EPA-650/2-74-001
January 1974
Environmental Protection Technology Series
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EPA-650/2-74-001
A REGENERATIVE LIMESTONE PROCESS
FOR FLUIDIZED-BED COAL
COMBUSTION AND DESULFURIZATION
by
R. C. Hoke, M. S. Nutkis, L. A. Ruth, and H. Shaw
Esso Research and Engineering Company
P.O. Box 8
Linden, New Jersey 07036
Contract No. CPA 70-19
ROAP No. 21ADB-13
Program Element No. 1AB013
EPA Project Officer: S. L. Rakes
Control Systems Laboratory
National Environmental Research Center
Research Triangle Park, North Carolina 27711
Prepared for
OFFICE OF RESEARCH AND DEVELOPMENT
U.S. ENVIRONMENTAL PROTECTION AGENCY
WASHINGTON, D.C. 20460
January 1974
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This report has been reviewed by the Environmental Protection Agency and
approved for publication. Approval does not signify that the contents
necessarily reflect the views and policies of the Agency, nor does
mention of trade names or commercial products constitute endorsement
or recommendation for use.
11
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AB STRACT
An experimental study was begun of the pressurized combustion of coal
in a fluidized bed of limestone and regeneration of sulfated limestone.
The work is part of an overall program aimed at developing fluidized
bed coal combustion as a means of desulfurizing flue gas in—Situ and
generating clean power at low cost. The process includes regeneration
of spent limestone by reduction to lime. This produces a gas stream
containing a sufficient concentration of SO 2 to be fed to a by product
sulfur recovery unit.
The regeneration step was studied at pressures up to 10 atm and tempera-
tures up to 2100°F. S02 concentrations measured.lin the product gas
averaged about 2% at 10 atm pressure and 2100°F, about 40% of the
concentrations calculated by assuming equilibrium between the solids
and regenerating gas. High conversion of sulfated material to lime was
achieved by injecting air into the bed, forming adjacent reducing and
oxidizing zones, and minimizing formation of undesired CaS.
Combustion studies also began. The combustion runs were limited by
operating problems, especially plugging in the coal injection line.
Initial SO 2 removal rates ‘were about 85%. However attrition rates were
high with one SO 2 sorbent, Tyinochtee dolomite.
This report was submitted in fulfillment of Contract CPA 70—19 by Esso
Research and Engineering Company under the sponsorship of the Environ-
mental Protection Agency. Work was completed as of July 31, 1973.
iii
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CONTENTS
Page
Abstract iii
List of Figures v
List of Tables vii
Acknowledgments viii
Sections
I Conclusions 1
II Reconiinendations 2
III Introduction 3
IV Experimental Equipment, Materials, Procedures 5
V Development of Experimental Equipment and Procedures 30
VI Experimental Results 53
VII Discussion of Results 87
VIII Program for Operation of Niniplant 90
IX References 92
X List of Publications 93
XI Glossary 94
XII Appendix 96
iv
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FIGURES
No. Page
1 Cold model test unit 6
2 Cold model test unit solids reservoir 7
3 Cold model test unit simulated heat transfer coils 8
4 Fluidized bed coal combustion unit 10
5 Petrocarb coal injector 11
6 Fluidized bed coal combustor 12
7 Combustor fluidizing grid 14
8 Coal feeder test unit 16
9 Fluidized bed regeneration unit 18
10 Regenerator burner 19
11 Fluidized bed regenerator 20
12 Fluidized bed combustion and regeneration units 22
13 Fluidized bed unit control panel 23
14 Coal feeder orifice 32
15 Effect of injection air flow rate on coal feed rate —
low pressure 40
16 Effect of feed tank pressure on coal feed rate 41
17 Effect of air flow rate on coal feed rate —
high pressure 43
18 Effect of feed tank pressure on coal feed rate 44
19 Fused regenerator bed material and fluidizing grid 51
20 Effect of superficial bed velocity on expanded
bed height — 2 ft. settled bed, simulated heat
transfer coil 54
21 Effect of superficial bed velocity on expanded
bed height — 3 ft. settled bed, simulated heat
transfer coil 55
V
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FIGURES (Cont d)
No.
22 Effect of heat transfer coil on expanded bed height 56
23 Solids transfer rate 58
24 Adiabatic flame temperature versus % (CO + 112) at
8 atm total pressure 72
25 Regeneration studies — effect of pressure on
approach to equilibrium SO 2 concentration 74
26 Regeneration studies — effect of pressure on SO 2
concentration 75
27 Regeneration studies — effect of temperature on SO 2
partial pressure at equilibrium 76
28 Regeneration studies — effect of temperature on SO 2
concentration 78
29 Regeneration studies — effect of temperature on
approach to equilibrium SO 2 concentration 79
30 Regeneration studies — effect of temperature on
CaO/CaS ratio in regenerated solids 86
vi
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TABLES
Page
1 Coal Particle Size Distribution 25
2 Composition of Coal Used in Esso Batch - Fluidized
Bed Combustion Program 26
3 Properties of Limestone and Dolomites 26
4 Coal Feed Rates With a Modified Petrocarb Model
16-1 ABC Injector 33
5 Results of Regression Analysis of Coal Feed
Test Data 45
6 Operating Parameters - Coal Combustion Runs 60
7 Composition of Effluent Streams - Coal Combustion RUnS 61
8 NO-CO Reactions 65
9 NO-CO Reactions - Effect of Pressure and Residence
Time 65
10 NO-CO Reactions - Effect of 02 66
11 Run Summary - Fixed Bed Regeneration Studies 69
12 Summary of Runs With Batch Fluidized Bed Regenerator 70
13 Analysis of Solids from Regeneration Runs 82
A-i Fixed Bed Simulated Combustion Runs 97
A- 2 Air, Fuel and Solids Inputs for Regeneration Runs 99
A—3 Composition of Effluent Stream for Regeneration Runs 100
vii
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ACKNOWLEDGEMENTS
The authors wish to express their appreciation to Messrs. H. R.
Silakowski and W. H. Reilly for performing the laboratory work
described in this report.
We also wish to ackowledge the efforts of the personnel of the
Mechanical Division who were responsible for design and construction
of the pilot plants and who assisted in the startup of the units,
particulary Messrs. E. C. Vath and R. A. Van Sweringen.
We wish to express our appreciation to Mr. S. L. Rakes, the EPA
Project Officer, for his many and varied contributions to the conduct
of this program.
viii
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SECTION 1
CONCLUSIONS
The regeneration of CaSO4 to CaO and S02 appears to be limited by
reaction rates. The measured SO 2 concentration in the product gas from
the regenerator averaged about 40 to 50% of the maximum concentration
calculated by assuming the solids and gases were in chemical equilibrium.
At 10 atm* pressure, the measured SO 2 concentration averaged around 27..
Additional work will be necessary to determine if conditions can be
found to allow a closer approach to equilibrium. T o routes could be
followed — improving the quality of fluidization and increasing the
gas/solids residence time.
High conversion of CaSO4 to CaO was obtained by adding auxiliary air
directly to the fluidized bed. This created adjacent reducing and
oxidizing zones in the bed and reduced the tendency to form undesired
CaS. Addition of a portion of the fuel normally used in the regenerator
reducing gas generator directly to the fluidized bed produced higher
and more uniform bed temperatures without causing bed agglomeration.
Of the S02 released by the combustion of coal, 85 to 90% was retained
by the fluidized bed when calcined limestone or half—calcined dolomite
was used as the bed material at 5—8 atm pressure. However, the runs
were very short because of plugging in the coal injection probe.
Longer runs are necessary before definite conclusions can be drawn
about the suitability of the limestone or the dolomite. The dolomite
(Tymochtee) gave very high attrition rates and may not be suitable for
once—through operation.
Feeding of coal using a modified Petrocarb pneumatic injector at 10 atm
pressure and low feed rates is possible. However, careful attention
must be paid to operating conditions to prevent plugging of the injec—-
tion lines. Additional modifications to the probe design are being
made to reduce further coal plugging in the injection probe.
A number ofcequipmeiit developments have been made incidding the design
and construction of a very versatile burner used to generate heat and
reducing gas for the regenerator. The burner can be operated over a
wide range of pressures, flow rates and oxidizing or reducing condi-
tions. A water cooled fluidizing grid was also developed which
operates in the presence of very high gas temperatures in the regenerator.
* EPA policy is to express all measurements in Agency documents in
metric units. Because implementing this practice will result in
undue cost, NERC/RTP is providing conversion factors for the
particular non—metric units used in this document. For this report
these factors are located on page 95.
1
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SECTION II
RECc MENDATIONS
Work in the area of fluidized bed coal combustion should proceed along
two lines. First, work on a non—regenerative pressurized combustion
system should proceed at this time to speed application of fluidized
bed combustion and prevent any delays which could result from the
development of a regenerative system. Second, work on the regeneration
step should proceed as rapidly as permitted by the availability of
funds, to develop a process which would minimize the disposal burden
of spent limestone.
It is recommended that the pressurized coal combustion program should
first concentrate on the solution of the remaining operational problems.
The effect of precalcination of a typical limestone and dolomite on S02
removal should then be studied. This would be done to study further
the effects of precalcination conditions on stone capacity recently
reported by Westinghouse Research Laboratory. A study would then be
made of the effect of operating conditions on desulfurization effective-
ness, NOx and trace emissions, attrition and utilization of the sorbent,
combustion efficiency and heat transfer rates. The operating conditions
studied would include pressure, temperature, bed depth, fluidizing
velocity and excess oxygen. This study should be made with several
combinations of coals and sorbents (stones). Suggested coals are
Eastern, Western and possibly lignite.
The regeneration program should be based on the use of sulfated lime-
stone or dolomite rather than pure CaSO4. The study should determine
the effects of improving fluidization quality and increasing gas
residence time on the S02 product concentration and the effect of
repeated combustion/regeneration cycles on the activity of the recycled
limestone or dolomite.
The fluidized bed combustion unit could be used further to screen other
coals and sorbents and possibly even shale. It could also be run in
conjunction with the fluidized bed Miniplant to guide the selection of
operating conditions for the Miniplant and to help resolve problems
arising from the operation of the Miniplant.
2
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SECTION III
INTRODUCTION
The fluidized bed combustion of coal is a new combustion technique
which can reduce the emission of SO 2 from the burning of sulfur—con-
taining coals. This is done by using limestone or dolomite as the bed
material. This technique has other potential advantages over conven-
tional coal combustion systems which could result in a more efficient,
less costly method of power generation. High heat transfer coef-
ficients between the fluidized bed and immersed steam generation sur-
faces reduce the steam tubing requirements, and also permit opera-
tion at lower and more uniform bed temperatures, In the vicinity of
1500 to 1700°F. The lower temperatures also reduce NO emissions and
decrease steam tube corrosion. Lower grade coals can also be burned
since the bed temperatures are lower than ash slagging temperatures.
In the fluidized bed boiler, limestone or dolomite is calcined and
reacts with SO 2 and oxygen in the flue gas to form CaSO 4 as shown in
reaction (1). When used on a once—through basis, high limestone or
CaO + SO 2 + 1/202 ‘ CaSO 4 (1)
dolomite feed rates to the boiler are required if SO 2 removal of 90%
or more is to be achieved. In order to reduce the solid waste disposal
burden created by high limestone feed rates, a system is now under
study in which the CaSO4 would be regenerated back to CaO in a separate
fluidized bed reactor by reaction with a reducing gas at a temperature
of about 2000°F as shown in reaction (2). The regenerated CaO would be
CO CO 2
CaSO 4 + H 2 ) CaO + SO 2 + H 2 0 (2)
returned to the boiler where it would again react with SO 2 and 02.
Engineering and cost analyses carried out by Westinghouse Research
Laboratory for the Environmental Protection Agency (EPA) Indicated a
greater commercial potential for a pressurized combustion system when
used in conjunction with a combined gas—steam turbine power generating
plant. Based on theseanalyses, EPA requested Esso Research and
Engineering Company to study the combustion and regeneration steps at
pressures up to 10 atm. A new pilot plant capable of operating at
these conditions was built and operated. The primary purpose of the
pilot plant study was to explore the effect of operating parameters on
combustion and regeneration at pressures up to 10 atm. In addition,
the study was to include measurement of the activity maintenance of
limestone or dolomite due to repeated combustion, regeneration cycles.
Measurement of emission of trace elements from coal combustion was
also to be made. Development and testing of equipment and operating
procedures was also carried out as part of the program A test program
3
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for the operation of the large fluidized bed coal—combustion Miniplant
was also developed as part of the program. In addition, a small amount
of work was done in the cold model fluidization test unit to study
fluidization characteristics and transfer of solids between the com-
bustion and regenerator vessels. This was done in support of the
Miniplant unit construction effort. A short laboratory program was
also carried out in small fixed bed units studying the reduction of
NO by CO under simulated combustion conditions. Some fixed or semi—
fluidized bed regeneration studies were also carried out in small
laboratory equipment.
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SECTION IV
EXPERIMENTAL EQUIPMENT, MATERIALS, PROCEDURES
COLD MODEL TEST UNIT
For the purpose of verifying the design of the pulsed air solids
transfer system used in the design of the fluidized bed coal combustion
Miniplant, and to permit visual observation of its fluidization
characteristics, a cold model test unit (CMTU) was built. The CMTU
is a two vessel fluidized solids system constructed from Plexiglas,
a transparent acrylic plastic. The CMTU was designed to operate at
ambient temperature and pressures up to 60 psig.
Figure 1 is a picture of the CMTU. The unit consists of two 5.5 inch
I.D. vessels simulating the combustor and regenerator of the Miniplant.
The simulated combustor (i.e., the vessel on the left in Figure 1) is
18 feet tall and was assembled from three flanged sections. The regener-
ator is made of two flanged sections and stands 12 feet tall. Both
vessels have bottom plenum chambers and grids for distributing the
fluidizing air and single stage aluminum cyclones for removing entrained
solids. These solids can be returned to the fluidized beds or collected
as desired.
Nylon reinforced transparent PVC hose, 1.5 inch I.D. is used for solids
transfer between the two vessels. This permits visual observation of
solids movement in these lines which connect the upper part of the
vessel transferring the solids to the lower part of the vessel receiv-
ing them. The small receiving—injection pots to which the bottom of
these lines are attached are equipped with fluidizing and pulsed air.
In some of the solids transfer studies, the use of an overflow collector
was investigated as a means of creating an enlarged solids reservoir
in the transmitting vessel. In the Miniplant, such a reservoir would
serve to insure a continued solids seal in the transfer legs and
minimize backflow of gases between reactors. The enlarged reservoir,
shown on the regenerator vessel in Figure 2, was constructed of a
6 inch I.D. aluminum tee with a plate on the inside to serve as a
solids baffle and a Plexiglas window on the outside for observation
of the solids.
In addition to studying solids transfer techniques, the CMTU was used
to determine what effect the hair pin heat transfer ioops designed
for the Miniplant would have on the fluidization characteristics of
the combustor. Figure 3 shows a simulated heat transfer loop section
that was used in the CMTU. It was fabricated of 3/8 inch aluminum
as a prototype half—scale version with the same configuration and
pitch as the Niniplant design. It is 68 inches in vertical length
starting at a point 13.5 inches above the distributor grid. Inser-
tion of the coil reduces the average cross section area by 8%.
5
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Figure 1.
Cold model test unit
6
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ill I
Figure 2.
Cold model test unit solids reservoir
7
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I
Figure 3. Cold model test unit simulated tieat transter coils
8
1
*
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Gas flow rates to each of the two reactots of the CNTU could be varied.
A control panel situated to the left of the Plexiglas vessels per-
mitted measurement of their flow rates and of their individual bed
pressure drops. It also controlled and measured the pressure differ-
ential between beds, and provided a means of varying the on-off pulsing
times and the air flow rates for the solids transfer pots.
FLUIDIZED BED COAL CONBUSTION UNIT
A schematic diagram of the Esso fluidized bed combustion unit is shown
in Figure 4. The primary components of the unit are (1) the coal
feeding system, (2) the fluidized bed combustor, (3) and the gas handl-
ing and analytical equipment.
Coal Feeding Equipment
Figure 5 shows the Petrocarb Model 16-1 ABC injector. The main fea-
tures are a conical-bottom tank that holds solids to be fed and an
orifice and mixing valve assembly that mixes solids with carrier gas.
Solids in the tank are aerated by a controlled stream of air at a
selected pressure. Aerated solids flow through tl orifice at the
bottom of the tank into the mixing valve assembly and are picked up by
a controlled stream of carrier gas (air). Solids are pneumatically
conveyed through a transport line into the receiving vessel. The feed
rate of solids is controlled by pressure in the feed tank, carrier or
injection air flow rate, and pressure differential between the feed
tank and receiving vessel.
The Petrocarb solids feeder shown in Figure 5 was modified to feed
powdered coal (-30 mesh) at rates of 6—28 lbs/hr into the combustor
against pressures of up to 10 attn. The feeder, as ut is supplied by
Petrocarb, can handle only much higher feed rates. The diameter of the
orifice was reduced to 3/16 inch and the Petrocarb injection hose was
replaced with a 1/4 inch diameter (O.D.) x 15 feet long stainless steel
tube. In order to make the feeder work satisfactorily with the batch
combustor, the feeder to combustor presssure differential had to be
held constant. This was accomplished with automatic controls which
maintained the pressure in the feed tank above the pressure in the corn-
bustor by the desired amount. A summary of the coal feeder test pro—
gram and a detailed description of the modifications made on the
Petrocarb injector are given in Section V.
Fluidized Bed Combustor
A schematic diagram of the fluidized bed coal combustor is given in
Figure 6. The vessel was constructed from four sections of 10 inch
standard wall carbon steel pipe and refractory lined with Gref cc
Litecast No. 7528 to an inside diameter of 4 inches. The height of the
vessel, above the fluidizing grid, was about 16 feet. Below the grid
was a 2 foot long burner section lined with Bubbalite. The fluidizing
grid, which was made of stainless steel, had seven air distributor caps
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VENT
CYCLONES
DRAIN
REFRIGERATOR
C
COOLER
COOLING
COILS
WATER
WATER
DEMINERALtZER
FLUIDIZING
GRID
AIR FROM
COAL
FEEDER
PROPANE
WATER
DRAIN
INJECTION AIR
Figure 4. Fluidized bed coal combustion unit
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EXHAUST VALVE
Figure 5. Petrocarb coal injector
PRESSURE RELIEF
VALVE
LINE PRESSURE
GAUGE
DILUTER PRESSURE
GAUGE
DILUTER PRESSURE
REGULATOR
FEED VALVE
FILLING VALVE
TANK PRESSURE
GAUGE
REGULATOR
VALVE
GAS
SHUTOFF
VALVE
MIXING
-DILUTER HOSE
INJECTION HOSE
TO RECEIVING VESSEL
BUSHING
‘QUICK DISCONNECT
COUPLING
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GRE FCO LITECAST #7528
COOLING COILS (3 PAIR — EACH PAIR
CENTERED ABOUT FLANGE)
COOLING
WATER
INLET
2’ 0”
2’ 3’
COOLING WATER OUTLET
PORT
FLUIDIZING GRID (WATER COOLED)
BUBBALITE
BURNER
Fluidized bed coal combustor
THERMO-
COUPLE
8’lO”
10” STD. WALL STEEL PIPE
SOLIDS CHARGING PORT
F LAN GE
Figure 6.
