installed around the outside
of the water-cooled region. The reducer heat loss would then be slightly higher,
but could still be kept to a heat flux value of less than 2, 000 Btu/ft2 hr.
It can therefore be concluded that:
(1) The alumina-liner (hot wall) reducer design concept appears to
be well suited for the molten carbonate reducer, with achievable heat flux levels
of less than 2, 000 Btu/ft^ hr at the surface of the walls in contact with the melt
bed.
(2) The reducer heat loss can be minimized with natural convection
air cooling of the outside walls of the reducer vessel, which will then operate
at relatively high temperatures (> 500 °F).
(3) Water cooling of the walls of the reducer vessel can be used
to lower the outside temperature of the vessel, at a penalty, however, of a
small increase in heat loss.
(4) The frozen melt skull (cold wall) reducer design concept is not
applicable to a molten carbonate reducer using a. relatively low melting point
salt mixture such as a lithium-sodium-potassium carbonate-based melt. It
\
may, however, be well suited for application to a molten carbonate reducer
using a high melting point salt mixture such as a sodium carbonate-based melt.
e. Internal Melt Recirculation
The analysis of Section VIII, though based on very rough assumptions,
indicates that there should be no problem in achieving adequate recirculation
of the melt between the oxidation and reduction regions of the reducer. The
difference in effective densities between the two regions, resulting from a
slight difference in superficial gas velocities and, to a lesser extent, a
difference in average temperatures, is sufficient to provide the required
recirculation. The most important consideration in this respect may involve
control of the recirculation rate by correct sizing of the orifices at the top
and bottom of the baffle separating the oxidation region from the reduction region*
A-83
-------
These orifices must be small enough to regulate the melt flow rate to maintain
the minimum residence time of the melt in the reduction region required for
carrying out the reduction reaction to the desired degree of completion. Typical
orifice dimensions are shown in Figure 21*
4. Conclusions
A preliminary process analysis of a two-region molten carbonate reducer
design concept shows that:
1) The heat loss from the melt bed must be minimized in order to
minimize the coke and air requirements of the reducer and its physical
dimensions.
2) The sulfur compound concentration in the carbonate melt must
be as high as feasible subject to solubility limitations, thus minimizing
melt flow rate and melt preheat requirements.
3) The sulfate fraction of the sulfur compounds in the melt leaving
the scrubber must be minimized in order to minimize the reduction duty
of the reducer.
4) The coke utilization in the reducer must be maximized in order
to minimize the melt recovery and make-up requirement resulting from
the melt loss associated with the filtration of the unreacted coke.
The controlling constraints in the design of the reducer are the maximum
allowable superficial air velocity in the oxidation region and the minimum re-
quired residence time of the melt in the reduction region. A decrease in re-
duction region residence time from 30 to 15 minutes results in a more significant
reducer design improvement than an increase in superficial air velocity of 3 to
5 ft/sec.
Typical reducer dimensions have been determined as a function of processing
capacity and heat loss. Coke and air requirements have been estimated as a
function of heat loss. For a sulfur compound mole fraction in the melt of 0. 30,
one-third of which is sulfate, the coke and air requirements of a reducer with
a negligible heat loss amount to approximately 1. 30 Ib and 73 scf, respectively,
per Ib of sulfur fed into the scrubber.
*Page 64
A-84
-------
The heat loss from the melt bed in the reducer can be minimized through
use of an alumina-lined internally insulated (hot wall) reducer design. If it
is air-cooled the outside wall of the reducer vessel in this design cannot be
insulated and operates at a temperature of around 500 °F. Water cooling
surrounded by insulation would allow operation of the outside of the reducer
at a lower temperature.
Because of its inherently high heat loss a frozen melt skull (cold wall)
reducer design does not appear feasible when a low melting point carbonate
mixture such as the lithium-sodium-potassium carbonate eutectic is used as
the carrier melt.
Internal melt recirculation between the oxidation and reduction regions
of the reducer can be controlled through sizing of the orifices in the baffle
between the two regions. These orifices must be small enough to assure
satisfaction of the minimum residence time requirement in the reduction region
and large enough to allow sufficient flow to transport the required heat without
exceeding melt bed temperature rise limitations.
The results obtained are based on the technical information presently
available and on various assumptions and approximations which jiad to be made
in lieu of required experimental data. They are therefore preliminary and must
be confirmed through a development program. They are, however, believed to
provide a sound basis for the conceptual design of a molten carbonate reducer
and the planning of the required experimental program.
