EPA-650/2-73-039
November 1973
Environmental Protection Technology Series
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EPA-650/2-73-039
CHEMICALLY ACTIVE FLUID-BED PROCESS
FOR SULPHUR REMOVAL
DURING GASIFICATION
OF HEAVY FUEL OIL
SECOND PHASE
bv
J.W.T. Craig, G.L. Johnes, G. Moss,
J.H. Taylor, and D.E. Tisdall
Esso Research Centre
Abingdon, Berkshire, England
Contract No. 68-02-0300
ROAP No. ADD-BE
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
November 1973
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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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ABSTRACT
This report describes the second phase of studies on the CAFE
process for desulphurising gasification of heavy fuel oil in a
bed of hot lime.
The first test of the continuous pilot plant with U.S. limestone
BCR 1691 was hampered by local stone sintering and severe
production of a sticky dust during start up conditions. Batch
tests confirmed that BCR 1691 produced more dust than either of
the higher purity Denbighshire or U.S. BCR 1359 stones. With
5CR 1691 dust production rate was tenfold higher during kerosene
combustion at 870 deg. C than during gasification/regeneration
cycles.
Modifications were made to the continuous pilot plant to
improve operability under dusty conditions, and a total of 332
gasification hours was made in a second run with Denbighshire
and BCR 1691 stones in six operating periods, the longest being
109 hours.
Sulphur removal efficiency was comparable for the two stones,
ranging from 60 to over 95%. Regenerator performance was less
satisfactory than in earlier tests, and a poor sulphur material
balance indicates need for improved analytical procedures.
An engineering scoping study estimate that total CAFB development
through a large demonstration test will take about 6-7 years
and require $3,320,000 in engineering effort.
This report was submitted as a requirement of Contract No. 68—02—
0300 by Esso Research Centre, England under the sponsorship of
the Environmental Protection Agency. Work during the first half
of the above contract is included. This was completed in April
1973.
iii
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CONTENTS
Page
Abstract
List of Figures
List of Tables
Acknowledgments
Sections
I Conclusions 1
II RecommendationS 4
III Introduction 5
IV Experimental Equipment 10
V Oil and Limestone 18
VI Pilot Plant Operation - Run 4 20
VII Batch Unit Studies 29
VIII Run 5 Pilot Plant Modifications 46
IX Pilot Plant Operation — Run 5 50
X Scoping of Engineering Effort 82
XI References 85
XII List of Inventions 86
XIII Glossary 87
XIV Appendices 89
iv
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F IGURES
P age
1. Batch Unit Flow Plan 10
2. Batch Unit Reactor 11
3. CAFB Pilot Plant Flow Plan 13
4. Revised Main Burner 17
5. Photomicrographs - Regenerator Fines 26
5A Combustion Conditions SB Gasification Conditions
6. Photomicrographs of cyclone solids - Batch Unit 40
7. 40
8. 41
9. 41
10. 42
11. CAFB Revised Pilot Plant Flow Plan 47
12. CAFB Pilot Plant Sulphur Removal Efficiency 59
13. Vanadium Retention Run 5 61
14. Sodium Retention Run 5 62
15. Nickel Retention Run 5 63
16. Distribution of Fluid Beds in Run 5 69
17. Gasifier and Regenerator Fines below 600 microns 70
during Run 5
18. Average Size of Bed Particles Run 5 71
19. Heat Release vs. Air Fuel Ratio - During Gasification 75
20. Regenerator Selectivity - Conversion Data 80
V
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TABLES
P age
1. Batch Unit Gas Analysis Equipment 12
2. Properties of CAFE Test Fuel Oils 18
3. Chemical Properties of Test Limestones 19
4. GasificatiOn Summary Run 4 21
5. particle Size Analysis Run 4 22
6. Chemical Analysis of Lime Samples Run 4 23
7. Test Program for CAFE Batch Units 32
8. Summary of Batch Unit SRE Results 34
9. Fines loss 35
10. Loss Rates 37
11. Calcination Losses 39
12. Nature of Cyclone Fines from Batch Unit Studies 43
13. Summary of Run 5 Operating Periods 51
14. Operating Conditions During Run 5 Test Periods 58
15. Summary of Solids Losses Run 5 64
16. Silica/Calcium Oxide Ratios Run 5 67
17. Nitrogen Oxides in CAFB Boiler Flue Gas 73
18. Heat Release in CAFB Gasifier Run 5 74
19. Product Gas Composition Run 5 76
20. Summary of Gasifier Component Distributions 78
21. Summary of Regenerator Performance Run 5 79
22. CAFB Development Programme. Summary of Engineering 64
Effort.
vi
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ACKNOWLEDGMENTS
The support of Esso Petroleum Company Limited is acknowledged
for the construction of a 10 million BTU/hr Chemically Active
Fluid Bed Gasifier facility at the Esso Research Centre,
Abingdon, Berkshire, England, and for its use in the
generation of continuous gasification data for this project.