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and was water-cooled. it is described in Figure 7. A propane burner,
located at the bottom of the burner section, was used to preheat the
unit to above the self—ignition temperature of coal. The burner is
described in more detail in Section V. The maximum operating tempera-
ture and pressure of the batch combustor were 1900°F and 10 atm.,
respectively. Two air compressors supplied sufficient air for opera-
tion at a superficial velocity of 6 ft/sec at 10 atm and 1600°F (about
80 SCFM).
Bed temperature was controlled by three pairs of serpentine wound
cooling coils made of 1/4 inch diameter stainless steel tubing.
The total surface area was about 6.3 sq. ft. The center of each
pair of coils was located at a flange in the shell of the corn—
bustor. Water flowed in parallel to each of the six coils; flow rate
to the lower and upper pair of coils was manually controlled but flow
to the center pair was automatically controlled to give the desired
bed temperature. Thermocouples were located 6 inches apart in the
lower sections of the combustor and 12 inches apart in the upper
section.
Solids were loaded into the combustor through a charging port located
in the upper section. Solids could be removed through a port in the
lower section or, alternatively, transferred directly to the regenera-
tor by blowing them through a 2 inch pipe supplied for this purpose.
Coal entered through a water-cooled coal injection probe to which was
connected the end of the 15 feet x 1/4 inch diameter coal injection tube.
The probe was a 14 inch length of the 1/4 inch diameter stainless steel
tubing surrounded by a jacket through which cooling water flowed.
Gas Handling and Analytical System
Flow of air and fuel into the combustor, and system pressure, were under
automatic control. Gases leaving the combustor first passed through
two cyclones, which removed entrained solids and fly ash. An off-gas
cooler, which followed the cyclones, reduced the temperature of the
off-gas to the desired level. The off-gas then entered an expansion
coil of 1 inch diameter stainless tubing which was electrically heated
during startup to raise the temperature of gases above the dew point.
A 1.5 inch Aerotec cyclone, following the heater, was used to remove
particulates during startup of the combustor, when water vapor condens-
ing in the first two cyclones caused them to operate at reduced effi-
ciency. Fine particulates were removed with a Pall Model MD-600 filter,
located upstream of the back-pressure control valve. This filter had a
mean pore size of 165 microns and an area of 0.73 sq. ft. Before being
vented, the off-gas entered a chiller and knockout to remove moisture
so that the water content of the gas could be determined. Alternatively,
the gas could be sent directly to a second knockout at room temperature
and vented. A small portion of the off-gases were diverted to a refrig-
erator which lowered the dew point of the gas to about 35°F before it
was sent to the gas analysis equipment.
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WATER OUTLET
5 ”— 5 ’”SLOT
16 16
WATER-COOLING
CHANNELS 5,I
(Sj w\ 16
‘16)
#47 DRILL, 8 PLACES
EQUALLY SPACED
32
AIR DISTRiBUTOR CAP
304 S.S., 7 REQ’D.
THERMOWELL(” ”)
GASKET SEALING GROOVES
(BOTH SIDES)
WATER COOLED GRID PLATE
304 S.S., 12 ” DIA., “ THICK
Figure 7. Combustor fluid izing grid
WATER tNLET
14
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Gas analysis equipment included Beckman Model NDIR 315 analyzers for
SO 7 , NO, and CO, a Beckman Model 715 polarographic analyzer for 02,
ana a LIRA Model 300 analyzer for CO 2 . The concentration ranges of the
analyzers used with the combustor are given below.
So 2 0-3000 ppm
Instrument No. 1 Instrument No. 2
CO 0-1250 ppm 0-10 per cent
0-6250 ppm 0-30 per cent
NO 0-500 ppm
0-1000 ppm
0-2500 ppm
02 0-5 per cent
0-25 per cent
Co 2 0-25 per cent
Coal Feeder Test Unit
Figure 8 is a diagram of the apparatus used to test the Petrocarb
coal feeder before use with the combustor. Air was used to pressurize
the Petrocarb coal feeder and transport coal through the injection tube
to the receiving vessel. The Petrocarb unit was equipped with pressure
regulators for both the feed tank and injection air, a control valve
for injection air, and a rotameter to indicate total air flow. The
receiving vessel was mounted on a scale which indicated weight changes
as small as 001 lb. Pressure in the receiving vessel was set at the
desired value by means of a back pressure regulator. A differential
pressure cell measured i P between the coal feed tank and receiving
vessel. House air for the feeder first entered a bed of Drierite to
reduce its moisture and thus keep the coal dry. Air leaving the receiv-
ing vessel was passed through a 65 filter to remove coal fines.
Experimental runs were first made at pressures of 70 psig and below in
the feed tank. After some operating experience was obtained, modifica-
tions wa made in the equipment to permit operation at pressures of 135
psig in the feed tank or about 10 atmospheres (absolute) in the receiv-
ing vessel. For operation at higher pressures, the Petrocarb unit was
connected to an air compressor. Air from the compressor passed through
two silica gel driers to keep moisture reaching the coal to a minimum.
The original receiving vessel, rated for 100 psig, was replaced with
one good for 250 psig. Heavy rubber hose was substituted for Polyflow
tubing for runs made at higher pressures. Pressure regulators supplied
with the Petrocarb feeder had a 125 psig limit and were replaced with
regulators designed for higher pressures. Air from the compressor was
stepped down in pressure with an additional regulator.
15
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Figure 8. Coal feeder test unit (section within dashed lines provided by Petrocarb)
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FLUID IZED BED REGENERATION UNIT
Figure 9 is a schematic diagram of the Esso fluidized bed regeneration
unit. The main components of this unit are (1) the burner, (2) the fluid
bed regenerator, and (3) the gas handling and analytical equipment.
Burner
The burner, shown in Figure 10, was located at the bottom of the regen-
erator, under the fluidizing grid. It produced heat and a reducing gas
mixture containing CO and H 2 . The main parts were a tube to mix the
fuel (propane + propylene) with air and a water-cooled grid to hold the
flame.
The tubular part of the burner was made of 304 stainless pipe, 2 inches
in diameter and 5.5 inches long. At the bottom of the tube were two
baffles which supported a bed of 1/8 inch alumina beads and also helped
to prevent channeling of the air-fuel mixture up the sides of the tube.
Air and fuel entered at a tee below the burner tube; fuel entered the
tee through a sparging tube containing eight—3/64 inch diameter holes at
its tip to promote even distribution. The alumina beads further mixed
the air-fuel mixture.
The burner grid was a 3 inch diameter x 1/4 inch thick brass plate
which screwed onto the top of the burner tube. The grid was water-
cooled and contained about 400—1/32 inch holes through which the air—
fuel mixture flowed. On the underside of the grid was a disk of porous
stainless steel which acted as a flame arrestor and prevented flashback
into the burner tube.
The burner was ignited by a pilot light which was positioned near the
side of the burner tube below the burner grid. The pilot light, in
turn, was ignited by an electrode. This arrangement kept the ignition
equipment away from the flame of the burner, where it would be damaged
by excessive heat.
Batch Fluidized Bed Regenerator Vessel
Figure 11 shows the batch fluidized bed regenerator. The vessel was
constructed from three sections of 12 inch carbon steel pipe and refrac-
tory lined with Grefco Litecast No. 7528. The interior of the regen-
erator was lined with a 3.25 inch diameter alumina tube. The height
of the unit above the fluidizing grid was about 16 feet. A burner sec-
tion, about 2 feet long, and lined with Bubbalite, was located under the
grid. The burner was mounted on a flange which attached to the bottom
of the burner section. Air and fuel required for regeneration entered
the regenerator at the bottom of the burner. However, for some runs,
a portion of the fuel and secondary air were added at various points
above the fluidizing grid.
17
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VENT
REFRIGERATOR
COMPRESSED AIR
(FILTERED, DRIED)
INFRARED ANALYZERS
FRC
OFF-GAS
COOLER
AIR
CONTROL
VALVE
OFF-GAS
CHILLER
DRY
GAS
METER
FLUIDIZING GRID
BURNER
PROPANE
KNOCKOUT
DRAIN
Figure 9. Fluidized bed regeneration unit
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3” DIA.
BURNER HEAD
(BRASS)
GASKETS
COOLING COIL
SCREEN
400 — HOLES
WATER COOLING
CHANNEL
POROUS METAL DISK
2” DIA. S.S. TUBE, 2 LONG
ENTIRE TUBE FILLED WITH
ALUMINA BEADS
4 + 4 4 4 AIR&FUEL
Figure 10. Regenerator burner
FLANGE
BAFFLES
19
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GREFCO LITECAST #7528
5’lO” ____ — LD.
THERMOCOUPLE OLIDS CHARGING PORT
_— 12 x 0.375 WALL STEEL PIPE
— = = = — FLANGE
=
5’ 0
-
5’ O
SOLIDS REMOVAL PORT
— = = ..— FLUIDIZ1NG GRID (WATER-COOLED)
2’ 3”
— BUBBALITE
SIG HTG LASS
Figure 11. Fluidized bed regenerator
20
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The maximum operating temperature and pressure of the batch regenera-
tor were 2100°F and 10 atm., respectively. Thermocouples, spaced 6
inches apart, were used to judge bed temperature. Besides air and
fuel, N 2 and CO 2 could be fed into the reactor. For runs in which
secondary air was not added above the fluidizing grid, °2 was added
near the top of the regenerator to convert elemental sulfur, which
might plug lines, to So 2 . All gas flows and system pressure were under
automatic control.
Air could be preheated before entering the burner by using an electri-
cal air preheater with 7.5 kilowatts capacity. Kanthal wire was wound
around the outside of the ceramic tube in each of the two lower sec-
tions of the regenerator. Each heater had a capacity of 7.5 kilowatts
and could be used to provide extra heat to the bed.
The fluidizing grid for the regenerator was made of 304 stainless
steel and contained 69-5/64 inch holes for the passage of fluidizing
gases. A five inch square at the center of the grid was water-cooled.
Solids were charged into the regenerator through a port located near
the top of the unit. Solids could be removed through another port
located just above the fluidizing grid.
Gas Handling and Analytical System
Gases leaving the regenerator were passed through a cyclone to remove
entrained solids and then to an off-gas cooler which reduced tempera-
ture to the desired level. Downstream of the off-gas cooler, the
regenerator shared equipment with the batch combustor. Thusly, the gas
stream entered a solids filter, back pressure control valve, chiller
and knockout before being vented. The analytical instruments, too, were
shared with the combustor. However, an additional Beckman Model
NDIR 315 analyzer was used to determine the concentration of SO 2 in the
0-157. range.
Air for the regenerator was supplied by a compressor. Other gases came
from cylinders. Fuel cylinders (propane-propylene) had to be kept in a
steam-heated hot-box at 120°F in order to maintain sufficient vapor
pressure. In addition, fuel lines near the regenerator were heated to
prevent fuel from liquefying. 1 igure 12 is a photograph of the fluidized
cotnbustor and regenerator. Figure 13 is a photograph of the control
panel used to operate the units.
LABORATORY FIXED BED U TS
Steel Reactor
A 1 inch 0.D. stainless steel reactor (0.75 inch I.D.) was used for
fixed bed studies examining reactions between NO and Co under simulated
combustion conditions. The reactor was capable of operating at temper-
atures up to 1600°F at pressures up to 10 atm. Feed gases were purch-
ased with various concentrations of NO, CO, C0 2 , and 02 in N 2 or A
21
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Figure 12.
Fluidized bed combustion and regeneration units
COMBUSTOR r
REGENERATOR I
T 4
i — W I —
‘WV
22
-------
- —___p — I I
ci
-
#lI
It . 1
‘ v -s’ S
S S •
L .J.
I ‘. \ !• l \
I
Figure 13.
FLuldized bed unit control panel
-------
diluent. Provisions were made to blend two gases before passing the
gas mixture through the reactor. Rotameters were used to meter the
gases. The reactor was packed with alumina beads which acted as a pre-
heat section, followed by the bed material. The bed material was
calcined limestone (1359) or dolomite (1337). The entire reactor
was held in an electrically heated furnace. The off gases were anal-
yzed for NO and CO by Beckman Model 315 NDIR analyzers. CO 2 was measured
by an MSA Model 200 NDIR analyzer. The gases were cooled to 35°F before
analysis to remove water. A Beckman Model 715 polarographic analyzer was
used to measure 02 concentrations in the product. The blended feed gas
could also be diverted to the analyzers for measurement.
Ceramic Reactor
A one inch I.D. ceramic reactor was used to study regeneration of
CaSO 4 at temperatures up to 2000°F and pressures up to 9.5 atm. The
reactor consisted of a one inch I.D. ceramic tube mounted axially in a
3 inch Schedule 5 304 stainless steel pipe.
The reactor was heated electrically by resistance heater windings on
the external surface of the ceramic tube. The overall length of the
ceramic tube was 28 inches but only the mid 24 inch section was heated.
For most of the runs, a 40 gin charge of CaSO4 (anhydrite) or partially
sulfated lime was used and was placed about 17 inches above the bottom
of the ceramic tube. The lower section of the ceramic tube was filled
with ceramic beads and served as a preheat section for the feed gases
flowing up through the beads. Temperature was measured by a therrno-
couple inserted in a 1/4 inch O.D. ceramic well located axially in the
center of the 1 inch reactor. A second thermocouple in this well oper-
ated the temperature controller.
The feed gas was produced by metering a CO/N 2 blend and a C0 2 /N blend
to give the desired feed gas composition and flow rate. Rotame ers
were used for gas flow measurement. The product gas was then cooled,
passed through a filter, back pressure regulator, and chiller before
analysis for SO 2 content in a Beckman Model 3l5A infrared analyzer.
24
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NATERIALS
C oa 1
The coal used in the batch fluidized bed combustor, and in the test
program for the coal feeds, was a high volatile (A) bituminous coal
from Consolidation Coal Company’s Arkwright mine in West Virginia.
It was ground to —30 mesh by Peun—Riliton Co. The specified particle
size distribution is given in Table 1. A proximate and ultimate
analysis was made on each of two samples of coal. Sample No. 1
was taken directly from the fifty pound bag in which it was supplied;
sample No. 2 as taken from the receiving vessel used in the test
program for the coal feeder, after it had been fed from the Petrocarb
injector. The results of the analyses of both samples are given in
Table 2.
TABLE 1. COAL PARTICLE SIZE DISTRIBUTION
Penn—Rillton Co. Grind B—2 Specification
U. S. mesh size 10 20 30 40 100 200 pan
wt. fraction on screen 0 4.5 15.5.14 35.5 12.5 18
Limestone and Dolomites
A limestone and two dolomites were used in the experimental studies.
The stones and their properties are given in Table 3. Various particle
size ranges of these materials were used and are specified in the sections
whera the experimental results are described.
Ca S 04
A high purity form of anhydrous CaSO 4 was used in the regeneration
studies. The material was purchased from W. A. Hammond Co. under the
trade name “Drierite” in 10 x 20 mesh particle size range.
EXPERIMENTAL PROCEDURES
Cold Model Test Unit
The measurements of bed expansion ratios, quality of fluidization and
minimum fluidization velocity were made visually.
The rate of solids transfer was evaluated by two methods. In one
method solids transfer was permitted only from the prototype regenerator
to the prototype combustor, and the increase in solids inventory in the
combustor was measured over a timed interval. In the second method,
transfer was accomplished in both directions, and a charge of painted
25
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TABLE 2. COMPOSITION OF COAL USED IN ESSO
BATCH—FLUIDIZED BED COMBUSTION PROGRAM
Source : Consolidation Coal Co. - Ark right Mine
Proximate Analysis
Wt. Percent
Sample #1 Sample #2
Component ( bag) ( receiving tank )
Moisture 1.00 1.03
Ash 8.11 8.01
Volatile matter 36.86 36.69
Fixed carbon 54.03 54.28
Ultimate Analysis
Wt. Percent
Sample #1 Sample #2
Component ( bag) ( receiving tank )
Moisture 1.00 1.03
Ash 8.11 8.01
Total carbon 76.26 76.67
Hydrogen 5.30 5.30
Sulfur 2.66 2.49
Nitrogen 1.49 1.47
Chlorine 0.07 0.08
Oxygen (by difference) 5.11 4.95
Higher Heating Value, Btu/lb.
Sample No. 1 Sample No. 2
14,045 14,070
TABLE 3. PROPERTIES OF LIMESTONE AND DOLOMITES
Quarry Stone Chemical Analysis, Wt.
Designation source — _ type 2 1 P Q2 _ 223 223
1359 Grove Lime Co. Limestone 97.0 1.2 1.1 0.3 0.2
(Stephen City, Va.)
1337 Chas. Pfizer Co. Dolomite 54.0 44.0 0.9 0.2 0.3
(Gibsonburg, 0.)
Tymochtee C. F. Duff & Sons Dolomite 53.8 38.7 5.3 0.9 1.2
(Huntsville, 0.)
26
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limestone was added just after the combustor transfer catch pocket to
serve as a tracer. The time required for the painted limestone to
reach the regenerator was measured, and knowing the volume of the transfer
line, the transfer rate was calculated.
The stone used in the studies was a wide particle size distribution of
calcined limestone (-6 to +20 U.S mesh) representative of solids from
the limestone grinding equipment.
Coal Feeder Test Unit
Before a run, the coal feeder was checked to see that it contained 50-100
lbs. of coal. The feeder and receiving vessel (see Figure 8) were then
pressurized and the back pressure regulator located downstream of the
receiving vessel was adjusted so that the pressure in the vessel was less
than the pressure in the feeder by the desired amount (routinely 3 psi)
Flow of injection air was started by setting the injection air pressure
regulator to about ten psi above the feed tank pressure, opening the
injection air shutoff valve, and adjusting the metering valve until the
desired air flow was indicated on the rotameter. The vibrator, which
was attached to the feeder directly above the mixing tee, was turned on.