A-85
-------
REFERENCES
1. G. J. Janz, "Molten Carbonate Electrolytes as Acid-Base Solvent
Systems," AD 651604, Rensselaer Polytechnic Institute, Troy,
New York, February 1967.
2. H. C. Weber, "Thermodynamics for Chemical Engineers," John Wiley
& Sons, Inc., New York, 1939, p. 52.
3. Ibid. , p. 54
4. PB-168370, "JANAF Thermochemical Tables," Dow Chemical Company,
Midland, Michigan, August 1967.
5. J. H. Perry, "Chemical Engineers' Handbook, " McGraw Hill Book Co. ,
Inc., Third Edition, 1950, pp 239-243.
6. G. T. Skaperdas, "Commercial Potential for the Kellogg Coal Gasification
Process," PB 180358, M. W. Kellogg Company, Piscataway, N. J. ,
September 1967, pp 87-92.
7. "Development of a Molten Carbonate Process for Removal of Sulfur
Dioxide from Power Plant Stack Gases, Progress Report No. 2,
Part I. Process Chemistry - Radiation. April 1, 1968 to October 27,
1968." AI-70-5, pp. 48-50.
8. W. H. McAdams, "Heat Transmission, " McGraw Hill Book Co. , Inc.,
Third Edition, 1954, p. 172.
9. "Chemical Recovery in Alkaline Pulping Processes, " Roy P. Whitney,
Editor, TAPPI Monograph Series No. 32, 1968, pp. 59-99
A-86
-------
APPENDIX B
PRELIMINARY PROCESS ANALYSIS OF A MOLTEN
CARBONATE REGENERATOR
-------
I. INTRODUCTION
In the molten carbonate regeneration step, the alkali metal carbonate-
sulfide melt from the reduction step is contacted with a carbon dioxide-steam
mixture. The reaction converts the sulfide into carbonate and releases hydrogen
sulfide. This Appendix describes in detail the results of an analytical study of
the regeneration step. The study was conducted to determine the hydrogen sul-
fide concentrations achievable in the regenerator off-gas and to establish the oper-
ating conditions required to attain these concentrations. Specific emphasis was
given to the constraints imposed by the thermodynamic equilibrium and heat gen-
eration of the regeneration reaction, by the alkali metal sulfide solubility in the
melt, and by the composition of the available regeneration feed gas.
II. CONCLUSIONS
1. The operating parameters of the regenerator must be optimized to
satisfy the following constraints:
a. A thermodynamic equilibrium strongly favored by low temperatures
b. A highly exothermic heat of reaction
c. A solubility of sulfide in the carbonate melt which increases sig-
nificantly with increasing temperature
2. The regeneration equilibrium is strongly favored by a high steam con-
centration in the feed gas, and to a lesser extent by increased total pressure or
CO concentration (especially at low temperatures or high CO concentrations).
2 *-
At a given steam concentration the effect of CO- concentration is more important
than that of total pressure.
3. The maximum hydrogen sulfide concentration achievable in the regen-
erator off-gas is for all practical purposes independent of the M2S concentration
in the carbonate feed melt.
4. A stoichiometry limitation combined with equilibrium considerations
makes it essential that the regeneration gas have as high a concentration as
possible of CO and HO. This requirement can be adequately satisfied through
2 £
use of the reducer off-gas from a fluid coke reduction system (about 40 mole
B-l
-------
5. As a result of potential solubility limitations, the operating tempera-
ture of the regenerator may be determined by the M_S concentration in the car-
£*
bonate feed melt. For operation in the temperature range below 1000 F, the
MS concentration in the carbonate feed melt may have to be kept under about
17 mole %.
6. At atmospheric pressure a typical fluid coke reducer off-gas contain-
ing about 25 mole % CO and 25 mole % HO (through addition of steam)* allows
w L*
the production of an off-gas containing a maximum of about 10 mole % H S.
L*
Adding more steam, available as a by-product of the process, to increase the
HLO content of the gas to 40 mole %, concurrently reducing its CO content
Lt t*
to 20 mole %, increases the achievable H S concentration in the product gas to
about 15 mole %.
7. Approximately nine theoretical plates are required to regenerate 95%
of the MS in the temperature range of 950 to 1000 F with a feed gas containing
4*
20 mole % CO and 40 mole % HO and a feed melt containing 15 mole % MS.
L* CM £*
The corresponding H S concentration of the product gas is about 13. 7 mole %.
Lm
Increasing the total pressure to two atmospheres would allow production of a
16 mole % H S product gas or a decrease in the number of theoretical plates
C*
from 9 to 6.