The authors would also like to thank Mr. O.R. Priestnall, Mr. J.
Buzzacott, Mr. D. Storms and Mr. J. Cocker for their strenuous
efforts in generating the experimental data on which this report
is based.
Finally, they acknowledge the assistance given by other Esso
Petroleum Company personnel in operating the pilot plant, in
maintenance of equipment, and in chemical analysis of test samples.
vii
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SECTION 1
CONCLUS IONS
TASK I
1. Stability and quality of continuous pilot plant operations
were greatly improved by modifications which included
uninterrupted limestone addition, regenerator back pressure
control, flue gas recycle scrubbing, nitrogen quench for
regenerator over temperature protection, and a positive
pressure pilot flame for the main burner.
2. No significant difference was found between the sulphur
removal ability of BCR 1691 and Denbighshire limes at
comparable operating conditions in the continuous pilot
plant.
3. Lime replacement rate has a significant effect on sulphur
removal efficiency with stoichiometric Ca/S replacement
giving about 80% sulphur removal at 18 to 25 inch water
gauge (w.g.) bed depth.
4. Increasing gasifier bed depth increases the rate of lime
loss into the boiler and thus may obscure any effect of
depth on sulphur absorption. No correlation of sulphur
removal efficiency with bed depth was found in the region
of 18 to 25 inches w.g.
5. Regenerator performance in the continuous pilot plant
was not as good as experienced using Denbighshire stone in
Phase I studies. Quality of fluidisation was poorer, and
selectivity of oxidation of CaS to CaO plus SO 2 was lower.
Regenerator off gas concentrations remained below
4% for much of the test. Also, less than half the sulphur
removed from the fuel (as measured by flue gas analysis)
was recovered as SO 2 by the regenerator (as indicated by
regenerator gas analysis) . The reduction in regenerator
performance may be due to the absence of preferential fines
circulation from gasifier cyclone to regenerator in this
test.
6. Regenerator SO output as measured by gas analysis was
significantly lower than values calculated from solids
analysis.
7. In the continuous pilot plant test, metals removal from
the fuel by the lime averaged 36% for sodium, 75% for
nickel, and essentially 100% for vanadium.
8. Cyclone fines return performance deteriorated during the
third day of the second continuous pilot plant test in this
phase of studies, and subsequent periods experienced high
loss rates of lime from the gasifier into and through the
boiler.
1.
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9. In the continuous pilot plant with Denbighshire lime, most
of the gasifier solids losses were recovered in the boiler
or in the external flue gas cyclone. With BCR 1691, most
of the lime losses passed through both boiler and external
cyclone.
10. Variations in cyclone fines return system performance
during the second continuous pilot plant run were reflected
by changes in bed particle size distribution.
11. Stainless steel cyclone liners are not satisfactory as a
means of providing a smooth cyclone surface. In the pilot
plant test they failed under decoking conditions and
provided a surface for increased carbon deposition.
12. Retention of fine lime within the gasifier and avoidance
of deterioration in cyclone performance are the major
operating problems remaining in the CAFB development.
13. Operation of the gasifier bed at high enough levels to
assure a high concentration of solids in the gas entering
the cyclones reduces the rate of carbon deposition in
these ducts and increases the operating interval
between decoking operations.
14. Combustion of CAFE gasifier product in the pilot plant
burner produces less nitrogen oxides (166 ppm average)
than direct combustion of the fuel oil (263 ppm average)
15. The heat release by fuel partial combustion in the
gasifier is approximately 3100 Btu/lb at 20% of
stoichioxnetric air. This value agrees with thermodynamic
estimates based on the fractions of carbon and hydrogen
oxidised and the amount of CO produced.
16. Sintering of 3CR 1691 stone during unit startup is possible
but can be avoided by adjusting the gasifier/regenerator
pressure balance to avoid direct impingement of the startup
burner flame into solids transfer ducts.
TASK II
1. Lime from 3CR 1691 stone produces much more dust under
ccmparable CAFE fluidisation conditions than either of the
higher purity limes 3CR 1359 and Denbighahire.
2. Dust production with 3CR 1691 lime is particularly severe
during combustion with kerosene at 870 deg. C, the normal
CAFE start-up condition.