If necessary, the back pressure regulator on the receiving vessel was
readjusted to bring the feed tank to receiver t P to the desired value.
Coal feed was started by opening the coal feed valve. Weight of the
receiving vessel was recorded, typically every minute, so that the coal
feed rate could be determined.
To stop the flow of coal, the coal feed valve was shut. Injection air
was left on for several minutes after the feed valve was closed so that
the mixing tee and injection tube could empty. If injection air was
stopped with coal remaining in the tee or injection tube, a plug
usually formed when the feeder was restarted.
Fluidized Bed Coal Combustion Unit
Operation of the batch fluidized bed combustor can be divided into four
phases: startup, ignition and pre-heating, coal feeding, and shutdown.
Startup consisted of those activities preliminary to ignition of the
propane burner. These activities included checking equipment to make
sure it was ready for a run, checkout of the analyzer calibratLon,
charging solids, turning on electrical circuits and the air compressor,
turning on all cooling water systems (fluidizing grid, burner, coal
probe, steam coils, condenser) and purge nitrogen systems (pressure
taps, sight—glasses, d/p cells).
To ignite the propane burner, air and fuel flows were set and the
ignition electrode was activated. Safety devices shut down all flows
if ignition was not obtained within ten seconds or if a flame—out
occurred afterwards. A safety interlock prevented startup for 3 min-
utes after an automatic shutdown to assure adequate purging of the
combustor. Subsequent to ignition, air flow and coinbustor pressure
27
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were adjusted to the values desired for making the run. All gas flows
and pressure were controlled automatically.
Preparation of the coal feed system for a run consisted of setting the
flow of injection air and activating and adjusting the coal feeder to
combustor t P control system. Coal injection could be started only
after the temperature in the combustor was high enough for self—ignition
of the coal to occur. Propane flow to the burner was stopped automat-
ically at the same time that feeding of coal was started. An automatic
safety circuit would shut down coal injection if the combustor tempera-
ture dropped too low to ensure combustion of the coal or if the feeder
to combustor Al ’ dropped below a pre—set minimum (about one psi).
Data on the weight of the coal feeder vs. time was taken so that the feed
rate of coal could be determined. Another method of estimating the feed
rate was to observe the oxygen concentration in the off-gas from the
combustor. A rapid rise in oxygen concentration was usually the quickest
way of determining that a problem was developing with the coal feeding
System.
Temperatures in the combustor could be adjusted by regulating the amount
of water entering the three pairs of cooling coils. The feed rate of
coal could be adjusted by changing the flow of injection air.
To shut down the combustor routinely, the coal feed valve was shut,
fluidizing air was stopped, and a nitrogen purge was started to preserve
the solids. Flow of injection air was kept on for several minutes so
that the coal feed line could be cleared of coal. All water flows were
reduced. Solids could be discharged from the reactor (by blowing them
out of a port located just above the fluidizing grid) after the unit
bed cooled overnight.
Fluidized Bed Regeneration Unit
Operation of the regenerator was similar to the combustor except that no
coal feeding was involved. Preliminary startup of the unit included
checkout of the calibration of the analytical instruments, activation of
electrical systems and air compressor, check of the propane system, and
turning on all cooling water systems.
The empty regenerator was first preheated to minimize condensation of
water after solids were added. The burner was then shut off and solids
were charged through a port located near the top of the unit. Reheat-
ing of the bed was then begun and pressure and air flow were adjusted
to meet the conditions selected for the run. If desired, air could be
electrically preheated. When the bed was near the temperature of the
run, reducing conditions were established by increasing the flow of
fuel. Additional fuel could be added through the burner or, alterna-
tively, directly into the bed. Auxiliary air could also be added,
higher in the bed, to reduce formation of CaS. When auxiliary air was
not used, it was necessary to provide a small stream of 02 near the top
of the regenerator to oxidize any sulfur which formed and which might
otherwise plug lines.
28
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Shut down was accomplished by stopping the flow of fuel and air and
reducing the pressure in the system to atmospheric. A nitrogen purge
was applied to prevent combustion of any carbon deposits in the reactor
and to prevent moisture from reaching the solids until they had cooled
sufficiently to be removed (overnight).
Laboratory Fixed Bed Units
Steel Reactor —
The reactor was packed with 40 gm of calcined limestone or dolomite and
inert alundum beads. The bed was heated to the operating temperature
in the presence of N 2 . The proper blend of inlet gases was then made,
analyzed in the continuous IR analyzers and then admitted downf low into
the reactor. The outlet gases were continuously analyzed.
Ceramic Reactor —
The lower section of the reactor was first packed with Alundum beads to
serve as the preheat section. For most runs, 40 gm of bed material was
then placed above the preheat section. In some runs, the bed charge was
increased to 60 gin. After heating to the proper temperature in the pre-
sence of a stream of N 2 passing upf low through the bed, the feed gases
were blended to give the proper composition and then admitted upf low
into the reactor. The product gases were analyzed continuously. Since
this unit was used to study regeneration of sulfated material, the pri-
mary component in the product gas was S02 formed by the reduction of
the sulfated bed material. The S02 concentration in the product gas
decreased as the bed material was reduced, and the runs were terminated
when the SO 2 concentration decreased to a very low level. The time of
the runs ranged from 1 to 2 hours, during which time five to ten times
the stoichiometric amount of reducing gas was passed over the bed. After
the bed cooled, the charge was removed, weighed and analyzed for total
sulfur content.
29
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SECTION V
DEVELOPMENT OF EXPERIMENTAL EQUIPMENT AND PROCEDURES
COAL FEEDER
Introduction
The coal feeding system for the pressurized batch—fluidized bed combus—
tor must be capable of metering 6—28 lbs./hr. of powdered coal (—30
mesh) into the combustor against pressures of up to 10 atmospheres.
A Petrocarb ABC injector, which is a small version of the solids
feeder to be used with the combustor of the Miniplant, was chosen
to feed coal into the batch combustor. An experimental program
was undertaken to test the suitability of the Petrocarb feeder
f or doing this. The purpose of the program was to determine what
modifications of the Petrocarb system were needed to make It operate
satisfactorily and to gain experience in feeding coal prior to startup
of the combustor.
Argonne National Laboratory had tested the suitability of the
Petrocarb ABC injector for feeding 20—40 lbs./hr. of coal, limestone,
and mixtures of coal and limestone, into a bench scale combustor—
regenerator (1). Coal feed rates averaged 31.6—45.7 lbs./hr. for
runs at atmospheric pressure to 97.8—179.3 lbs./hr. for runs with
75—82 psig pressure in the receiving vessel. Many runs were character-
ized by feed rates that fluctuated widely. Argonne terminated work
with the Petrocarb feeder when they concluded that it could not feed
coal at steady rates in the flow ranges they required. Argonne made
only minor modifications to the system supplied by Petrocarb. At
Esso, it was hoped that the Petrocarb system could be made to perform
satisfactorily if additional modifications were made.
Theory of Operation
Operation of the Petrocarb solids feeder depends on gravity flow of
solids through an orifice and subsequent pneumatic conveying of the
solids to the receiving vessel. These two operations have been studied
by others with many different types of solids (2). For example, it
is known that the reduction in “viscosity” of a bed of powder will
result in an increased flowrate through an orifice. This reduction
in viscosity can be accomplished by pressurizing the bin by means of
a connection atop a closed vessel or by adding an aeration bleed at
the base of an open bin. Either procedure results in a small stream
of gas leaving with the solids through the orifice. This additional
stream of gas “lubricates” the flow at the orifice and reduces the
viscosity at this point. By means of this principle, solids can
30
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be made to flow through smaller openings than would be possible by
simple gravity flow. Increasing the pressure, as can be done in the
Petrocarb solids feeder, increases the lubricating effect at the
orifice and permits solids to flow at a greater rate.
Solids flow through the orifice at the bottom of the Petrocarb feeder
and enter a mixing tee where a moving stream of air picks up and
conveys the solids through a tube to the receiving vessel. The
relationship between pressure drop per unit length of tube and
superficial velocity is similar to that obtained for a tube with gas
only flowing. However,pressure drop is higher due to the force
required to maintain the solids in suspension and move them along with
the gas stream. At constant superficial velocity, pressure drop
increases with increasing solids loading. In order to maintain
a given feed rate of solids through the tube, the superficial gas
velocity must not fall below a certain critical value, called the
saltation velocity. Below the saltation velocity, which is a function
of solids flow rate for any particular gas—solid system, particles
settle out in the tube.
It is interesting to compare the saltation velocity for horizontal
transport with the minimum velocity required to transport solids
vertically, called the choking velocity. Experimental data indicates
that saltation velocities and choking velocities are identical for
uniform sized particles. However, for mixed size material saltation
velocities are three to six times as great as choking velocities.
This is in accord with numerous qualitative statements appearing
in the literature on pneumatic transport to the effect that the
velocity required to convey materials horizontally is several times
as great as that required to convey vertically.
Experimental Results
Work with the Petrocarb solids feeder was divided into two sections.
The purpose of the initial work was to modify the feeder so that
it could deliver coal smoothly and dependably at the required rates.
The modified equipment would be used to collect data on feed rate
of coal vs. feed tank pressure, feed tank—receiving vessel pressure
differential, and injection air flow rate. Table 4 gives results of
68 runs made with the modified Petrocarb Model 16—1 ABC injector.
A 3/16 inch orifice was used for all runs except no. 22, in which
a 7/32 inch orifice was used. After run no. 22, the taper of the
original 3/16 inch orifice was made steeper (see Figure 14).
Modification of Petrocarb Feeder
The Petrocarb ABC injector was designed for much higher solids feed
rates than 6—28 lbs./hr., which is required for use with the combustor.
Therefore, modification of the unit to permit lower feed rates was
31
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11716” DIA.
3/16” DIA.
MATERIAL: 316 STAINLESS
STEEL
Figure 14. Coal feeder orifice
3/64”
32
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Table 4. COAL FEED RATES WITH A MODIFIED PETROCARB MODEL 16—i ABC INJECTOR
Orifice size = 3/16 inch
Coal = —30 mesh
Receiving
Feed tank vessel Air flow Average
Run pressure, pressure, rate, coal feed
No. Injection Hose psig psig SCFM rate, lb/hr. Comments
1 1/4 in. x 25 ft. 5 0 1.0 1.0 Injection hose wrapped around receiver.
Polyflow Plugged after 15 mm.
2 20 0 2.0 7.8 Injection hose plugged after 35 mm.
3 1/2 in. x 25 ft. 5 0 1.0 10—44 Hose wrapped around receiver. Erratic
Rubber (Petrocarb flow. Plugged after 13 mm.
supplied)
4 5 0 1.0 20
5 5 0 1.6 6.5 Hose in 30 inch loops on floor. Ran
well.
6 10 0 2.3 45
7 20 0 2.3 125
8 5 0 2.3 7.6 Conditions of Run No. 5 repeated.
9 1/4 in. x 25 ft. 20 0 2.3 8.0 Hose in 30 inch loop on floor. Flow
Polyf low stopped after 19 mm.
10 20 0 2.3 21 Flow stopped after 9 mm. High static
charge on injection hose.
11 20 0 2.6 17 Flow stopped after 10 mm.
12 1/4 in. x 15 ft. 10 0 2.3 12.3 Injection hose in one large loop on
copper floor.
13 5 0 1,6 7.3 Coal flow increased when injection air
1,0 28 decreased.
-------
19
20
21
22
40
60
40
40
70
70
60
70
36 2.0
56 2.0
38 2.0
1.3
1.0
0.7
36 2.0
67 2.0/1.3
67 1.3
57 1.3
67 2.0
21.1
24.0
7.2
22.3
25.1
34
12.8
16.8/37.0
32.0
26.6
20.2
Table 4 (continued). COAL FEED RATES WITH A MODIFIED PETROCARB MODEL 16—1 ABC INJECTOR
Receiving
Feed tank vessel Air flow Averaged
Run pressure, pressure, rate, coal feed
No. Injection Hose psig psig SCFM - rate, lb/hr. Comments
14 1/4 in. x 15 ft. 20 16 1.6 18.0
copper
15
16
17
18 1/4 in. x 15 ft.
copper
1/4 in. x 15 ft.
Stainless steel
23
24
25
26
27
28
29
Vibrator used.
Vibrator used.
Coal flow increased as injection air
decreased. Vibrator used.
Ran only 10 minutes.
7/32 inch orifice used for this run.
Taper of orifice made steeper.
Higher coal rate ran only 3 mm.
70
70
60
40
40
70
70
67
67
57
37
37
67
67
1.3
2.0
2.0
1.3
2.0
2.6
1.6f1.3
28.7
18.6
16.5
19.8
8.2
9.8
24/31.6
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Table 4 (continued). COAL FEED RATES WITH A MODIFIED PETROCARB MODEL 16—1 ABC INJECTOR
Run
No.
30
Injection Hose
1/4 in. x 15 ft.
stainless steel
Receiving
Feed tank vessel Air flow
pressure, pressure, rate,
psig psig SCFM
70 67 2.6
Average
coal feed
rate, lb/hr.
8.3
Comments
31
70
67
2.6
13.8
Bored out tee—injection tube connector.
32
40
37
2.0
5.8
33
60
.
57
2.0
15.4
34
60
57
1.3
30.6
35
40
37
1.3
21.9
36
60
57
2.3
8.0
L i i
37
38
70
20
67
17
2.0
1.3
24.3
10.1
39
20
17
0.65
29.9
40
60
57
1.6
28.8
41
40
37
1.0
35.4
Frequent plugging
42
40
37
1.6
16.3
43
70
67
2.3
15.8
New Syntron vibrator
installed.
44
70
67
1.6
25.9
45
70
67
1.3
40.9
46
60
67
1.6
28.6
47
70
67
1.6
29.1
48
60
57
1.3
34.6
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Table 4 (continued). COAL FEED RATES WITH A MODIFIED PETROCARB MODEL 16—1 ABC INJECTOR
Receiving
Feed tank vessel Mr flow Average
Run pressure, pressure, rate, coal feed
No. Injection Hose psig psig SCFN rate, lb/hr. Comments
49 1/4 in. x 15 ft. 121 118 2.6 27.5 Pressure at rotameter = 155 psig
stainless steel
50 135 132 3.5 15.7
51 135 132 4.0 6.1
52 135 132 3.1 25.6
53 106 103 2.6 19.4
54 106 103 2.6 26.3 Adjusted calibration of DP cell—was off
slightly for runs 49—53.
55 135 132 3.5 12.6
56 121 118 3.1 17.7
57 106 103 3.1 17.1
58 91 88 2.2 24.9 Cleaned filter before run.
59 121 118 3.1 17.4
60 135 132 3.1 23.0
61 135 132 4.0 6.9
62 121 118 2.6 26.1
63 121 118 3.5 9.7
64 106 103 3.5 4.4
65 106 103 3.1 8.8—12.7, Problems with air flow.
avg. 9.9
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Table 4 (continued). COAL FEED RATES WITH A MODIFIED PETROCARE MODEL 16—1 ABC INJECTOR
Receiving
Feed tank vessel Mr flow Average
pressure, pressure, rate, coal feed
_______________— psig psig SCFM rate, lb/hr. Comments
91 88 2.6 15.4 Minor problems with air flow.
135 132 3.7 11.3 Minor problems with air flow.
135 132 3.3 20.9 Minor problems with air flow.
Run
No. Injection Hose
66 1/4 in. x 15 ft.
stainless steel
67
68
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necessary. Factors which affect feed rate are size of the orifice
(5/16 inch supplied) and size of the tube connecting the feed tank
with the receiving vessel (1/2 inch dia. x 25 feet hose supplied).
An important part of the experimental program was to find that
combination of orifice and injection hose size which would permit
satisfactory operation over the range of required flowrates.
The first task was to establish a satisfactory size for the orifice.
Orifice sizes of 1/8 inch, 5/32 inch, 3/16 inch, and 7/32 inch were
tried but, with the two smaller orifices, coal would either not flow
at all or flow intermittently upon tapping the orifice assembly.
The two larger orifices plugged occasionally, but the 7/32 inch size
produced coal feed rates which were too high. Thus, 3/16 inch was
the optimum size to be used with the coal on hand (—30 mesh). After
run no. 22, the taper of the orifice was made steeper. This reduced
plugging.
Several injection tubes were tested. The first was a 3/16 inch x 25 feet
copper tube (I.D. = 0.1275 inch) which was coiled around the receiving
vessel so that the diameter of each coil was about 18 inches. The
tube plugged frequently, and it was difficult to remove the plugs
as this required disconnecting the tube, straightening it, and tapping
an end to remove the plug. Polyflow tubing of 1/4 inch x 25 feet size
(I.D. = 0.17 inches) was subsequently tried and had the advantage
that it was translucent, so that plugs could be easily spotted,
and that it could be readily put into any desired shape. When this
tubing was wrapped around the receiving vessel, plugging usually
occurred after only a few minutes. However, when placed in large
(30 inch) loops on the floor, plugging occurred much less frequently.
A 12 foot length of Polyf low tubing was also tested and gave
considerably higher coal feed rates than did a 25 foot length. Use
of Polyf low tubing was terminated because a static charge developed
on the tubing during operation, which contributed to plugging.
A copper tube, 1/4 inch x 15 feet (0.19 inch ID.), formed into one
large loop was tried next. Decreasing the tube length made it easier
to make a single loop and less plugging occurred than in longer tubes.
After run No. 14, a vibrator was attached to the mixing assembly at
the bottom of the feed tank. The vibrator caused both the mixing
assembly and the entire injection tube to vibrate and reduced plugging
at both the orifice and injection tube. Plugging had continued to
be a problem with the 1/4 inch x 15 feet (I.D. 0.19 inch) copper
tube, and it was replaced with a 1/4 inch x 15 feet (I.D. = 0.18 inch)
tube of stainless steel. The stainless steel tube proved to be
a significant improvement over the copper since it plugged less
frequently. The improvement is probably due to the smoother surface
of the stainless tube, which, reduces the tendency of the coal to
stick.
Plugging of the injection tube is a serious problem because to remove
38
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the plug it is usually necessary to disconnect the injection tube;
hence, plugging of the tube during a run with the combustor would
make shutdown of the combustor necessary. Plugging of the orifice
in the mixing assembly is less serious as tapping in the area of
the orifice usually removes the plug and coal flow resumes.
S fl additional modification made in the equipment was to bore out
the inside of the tube fitting which connects the 1/2 inch diameter
mixing tee located at the bottom of the feeder to the 1/4 inch
diameter injection tube. The abrupt reduction in diameter had
caused coal to accumulate in the tee and eventually plug. By
boring out the connector, so that a gradual transition was made from
1/2 inch to 114 inch diameter, the plugging problem was eliminated.