8. The number of theoretical plates required, and the maximum achiev-
able H S concentration in the product gas, are controlled by a pinch point
£*
between the operating line and the equilibrium line at the maximum allowable
operating temperature of the regenerator. This pinch point is approached
as the heat of regeneration raises the temperature of the melt, requiring
intercooling to allow regeneration to proceed. Several stages of external(
cooling, or cooling on every plate, would practically eliminate the pinch point
limitation and thus appreciably decrease the number of plates required and
increase the H S concentration achievable in the product gas. However,
a single stage cooling the melt taken off a plate slightly above the middle
*The reducer analysis conducted after completion of the regenerator study has
shown that a regeneration gas containing- about 40 mole % CO^ and 40 mole %
H~O will be available from a two-region molten carbonate reducer.
B-2
-------
of the regenerator from about 1000 F down to about 900°F appears to be ade-
quate. Approximately 60% of the heat of reaction can thus be removed. The
remainder goes into the product gas and the regenerated carbonate melt.
HI. BASIS
A. THERMODYNAMIC EQUILIBRIUM
The effective regeneration reaction was assumed to be
C02(g) + H20(g) * M2C03(4) + H2S(g) (1)
(where M = mixture of Li, Na, and K cations)
and consideration was not given to any actual reaction mechanism and formation
of intermediate compounds such as the thiocarbonate M CO S.
L* L*
The equilibrium data used were those obtained experimentally and
reported in reference (1). They can be formulated as follows:
[H2S]
[C°2] [H2°] [MZS] * + V [C°2J
with: KI - 1.90xlO-6e5°'400/RT
K, = 1.24x10-3327'720/RT
Lf
rr = Total pressure, atmospheres
R = Gas constant, Btu/ R Ib mole
T = Absolute temperature, R
Figure 1 presents an arithmetic scale plot of K^ and K2 as a function of tempera-
ture over the temperature range considered to be of interest, 800 to 1050 F.
B. HEAT OF REACTION
A value of -25.8 kcal/g mole = 46,400 Btu/lb mole, based on estimates from
reference (1), was used. While this value had been estimated for a temperature
of 850°F, no attempt was made in the present calculations to include the effect
B-3
-------
1000
7T = TOTAL PRESSURE, atm
R = GAS CONSTANT, 1.987 Btu/°R-lb-m
T = ABSOLUTE TEMPERATURE, °R
850
900
950
TEMPERATURE (°F)
1000
1050
1100
Figure 1. Regeneration Equilibrium Constants
B-4
-------
of temperature upon the heat of reaction since this value was used only in the
approximate heat balances carried out to establish the number of theoretical
plates required for efficient regeneration.
C. SULFIDE SOLUBILITY IN CARBONATE MELTS
The solubility of sulfide (MS) in the alkali metal carbonate (M CO ) melts
L 23
was determined by measuring the initial freezing temperature of fused salt mix-
tures. The first measurements were done with mixtures made up by combining
purified sodium sulfide with lithium, sodium, and potassium carbonate in the
proper proportions. The freezing points of these mixtures are shown plotted in
Figure 2; the data indicate that about 18 mole % is soluble at 950 F, and that a
32 mole % sulfide mixture started to freeze at 1060 F. The 32 mole % mixture
was then treated by bubbling CO through it to simulate melt from the reduction
L*
step of the process. This treatment lowered the freezing point to 770 F, probably
by converting part of the sulfide to thiocarbonate (M CO S). The regeneration
L* C*
tests used to determine the equilibrium and heat of reaction data were all done
with reduced melts, not synthetic mixtures. Therefore, it is believed that the
regeneration tests actually involved carbonate-sulfide-thiocarbonate mixtures;
This was discussed in reference (1). Because of this, the initial freezing point
of the melts to be regenerated will probably not exceed 770 to 800 F. However,
the freezing points and solubilities of the synthetic mixture were used as design
constraints in this study, to be conservative.
IV. METHOD OF CALCULATION
The basic thermodynamic equilibrium equation, equation (2), was written
in the form:
« P / CH2S1
TT ^
B-5
-------
w
1100
1050 —
750
10
20
[M2S)
30
Figure 2. M-,S Solubility in Molten Carbonate Melts
MELT TREATED _
WITH CO2
40
-------
with: y = rM g-r > ratio of mole fractions in the melt on a regenerator
plate in equilibrium with the gas leaving the plate,
<* = Lco2J
3 = [H20]
[H2S]
mole fractions in the gas leaving a plate
in equilibrium with the melt on the plate,
K^ and K^ = regeneration reaction equilibrium constants, from Fig. 1,
TT = total pressure, atmospheres.