3. Kerosene combustion at 1050 deg. C causes less dust production
than combustion at 870 deg. C both with BCR 1691 and with
Denbighshire lime.
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4. The fine dust produced during 870 deg. C combustion with
BCR 1691 lime is sticky in nature and clings to pipe walls
and cyclone internals unless mechanical force such as
rapping is employed to dislodge it. Under gasification
condition3 the dust is not sticky, and is produced at a
lower rate.
5. Sulphur removal activity of BCR 1359 appears to be iritermed-
iate between Denbighshire and BCR 1691 limestones.
TASK III
An engineering scoping study by Esso Engineering indicates that
total CAFE development through a 100 + MW demonstration test
period is expected to take about 6¼ years and require $3,320,000
in engineering effort. Optimistically the development time
might be reduced to 4-¼ years with a cost of $2,520,000, but
risks associated with the large unit would be correspondingly
increased.
3.
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SECTION II
RECOMMENDATIONS
1. Pilot plant efforts should be shifted from BCR 1691 stone
to BCR 1359 which poses much less severe operating problems
from a dust production and potential sinterthg stand point.
2. Future tests of potential CAFE limestones should include
measurements of their dust producing characteristics under
the fully cornbusting conditions encountered during startup
and hot standby operations.
3. The pilot plant should be modified to permit reexamination
of the effect of preferential circulation of gasifier fines
direct to the regenerator.
4. Any CAFE installation should provide steady addition of
limestone to prevent temperature variations caused by the
changing heat load.
5. An emergency regenerator quench system should be included
in CAFB installations to prevent sintering and agglomeration
by temperature upsets.
o. A scrubber is needed in CAFB installations for the flue
gas recycle to prevent fouling the recycle system, blowers,
and air distributors.
7. Means of ensuring a high circulation rate of solids through
the gasifier cyclones should be provided to decrease carbon
deposition rate and to increase run length between decoking
steps.
8. Investigation of a means to provide smooth cyclone interior
surfaces should be continued.
9. Boiler gas and regenerator gas rate measurements, sampling
systems, and analytical procedures should undergo contin-
uing study and revision until discrepancies in sulphur
material balance values are resolved. In particular, a
high velocity, hot flue gas sample system with hot cyclone
and filter is recommended for the flue gas analysis
procedure.
10. Regenerator operation should be tested at lower air rates
to confirm if reduced CaS conversion level will improve
selectivity of CaS oxidation to CaO and reduce the quantity
of Ca50 4 returned to the gasifier.
11. An increased level of effort on data analysis and engineering
study of various features of the process is recommended
to improve understanding of results already obtained.
4.
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SECTION III
INTRODUCTION
GENERAL
The Chemically Active Fluid Bed process is a means of avoiding
sulphur oxide pollution while using heavy fuel oil for production
of power. The process uses a fluidised bed of lime particles
to convert the oil into a hot, low sulphur gas ready for combustion
in an adjacent boiler. Sulphur from the fuel is absorbed by the
lime which can be regenerated for reuse. During lime regeneration
the sulphur is liberated as a concentrated stream of SO 2 which may
be converted to acid or elemental sulphur.
Exploratory work on the CAFB began at the Esso Research Centre,
Abingdon (ERCA) in 1966. Batch reactor fuel and limestone
screening studies, a variable study with U.S. limestone BCR 1691,
and initial operation of a pilot plant incorporating continuous
gasification and regeneration were conducted from 1970 to 1972
under Office of Air Programmes (OAP) contract 70—46 and were described
in the final report (1) for that contract, dated June 1972.
The current contract represents the second phase of work on this
project. This interim report covers work on this phase for the
period July 1, 1972 through April 1973.
GASIFIER CHEMISTRY
When heavy fuel oil is injected into a bed of fluidised lime under
reducing conditions at about 900 deg. C, it vaporises, cracks,
and forms a series of compounds ranging from H, and CHA through
heavy hydrocarbons to coke. The sulphur contained in he oil
forms compounds such as H 5, COS and CS with H 2 S predominating.
The sulphur compounds reagt with CaOto orm CaS and gaseous
oxides.
For example:
CaO + H 2 S ) CaS + H 2 0
The equilibrium (1) for this reaction is far to the right. With
a fuel containing 4% sulphur the equilibrium permits a desuiphurising
efficiency greater than 90% up to 1100 deg. C. Other factors
however limit gasification temperature to the range of about 850
to 900 deg. C. where the equilibrium sulphur removal would be about
99%.
In the shallow fluidised bed of the gasifier there is a rapid
circulation of lime between top and bottom. Indications are that
coke is laid down on the lime in the upper portion of the fluid
bed by oil cracking and coking reactions and that this coke burns
off in the lower portion where oxygen is supplied by the air
distributor.