This change caused also a moderate increase in coal flow.
Calibration of Modified Feeder
A program was undertaken to determine how feed tank pressure, feed
tank—receiving vessel pressure differential, and air flow rate affect
the feed rate of coal in the modified Petrocarb injector. Data col—
lected would make it possible to pre—set operating conditions needed
to produce the desired coal feed rate.
For a given orifice, coal feed rate appears to be a function of
pressure drop across the orifice of the coal feeder and the pressure
in the feeder. The size of the injection tube and the injection air
flow affect coal flow indirectly by causing more or less of the total
pressure drop from feeder to receiving vessel to occur across the
orifice rather than through the tube.
Figure 15 shows the effect of air flow rate on coal feed rate for
four pressures of up to 70 psig in the coal feed tank. AU data
shown are for a 3/16 inch orifice, 0.18 inch I.D. x 15 feet long
stainless steel injection tube, and a feed tank to receiving vessel
pressure differential of 3 psig. Increasing the air flow rate
through the injection tube causes a higher pressure drop across the
injection tube and a correspondingly lower pressure drop across the
orifice. Hence, coal flow decreases with increasing air flow.
It should be noted that not all of the air which passes through
the air rotometer enters the injection tube. A fraction of the total
air flow enters the feed tank to compensate for the small flow of air
leaving the feed tank with coal through the orifice. Thus, the
actual air flow in the injection tube is slightly less than that
indicated by the air rotometer setting.
Figure 16 shows the effect of feed tank pressure on coal flow rate
at constant air flow rate and feed tank to receiver P. An increase
in feed tank pressure permits coal to flow more readily through the
orifice.
39
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F I g u re
Effect of injection air flow rate
15
on coal feed rate — tow pressure
I I
3/16 INCH ORIFICE,
1/4 INCH x 15 FT STAIN-
LESS STEEL INJECTION
TUBE, P (FEEDER-
RECEIVER) 3 PSI, —30
MESH COAL
-
0
20 30 40 50 60 70
AIR ROTAMETER SETTING, 010 (10O% = 3.3 SCFM)
40
80
A
42
40
38
36
34
32
30
W26
24
w
22
-J
< 20
0
C -)
18
16
14
12
10
8
6
4
U
I I I 1
-------
Figure 16
Effect of feed tank pressure on coal feed rate
3/16 INCH ORIFICE, 1/4 INCH x 15 FT STAINLESS STEEL
INJECTION TUBE,
AIR ROTAMETER = 40% (1.3 SCFM),
t P (FEEDER-RECEIVER) = 3 PSI,
-30 MESH COAL
20 30
FEED TANK PRESSURE, psig
40 50 60 70 80 90 100
100
90
80 —
70
60 —
50 —
40
30
20
-o
w
I-
w
U i
U-
-J
0
C —)
I I I I I_
0/0
I I I I I 1_
10
41
-------
Figure 17 shows the effect of air flow rate on coal feed rate at four
levels of feed tank pressure from 91 to 135 psig. At any given
pressure, small changes in air flow rate change the coal feed rate
considerably. The effect is similar to that shown in Figure 15 for
data at lower pressures.
Figure 18 shows the effect of feed tank pressure on coal feed rates
for three different air flow rates.
Plugging at the orifice and in the injection tube occurred much less
frequently in runs made at higher pressures (above 90 psig) in the
coal feed tank than at lower pressures. For example, during runs
49—64 (See Table 4) the injection line plugged only twice. The orifice
plugged during one run, but the vibrator was not turned on at the time.
These runs represent a combined total of about eight hours of operation.
The flow of coal was also steadier at higher pressures.
A regression analysis was made on the data of runs 49—68 with
receiver pressure and rotameter setting as the independent variables.
The least squares correlation is
F = 36.0 + O.37lP — 0.873R
where F = coal feed rate, lb./hr.
P = pressure in receiving vessel, psig
R = air rotameter reading, % of scale
This correlation is valid only over the range of Pand R in the experi-
mental work (runs 49—68). These limits are 88—132 psig and 50—90%,
respectively. The least squares lines are shown as the solid lines
of Figures 17 and 18. Details of statistical tests of the model are
given in Table 5.
BURNER
The burner for the batch fluidized bed regenerator had to handle a
wide range of flowrates, up to a maximum of about 30 SCFM air and 1.5
SCFM fuel (60% propane + 40% propylene), in proportions ranging from
strongly oxidizing to strongly reducing. Furthermore, it had to
operate at pressures of 1—10 atm. and accept preheated air at
temperatures up to about 800°F. As pressure, temperature, and
fuel/air ratio increase, the tendency of a flame to flash back
from the supporting grid into the burner tube also increases. These
factors, combined with the need to produce a burner with a reasonably
long service life, made the development of the burner a formidable
task.
All burners that were designed and tested consisted of two main
42
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Figure 17
Effect of air flow rate on coal feed rate — high pressure
3/16 INCH ORIFICE
1/4 INCH O.D. x 15 FT STAINLESS STEEL INJECTION TUBE,
P (FEED TANK-RECEIVER) = 3 PSI
1
118 PSIG
40 50 60 70 80 90
AiR ROTAMETER SETTING, % (100% 4.4 SCFM)
100
RECEIVER PRESSURE
= 132 PSIG
U
LU
I—
U i
LU
LL
-I
C
C)
32
30
28
26
24
22
20
18
16
14
12
10
8
6
4
I I I I
43
-------
Figure 18
Effect of feed tank pressure on coal feed rate
3/16 INCH ORIFICE,
1/4 INCH x 15 FT STAINLESS STEEL INJECTION TUBE,
L P (FEED TANK-RECEIVER) 3 PSI
FEED TANK PRESSURE, psig
140
32
30
28
26
24
-
w
I —
0
U
U
-J
Q
C)
22
20
18
16
14
12
10
8
6
44
2
0
70
80 90 100 110 120 130
44
-------
Table 5. RESULTS OF RE(RESSION ANALYSIS OF COAL FEED TEST DATA
Dependent variables = F (coal feed rate, lb../hr.)
Independent variables = P (receiver pressure, psig)
R (rotarneter reading, %)
Correlation Matrix
P R F
P 1.000
R 0.687 1.000
F —0.126 —0.781 1.000
Standard error of estimate = 2.033
F ratio for regression 114.1
Fraction of explained variance = 0.931
Determination of correlation matrix = 0.528
- Residual degrees of freedom = 17
Std. error Co—variance
Variable Coefficient of coeff. T—ratio F—ratio ratio
Constant 36.0
P 0.371 0.042 8.86 78.46 0.47
R —0.873 0.058 14.98 224.3 0.47
45
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parts: the burner tube and the grid. (See Figure 10). The burner
tube was essentially the same in all designs; it was a stainless
steel tube about 6 inches long with a diameter of 2 inches. The
burner tube was welded to a flange which attached to the bottom
of the burner section of the regenerator. The burner tube contained
baffles and was filled with 1/8 inch alumina beads to adequately
mix fuel and air, which entered under the flange. Fuel entered
through a sparging tube Which distributed the fuel radially into the
surrounding air. The position of the sparging tube in the burner tube
could be adjusted to achieve optimum mixing. The best mixing occurred
when the openings of the sparging tube were positioned at the tee
where the air and fuel lines joined. It was also found that the
number of baffles in the burner tube could be reduced from nine to
two without any noticeable change in the appearance of the flame.
The first burner grid tested was a 1/8 inch thick plate of micro—
metallic mesh, which was welded to the top of the burner tube.
Directly under the grid was a water cooling channel in the shape of
a circle with four spokes. ttThermonfl cement was placed between the
mesh and the cooling tubes to promote good thermal contact. This
burner assembly performed well in terms of its range of applicability
for burning very lean to very rich. The reliability and durability
of these grids was a problem, however. Three grids had been tried;
the first and third lasted for only one to two runs while the second
lasted considerably longer. Typically, cracks would develop in
the mesh or around the weld. Variability in the service life of
different grids was attributed to slight differences in construction.
In order to learn more about the cause of the loss of burner grids,
the fine micrometallic mesh was replaced with coarser “Rigimesh”.
However, the Riginiesh grid, which caused a considerably lower pressure
drop, burned out after only a few hours of operation.
Since no clear indication was then available on how to design a
burner with a reasonable life expectancy, a new burner assembly was
built to allow easy replacement of burner grids. This new burner
had a grid Which, rather than being welded, was attached to the top
of the tube with a single bolt. This burner has been used successfully
with the combustor for some time, perhaps because it has been used
only for oxidizing flames.
A new burner grid for the regenerator burner was designed and built.
It was a water cooled stainless steel plate containing 48—1/16” holes.
The grid worked well at 10 atm with an oxidizing flame, but a flashback
occurred with a reducing flame.
Another water cooled burner grid was built, this time of brass, to
take advantage of the high thermal conductivity of this material.
To help prevent flashback, the grid had small holes (400—l/32 inch),
and a layer of micrometa].ljc mesh was placed under the grid to act
46
-------
as a flame arrester. This sandwich was held onto the burner tube by
eight small studs. The burner performed very well for two runs
(115R, 116R) but a flashback occurred during preparations for an
additional run. The brass grid was not damaged but the baffles
in the tubular section of the burner were destroyed. The probable
cause of the flashback was a failure of the gasket sealing the metal-
lic mesh flame arrester to the burner tube.
The design was modified to provide a more positive seal of the mesh
to the burner tube. The mesh was welded onto the top of the burner
tube and the brass grid was placed over the mesh and attached with
studs. A gasket was still needed under the brass grid but, since
it was directly under the cooling channel of the grid, it was kept
cool. This design eliminated the second gasket, which would be
exposed to higher temperatures. A radiation shield for the tubular
section of the burner, made out of copper tubing (water—cooled),
was also added to preclude the possibility of exceeding the self—
ignition temperature of the fuel—air mixture before it reached the
grid. The modified burner worked very well except for the breakage
of a stud on two occasions. Disassembly and inspection of the burner
after a series of seven runs showed it to be in excellent shape.
The last burner grid developed was of similar design but screwed
onto the burner tube, so that the delicate studs could be eliminated.
The mesh flame arrester was not welded to the burner tube but was
seated on a shoulder cut into the wall of the burner tube at its top.
Performance of this new burner has been excellent. It ignites easily,
even at elevated pressures, the flame is stable over almost all
operating conditions, and it should have a very long life expectancy.
COOLING COILS
The combustor contained three pair of serpentine wound cooling coils
which extended from about 6 inches above the fluidizing grid
upwards for about 6 feet. The coils were made from 1/4 inch diameter
stainless steel tubing and had a total surface area of about 6.3 sq.
feet. Each pair was centered at a flange so that one coil could be
removed without disturbing the others.
Water flow into each pair of coils was independently controlled.
House water was denjineralized before its pressure was increased
with a pump. The water flow then split into three branches which
connected with the cooling coils. Each branch contained a rotanieter
followed by a metering valve. The branch connecting with the middle
pair of coils also contained an automatic control valve which could
be used instead of the manual valve.
47
-------
During the first runs with the combustor, house water entered the
cooling system at about 40 psig and no pump was used to increase
pressure further. With this arrangement, control over the amount of
water entering the coils was unsatisfactory. Because of expansion
produced during the formation of steam, back pressures greater than
the inlet pressure of the water developed in the coils.
This resulted in water entering the coils one slug at a time, a limit
to the maximum amount of water that could be fed into the coils, and
very little control over the flow rate of water.
It should be noted that it was not feasible to control heat removal
from the conibustor by using enough water to maintain single phase flow.
This is because the rate of heat removal is proportional to the log
mean T, which would be nearly constant over the range of flowrates
required to maintain single phase flow. Also, the use of air as
a cooling medium had been investigated, but the flow of air required
to keep the coil temperature at a safe level would be so high as to
remove an excessive amount of heat from the bed.
Control over water flowrate was improved by increasing the inlet
pressure of the water by means of a pump. This created a much larger
pressure drop across the control valve than existed previously. With
increased pressure drop, downstream pressure fluctuations affected
flowrate less. Initially, inlet pressure was increased to 150
psig; but this pressure was not high enough and a Delavan Model
20001 piston pump was then installed. This pump can deliver two
gallons/mm. at pressures up to 500 psig; however, because of
pressure limitations on the rotameters, operation was limited to about
300 psig. Fluctuations in water flow were greatly reduced and control
over water flowrate was satisfactory with water inlet pressure at
this value.
COMBUSTOR CYCLONES AND FILTER
Two cyclones normally removed flyash and entrained solids from the
combustor. It was found; however, that during the first several
minutes of some combustion runs, temperatures at the cyclones were
not high enough to prevent condensation of steam from the off—gas.
When this occurred, efficiencies of the cyclones were reduced
sufficiently to allow flyash to reach the system filter. The filter
then plugged rapidly, disturbing the pressure in the combustor,
the pressure drop from the coal feeder to combustor, and the flow of
coal.
The problem to be solved was to prevent the filter from plugging with
flyash. The easiest solution was to add an electrical heater and a
third cyclone downstream of the first two cyclones. The heater
was used during the early stages of runs to keep the temperature of
48
-------
the off—gas entering the third cyclone above its dew point, so that
this cyclone could effectively remove flyash from the off—gas. This
system worked well although it was needed on only several occasions
when steam condensation was a problem.
REGENERATOR BED HEATERS
The lower section of the regenerator contained two 15 KW electrical
heaters which were used during startup to bring the bed up to the
desired temperature more rapidly than was possible with the propane
burner alone. The heaters also smoothed the temperature distribution
over the fluidized bed by increasing temperatures in the upper portion
of the bed.
The heaters were constructed from resistance wire which was wrapped
around the alumina reaction tube located in the center of the shell
of the regenerator. Two layers of ttFiberfraxU insulation surrounded
the windings to provide a cushion for thermal expansion. The annular
space between the alumina tube and the inside of the steel shell was
filled with castable refractory insulation.
A temperature runaway in the bed of the regenerator occurred during
the first run and the heater windings were destroyed. New heaters
were installed which lasted for eight runs. The heaters were examined
and it was concluded that failure was caused by a combination of
corrosion and overheating. Corrosive gases could have worked their way
through cracks in the alumina reaction tube and contacted the heater
windings. A new alumina tube plus a new set of heater windings were
installed. The lower 10 inches of the tube was made removable so
that it could be easily replaced if it cracked. It was decided to
operate the new heaters at reduced power to minimize the possibility
of overheating.
Unfortunately, the new heaters failed after only two runs. Failure
was attributed to chemical attack by reducing gases and SO 2 .
It was decided not to re—install heaters as no simple way was known
to prevent corrosive gases from reaching the heater windings.
OFF—GAS COOLERS
Heat exchangers were used to cool off—gases from the regenerator and
combustor to temperatures just above their dew points. The original
unit used with the regenerator was an American Standard shell—and—tube
heat exchanger with ninety—l/ 4 inch stainless steel tubes through which
gas flowed. The off—gas cooler had to be removed from the system when
it was discovered that a number of tubes had cracked. Examination of
the unit showed that some tubes had developed radial cracks and that
many of the tubes were plugged with solids. The cause of the cracks
was attributed to differences in thermal expansion between the tubes
and a ceramic header, located near the gas inlet.
49
-------
Heat transfer calculations, based on experimental data, indicated that
the original exchanger had too much area. As a result, the exchanger
was rebuilt using three 1/2 inch stainless steel tubes inside the
original shell. No problems have since developed in over eight months
of operation.
The heat exchanger used with the combustor contained nine 1/4 inch
diameter (.049 inch wall) tubes of 304 stainless steel. Gas flowed
through the tubes. No operating problems have developed in about
five months of operation.
MATERIALS OF CONSTRUCTION
After the fluidized bed regenerator was started up, a number
of materials problems developed caused by the combination of high
temperatures and the chemical environment.
Fluidizing Bed Grids
The original fluidizing bed grid used in the regenerator was
made of cast alumina. In the first few runs made in the regenerator,
high temperatures occurred which resulted in melting of the bed charge,
apparent reaction of the molten charge with the grid and cracking of the
grid. The grid was destroyed in these runs. A photograph of the fused
material and the grid is shown in Figure 19. Analysis of the molten
material indicated the presence of CaSO 4 , CaO and Ca 3 A120 6 .
Examination of the phase diagram of the CaO—A1 2 0 3 system indicated that
a melt of calcium oxide and calcium aluininate could exist at temper-
atures as low as 2550°F. Apparently, CaO was reacting with the alumina
grid to form the oxide—alurninate mixture which melted at the lower
temperatures. An additional run was made with an alumina grid under
lower and better controlled bed temperatures conditions. Although
the damage to the grid was much less, it still showed some signs of
reaction with the bed charge,was bowed upward and cracked. A cast
zirconia grid was then fabricated and tested. Zirconia not only has a
higher melting point than alumina but also forms no low melting point
mixtures with CaO. The zirconia grid showed no signs of reaction with
the bed material, but was still slightly bowed and cracked. In order
to impart maximum strength to the zirconia it must be cured at
3000°F. Available curing ovens could only cure at 400°F and the grids
were therefore, not cured to maximum strength. Also, no assurance
could have been given that the fully cured zirconia could stand up to
the repeated oxidizing—reducing cycles in the regenerator. Therefore,
it was decided to abandon the concept of cast ceramic grids and
fabricate the water cooled stainless steel grid described in Section IV.
The water cooled steel grid has worked very well after some early pro-
blems of water leakage at a welded joint were solved. The grid was
examined after being in operation for a number of months. The only ap-
parent sign of damage was a small raised section which may have been
caused by a blocked cooling channel.
50
-------
.
U,
Figure 19.
Fused regenerator bed material and fluidizing grid
-------
Ceramic Reactor Liner
Melting and fusion of the bed charge in the regenerator which
destroyed the alumina fluidizing grids also severly damaged the mullite
ceramic bed liner. Mullite is an impure alumina, containing a signifi-
cant amount of silica. It was-believed that lo :me tiiig ilicate
systems could have formed by reaction of the liner with the molten bed
charge. Since the silicates have even lower melting points than
aluminates, it was decided to replace the mullite liner with an alumina
liner. The alumina reactor showed no signs of reaction with the bed
material during succeeding runs. However, due to its poorer resistance
to thermal shock, the alumina tube cracked and had to be periodically
patched with ceramic cement.
Refractory Insulation
The Litecast 7528 refractory insulation used in the regener-
ator and combustor vessels has held up very well with use. The only
damage sustained by the refractory occurred in the regenerator when the
bed charge melted and attacked the ceramic liner and portions of the
refractory also. Since the problem of bed charge melting was solved, no
further damage to the refractory has occurred.