If it is assumed that the regeneration feed gas introduced at the bottom of
the regenerator contains only CO , HO, and N , without any H S, the stoichi-
£* L* Lt L*
ometry of the regeneration reaction, Eq (1), is such that:
J r^o i /11 r<"><"\ i i I TTJ ci
1 L^*-'o -ITT / V'-'-'^-'^JT-,] f ~ L-'-^o" J
a =
1 + {[co2JF/(1-[co2JF)}
(CH20]F/(l-[H20]F)}- [H2S]
1+{[H20]F/(l-[H20]F)}
with the subscript F being applied to mole fractions in the feed gas.
It is important to note that stoichiometry thus limits the maximum
achievable concentration of H S in the product gas to the lesser of
or [H20]F/(l-[H20]F).
Given the composition of the regenerator feed gas, and the H2S concen-
tration in the gas in equilibrium with the melt on a plate, one calculates a and
P by means of equations (4) and (5), and then y by means of equation (3).
Knowing Y, one obtains the MS concentration in the melt in equilibrium
with the H S concentration in the gas:
L*
B-7
-------
20
W
I
00
5-10
REGENERATOR GAS
INLET FLOW RATE
INLET COMPOSITION
EQUILIBRIUM LINES
OPERATING LINES
M2S SOLUBILITY LIMIT
Figure 3. Molten Carbonate Regenerator Equilibria (A)
-------
with [I] = concentration of inerts, such as M SO , in the melt. In general [I]
Cf 4
will be small. It has been neglected in all further calculations, but could be cor-
rected for very simply by multiplying the [MS] scale (abscissa) values on all
Lt
the graphs by the constant factor 1- [I] .
Values of [M^S] in the melt in equilibrium with [H S] in the gas were thus
obtained at temperatures ranging from 800 to 1050 F, and plotted as a function of
each other for various regeneration feed gas compositions. The compositions
selected were:
Case A 10% CO , 10% HO, 80% N, Figure 3
c* £ Figure 10
Case H 40% CO^ 40% H2O, 20% NZ> Figure 11
Case I 50% CO2> 50% H2O, 0% NZ, Figure 12
With the exception of compositions B and E, all these compositions had
equal mole fractions of CO2 and H2O, ranging from the concentrations achiev-
able from the combustion of natural gas to those achievable through the com-
bustion of pure carbon or calcination of limestone.
Figures 4, 7 and 8 show the effect of addition of extra quantities of steam
to the regenerator feed gas (compositions B and E).
Figure 8 shows the effect of operation at a total pressure of two atmos-
pheres as compared with the atmospheric pressure on which all the other calcu-
lations were based. Gas composition E was used.
B-9
-------
M_S solubility v£ temperature values from Figure 2 were then superim-
b
posed on the equilibrium curves of Figures 3 through 12, yielding the curves of
"MS solubility limit" for each of these figures.
£*
Operating lines were also included in these figures to show the H S-M S
* £* L*
mass balance relationship from plate to plate in the regenerator. The lines
shown are ideal lines assuming complete (100%) stripping of the hydrogen sulfide
from the melt and represented by the following equation which takes into account
the stoichiometrically decreasing mole flow rate of gas as it passes through the
regenerator:
L/G [MS]
[HS] =
2 1-L/G [M2S]
with: L = liquid flow rate, Ib moles/hr
G = gas flow rate at regenerator inlet, Ib moles/hr
Actual operating lines follow the relationship:
-[M S] -(1-T))[M S]
s] = G L 2 - 5-J — (8)
with: 7| = hydrogen sulfide stripping efficiency
and [MS] = metal sulfide concentration in the feed melt to the regenerator.
M J?
The three cases of Figures 7, 8 and 9 (gas compositions E and F) were
selected for layouts of theoretical plate diagrams of the regenerator, since all
three apply directly to gas compositions achievable through use of the fluid
coke reducer off- gas, with various steam additions, as the regenerator feed
gas.* The regenerator feed melt was assumed to have an MS concentration of
^
15 mole % and be at a temperature of 950 F.
The plates were laid out in Figures 13, 14, and 15 in a standard McCabe-
Thiele diagram fashion, including a very rough heat balance around each plate.
This heat balance determined the temperature on the plate and therefore the
*As mentioned in an earlier footnote, gas,composition H (Fig. 11) may actually
be available from a two region reducer.