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Gasification conditions of temperature and air—fuel ratio must be
chosen to maintain a balance between the rate of coke and carbon
deposition and the rate of carbon burnoff. Broadly, this balance
is met at gasification temperatures in the range of 650 to 900
deg. C. and air—fuel ratios around 20% of stoichiometric. Lower
air fuel ratios are operable at the upper end of the temperature
range, and higher air-fuel ratios are needed as temperature is
reduced.
Much of the oxygen entering the gasifier is consumed in oxidising
coke to Co and CO near the distributor. Of course, some enters
other regions of he bed where it reacts with H and hydrocarbons
to form water and more carbon oxides. The fina product from the
gasifier is a hot combustible gas containing H hydrocarbons
CO, CO , H 2 0, and N 2 . Most of the energy rele ed by partial
combus ion of the fuel is retained by this gas as sensible heat.
Only a portion of the CaO in the lime is reacted on each pass of
solids through the gasifier. Good sulphur absorption reactivity
has been obtained with up to 20% of calcium reacted in single
cycle batch reactor tests, but in the continuous unit, the
average extent of calcium conversion to suiphide is held to less
than 10%.
When a single batch of lime is cycled between gasification and
regeneration conditions it gradually loses activity for sulphur
absorption. The activity of the bed can be maintained if some
of the lime is purged each cycle and replaced by fresh material.
Reactivity of the bed is therefore a function of the lime replace-
ment rate. The replacement lime is usually added to the gasifier
as limestone which calciries in situ.
Vanadium from the fuel oil deposits on the lime during gasification.
Experimental evidence is that practically all of the fuel vanadium
can remain fixed with the lime.
REGENERATOR CHENIS TRY
Calcium suiphide is regenerated to Calcium oxide by air oxidation.
CaS ÷ 3/2 02 CaO + SO 2
= -109.5 K cal/mole
A competing reaction also consumes oxygen and forms calcium
sulphate.
CaS + 2 02 CaSO 4
LH -220.2 1< cal/mole
Both reactions are strongly exothermic. A third reaction between
the solid species is also possible.
CaS + 3CaSO 4 4CaO + 4S0 2
221.5 K cal/mole
This reaction is strongly eridothermic.
6.
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The equilibrium constants (Appendix A, Reference 1) for these
reactions determine the maximum partial pressure of SO 2 which
can exist in equilibrium with mixtures of CaS, CaO, and CaSO 4
at any given temperature. These equilibria also determine a
relationship between regenerator temperature and the maximum
theoretical selectivity of oxidation of CaS to CaO.
At low oxidation temperature, the equilibrium SO 2 partial
pressure is too low to permit all the oxygen supplied to leave
in the form of SO 2 . The excess oxygen then goes to form CaSO 4 .
Experimental oxidation selectivities are lower than the theore-
tical maximum, probably because of contacting and kinetic
factors.
Since each sulphided lime particle passes through a range of
temperatures and oxygen concentrations during its transit through
the regenerator, it is exposed, on average, to less favourable
selectivity conditions than those at the top of the bed.
Calcium sulphide oxidation selectivities to CaO of 70 to 80% and
regenerator SO 2 concentrations of 8 to 10% have been achieved
in pilot plant operations at regenerator temperatures in the
range of 1040 to 1070 deg. C.
During the conversion of CaS to CaO and CaSO4 there is evidence
for existence of a transient liquid state (2) . If air is passed
through a hot static bed containing CaS, some of the particles
will agglomerate into lumps during the regeneration reaction.
Agglomeration does not occur if the bed is vigorously fluidised.
PREVIOUS EXPERIMENTAL WORE
A basis for the CAFB process had been established by experiments
in 7-inch i.d. batch reactors at ERCA prior to 1970.
During 1970, two new batch reactor units were constructed for
the OAP contract. Work in these batch units established the
suitability for CAFB of a Venezuelan fuel oil available in the
U.S., and selected the better of two U.S. limestones suggested
by OAP. Both of the U.S. stones, BCR 1690 and 3CR 1691, were
lower in CaO content than the U.K. stones tested previously.
In cycle tests the 3CR 1691 stone gave sulphur removal activity
comparable to that of the higher purity U.K. stones at equal Ca/S
ratios. The 3CR 1690 stone was found to be unsuitable in three
respects. It gave lower sulphur recovery at equal Ca/S ratio; it
attrited badly; and it sintered. and agglomerated during regener-
ation. The 3CR 1691 stone therefore was selected for further study.