The Bubbalite high temperature refractory used in the burner
zones of the combustor and regenerator has also held up very well.
The only damage sustained so far occurred when a failure in the
regenerator burner caused direct flame impingement on the refractory.
This resulted in the melting of a small portion of the refractory
but was easily repaired by patching.
52
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SECTION VI
EXPERIMENTAL RESULTS
COLD MODEL TEST UNIT
Effect of Simulated Steam Coils onFluidization Characteristics
Tests were conducted to determine the internal baffling effect of the
simulated coil upon the fluidizing and slugging character of the bed,
and the anticipated particle entrainment. The tests were conducted in
the combustor section of the CMTU over the following range of parameters:
Superficial Bed Velocity (ft/sec) 4, 6, 8, 10, 12
Settled Bed Height (ft) 2, 2.5, 3, 3.5, 4
Vessel Pressure (psia) 20, 30, 40, 60
It was visually observed that the simulated coils inserted in the acrylic
model produced a well fluidized bed, especially in the region where the
expanded bed is below the upper coil level. It is therefore anticipated
that this type of coil configuration will promote very satisfactory
mixing and heat transfer in a high temperature combustor.
Test results indicated that expanded bed height increases with increasing
settled bed depth, pressure and superficial bed velocity (Fig. 20 and 21).
Thus the tendency toward slugging and high entrainment rates is increased
by operating with deeper beds, higher velocities and higher pressures.
The insertion of the simulated coils, in addition to improving the bed
fluidization behavior, also significantly reduced the slugging nature of
the fluid bed (Fig. 22). For example, with a 3 feet settled bed at
40 psia and a superficial bed velocity of 6 feet/sec., the average expanded
bed height was reduced from 132 inches to 78 inches when the simulated
coil was inserted.
This implies that the coils provide a modulating or baffling effect and
that combustor outage height can be reduced without an increase in
particle entrainment.
Determination of the minimum fluidizing velocity was also obtained by
increasing the fluidizing air velocity until the bed particles were just
visually fluidized. The results indicated that the minimum fluidizing
velocity is not affected by bed depth and decreases with increasing reactor
pressure (see Fig. 20 and 21).
53
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Figure 20
Effect of superficial bed velocity on expanded bed height
2 ft. settled bed, simulated heat transfer coil
4-
. ., —
“I
I I I
SETTLED BED HEIGHT (24 IN)
I I I I I
4 6 8
SUPERFICIAL BED VELOCITY, ft/sec
10
140
120
100 —
80 —
60 —
40—
20
U)
U
=
w
=
LiJ
=
w
uJ
0
x
w
4u
0
2
12
54
-------
Figure 21
Effect of superficial bed velocity on expanded bed height
3 ft. settled bed, simulated heat transfer coil
U)
c i )
C)
C
w
I
Ui
=
LU
LU
x
LU
140
120
100
80
60
40
20
0
0
2 4 6 8 10
SUPERFICIAL BED VELOCITY, ft/sec
12
55
-------
Figure 22
Effect of heat transfer coil on expanded bed height
40 psia Pressure
I I I
ct,,
‘c /
(<,
S
6 8 10
WITHOUT
COIL
WITH
SIMULATED
COIL
(68 IN.
LONG)
SUPERFICIAL BED VELOCITY, ft/sec
180
.
/
a)
0
w
F—
w
w
w
0
x
I J
/
/
/
160 —
140—
120 —
100 —
80
60
40
/
/
/
S
F
/
/
/
/
/
/
/
I
.
I.
/
2
4
I I
I
12 14
56
-------
The minimum fluidizing velocity is considered to be most likely a function
of fluidizing gas density rather than just pressure. Since the tests
were run at the same temperature, the density varied directly with the
pressure. At 20 psia, for the stone particle size distribution used, the
minimum fluidizing velocity is 2.4 ft/sec; at 60 psia it is 1.6 ft/sec.
It should be noted that the same density effect is also most likely true
for the cases involving the expanded bed height. That is, it is probably
the density of the fluidizing air, rather than just the reactor pressure,
that influences the expanded bed height of the fluidized system.
Continuous Transfer of Solids
Studies were conducted in the Cold Model Test Unit to determine the effect
of various parameters upon the solids transfer rate between prototype
reactors and to measure this transfer rate.
Tests were conducted at reactor vessel pressures to 40 psia, with settled
bed heights of 3.0 feet. Since the solids were transferred using the
pulse air technique, the main parameters varied were pulse air flow rate
and pulse frequency, i.e., pulse on/off cycle. The results of the two-
way transfer runs at 40 psia are presented in Figure 23.
The tests verified the ability of the continuous system to transfer solids
in both directions while operating under pressure and maintaining good
bed level control in both vessels. Transfer rates of 6.1 lb/mm were
attained while transferring solids continuously in both directions at
40 psia vessel pressure. The rate of solid transfer between vessels
was controlled by varying the pulse cycle and also the pulse air flow
rate.
Observations of the continuous transfer operations in the CMTU indicated
that the 1-inch I.D. catch tube inserted in the prototype combustor was
a limiting or choking restriction on the maximum transfer rate attainable.
However, the solids reservoir installed in the prototype regenerator
(Figure 2) functioned satisfactorily in that it permitted freer solids
flow and a superior transfer line seal than the catch tube. When
transferring solids in one direction, from regenerator to combustor,
solids transfer rates as high as 9.9 lb/mm were attained. This is
attributed to the ample sizing and configuration of the regenerator
transfer reservoir.
Transfer line fittings and restrictions not only limit the solids flow
rate, but induce hold-ups and plugs in the transfer line. As these line
restrictions in the CMTU were systematically eliminated, improved solids
transfer rates were achieved.
It is anticipated that with adequate pipe line sizes and a minimum of
restrictions, the Miniplant will be able to achieve satisfactory solids
transfer with this pulse technique. The Miniplant transfer rate design
is 680 lb/hr (11.3 lb/mm) under maximum conditions.
57
-------
Figure 23
Solids transfer rate
10 I
REACTOR VESSEL PRESSURE, 40 psia
SUPERFICIAL BED VELOCITY, 8 ft/sec
E
8 PULSE AIR FLOW RATE, 2 SCFM
LU
I— _________
6
LU
.
LL
2
4
/
(I,
-J
2—
0 I I
3 2 1 2 3
OFF/ON RATIO -p ON/OFF RATIO
PULSE CYCLE
-------
COMBUSTION STUDIES
Fluidized Bed Combustion Unit
Because of operating problems coal has not been fed continuously into
the batch combustor for more than about twenty minutes at a time. These
problems involved poor control over the flow of water entering the
cooling coils, heavy accumulation of flyash in the off-gas filter, and
unsatisfactory operation of the coal feeding system. The first two
difficulties have essentially been resolved and progress has been made
toward improving the operation of the coal feeding system. The last
major obstacle appears to be preventing plugging in the coal probe and
orienting the probe so that coal burns uniformly throughout the
fluidized bed.
Table 6 presents operating parameters for coal combustion runs. The
compositions of effluent streams are given in Table 7. When evaluating
this data, it is important to recognize that, because coal was fed
only for brief periods, conditions in the coinbustor (e.g. coal feed
rate, temperature, off-gas composition) were unsteady. Because there
were problems with control of temperature, and also with the feeder,
coal rates had to be kept low. As a result, excess air levels were
far higher than they would normally be.
Calcination of Limestone and Dolomite
A 23 lb. charge of Grove #1359 limestone was calcined in the combustor
in about one hour at an average temperature of 1560°F. It was desirable
to calcine in the combustor because this eliminated the need to transfer
stone to the combustor from the vessel in which it was calcined.
A nearly stoichiometric air-propane mixture was used and the superficial
velocity was 8-10 ft./sec. The total pressure was about 2.6 atm; the
Co 2 concentration in the off-gas was approximately 2l7 ,. At the
conclusion of calcination the bed temperature rose to 1700°F and the
level of C02 in the off—gas fell to 15%, about the same level of CO 2
as observed before calcination. The combination of temperature rise
(due to the end of endothermic calcination) and drop in CO 2 concentration
gave a definite indication of the completion of calcination. Subsequent
chemical analyses showed that more than 99.9% of the carbonate had
been converted to oxide.
Subsequently, a 16.6 lb. batch of -8 + 25 mesh Tymochtee dolomite was
haif-calcined by heating in the combustor for about six hours at
1200-1400°F. Proper haif-calcination should have produced a material
containing only CaCO 3 and MgO. Subsequent analysis of solids showed
that some decomposition of the calcium portion of the stone had also
occurred
83 wt. L (CaCO +MgO)
17 wt. 70 (CaO MgO)
59
-------
Table 6. OPERATING PARAMETERS - COAL COMBUSTION RUNS
Coal: Arkwright Mine, —30 mesh
Limestone: Grove #1359, —7 mesh, calcined.
Dolomite: Tymochtee, —8+25 mesh, haif—calcined.
Average coal rate varied over range of 10.0—18.0 lb./hr.
Average coal rate varied over range of 15.0—17.1 lb./hr.
C. Actual air — stoichiometric air
/. Excess air = x 100
Stoichiometric air
Charge,
Run No. material,
lbs.
Settled
height,
bed
ft.
Press.,
psig
Superficial
velocity, ft./sec.
Input
Coal, Air,
lb/hr SCFM
66—2690—ll2a limestone,
23
2.0
103
3•9
13 • 3 a
66—2690—112b limestone,
23
2.0
103
43
15 5 b
66—2690—127 dolomite,
16.6
2.0
63
5.4
8.6
40.0
66—2690—128a dolomite,
16.6
2.0
63
5.8
7.2
41.5
66—2690—128b dolomite,
16.6
2.0
104
5.6
8.8
62.0
66—2690—l28c dolomite,
16.6
2.0
104
6.1
9.2
63.5
Average bed
temperature, °F
1480 + 100
1660 + 100
1480 + 100
1560 + 100
1510 ± 100
1640 + 100
% Excess
airc
43
23
104
153
209
203
-------
Table 7. COMPOSITION OF EFFLUENT STREAMS - COAL COMBUSTION RUNS
SO 2 , ppm
NO, ppm
C0 2 , %
CO, %
H 2 0, %
N 2 ,
SO 2 Removald(%)
SO 2 , ppm
NO, ppm
C0 2 , %
CO, %
1120 , %
N 2 , 7.
SO 2 Removal (%)
66—269 0—112 a
Measuredc
Avg. Range
240 steady
—— 310 230—370
12.0 10.6 6.5—19.0
—— 0.9 0.1—2.0
not collected
9.6 3.3—15.5
76.8
Avg.
145
— — 340
8.5 7
3.5
10.4
77.5
Range
100—140
160—550
3—11
.02—3.0
66—2690—112b
Measured
______ Range
steady
150—380
5. 2—23. 6
0.1—2.7
5.8 not collected
3.7 7.6 1.9—15.4
76.5
2.9
12.4
77.7
86.0
66—2690—128a
Measured
Range
125—550
250—550
4.5—9.5
.03 0.01—1.0
not collected
10 3—15
72.2
Run No.
Caic .
Caic. Avg .
1780 250
—— 290
13.9 13.5
—— 0.7
5.0
6.
Run No.
84. 4
66—2690—127
Measured
Caic .
1100
Caic. Avg .
900 250
—— 450
6.9 7.0
.03
not collected
12 2—17
86.8
61
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Table 7 (continued). C0?JIPOSITION OF EFFLUENT STREA1 IS —
COAL COMBUSTION RUNS
aA T O SO removed in bed.
bA all C 2 —) CO 2 and H —v H 0,
Corrected for water condensed pri r to analysis.
Calculated S02 — observed SOl
x 100.
SO 2 removal = Calculated SO 2
Run No.
66—2690— 128b
Measured
502,
NO, ppm
C0 2 , ppm
CO, %
H 2 0, %
N 2 , %
SO 2 Removal
66—269O—128c
Measured
______ Range
100—600
170—490
5.0—11.5
.01—.4
Caic.
Avg.
- Range
Calc.
Avg.
700
320
225—550
700
320
——
380
350—430
——
340
5.7
8.0
7.5—9.5
5.7
8.0
——
.03
.01—0.4
——
.02
2.4
not
collected
2.4
13.9
9
5—14
13.9
78.0
78.0
not collected
11 5—17
(%) 54.3 54.3
62
-------
Removal
In two runs (11 2 a, ll2b) made with calcined Grove limestone, SO 7 removal
was 84.4 and 86.O7 . It was estimated that conversion of CaO to CaSO
in the bed was 2.O7 prior to run No. ll 2 a and 5.47,, after run No. ll2 .
The combined coal feeding time for these runs was 47 minutes; 11.5 lbs.
of coal were burned.
Several brief runs were made with half-calcined Tymochtee dolomite; a
total of about 15 lbs. of coal were burned, which corresponds to a
conversion to sulfate of about l07 . It can be noted in Table 7 that
the percentage of sulfur removal decreased with each subsequent run.
This is probably due to loss of dolomite through attrition, which appeared
to be very high, rather than to deactivation caused by sulfate loading.
Even during the first combustion run with dolomite (127) sulfur removal
was only 877g. This is not surprising in light of poor temperature
control. Table 6 shows that the bed temperature for this run was
1480 ± 100°F. It was not possible to determine an average bed
temperature more precisely because of large temperature changes during
the run and steep temperature gradients in the bed. Poor temperature
control was due to inadequate flow of water through the cooling coils,
unsteady coal flow, and combustion of most of the coal in the immediate
vicinity of the coal inlet point. Progress has been made to correct
these problems.
NO Levels
—x
Average concentrations of NO were typically 300-400 ppm although the
instantaneous concentration fluctuated widely because the coal feed
rate was unsteady. NO concentrations would undoubtedly have been
considerably lower hadXit not been for the high levels of excess air,
which averaged between 23 and 209 percent.
Attrit ion
Attrition during runs with calcined Grove limestone appeared to be
fairly low; however, attrition was very high during runs made with
half—calcined Tymochtee dolomite. With the latter material, about
one-half of the bed wound up as fines in the cyclone diplegs. Reasons
for the high attrition rate are being examined to determine if the
stone, as it appears to be, or equipment (or both) was responsible.
63
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Fixed Bed Units
Reaction of CO and NO under
Simulated Combustion Conditions —
The reaction of CO and NO under simulated combustion conditions was
studied in the stainless steel fixed bed reactor. This study was car-
ried out to extend earlier work which had identified this reaction as
one which could be responsible for reduced NOx emissions observed in
the fluidized bed combustion of coal. Calcined limestone (No. 1359)
or dolomite (No. 1337) was used as bed material. In the study, pres-
sures were held at 1 or 10 atm, temperature was 1600°F. Premixed
gases were blended to maintain the CO and NO levels at either approx-
imately 1000 or 2000 ppm each. The C02 level was held at 0 or about
15%, the 02 level was held at 0 or about 2.5%. Residence time at 1 atm
pressure was held constant at 0.3 sec. At 10 atm pressure, residence
time was varied from 0.4 to 3 sec by varying the feed gas flow rate.
However, at residence times below about 0.5 sec, the flow rate was
too high to maintain the temperature at 1600°F and the temperature fell
to 1400°F at the highest flow rate studied. A summary of the runs is
given in Appendix Table A—l.
In the absence of C02, the reaction between CO and NO proceeds to 95—99%
of the limiting reactant over calcined limestone or dolomite. This is
shown in Table 8. The reaction proceeds in a 1/1 molar ratio of CO
and NO, suggesting the reaction
2C0 + 2N0 2C0 2 + N 2 (3)
As seen in Table 8, when C02 was added, conversion of the limiting
reactant decreased to 19—26%. The effect of CO 2 was almost instanta-
neous. In one case, the C02 blend was cut out of the feed and the flow
of the CO/NO blend was increased to maintain a constant feed rate. The
NO concentration fell from 625 ppm to 63 ppm in about 1 minute and then
decreased more slowly to 22 ppm over the next 10 minutes. The lag time
for the sampling and analytical system was measured at 50 sec, so the
response of the system to the removal of C02 was virtually instantaneous.
64
-------
TABLE 8. NO—CO REACTIONS
Bed Source CALCINED CALCINED
LIMESTONE #1359 DOLOMITE #1337
Inlet Gas Comp .
NO, ppm 1400 1800 860 1400 1990 840
CO, ppm 940 1870 990 900 2080 980
CO 2 ,% 0 0 17 0 0 16
Outlet Gas Comp .
NO, ppm 400 20 640 350 240 680
CO, ppm 10 160 770 20 100 830
C0 2 ,% 0 0 17 0 0 17
Cony., , 99 99 26 98 95 19
Temperature : 1600°F
Res. Time : 0.3 Sec.
The effect of pressure was studied and, although increasing the pressure
to 10 atm in the presence of CO 2 apparently increased the conversion,
the increase was probably due to increased residence time. At equivalent
residence times, increasing pressure appeared to decrease the conversion
slightly. This is shown in Table 9.
TABLE 9. NO-CO REACTIONS
EFFECT OF PRESSURE AND RESIDENCE TINE
Bed Source CALCINED CALCINED
LIMESTONE #1359 DOLOMITE #1337
Press. (Atm. ) 1 10 10 1 10
Res. Time (Sec. ) 0.3 3 0.3 0.3 3
Inlet Gas Comp .
NO, ppm 860 890 1150 840 840
CO, ppm 990 980 1240 980 980
CO 2 , % 17 18 13 16 16
Conv. 26 82 16 19 88
Temperature : 1600°F
65
-------
Changing the background gas from N 2 to argon appeared to increase the
conversion very slightly, but the effect may not be significant. was
added to the feed and appeared to increase conversion slightly. Thts
is shown in Table 10.
TABLE 10. NO—CO REACTIONS
EFFECT OF 02
Bed Source CALCINED CALCINED
LIMESTONE #1359 DOLOMITE #1337
Inlet Gas Comp .
NO, 1150 970 840 980
cö, ppm 1240 1060 980 1080
C0 2 , % 13 16 16 15
02, % 0 2.4 0 2.3
Cony., % 84 95 88 91
Temperature : 1600°F
Pressure : 10 Atm.
Increasing the temperature from 1600 to 1700°F in the presence of CO 2
at 1 atm pressure had essentially no effect on conversion.
The reason for the large CO 2 inhibiting effect is not understood.
Carbonation of lime was considered but does not explain the effect since
CaO does not carbonate at the conditions used in the 1 atm runs. At these
conditions the inhibiting effect was pronounced. Carbonate is the
stable form at 10 atm pressure and 1600°F but the oxide is the stable
form at 1700°F. However, changing reaction temperature from 1600 to
1700°F at 10 atm had no effect on conversion. Inhibition by chemical
reversibility was also considered, but does not explain the results.