B-10
-------
30
20
REGENERATOR GAS
INLET FLOW RATE
INLET COMPOSITION
-EQUILIBRIUM LINES
—OPERATING LINES
G Ib-m/hr
CO2 10 m%
H20 25 m%
No 65 m%
^ M2SSOLUBILITY LIMIT
950°F
10
20
30
[M9S]
40
Figure 4. Molten Carbonate Regenerator Equilibria (B)
-------
w
I—1
to
a?
J,
REGENERATOR GAS
INLET FLOW RATE
INLET COMPOSITION
G Ib-m/hr
C02 15m%
H2O 15 m%
EQUILIBRIUM LINES
—. — — — — OPERATING LINES
STOICHIOMETRY LIMIT
M,S SOLUBILITY LIMIT
Figure 5. Molten Carbonate Regenerator Equilibria (C)
-------
30
bd
i—•
UJ
20
CM
X
10
REGENERATOR GAS
INLET FLOW RATE G Ib-m/hr
INLET COMPOSITION CO2 20 m%
H20 20 m%
N- 60 m%
STOICHIOMETRY LIMIT
M2SSOLUBILITY LIMITv
10
20
30
1000°F
1050°F
40
[M,S]
Figure 6. Molten Carbonate Regenerator Equilibria (D)
-------
w
I
REGENERATOR GAS
INLET FLOW RATE
INLET COMPOSITION
EQUILIBRIUM LINES
OPERATING LINES
Figure 7. Molten Carbonate Regenerator Equilibria (E)
-------
td
[MS]
Figure 8. Molten Carbonate Regenerator Equilibria (E, 2 Atmospheres)
-------
30
W
20
M2SSOLUBILITY LIMIT
10
REGENERATOR GAS
INLET FLOW RATE
G Ib-m/hr
INLET COMPOSITION C02 25 m%
H2O 25 m%
N2 50 m%
" EQUILIBRIUM LINES
— OPERATING LINES
10
20
[M2S]
30
40
Figure 9. Molten Carbonate Regenerator Equilibria (F)
-------
w
I
M2S SOLUBILITY LIMIT
REGENERATOR GAS
INLET FLOW RATE
INLET COMPOSITION
G Ib-m/hr
C02 30 m%
H,0 30 m%
EQUILIBRIUM LINES
OPERATING LINES
Figure 10. Molten Carbonate Regenerator Equilibria (G)
-------
w
H-
oo
800°F 850°F 900°F
Figure 11. Molten Carbonate Regenerator Equilibria (H)
-------
REGENERATOR GAS
INLET FLOW RATE
INLET COMPOSITION
Glb-nvhr
C02 50 m
H20 50 m%
[M2S| l
Figure 12. Molten Carbonate Regenerator Equilibria (I)
B-19
-------
950°F
10QO°F
1050°F
[M2S] (
Figure 13. Theoretical Plate Diagram for Molten Carbonate
Regenerator (F, 1 Atmosphere)
B-20
-------
950°F
1000°F
1050°F
(M2S) (
Figure 14. Theoretical Plate Diagram for Molten Carbonate
Regenerator (E, 1 Atmosphere)
B-21
-------
w
I
Figure 15. Theoretical Plate Diagram for Molten Carbonate Regenerator
(E, 2 Atmospheres)
-------
applicable equilibrium line. In a simplified manner it was assumed that for
each percent of sulfide regenerated in the melt 46, 400 x 0.01 = 464 Btu/lb mole
of heat would be generated on a plate, which with an approximate specific heat of
the melt of 41.4 Btu/lb mole F, would raise the temperature of the melt by
o
about 11 F. In actuality the melt would receive heat not only from the chemical
reaction on the plate but also from the sensible heat brought by the hotter gas
rising from the plate below. A value of 13°F was estimated for the temperature
rise of the melt per percent of the melt regenerated to take into account this
sensible heat of the gas. Exceptions were made for the plate from which the
melt was taken to the intercooler and the two bottom plates to take into account
the effect of the colder gas rising to these plates. More detailed heat balances
were performed for these plates (the regenerator feed gas was assumed to be
introduced at a temperature of 850 F).
Additional calculational background and data are presented in Section VI.
V. DISCUSSION
Figures 3 through 12 show that for a given regenerator feed gas composition
at a given total operating pressure the flatness of the MS solubility limit curve
for concentrations of MS above a relatively low value (generally less than
b
10 mole %) prevents improvement of the H S concentration ideally achievable in
the product gas by means of an increase in the MS concentration in the melt.
Increasing the M S concentration in the melt only raises the required operating
temperature in the regenerator without any H2S concentration increase in the
product gas. It does, however, decrease the melt flow requirements through
the overall SO removal system and the preheating requirements in the reducer.
Figure 2, and Figures 3 through 12 show that on the basis of the assumed
sulfide solubility data, operation of the regenerator at various temperature ranges
would impose the following constraints on the M2S concentrations in the regener-
ator feed melt (the subscript F here again applies to the feed):
B-23
-------
860 to 900°F [M _S]_ £ 6 mole %
Z Jc
900 to 950°F [M S]_ =£11 mole %
L* £
950 to 1000°F [M_.S]_ £17 mole %
Z r
1000 to 1050°F [MS] £ 22 mole %
2 Jc
These could be raised somewhat if one were willing to use several stages
of melt interceding, rather than a single stage, or cooling of the melt directly
o
on each individual plate. A feed melt temperature of 950 F and MS concentra-
tion of 15 mole % were arbitrarily selected in this study to allow operation of the
o
regenerator within the 950 to 1000 F temperature range of present day materi-
als compatibility technology.
As stated previously the stoichiometric limitation on the maximum H S
t*
concentration achievable in the product gas is equal to the lesser of
[CO J/fl- [CO ~\\ or [H OJ /(1- [HO]") in the feed gas. This value could be
z \ z/ z \ z/
achieved only in equilibrium with a melt with an MS concentration of unity. Con-
centrations actually achievable are generally limited by the approach to the pinch
point between the operating line and the equilibrium line at the maximum allow-
able operating temperature. For equal concentrations of CO and HO in the
regenerator feed gas (i.e., as in Figures 3, 5, 6, 9, 10, 11, and 12) they are
roughly estimated at
[CO. L, = [H_0]_ = 10 mole % [H.S ] * 3 mole %
L. J! c. E &
[CO 1 = [H_0] = 15 mole % [H,S] s 6 mole %
Z JC L, Jt? LM
[C02JF = [H20]F = 20 mole % [H^] £ 8 mole %
[C02JF = [H20]F = 25 mole % [H2S] S 10 mole %
[CO L = [H 0]_ = 30 mole % [H_S] <; 13 mole %
2 Jb 2 Jb Z
[C02]F = [H20]F = 40 mole % [HS] <. 19 mole %
[CO 1 = [H0O]_ = 50 mole % [H S] £ 26 mole %
2 F 2 F Z
It is to be noted that to be suitable for economic recovery of sulfur through
a Glaus-type recovery process the product gas should have an H_S concentration
(2)
of at least 10 mole %. For lower concentrations a Stretford-type process
may well prove to be more economical.
B-24
-------
The off-gas from a fluid coke molten carbonate reducer may contain up
to about 67 mole % of CO , yielding a 40% COo - 40% H0O mixture. This should
*• £. £.
be adequate for regeneration. An even better source would be a pure CO gen-
£*
erator, but this would add the requirement for an additional major piece of equip-
ment to the molten carbonate SO removal system. While pure CO is normally
not available at a molten carbonate SO removal plant, considerable quantities
L»
of excess steam are produced in such a plant. Since the thermodynamic equi-
librium relationship, equation (2) above, is such that the effect of steam con-
centration on the achievable H S concentration is much more important than
C*
that of CO concentration, it seems obvious that an appreciable improvement
should be achieved through the addition of excess steam to a CO containing
L*
regeneration feed gas.
Doubling the amount of steam in an originally 25 mole % CO , 25 mole %
HO (composition F) gas increases its HO concentration to 40 mole % while its
L* £*
CO_ concentration is decreased to 20 mole % (composition E). The net effect is
L*
shown in Figure 7. The H S concentration achievable in the product gas is raised
Lt
from approximately 10 mole % to approximately 15 mole %, a more than adequate
value for the economic recovery of sulfur through a Glaus-type process (though
the gas may have to be dried prior to or during its treatment in the Glaus plant).
A regeneration feed gas with a composition of 20 mole % CO_, 40 mole % HO
and 40 mole % N is easily obtainable in a molten carbonate SO recovery plant
2 ^
through addition of by-product steam to reducer off-gas.
Figure 4 shows a comparable situation with a low CO content regeneration
feed gas. The 10 mole % CO , 25 mole % H2O, 65 mole % NZ (composition B) gas
is representative of a gas obtainable from the combustion of natural gas (the
actual CO concentration in such a combustion flue gas would be about 8.8 mole %).