The most likely explanation for this effect is a kinetic limitation
caused by the presence of the CO 2 .
REGENERATION STUDIES
Introduction
A gas containing CO and I I2 can reduce calcium sulfate to the oxide at
high temperatures. The reactions are
CaSO 4 + CO 4 CaO + SO 2 + CO (4)
CaSO 4 + ) CaO + SO 2 + H (5)
66
-------
Formation of CaS, via side reactions, is undesirable because large
amounts of reductant are used and no S02 is produced:
CaSO 4 + 4C0 — CaS + 4C0 2 (6)
CaSO 4 + 4H 2 > CaS + 41120 (7)
CaSO 4 can react with CaS to produce CaO but this reaction does not occur
to a great extent:
3CaSO 4 + CaS > 4CaO + 4S0 2 (8)
Reactions (4), (5), and (8) are favored by high temperatures whereas
(6) and (7) are favored at lower temperatures. Gas composition has
an effect through CO/C02 and 1 12/H 2 0 ratios. High ratios promote
reactions (4) and (5) but promote reactions (6) and (7) to an even
greater extent.
There are at least nine process variables which can affect the perfor-
mance of the regenerator. These are:
1) pressure
2) temperature
3) air/fuel ratio (concentration of CO + 112 in reducing gas)
4) space velocity (superficial velocity/bed depth)
5) manner of fuel introduction and use of auxiliary fuel
6) use of auxiliary air
7) particle size of bed material
8) nature of bed material (pure CaSO 4 , sulfated limestone or dolomite)
9) quality of fluidization
To evaluate the overall performance of the regenerator several factors
must be examined. The most important is concentration of SO 2 in the
off—gas, which, of course, is to be maximized. Concentration can be
considered both as an absolute molar percentage or as a percentage of
the maximum concentration which is possible at equilibrium. Another index
of performance is the conversion of CaSO 4 . Even a high concentration of
is not alone sufficient to make performance of the regenerator accept-
able if only a small fraction of the CaSO 4 charged has been regenerated.
Moreover, CaSO4 can be reduced to CaS and CaO. It is desired to minimize
the ratio of CaS/CaO in the regenerated product. Other indices of
performance include the degree of attrition of the bed and the uniformity
of temperatures in the bed (an indication of the quality of fluidization
and gas—solids contacting).
Fixed Bed Regeneration Studies
Prior to the construction of the high pressure batch regeneration unit,
a few regeneration runs were made in the ceramic tube fixed bed regenera-
tion unit at pressures between 1 and 9.5 atm. This work was done primarily
to measure SO 2 product concentrations at the higher pressures. Tempera-
tures were held at 1900 or 2000°F. Anhydrous CaSO4 (Drierite) and
67
-------
sulfated limestone were used as bed materials. The reducing gas feed
was blended from C0/N2 and C0 2 /N 2 mixtures to permit variation of the
CO from 10 to 30% and the CO 2 from 10 to 20% in the feed. Gas linear
velocities ranged from 0.16 to 1.45 ft/sec. corresponding to 20 to
150% of the calculated minimum velocity required to fluidize the beds.
The results are given in Table 11. At lower pressures, 1 and 3 atm,
S02 concentrations in the product peaked at about 7% using CaSO4 beds.
These concentrations correspond to 35 and 50% respectively of the SO 2
concentration calculated at equilibrium. However, when the pressure
was increased to 9.5 atm, the SO 2 concentration dropped to 1.6% or 34%
of the equilibrium level.
Two runs were then made with sulfated limestone obtained from earlier
work at Esso and from the U.S. Bureau of Mines fluidized bed combustor.
In these runs, the measured SO 2 concentrations were only about 0.5%,
corresponding to about 10% of the equilibrium level. Moreover, the
solids discharged from the reactor in the runs using sulfated limestone
were partially agglomerated. This could have been due to softening
of fly ash present in the sulfated limestone at regeneration temperatures.
This could have then caused the bed to agglomerate in the absence of
vigorous fluidization. The agglomeration may also have caused poor
gas—solids contacting and low S02 product concentrations. Fixed bed
studies were terminated at this point to permit operation of the batch,
fluidized regenerator unit.
Fluidized Bed Regeneration Studies
Table 12 is a summary of runs made with the batch fluidized bed regener-
ator. A more complete listing of the operating parameters and results
for each run is given in Appendix Table A—2.
Agglomeration of Bed
The first runs made with the batch—fluidized regenerator resulted in
fusion of the bed. Heat transfer calculations showed that temperatures
on the top surface of the ceramic fluidizing grid could exceed the
melting point of CaSO4 (about 2650°F). Additional calculations showed
that radiation errors could result in thermocouples indicating tempera-
tures several hundred degrees below the true temperature of the reducing
gases. As a result, gas temperatures could have been high enough to
melt CaSO 4 .
It was decided to replace the ceramic fluidizing grid with a water—
cooled stainless steel grid. This would preclude the possibility of
localized melting of CaSO4 occurring when it contacted the hot surface
of the grid. To reduce the temperature of the reducing gas, diluent
nitrogen was added to the burner along with air and fuel.
68
-------
Table 11. RUN SUMMARY—FIXED BED REGENERATION STUDIES
Bed Temp., Press., Feed Comp. Flow, Vel. Peak % of
tnat’l. °F atm % CO % CO 2 SCFM ft./sec Vmf S0 2 % eguil .
CaSO 4 1900 1 21 15 0.1 1.45 1.4 7.1 0.33
CaSO 4 2000 3 31 10 0.1 0.51 0.5 7.2 0.48
CaSO 4 2000 9.5 10 20 0.1 0.16 0.2 1.6 0.34
Suif. (a) 2000 9.5 10 20 0.1 0.16 0.2 0.6 0.13
Lime
Sulf . 2000 9.5 10 20 0.19 0.31 0.3 0.4 0.08
Lime
(a) Bureau of Mines source
(b) Esso source
(c) Velocity/minimum velocity to fluidize
Charge wt. = 40 gm except 60 gm in run 36
Particle size = —18+20 mesh
-------
Table 12. SUMMARY OF RUNS WITH BATCH FLUIDIZED BED REGENERATOR
(—10+20 mesh CaSO 4 (Drierite) bed)
Avg.
Pressure,
Feed
Product
Settled bed Superficial
Observed
Equil.
%
of
Run No. temp., °F atm
%(C0-I-H 2 )
CO/CO2
height, ft.
velocity,
ft/sec
S02,
%
SO2,
%
eguil.
-.1
0
66—2690-. 71 2040 3.1 2
0.013 2.8 2.7 3.4
12.9
26
72 1990 3.2 15
0.013 5.6 2.9 5.2
9.8
53
75 2100 6.2 9
0.016 5.4 5.4 7.5
11.3
66
78 1950 6.0 11
0.021 4.0 2.4 2.0
5.0
40
86 1870 6.0 15
0.026 4.0 2.4 1.8
3.8
47
89 1970 6.1 15
0.023 4.0 2.5 1.7
5.1
33
98 2030 9.0 15
0.021 2.0 4.0 3.1
6.7
46
100 2080 10 15
0.023 2.0 4.3 1.6
9.0
18
101 1920 10 14
0.017 2.0 2.0 1.1
2.7
41
105 b 1800 10 14
109 2030 ± 40 10 7
0.013 2.0 1.8 0.7
—— 2.0 3.5 3.0
1.9
6.0
37
50
115 2050 ± 20 10 11
0.048 2.0 3.6 1.9
6.7
28
116 1840 + 20 4.9 13
119 c 2000 ± 20 10 10
2020 ± 30 10 10
121 e 1900 + 30 5.0 17
122A 1840 + 20 5.0 13.5
122B 2000 + 20 10 14
l23 1970 + 20 5.0 13.5
0.040 2.0 4.1 1.1
0.064 2.0 3.5 — 2.9 1.5
0.059 2.0 3.8 1.6
—— 2.0 3.2 1 3 a
—— 2.0 2.6 054 a
—— 2.0 2.8 2 • 0 a
—— 2.0 4.0 2 , 3 a
4.2
4.6
5.4
5.1
4.4
4.6
7.0
25
33
29
25
12
44
33
124 g 1940 + 40 10 23
0.061 4.0 2.6 20 a
2.8
72
l25 1950 + 30 10
0.16 2.0 3.0 062 a
3.0
21
126 gj 1880 + 80 10
0.062 4.0 3.0 12 a
2.4
50
Corrected for N2 added in au . air.
Run terminated after 5 mm.
fuel 6 inches above grid, aux. air
grid overall 1.02 for 122A, .99 for
42 inches
122B, 1.06
above
for
Aux. air injected 12 inches above grid.
Fuel injected 6 inches above grid.
eAUX. fuel 6 inches above grid, aux. air
34 inches above grid overall = 10.8.
123.
fuel 6 inches above grid, aux. air
rid.
.Sulfated limestone (Bureau of Nines)
—2O+40 mesh CaSO4 (Drierite)
54 inches
above
-------
Computer calculations (3) were made to determine the flow rates of air,
fuel, and nitrogen required to achieve the desired reducing gas com-
position and temperature. Equilibrium gas composition and adiabatic
flame temperaturewerefound as a function of equivalence ratio, pres-
sure, and per cent dilution of air with nitrogen. Calculations were
made over the range of 0.8—2.0 equivalence ratio, 1—10 atm, and 0—75%
nitrogen dilution (a 75% dilution means that 75% of the air is
replaced by N2). P n equivalence ratio of 0.8 is typical of that used
during pre—heating of the regenerator (oxidizing conditions). Under
reducing conditions, the equivalence ratio is greater than one.
Figure 24 presents the major results of the computer study. Adiabatic
flame temperatures are plotted as a function of the percentage of
CO + H 2 in the reducing gas mixture for nitrogen dilutions of 0, 10,
15, 20, 25 and 50%. Lines of constant equivalence ratio are also shown.
Curves indicating constant dilution are drawn as dashed lines between
= 0.9 and = 1.2 because there is insufficient data in this region
to draw the curves accurately. The actual maxima in temperature should
occur at an equivalence ratio of slightly greater than 1. All the
results of Figure 24 are for a total pressure of 8 atm. but no signi-
ficant error would result if the data were used over the pressure
range of 1 to 10 atm.
From experience with the regenerator, it was found that the adiabatic
flame temperature had to be below about 2900°F to be certain that no
melting of CaSO 4 would occur. Hence, Figure 22 was used to determine
the percentage of nitrogen dilution and the air/fuel ratio required to
prevent melting at whatever concentration of CO + H2 was desired. The
technique of diluting incoming air with nitrogen was found to be highly
successful in preventing melting and agglomeration of the CaSO 4 bed.
Pressure
Runs were made over the pressure range of 3—10 atm. As pressure
decreases, with superficial velocity constant, the amount of fuel
added to the regenerator decreases. Hence, a minimum operating pres-
sure exists which corresponds to a flow of fuel which is just sufficient
to maintain the desired bed temperature. This pressure was about five
atmospheres. The first runs were made at pressures of 3.1 and 3.2
atmospheres, respectively; however they were characterized by agglomera-
tion and melting of solids and, as a result, there were very non—uniform
temperature distributions in the bed. Thus, the average temperatures
for these runs are not directly comparable to those reported for later
runs, in which no agglomeration of solids occurred and in which
temperatures were much more uniform. Also, somewhat higher air pre—heat
temperatures were used in earlier runs (about 1000°F vs. 600—800°F for
runs made later).
71
-------
Figure 24
Adiabatic flame temperature versus
% (CO + H 2 ) at 8 atm total pressure
O 2 4 6 8 10 12 14 16 18 20 22 24 26 28
% CO + H 2
c = 1.0
3700
3600
ct=O.9
3500
3400
3300
3200
3100
3000
2900
2800
2700
2600
2500
2400
2300
2200
2100
2000
1900
1800
1700
LL
0
uJ
I-
LU
0
LU
I—
72
-------
Use of pre-heated air produces higher bed temperatures but increases
the likelihood of melting solids in the regenerator or causing a
flashback into the tubular section of the burner if the self-ignition
temperature of the fuel-air mixture is exceeded.
As the mole fraction of SO 2 in the regenerator off-gas increases, fuel
costs for regeneration decrease. This provides an incentive to operate
at low pressures because the mole fraction of SO 2 at equilibrium
rises as the total pressure of the system is reduced. Hence, runs
were made over a range of operating pressures. However, the percent
approach to equilibrium was less at low pressures and, consequently,
the mole fraction of SO 2 as not higher for runs made at low pressures
than for runs made at higher pressures.
Figure 25 is a plot of concentration of SO 2 , expressed as a percent of
calculated equilibrium concentratior vs. pressure. If several runs
are eliminated, for the reasons indicated in the figure, a trend can be
observed toward increased approach to equilibrium at higher pressures.
Figure 26 is a plot of concentration of SO 2 , expressed as a molar
percentage, vs. pressure. The three runs with the highest SO 2
concentration were made early in the program and agglomeration of the
bed occurred. Hence, temperatures in the bed were probably higher than
indicated, which could account for high concentrations of SO 2 . The
remaining data do not show any significant effect of pressure on
concentration. Presumably, at higher pressures, the closer approach
to equilibrium compensated for the lover equilibrium concentration.
Temperature
Runs were made over the temperature range of 1800-2100°F. Figure 27
shows that the equilibrium partial pressure of SO increases sharply
with temperature in this range. Unfortunately, ttie risk of deactivating
the solids by sintering, or deadburning, increases at higher
temperatures, especially when the solids must undergo repeated cycles
of sulfur absorption and regeneration. Hence, the optimum temperature
for regeneration is the one which produces the maximum concentration
of SO 2 over a number of cycles.
Temperatures in the fluidized bed were judged by thermocouples spaced
six to twelve inches apart. The height of bed must be known to determine
which thermocouples are located in the bed. Since this height could
only be approximated, the temperature of the bed had to be estimated.
Another reason why temperature could not be determined precisely was
that it often changed during a run. When this occurred, an average
with respect to time was taken to calculate an average temperature.
73
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Figure 25
Regeneration studies — effect of pressure on
approach to equilibrium SO 2 concentration
o RUNS WITHOUT SECONDARY AIR AND/OR FUEL
• RUNS WITH SECONDARY AIR AND/OR FUEL
o LAB RUNS
TEMPERATURES ARE GIVEN BELOW EACH POINT(°F )
80 i i i i i t i i i
70 1940 LOW
o SPACE
2100 HIGH VELOCITY
TEMPERATURE
50 1990 BED FUSION, HOT o 2035 —
SPOTS
1870 2O3Q s”
040 0
1950 0
1970—’• 18 0
30 ‘197O 2 OO —
2040— -o ‘ 1840 2020
(-) o” 1900
20 . 200O 0
2080 LARGE
0 • TEMPERATURE
10 — 2000 1845 CHANGES
DURING RUN
0 I I I I I I I I I I
1 2 3 4 5 6 7 8 9 10
PRESSURE, atm
74
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Figure 26
Regeneration studies — effect of pressure on SO 2 concentration
o RUNS WITHOUT AUXILIARY AIR AND/OR FUEL
• RUNS WITH AUXILIARY AIR AND/OR FUEL
TEMPERATURES ARE GIVEN BELOW EACH POINT (°F)
8
I I I I I I I I I
0
2100 HIGH TEMPERATURE
7
6—
C -)
1-
0 BED FUSION,
5 1990 HOT SPOTS
0
E
0
4
I—
U i 0
2040 o —205O
2—
o ‘ 2030 1935
C-)
2000,2000
0
v - i •
1970 o —l95O S4—194O
2
•4—1880
4—1870 2O80- - —_195O
1970 “ —2O20
1 1840’ ’ 19O0 0
1920
0
• 1800
1845
I I I I I I I I I
0
0 1 2 3 4 5 6 7 8 9 10
PRESSURE, aim
75
-------
Figure 27
Regeneration studies — effect of temperature on SO 2 partial pressure at equilibrium
E
C t
LsJ
1 -f )
—4 (1)
LU
-J
C
C
a-
(‘4
0
(1 )
1.0
.9
.8
.7
.6
.5
‘4
.3
.2
.1
0
1800 1900 2000
2100
TEMPERATURE, °F
-------
Runs made at temperatures below 1900°F generally resulted in low
levels of S02 in the regenerator off—gas. For example, of four runs
made at temperatures below 1875°F, only one resulted in an S02 con-
centration greater than 1.1% (volume) and two resulted in concentra-
tions under 0.8%. On the other hand, for nine runs made in the tem-
perature range 2000—2100°F, the lowest SO 2 concentration was 1.5% and
five runs had average concentrations of over 2.0%. Figure 28 is a
plot of mole percent of S02 vs. temperature. There is a definite
trend toward higher SO 2 concentration at higher temperatures.
In addition to the molar or volume percentage of SO 2 , the concentration
of SO 2 , expressed as a percent approach to equilibrium, also increased
with temperature. Figure 29 is a plot of the percentage of equilibrium
concentration vs. temperature. The figure shows a trend toward closer
approach to equilibrium as temperature increases. This trend is more
clearly seen if several data points are omitted (noted in figure).
Superficial Velocity and
Settled Bed Height
Varying superficial gas velocity and settled bed height can affect the
concentration of S02 in the regenerator off—gas by changing both the
contact time between gaseous and solid phases and the quality of fluid—
ization. Decreasing the space velocity, e.g., by increasing bed depth,
may not improve gas—solid contact if it causes a bed which is well
fluidized to begin slugging. Space velocity can be defined as volumetric
feed rate of gas based on average reactor conditions/volume of settled
bed. It is equal to superficial velocity/settled bed height.
Most runs, as Table 12 shows, were made with a settled bed height of two
feet. The greatest bed height used was 5.6 feet. Superficial velocities
ranged from 1.8 to 5.4 feet/second. Space velocities, based on condi-
tions inside the regenerator, ranged from 1850 to 7700 hr . It is uncer-
tain if changing settled bed height alone had a significant effect on
the observed concentration of S02. Several earlier runs were made with
different bed heights but these runs were characterized by poor fluid—
ization, caused by melting and agglomeration of the bed. Thus, it
is not possible to determine if different bed depths caused differences
in results of these runs. It appears that a modest change in bed height
affects regenerator prefortnance only to the extent that it affects
quality of fluidization. The same can be said of superficial velocity.
Toward the end of work with the regenerator, evidence was obtained which
indicated that reducing space velocity would improve performance. In
run No. 124 (see Table 12), the reactor space velocity was decreased to
about one—third of that of many previous runs. At an average bed tem-
perature of 1940°F, the concentration of SO2 was 2.0% or about 72% of
equilibrium. This was a closer approach to equilibrium than was
attained in any other run. From observations of the temperature
77
-------
Figure 28
Regeneration studies — effect of temperature on
SO 2 concentration
8
7—
• 3-6 ATM
t o1OATM
1800 1900 2000 208C
TEMPERATURE, °F
78
-------
Figure 29
Regeneration studies — effect of temperature on
approach to equilibrium SO 2 concentration
80
o 1OATM
• 5-6 ATM o LOW SPACE VELOCITY
70 —
E 60—
50 — HIGH TEMP. EXCURSION —
. DOWNSTREAM OF 02 0’
INJECTION POINT 0 .“
105 .0 •o 0
LARGE TEMPERATURE
10 — CHANGES DURING RUN
0— I
1800 1900 2000 2080
TEMPERATURE, °F
79
-------
profile in the bed, it was believed that fluidization was not markedly
different from earlier runs. Hence, the improvement in performance
was attributed to an increase in gas—solid contact time.
Concentration of CO + 2 in Reducing Gas
The air/fuel ratio determines the concentration of CO and H 2 in the
reducing gas. The fuel used was a mixture containing 60 weight per
cent propane (C 3 H8) and 40 weight per cent propylene (C 3 H 6 ). A com-
puter program (3) was used to estimate the air/fuel ratio necessary
to give the desired concentration of CO + H 2 . The program calculated
equilibrium concentrations as a function of equivalence ratio, ,
defined as
= (air/fuel) stoichiometric/(air/fuel) operating
An equivalence ratio of 1.4, for example, produces a gas product con-
taining about 14% CO + H 2 .
The percentage of CO + H 2 used in the reducing gas varied between 2 and
23%; however, the concentration most often used was about 15%. There
appeared to be no significant effect of the concentration of reducing
gas on the concentration of S02. Lower air/fuel ratios produce more
CO + 112 and consequently higher CO/CO 2 and 112/1120 ratios. However,
there was insufficient data to determine whether increased CO/CO 2 and
112/1120 ratios caused increased formation of CaS as would be expected
from equilibrium considerations.
Auxiliary Fuel
Run 120 (2020°F, 10 atm) was the first run made with a portion of the
fuel added above the fluidizing grid, directly into the bed of solids.
The remainder of the fuel and all of the air entered through the burner
and produced a strongly oxidizing flame. Additional fuel added above
the fluidizing grid produced a reducing gas containing approximately
10 CO + B 2 .
Injection of fuel above the fluidizing grid significantly affected the
temperature distribution in the bed. Thermocouples located at points
6, 12, and 18 inches above the grid indicated temperatures less than 10°F
apart. This is a much closer distribution than was ever achieved
without help from the electrical bed heaters, which were inoperative.
Not only is the uniformity of temperature significantly better when a
portion of the fuel is burned in the bed, but combustion of fuel in
this manner produces higher bed temperatures than introduction of the
same amount of fuel into a burner located beneath the fluidizing grid.
For example, run No. 123 (Table 12) was made at five atmospheres pressure
with auxiliary fuel added 6 inches above the fluidizing grid (auxiliary
air added 42 inches above grid). An object of this run was to determine
80
-------
the maximum bed temperature that could be attained at five atmospheres
pressure using auxiliary fuel and preheated air entering the burner.
The temperature achieved was 1970°F. This compares to a bed tempera-
ture Of 1840°F, attained in run No. 116, in which all of the fuel and
air entered the burner, with none entering the bed. Moreover, the
total flow of fuel was about 10% less in run No. 123 than in No. 116.
Another advantage of burning fuel directly in the bed is that the
tendency of the bed to agglomerate is reduced. No melting or
agglomeration of solids occurred in runs Nos. 120—124, made with
auxiliary fuel, even though nitrogen was used to reduce the temperature
of the burner flame in only two of these runs.
Auxiliary Air
A persistent problem with regeneration has been formation of large
proportions of GaS, instead of CaO. Elemental sulfur and reduced sul-
fur compounds have also formed and an oxygen stream near the top of the
regenerator (above the fluidized bed) was provided early in this work
to oxidize these substances to S02. For run No. 119, this oxygen stream
was removed and, instead, air was introduced into the bed at a point 12
inches above the fluidizing grid. The expectation was that any CaS which
formed near the bottom of the bed would be oxidized to CaSO4 or CaO
upon contacting the air. Reducing gas was produced at the burner, in
the usual manner. With good mixing in the fluidized bed, solids would
be alternately exposed to strongly reducing and oxidizing (or midly
reducing) environments. U-nfortunately, something less than this could
be expected to occur because mixing of solids in the regenerator is
poor. Poor mixing was indicated in previous runs where the temperature
distribution in the beds of solids was poor and also by results with
the cold model test unit which showed that slugging of solids occurred,
rather than uniform fluidization.
The results of run 119 (2000°F, 10 atm, air injected into bed) were
different from most runs made previously at higher pressures. The
average SO 2 concentration attained was 1.54% (corrected for dilution of
the off—gas by N 2 which is present in the auxiliary air), or 33% of
equilibrium, which is not unusual. However, the S02 concentration in
the off—gas was nearly constant over the duration of the run, in
previous work at 10 atm, the S02 concentration peaked sharply at the
start of a run and continually decreased as the run progressed. An
even more spectacular effect of auxiliary air was a sharp reduction in
the amount of CaS formed during regeneration. Without auxiliary air,
regenerated solids contained, typically, 40—70 mole percent GaS. With
the addition of auxiliary air, the CaS content of solids was reduced
to under one per cent in the majority of runs for which analyses were
made. This is shown in Table 13.
-------
Table 13. ANALYSIS OF SOLIDS FROM REGENERATION RUNS
Run No.
Temp., °F
Pressure, atm.
Sup. vel., ft./sec.
% Stoich. air
% Dii. N2
Set. bed ht., ft.
Aux. air
Aux. fuel
66—2690—98
2010—2060
9
4.0
66
16.3
2
NO
NO
66—2690—100
2080
10
4.3
66
16.7
2
NO
NO
66—2690—10 1
1910—1940
10
2.0
67
20
2
NO
NO
66—2690—105
1720—1880
10
1.9
69
16.7
2
NO
NO
Run No.
Temp., °F
Pressure, atm.
Sup. vel., ft./sec.
% Stoich. air
Z Dii. N2
Set. bed ht., ft.
Aux. air
Aux. fuel
66—2690—115
2050
10
3.6
77
0
2
NO
NO
66—2690—116
1840
4.9
4.1
70
0
2
NO
NO
66—2690—119
2000
10
3.7
78
5
2
12 in. above grid
NO
6 6—2 690—120
2020
10
3.8
79
0
2
Nob
6 in. above grid
Wt. % Mole %
Wt. % Mole %
wt. %
0.9
89.7
9.4
Solids Composition
Note %
Wt. %
Mole %
Wt. %
Mole 7.
Wt. 7.
Mole 7.
Wt.%
CaS
47.1
50.6
62.9
66.8
73.4
73.9
73.5
69.1
cc
CaO
CaSO4
47.8
3.2
39.9
6.5
34.5
0.8
28.6
1.6
21.3
3.4
16.6
6.5
14.9
9.6
10.9
17.1
CaCO 3
+
CaSO3
<1.8
<3.0
<1.8
<3.0
<1.9
<3.0
<2.0
<2.9
Solids composition
Ca S
CaO
CaSO 4
CaCO3 + CaSO 3
Mole % Wt. 7. Mole 7 .
41.3
45.2
57.7
60.5
0.2
0.2
0.7
54.7
4.0
46.6
8.2
38.0
4.3
31.0
8.5
98.4
1.4
96.5
3.3
95.2
4.1
a
-------
Table 13 (continued). ANALYSIS OF SOLIDS FROM REGENERATION RUNS
Run No. 66—2690—122b 66—2690—123
Temp., °F 2000 1970
Pressure, atm. 10 5
Sup. vel., ft./sec. 2.8 4.0
% Stoich. air 72 72
%Dil. N 2 0 0
Set. bed ht., ft. 2 2
Aux. air 42 in. above grid 42 in. above grid
Aux. fuel 6 in. above grid 6 in. above grid
Solids composition Mole % Wt. % Mole % Wt. %
CaS 8.3 10.0 0.2 0.3
CaO 88.5 82.8 ‘98.1 95.7
CaSO4 3.2 7.2 1.7 4.0
CaCO 3 + CaSO3 a
Neg11g1b1e
Solids exposed to oxidizing mixture (107% stoich. air) for several minutes after run.
-------
Particle Size
All but one run was made with —10 + 20 mesh material. To test the
effect of particle size, run No. 126. was made with —20 + 40 mesh
Drierite. Unfortunately, problems with equipment necessitated a
shutdown several minutes after the start of regeneration. The unit
was restarted several days later with the bed still in place. The
temperature profile in the bed indicated poor fluidization. It is
probable that moisture condensing in the reactor during pre—heating
had caused some of the bed to agglomerate. Hot spots developed and
the run had to be terminated. Average bed temperature was 1880 + 80°F.
Examination of solids indicated that the bed had agglomerated but lit-
tle fusion had taken place.
The average concentration of S02 in the off—gas was 1.2% although a
peak concentration of 1.9% (corrected for N2 addition in auxiliary air)
was observed. Based on an average bed temperature of 1880°F, these
levels represent approaches of 50 and 70 percent, respectively. How-
ever, when the uncertainty in the average bed temperature is taken into
account, the approach to equilibrium could be anywhere in the range of
35—100 percent. Hence, it cannot be clearly determined if SO 2 con-
centration was increased by reducing particle size. Nevertheless, if
poor fluidization is taken into account, an approach to equilibrium
of only 35% still appears high when compared to other runs in which
agglomeration of the bed had occurred.
Use of Sulfated Limestone
Drierite was the bed material for all runs except No. 125, in which
—10 ÷ 20 mesh sulfated limestone from the Bureau of Mines was used.
Early in the run, a hot spot developed near the point at which secondary
air entered the bed. For a brief period, temperatures at this position
approached 2370°F. The temperature profile in the bed was much poorer
than usually observed with Drierite. In addition, a poor temperature
profile was noted when the bed was preheated prior to regeneration.
Examination of solids showed that much of the bed had agglomerated in
the reactor. Fused pieces which were quite hard were also found. It
is likely that the bed had begun to agglomerate at the outset of the
run (possibly during preheating). Diluent nitrogen was introduced early
in the run; however, by this time, high bed temperatures were already
observed. The average SO 2 concentration was only 0.62% although a
peak concentration of 1.57% was observed, which corresponds to 52% of
equilibrium.
Quality of Fluidization
Quality of fluidization is probably the most poorly controlled and most
difficult to evaluate of all factors which affect performance of the
regenerator. The best means available to determine how well solids
84
-------
were fluidized was to check the closeness of the temperature distri-
bution in the bed. In several runs, where solids agglomerated, tem-
perature differences across the bed were as high as 500°F. A typical
span in temperature for beds in which no agglomeration occurred was
150°F. Where auxiliary fuel was used, a spread in temperature of
50°F was common, but the improvement in temperature distribution was
probably caused by the additonal fuel and not by better fluidization.
Analysis of Solids
Solids from some regeneration runs were analyzed for CaSO4, CaO, CaS,
CaCO3, and CaSO 3 . The analytical techniques that were used are given
in the Appendix. Table 13 gives results of the analyses. No signi-
ficant amounts of CaCO3 or CaSO 3 were found. The amount of unconverted
CaSO 4 was usually small. Run Nos. 119, 122b, and 123 were made with
auxiliary air; solids from these runs had levels of.CaS which were
sharply lower than solids from runs made without auxiliary air.
Figure 30 is a plot of the molar ratio of CaO/CaS vs. temperature for
runs made without auxiliary air. Increased temperatures appear to favor
higher CaO/CaS ratios. All data except for run 116 are for ten atm.
pressure. Run 116, made at five atm, has a higher CaO/CaS ratio than
would be expected at this temperature from data at ten atmospheres.
This might indicate that reduced pressure favors higher CaO/CaS ratios.
Attrition
Attrition in the batch fluidized bed regenerator was very low, typically
1—2% of the weight of Drierite charged.
85
-------
Figure 30
Regeneration studies - effect of temperature on CaO/CaS ratio in regenerated soI ds
2.0
1.8 — • 10 ATM RUNS
o 5 ATM RUNS
1.6 —
1.4—
.1...
E
1.0—
L)
V.L-)
0
0.6 —
0.4 —
0.2 —
0
1800 1900 2000 2100
TEMPERATURE, °F
-------
SECTION VII
DISCUSSION OF RESULTS
REGENERATION STUDIES
The S02 concentration measured in the product gas at 10 atm
pressure averaged about 2%. This corresponds to 40 to 50% of the
concentration calculated if the gas and solids were in chemical
equilibrium. This suggests that the SO 2 concentration is limited by
reaction rates. This is supported further by data which indicate that
higher temperatures, higher pressures, longer gas/solid residence
times and possibly smaller solids particles give SO 2 concentrations
which are closer to the calculated equilibrium levels. The reaction
rates can be increased by any number of changes in operating conditions.
However, only a few of these options are open because of certain
practical limitations. The two most likely ways of improving the con-
version of CaSO4 to SO2 and CaO are to increase the overall residence
time and to improve the gas/solids residence time distribution by
improving the quality of fluidzation.
For a given regenerator vessel diameter, the best way to increase the
residence time is to decrease the gas linear velocity. However, for
a given SO 2 production rate, a larger diameter regenerator vessel
with an increased inventory of solids would be required. Increasing
the residence time by increasing bed depth is not as attractive since
the bed pressure drop would be increased, solids entrainment from the
regenerator would be increased and the quality of fluidization would
probably be poorer, giving a less favorable residence time distribution.
Improvements in the residence time distribution could be made by
optimizing the distributor design and also using shallower beds. The
use of baffles to promote better fluidization does not appear to be
as likely since the very high temperatures occurring in the regenerator
pose a materials problem. Also, if the bed were to agglomerate,
baffles would complicate the job of cleaning out the bed.
Other ways of increasing the S02 concentration in the regenerator off
gas do not appear to be promising. For example, higher temperatures
will increase the probability of bed agglomeration, especially when
fly ash carried over from the combustor is present. Higher pressures
may increase the rate, but decrease the equilibrium concentration of
S02, so the net effect is small or negative. Using smaller particles
does not appear practical since it would require a larger combustor
diameter to prevent excessive solids entrainment. Dropping pressure
to 1 to 2 atm will result in an increase in SO 2 concentration by making
the equilibrium more favorable, even if rates are decreased. However,
this would require operating the regenerator at a lower pressure than
the combnstor, and devel3pment of high temperature lock hoppers and
valves would be necessary. In addition, it may not be possible to
provide enough thermal energy in the hot reducing gases fed to the
87
-------
regenerator to carry out the regeneration at low pressure without
diluting the product gas and decreasing the 502 concentration. If this
is the case, the SO 2 product concentration could be limited by the anlomit
of feed gas required to provide the needed energy and not by either
chemical equilibrium or reaction rates. Operating at lower pressure
would also require higher fluidizing velocities and/or larger reactor
diameters.
Introduction of auxiliary air directly into the regenerator bed pro—
motes high conversion of CaSO4 to CaO with very little CaS formation.
It is believed that air injection creates an oxidizing zone in the bed
above the reducing zone. The solids move back and forth between the
zones so that CaS formed in the reducing zone by over—reduction of
CaSO4 can be oxidized back to CaSO4 or directly to CaO and S02 in the
oxidizing zone. This ultimately gives high conversion of CaSO4 and
high selectivity to CaO.
Injection of at least a portion of the fuel directly into the bed also
improves operation by providing higher and more uniform bed tempera-
tures. There was also less tendency to agglomerate the bed when direct
injection was used. Higher bed temperatures are due to less heat
losses compared to the case where all fuel is burned in an adjacent
burner zone. Burning part of the fuel in the bed also causes the
thermal energy to be released over a greater volume and in direct con-
tact with the bed solids. This promotes more uniform bed temperatures
and probably results in the transfer of energy between the gas and
solids at a lower gas temperature. This, in turn, decreases the chance
for localized overheating of solids which can cause solids to soften
and agglomerate.
The best method of operating the regenerator appears to be to add both
auxiliary air and part of the fuel to the bed at separate points. The
optimum split between air and fuel added to the burner and directly to
the bed and the location of the injection points must be determined.
COMBUSTION STUDIES
Initial SO 2 removal of about 857. has been measured with both
—7 mesh calcined limestone and haif—calcined dolomite at combustion
pressures of 5—8 atm. However, the runs were too short to draw any
general conclusions about the suitability of either sorbent at these
conditions. Also, the Tymnochtee dolomite showed very high attrition
rates. Unless this is related somehow to the equipment or the manner
in which the dolomite was used, it would appear to rule out the use of
Tymochtee dolomite in a once through operation.
The NO levels measured in the flue gas are very high considering the
high combustion pressures. However, because coal feeding was very
erratic during the runs, the excess air level was much higher than
planned. It is believed that high NO emissions were due to the high
oxygen levels in the bed.
88
-------
After resolving the coal feeding problem, the combustion program will
continue testing other sorbents and measuring the effects of operating
conditions on SO 2 and NO emissions.
EQUIPMENT DEVELOPMENT
Development of equipment became a significant part of the experimental
program. The use of high temperature ceramic materials in the
regenerator did not prove successful due to poor resistance to thermal
shock and reaction with molten bed material. Water cooled stainless
steel components performed very well under the same conditions.
A burner was developed in this study to heat and provide reducing gas
feed for the regenerator and also to preheat the combustor. This
burner proved to be very versatile, operating over a wide range of
pressures, oxidizing and reducing conditions and flow rates.
Coal feeding at pressures up to 10 atm is possible with a modified
Petrocarb injector, but careful attention must be paid to the operation
to prevent plugging. Close control of the pressure differential between
the injector and the combustor, the flow of injection air, lack of sharp
bends or corners in the transfer line and coal particle size and dryness
must be maintained. Some problems still remain in the proper design and
location of the coal probe which injects the coal into the hot combustor.
The continuous transfer of solids between the coinbustor and regenerator
by the use of gas pulses is feasible. The system works well at ambient
temperature and based on other studies (4) should work even better at
higher temperatures due to increased gas viscosity.