Lf
The maximum achievable H2S concentration in the product gas is less than
7 mole %. This is probably inadequate for a Glaus process recovery plant, but
should be suitable for the Stretford process.
B-25
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Figure 8 shows the effect of increasing the total pressure in conjunction
with the use of the composition E regeneration feed gas. An increase from one
to two atmospheres increases the maximum achievable H S concentration in the
product gas from about 15 mole % to about 17 mole %. Such operation would,
however, imply appreciably pressurized operation of the reducer, which might
not be technically feasible or economically practical.
Figures 13 and 14 show typical McCabe-Thiele theoretical plate diagrams
for operation with a regeneration feed gas obtained from the reducer off-gas
through the addition of steam to form the 1:1 CO HO and 1:2 CO - HO gas
£ £m £» C*
compositions previously used (F and E). Figure 15 applies to the composition E
feed gas at a pressure of two atmospheres. MS concentration in the feed melt
CM
was 15 mole %. Feed melt and feed gas temperatures were 950 and 850 F,
respectively. A single stage of melt cooling was provided. The results can be
summarized as follows:
Operating Conditions Regeneration H2S Product No. of
Feed Gas Pressure Temp Effectiveness Cone. Theoretic
Composition (atm) (°F) L/G (%) (%) Plates
CO 25 mole % 1 950-1000 0.58 95 8.9 8
L*
HO 25 mole %
£*
N 50 mole %
L*
CO 20 mole % 1 950-1000 0.85 95 13.7 9
L*
HO 40 mole %
£
NZ 40 mole %
CO 20 mole % 2 950-1000 0.85 95 13.7 6
<£
HO 40 mole %
N 40 mole %
L*
It can be seen that 95% regeneration can be achieved with a reasonable
number of theoretical plates. The H S concentration in the product gas is com-
Lf
pletely adequate for economic recovery of sulfur through a Glaus-type process.
B-26
-------
In the 2 atmosphere process case a still higher concentration of H S (about
L*
16 mole %) could be achieved in the product gas but at the expense of an increase
in the number of theoretical plates required.
As a last item, it is important to emphasize that this whole analysis has
been based on a rather limited amount of analytical data. Especially with respect
to the MS solubility data in the carbonate melt there are indications that inter-
c*
mediate compounds formed in the regeneration process, such as the thiocarbonate
M _CO S, are considerably more soluble in the melt than the sulfide itself. How
L* Ct
much this can be taken advantage of in the process remains in question. It is
expected that the results are conservative, however, since any decrease in
initial freezing point temperature below that assumed will make the regeneration
easier.
VI. BACKGROUND DATA
A. EQUILIBRIUM CONSTANTS
K = 1.90*10e' K
1 "
X t^XXXIJC^J. CLtr^JLJ. \.
( F)
800
850
900
950
1000
1050
1100
1200
1300
1400
K
1
1051
487
239
123
66.6
37.5
21.9
8.22
3.45
1.59
K
2
79.8
52.3
35.4
24.6
17.5
12.76
9.49
5.54
3.44
2.24
B-27
-------
B. MS SOLUBILITY IN CARBONATE MELTS
v = £*
Melt Composition
(mole %)
— 2—
-
1.5
17.2
31.8
31.8*
M0CO0
100.00
98.5
75.2
81.4
66.9
66.9
1 77.>f .—."? . *3
L. '4 ' £ ' J
1.5
10.1 13.2
1.4
1.3
1.3
Liquidus Point
751
754
802
937
1059
769
#Treated with CO at
L*
(600°C)
Data from Figure 2
Temperature
800
850
900
950
1000
1050
MS Solubility
(mole %)
1.4
6.8
12.7
18.8
24.8
30.8
C. SOLUTION OF EQUILIBRIUM EQUATION
1. Determination of [MS] as a function of [H S]
(L &
As described in the text of the Appendix:
B-28
-------
with [I] being neglected for purposes of this calculation, and
Y • K! 1. .. <3>
Since use of a large computer was not available, the numerical calcula-
tions were carried out on the Commodore Electronic Desk Top Calculator
Model AL. 1000 using the following programs:
(a) Calculation of a
Memories: I 1 Program: 427352. 795324252
III OH S] *432639537.
IV
(b) Calculation of a ft/[H S]
Memories: I 1 Program: 427352.795324252
HI [H S] *4326395386. 39637.