89
-------
SECTION VIII
PROGRAM F0R OPERATION OF MINIPLANT
The overall objective of the Miniplant program is to demonstrate the
continuous fluidized bed combustion process at conditions anticipated
for commercial plants. The specific objectives are to
• Verify the reductions in SO 2 emissions
• Measure the reduction in NO emissions
x
• Verify particulate loadings in the flue gas
• Measure erosion of a stationary target simulating a
turbine blade surface in the flue gas
• Test various coals in combination with various sorbents
(limestones, dolomites). Select sorbents, based on location,
suitable for the 30MW demonstration unit
• Verify overall operability of the continuous system including
turndown ratio arid methods
• Test and develop (if necessary) system components
• Measure, verify and/or demonstrate other items needed for
the design and operation of the 30MW demonstration unit
At the present time, no plans are being made to operate the regenerator
section of the Miniplant. Therefore the operation will be based on the
once-through use of sorbent. The primary operating variables whose
effects will be studied are: combustion pressure and temperature,
sorbent/coal feed ratio, coal and sorbent type. Other variables which
cannot be as readily changed or whose effects may be less will be
studied to a lesser extent. These are: fluidized bed level, coal
and sorbent particle size, excess oxygen in the flue gas and coal rate.
Fluidizing velocity is fixed by the coal rate and excess oxygen.
The primary measured quantities will be:
• Flue gas composition
• Used solids composition
• Temperature distribution in the bed
• Solids loss and particle size change
90
-------
• Corrosion and erosion effects including the effects on the
stationary test piece in the flue gas
• Burning front location by measurement of intra bed gas
compos it ions
• Overall system response to changes in operating conditions,
including response to turndown and turnup measures
• General equipment performance
From these measurements, the following will be calculated:
• Fraction SO 2 reduction
• Combustion efficiency
• Heat transfer coefficients
• Fraction of bed attrited
• Heat losses
The initial operation of the combustor will consist of testing various
coal and sorbent combinations. Because of the somewhat limited objectives
of this initial study, only the effects of the primary operating
variables will be studied, the other variables ‘will be fixed. The run
plan will consist of a block of eight runs for each sorbent and coal
combination. The runs will be as follows:
Temp (°F) Ca/S Mole/Mole Press (atm )
1500 2/1 10
1600 2/1 10
1700 2/1 10
1700 4/1 10
1500 4/1 10
1600 4/1 10
1600 3/1 10
1600 3/1 5
The sorbents and coals tested will be chosen in cooperation with the EPA.
The choice will depend on the results of current experimental studies
sponsored by the EPA and the location of the 30MW demonstration plant.
Other test conditjOflS, such as the possible use of precalcined sorbent,
will be set before the tests begin.
After the initial program has been completed, the program will be directed
toward solution of problems which become evident during the initial
program. At this time, the effect of the secondary variables can also
be studied if warranted. Another option would be to study regeneration
of sulfated sorbent, if that appears to be warranted at that time.
91
-------
SECTION IX
REFERENCES
1. Jonke, A. A., et al., Reduction of Atmospheric Pollution by the
Application of Fluidized Bed Combustion. Argonne National
Laboratory, Monthly Progress Reports 32—35, June — September 1971.
2. Zenz, F. A., and Othmer, D. F., Fluidization and Fluid Particle
Systems, Reinhold Publishing Corporation, 1960, p. 136—174, 313—350.
3. Gordon, S., and McBride, B., Chemical Equilibrium Program. NASA
Lewis Research Center, November 1970.
4. Craig, J. W. T., et al., Study of Chemically Active Fluid Bed
Gasifier for Reduction of Sulphur Oxide flnissions. Final Report
OAP Contract CPA 70—46, Esso Petroleum Co., June 1972.
ADDITION REFERENCES
5. Archer, D. H., et al, Evaluation of the YIuidi.zed Bed Combustion
Process, Vol. I, 11, III, Westinghouse Research Laboratories,
November 1971.
6. Hammons, G. A., and Skopp, A., A Regenerative Limestone Process for
Fluidized Bed Coal Combustion and Desulfurization, Esso Research
and Engineering Co., February 28, 1971.
7. Skopp A., et al., Studies of the Fluidized Lime—Bed Coal Combustion
Desulfurization System, Esso Research and Engineering Company,
December 31, 1971.
92
-------
SECflON X
LIST OF PUBLICATIONS
1. Hoke, R. C., Shaw, H., and Skopp, A., A Regenerative Limestone
Process for Fluidized Bed Coal Combustion and Desulfurization.
Presented at the Third International Conference on Fluidized Bed
Combustion, College Corner, c io, cktober 29 — November 1, 1972.
2. Bertrand, R. R., Hoke, R. C., Shaw, H., and Skopp, A., Combustion
of Coal in a Bed of Fluidized Lime. Presented at the American
Chemical Society National Meeting, 1icago, Illinois, August
26—31, 1973, also submitted for publication in Hydrocarbon
Processing .
3. Hoke, R. C., Fluidized Bed Combustion of Solid and Liquid Fuels.
Presented at the Symposium on Modern Developments in Combustion
Technology, Pennsylvania State University, August 1, 1973.
PATENT MEMORANDA SUBMITTED
1. Ruth, L. A., Hoke, R. C., Reduction of CaSO4 by direct injection
of fuel into a fluidized bed.
2. Hoke, .R. C., Vath, E. C., Removal of SO 2 from flue gas followed
by sulfite precipitation.
3. Ruth, L. A., Shaw, H., Vath, E. C., High pressure gas generating
burner.
93
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SECTION XI
GLOSSARY
Abbreviation Definition
atm atmosphere — unit of pressure
BTU British Thermal Unit
CMTU Cold Model Test Unit
dia. diameter
d/p differential pressure
pressure difference
temperature difference
degree Fahrenheit
F coal feed rate, lb/hr
FRC flow recorder controller
ft. foot
gm. gram
hr. hour
I.D. internal diameter
IR infra—red
Kw kilowatt
lb. pound
mm. minute
NDIR non dispersive infra—red
O.D. outside diameter
P pressure in coal receiving vessel, psig
ppm parts per million
PRC pressure recorder controller
94
-------
psi pounds per square inch
psia pounds per square inch absolute
psig pounds per square inch gage
R injection air rotameter setting, % of scale
SCFM standard cubic feet per minute
sec. second
sq. ft. square feet
wt. weight
equivalence ratio
p micron — unit of length
Conversion Factors — English to Metric Units
English system tric equivalent
Length inch 2.54 centimeter
foot 0.305 meter
Area square foot 0.093 square meter
Volume gallon 3.785 liter
cubic feet 38.32 liters
Mass pound 453.6 grams
Pressure pound per square inch 51.70 millimeters Hg
atmosphere 760 millimeters Hg
Temperature ° Fahrenheit 1.8 (°Celsius) + 32
Energy British thermal unit 252 calories
95
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SECTION XII
APPENDIX
TECHNIQUES FOR ANALYSIS OF SOLIDS
Solids from some regeneration runs were analyzed for CaSO 4 ,
CaO, CaS, CaCO3, and CaSO3. The analytical techniques that were used
are described below.
SO 4 2 — The sample was treated with acidic BaC1 2 solution.
The BaSO 4 precipitate was weighed.
CO 3 2 — Acidic H 2 02 was added to the sample. The solution
was stripped with N 2 and the gas passed through
H 2 0 2 /H 2 SO 4 , Drierite, CuSO 4 , and Ascarite. C0 3 2
was determined from the weight gain of the Ascarite.
+2 +2
Ca — Ca was determined by atomic absorption.
SO 3 2 — The sample was treated overnight with HgCl , then
heated with acid and stripped with N2. The gas was
passed through excess 12, S 2 0 3 2 , and H 2 0. The
scrubbers were combined and the solution back titrated
with S 2 0 3 2 . SO 3 was determined from 12 consumed.
— Since no S0 3 2 was found, the sample was treated with
acidic 12 and the excess 12 titrated with S 2 0 3 2 . S 2
was determined from 12 consumed.
o 2 — CaO was determined by difference. The excess number of
moles of calcium over the moles of anions was assumed
to correspond to 2.
96
-------
Table A-i. FIXED BED SIMULATED COMBUSTION RUNS
Bed Mat’l.
Temp.,
°F
1600
Press.,
psig
0
Gas rate,
SCFM
0.11
Feed gas
comp.
Prod. gas comp.
Cony.,
%
CO,
ppm
945
NO,
ppm
1400
C02,
%
——
02,
%
——
CO,
ppm
10
NO,
400
C0 2 ,
%
——
02,
%
——
CO
99
of
NO
71
Limestone 1359 a
1600
0
0.10
988
862
17
——
770
635
17
——
22
26
1700
0
0.10
988
862
17
——
800
625
17
——
19
27
1700
0
0.10
1870
1800
——
——
160
22
——
——
92
99
1600
135
0.10
985
893
18
— —
300
165
13
——
70
82
1600
135
0.10
1870
1800
——
——
160
23
——
——
91
99
1600
135
0.10
900
1400
——
——
10
375
——
——
99
73
Dolomite
1600
0
0.1
900
1400
20
350
——
——
98
75
1600
135
0.1
900
1400
——
——
10
280
——
——
99
80
1600
0
0.1
983
840
16
——
830
680
17
——
16
19
1600
135
0.1
983
840
16
——
270
105
13
——
73
88
1600
135
0.10
1080
978
15
2.3
100
112
16
2.3
91
89
1600
135
0.10
2080
1987
——
——
85
180
——
——
96
91
1600
0 0.10 2080 1987 — — —— 95 243
—— — — 95 88
-------
Table A—i (continued). FIXED BED SINULATED COMBUSTION RUNS
a_ 18 ÷ 20 mesh, calcined.
b.. 16 + 25 mesh, calcined.
Bed Mat’l.
Limestone 1359 a
Feed gas comp. Prod. gas comp. Cony., %
Temp.,
°F
Press.,
ps ig
Gas rate,
SC 4 —
CO,
pp i
NO,
pp n
CO2,
%
02,
%
CO,
ppm’
NO,
ppm
C0 2 ,
Z
02,
%
of
CO
NO
1600
135
0.10
1245
1150
12.7
——
330
355
15.2
——
74
69
1400
135
0.82
1245
1150
12.7
——
1055
965
15.7
——
15
16
1600
135
0.40
1245
1150
12.7
990
888
15.7
——
20
23
1600
135
0.55
1245
1150
12.7
1040
905
15.7
——
16
21
1550
135
0.70
1245
1150
127
——
1080
940
15.7
——
13
18
1520
135
0.66
1245
1150
12.7
——
1095
935
15.7
——
12
19
1600
135
0.10
2080
1987
——
——
155
177
——
——
93
91
1600
135
0.10
1080
987
15.2
2.3
255
115
15.6
2.2
76
88
1600
135
0.10
1060
972
163
2.4
50
90
16.8
2.5
95
91
1600
135
0.10
2080
1987
——
——
38
163
——
——
98
92
1600
135
0.10
1245
1150
12.7
——
205
300
15.5
——
84
74
1600
1. 00
135
135
0.55
0.10
1245
965
1150
978
12.7
——
——
——
890
0
840
100
15.5
——
——
——
28
100
27
90
-------
Table A—2. AIR, FUEL AND SOLIDS INPUTS FOR REGENERATION RUNS
Injected above bed to prevent sulfur deposition
0.17 SCFM C02 also added.
C 20 40 mesh; —10+20 mesh used in all other runs.
in exit lines.
a
02
0.83
0.80
—— 2.34
—— 0.98
—— 1.00
—— 0.76
Run No.
Charge
material, weight,
lbs.
Fuel
0.21
0.32
Input in_SCF1 _ —
Air
5.0
4.9
Aux. fuel
Aux. air
N 2
0.67
0.64
66—2690—
71
72
Drierite, 9.2
Drierite, 18.4
75
Drierite, 9.8
1.11
20.8
78
Drierite, 13.1
0.46
8.4
86
Drierite, 13.1
0.52
8.4
89
Drierite, 13.1
0.53
8.9
98
Drierite, 6.6
1.35
20.5
100
Drierite, 6.6
1.58
24.0
101
Drierlte, 6.6
0.75
11.6
105
Drierite, 6.6
0.73
11.6
109
Drierite, 6.6
1.23
24.0
115
Drierlte, 6.6
1.38
24.3
116
Drierite, 6.6
0.93
15.0
119
Drierite, 6.6
1.00
18.0
——
3.46—7.52
0.82—2.32
120
Drierlte, 6.6
0.92
26.3
0.53—0.92
0—2.63
121
Drierite, 6.6
0.42
11.6
0.21—0.33
2.42—4.26
l22A
Drierite, 6.6
0.36
9.6
0.22
3.464.34
122B
Drierlte, 6.6
0.74
19.4
0.44
7.84--8.95
123
Drierite, 6.6
0.45
13.8
0.38
2.42—6.28
124
Drierite, 13.2
0.61
17.8
0.63—0.82
3.5—7.0
125
sulfated limestone,
9.35
1.2
20.0
0.12
7.4
2.7
126
Drieritec, 15.6
0.7
21.0
0.1
7.0
3.98
4.80
2.85
2.33
2.48
2.98
2.30
2.30
2.4
2.2
2.2
-------
Table A—i. COMPOSITION OF EF}1 UENT STREAM FOR RE(ENERATION RUNS
Concentrations have been
F un1d by difference.
Second column gives concentrations
dN2 which is contained in auxiliary air.
Calculated. Water not measured.
Run No. S02, % CO, ppm C02, %
Component, molar % or ppma
66—2690—
71
3.04
3 .. 37
1120, %
02, %
9.83
N 2 ,
NO,
ppm
1570
40 C
11.6
l 2 . 9
10.6
ll. 7
64.8 718 C
125
139 C
72
4.78
5.23
1750
1910
13.7
14.9
14.6
16.0
8.54
58.2 63.7
150
164
75
6.94
7.52
2360
2560
14.7
16.0
14.3
15.5
7.71
56.1 60.8
260
280
78
1.81
2.01
2590
2900
12.3
13.7
15.2
16.9
9.87
60.5 67.2
——
——
86
1.58
1.79
3630
4130
14.0
15.9
14.7
16.7
11.9
57.4 65.2
——
——
89
1.56
1.73
3000
3320
15.6
17.3
21.9
23.5
9.9
50.7 56.3
——
——
98
3.0
3.1
1200
1230
12.5
12.8
14.6
15.0
2.5
66.9 68.6
——
——
100
1.5
1.6
——
——
14.9
15.7
15.0
15.8
4.8
63.8 67.0
——
——
101
1.0
1.1
——
——
12.7
14.4
13.5
15.3
12.0
60.8 69.0
——
——
105
109
115
0.6
2.8
1.8
1.1
3.0
1.9
——
——
--
——
——
--
14.8
13.0
13.3
16.9
14.0
14.1
11.9
13.0
15.7
13 • 6 d
14.0
16 • 7 d
12.5
7.0
5.8
60.2 68.4
64.2 69.0
63.4 67.3
——
——
160
——
——
170
°
116
0.96
1.07
——
——
13.7
15.2
13.0
14.4
10.0
62.1 68.5
210
230
119
1.30
1.54
630
750
14.1
16.7
12.0
14.2
<1
72.6 67.6
190
220
120
1.57
1.57
9000
——
15.2
——
13.0
——
<1
69.3 ——
180
——
121
1.05
1.29
870
1070
13.1
16.1
12.7
15.6
<1
73.2 67.0
105
130
122A
0.44
0.54
17
21
11.8
14.6
12.6
15.6
<1
75.2 69.3
100
120
122B
1.60
2.02
61
77
12.7
16.0
12.4
15.6
<1
73.3 66.4
110
140
123
1.90
2.30
260
310
10.5
12.5
15.0
17.9
<1
72.6 67.3
100
120
124
125
126
1.67
0.42
1.0
2.05
0.62
1.2
880
2.0%
0.8%
1080
2.5%
1.0%
16.7
12.0
13.0
20.5
14.9
15.9
12.3
15.6
13.9
15.1
19 • 5 d
l 7 .O
<1
<1
——
69.3 62.3
70.0 62.5
71.3 64.9
110
110
190
140
140
230
corrected for condensation of water prior to analysis.
corrected also for injection of oxygen above bed OR addition of
-------
BIBLIOGRAPHIC DATA 1. Report No. 2.
SHEET E PA-650/2-74-001
3.’. ecipient’s Accession No.
4. Title and Subtitle
5. Report Date
A Regenerative Limestone Process for Fluidized-Bed
Combustion and Desulfurizat ion
Coal
January 1974
6.
7. Author(s)
R. C. Hoke, M. S. Nutkis, L. ARuth, and H. Shaw
8. Performing Organization Rept.
No. GRUS. 1 4GFGS. 74
9. Performing Organization Name and Address
10. Prpjcct/Task/Work Unit No.
Esso Research and Engineering Co.
P.O. Box 8, Linden, NJ 07036
ROAP 2IADB-13
U. Contract/Grant No.
CPA 70-19
11 Sponsoring Organization Name and Address
EPA, Office of Research and Development
NERC-RTP, Control Syst ems Laboratory
Research Triangle Park, NC 27711
13. Type of Report & Period
Cot ered
Final
14.
15. Supplementary Notes
16. Abstracts The report gives results of an experimental study of the pressurized comb-
ustion of coal in a I luidized bed of limestone and regeneration of sulfated lime-
stone. The study is part of a program to develop fluidized—bed coal combustion as a
means of desuif urizirl.g flue gas in-situ and generating clean power at low cost. The
process, including regeneration of spent limestone by reduction to lime, produces
a gas stream containing a sufficient concentration of S02 to be fed to a by—product
sulfur recovery unit. The combustion runs were limited by operating problems,
es ecia1ly plugging in the coal injection line. Initial S02 removal rates were about
85%; however, attrition rates were high with one S02 sorbent, Tymochtee dolomite.
The regeneration step was studied at pressures up to 10 atm and temperatures up to
2100°F. S02 concentrations measured in the product gas averaged about 2% at 10 atm
and 2100°F, or about 40% of the concentrations calculated by assuming equilibrium
be een the solids and regenerating gas.
High conversion of sulfated material to
lime was achieved by injecting air into
the bed, by forming adjacent reducing
and oxidizing zones, and by minimizing
formation of undesired CaS.
17. Key words and Document Analysis. 170. Descriptors
Air Pollution
Des ul.furization
Flue Gases
Re generation
(Engineering)
Limestone
Calcium Oxides
Fluidized-Bed Processors
Combustion
17b. identifiers Open-Ended Terms
Air Pollution Control
Stationary Sources
Regenerative Limestone Process
17c. COSATI Field/Group
13B, 2lB
Report 1
Unlimited UNCLASSIFIED j 101
18. Availability Statement r 9 - Security Class (This I2LNo. of Pages
20. Security Class (This 22. Price
Page
UNCLASSIFIED
USCOMMDC r49 52- 0 72
FORM W V 5-35 REV. 3-72)
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101
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