IV [HO] 6. - or
(c) Calculation of [M Si
Note that both previous sets of calculations were purely
stoichiometric and therefore independent of temperature and pressure. It is
only in this calculation that the effects of temperature and pressure are intror
duced.
Memories: I 1 Program: 87326.252.6.953863
in K *24252.49537.
•" 2 Start program with o
6. - I/TT
6. - a p/[H7S]
B-29
-------
2. Alternate Determination of [H0S] as a Function of [M0S]
~=== ' " " Lt ~r C* ~
Equation (2) can be solved for [H S]. Define:
a =
b =
(9)
K
K,
d =
b)
[M2C03J
IM2s]
2(1+c)
(10)
(11)
Equation (2) is then rewritten in the quadratic form:
- 2d [HS]
= 0
(12)
which yields:
[H2S] = d
1-1-
ab
d (1 + c)
(13)
D. EQUATION OF OPERATING LINE
Let [MS] = x
Lf
[H2S] = y
From the stoichiometry of equation (1), the gas flow rate G leaving
plate n can be expressed as follows:
G
G =
n 1 + y.
'(14)
n
where G is the feed gas flow rate.
B-30
-------
Since the regenerator feed gas does not contain any H S, a sulfide mass
Li
balance below plate (n - 1) from the top of the regenerator can be written as
follows:
where 7) = efficiency of regeneration and :x = [M S] in the feed melt
F 2 F
2"
n - r (15)
n
L/G(X - (I-TI)X ]
y = LJlLl LL_ (16)
L/G-f [M?S]
[H0S] =
1 -L/G
It is to be noted that in the molten carbonate SO removal process using
C*
a fluid coke reduction system the amount of regeneration gas available from the
reducer is limited, so that the ratio L/G has a definite lower boundary.
If the total amount of sulfur processed by the system is S Ib-mole/hr:
. L 1 (17)
2 . . 12 [M
L*
G <~ 12S
Therefore if: [M2S] = 10% ^- > 0.83
= 15% £- > 0.56
= 20% r > 0.42
G
B-31
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E. HEAT BALANCE DATA
Specific heat of melt at 950°F:
C_ = 28.41 + 16.56 x 10"3 -M^cT - 41.4 Btu/lb mole °F
J_i 1 . O
Specific heat of regeneration gases at 900 F:
CO 1-2.3 Btu/lb-mole °F
L*
HO 9.0 Btu/lb-mole °F
L*
N 7. 6 Btu/lb-mole °F
H S 10.2 Btu/lb-mole °F
which yields for typical gas compositions:
C = 9.1 Btu/°F per Ib-mole of feed gas
C = 8.8 Btu/ F per Ib-mole of product gas
For regeneration of 1 mole % of the melt the heat of reaction amounts to:
46, 400 x 0. 01 = 464 Btu/lb-mole of melt
which would correspond to a (464/41.4) = 11.2 F temperature rise if all this
heat were to go into the melt.
For most plates in the regenerator the gas rising from the plate below is
hotter than the liquid on the plate and therefore adds heat to the melt. Check-
ing a few actual heat balances indicated that the sensible heat thus added might
amount to about 20% of the heat of reaction. A constant value of 13 F per
percent of the melt regenerated was therefore used for the temperature rise on
a plate not affected by incoming colder regeneration gas.
An approximate heat balance was carried out for the plate above that onto
which the cooled liquid from the intercooler was returned.
For the bottom two plates, plates (n-1) and n, the following heat balance
equations were obtained (neglecting the change in gas flow rate):
B-32
-------
c
? 7^- AHAR +F AH(AR + AR \
n-2 G/ C n G V n n-1/
t -t = t (18)
n Lr
m + pC
C
C ft - t_) + 7^- AH AR ^ AHAR
G\ n-2 G/ C n-1 G n
t -t = ±— (19)
n-1 n m + pC
with t_ = regeneration feed gas temperature
G
t = temperature on plate n
n
t - temperature on plate n- 1
n- 1
t = temperature on plate n-2
n-2
AH = heat of regeneration, Btu/lb-mole
AR = fraction of melt regenerated on plate n
n
AR = fraction of melt regenerated on plate n- 1
n-1
m
B-33
-------
REFERENCES
1. AI-70-29, "Development of a Molten Carbonate Process for Removal of
Sulfur Dioxide from Power Plant Stack Gases - Progress Report No. 3,
October 28, 1968 to July 31, 1969, " pp 41-43
2. The Stretford Process. The North Western Gas Goard (U.K.). October,
1967
B-34
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