During 1971 an intensive study of gasification variables was
conducted in the batch reactors with this stone. Tests with fresh
beds screened the effects of major variables including air fuel
ratio, gasification temperature, bed depth, lime particle size,
and gas velocity in the bed. The variables of lime replacement
rate and extent of calcium reaction between regenerations were
probed in cyclic tests where the lime was cycled between gasifying
and regeneration conditions in the batch reactor.
7.
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These studies provided the basis for a number of guidelines and
process correlations. The effects of bed depth (10 to 20 inches)
and fluidisation velocity (4 to 8 ft/sec) were correlated as
gas residence time in the bed, giving an approximately first
order sulphur removal rate expression. Sulphur differential,
the quantity of sulphur to which the lime was exposed in each
gasification cycle, emerged as an important variable. As an
approximation lime reactivity varied inversely with the square
of this differential and increased directly with lime to sulphur
replacement ratio.
In parallel with the batch unit experiments, ERCA constructed
a pilot plant in which the gasification and regeneration
reactions could be studied under continuous operating conditions.
A 10 million Btu/Hr water cooled boiler was included in the
system to burn the gasifier product and dispose of the heat.
Three tests designated CAFE Runs 1, 2, and 3 were made in this
pilot plant during 1971. Run 3 lasted 230 hours of which 204
were at gasifying conditions. Denbighshire limestone (UK) was
used throughout these runs together with Venezuelan fuel oil
containing 2.5% sulphur. The pilot plant successfully demon-
strated many features of the process including sulphur removal,
lime regeneration, temperature control, start up, shut down,
solids circulation, and release of the sulphur as a rich (8—10%)
stream of SO . It also pinpointed areas for improvement which
included redaction or elimination of carbonac oua deposits in
cyclones and gas transfer ducts, ininimisation of fines production
and losses into the boiler, and improvement of regenerator
oxidation selectivity.
CURRENT WORK OBJECTIVES
Work on this current contract constitutes the second phase of a
six phase programme to demonstrate and evaluate the CAFB gasifier
on a commercial scale with a power plant boiler. Work on this
phase consists of three tasks.
Task I is operation of the CAFE pilot plant with the following
set of objectives.
a. Verify that continuous gasification, sulphur and metal
removal and lime regeneration results are as batch studies
have indicated and evaluate the effects of bed depth,
velocity, fuel/air ratio, lime make-up and fuel rate.
b. Determine minimum excess air requirements for operation of
a continuous regenerator with good temperature control and
maintenance of a low residual concentration of sulphur on
the lime bed.
c. Demonstrate operability of the process over a prolonged
period of time to show that accumulation of fines, agglo-
merates, carbon or other deposits do not interfere with
continuous operation.
8.
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d. Demonstrate means of preventing or removing deleterious
accumulations of tar or carbon from gasifier and transfer
duct internals.
e. Determine effects of number and location of fuel injectors
on gasification, sulphur removal, and carbon content of
gasifier lime. Include operation with single oil injector
passing through the air distributor.
f. Test and demonstrate means of process start-up, shut—down,
turndown, and control. Determine maximum turndown ratio with
independent control of gasifier and regenerator variables.
g. Determine effect of regeneration temperature on the
maintenance of lime activity.
h. Study the existing burner operation with CAFB gasifier
product. Establish operability with high gas velocity,
measure flame characteristics, efficiency of combustion,
production of NO, and flame stability.
i. Under conditions of lined out operation with equilibrated
lime, measure SO removal in the regenerator and determine
rate of lime att ition and particle size distribution of
solids carried over from the gasifier and regenerator.
Determine engineering properties of equilibrium solids
such as fluidized bed density, minimum fluidization velocity
and particle size distribution.
Four pilot plant runs were planned to accomplish these objectives.
To continue the numbering system begun in Phase I, these runs are
designated runs 4,5,6 and 7.
Task II is an evaluation of additional limestones with two fuel
oils in batch reactor experiments conducted between continuous
unit runs. One of the test oils is to be high sulphur residue
by—product of gas oil desulphurisation. Originally, four new
stones were to be studied. Because of factors uncovered during
the first pilot plant run, the batch unit programme has been
modified to include measurements of dust production tendencies
of the stones under combustion conditions and to reduce the
number of stones investigated to three.
Task III is a definition and assessment of the scope of engineering
effort required to move the CAFB from the pilot plant stage
through the development stage including the demonstration unit.
Tasks I and II are in progress at the Esso Research Centre,
Abingdon, England. Task III was conducted by the Esso Research and
Engineering Co., Florham Park, N.J., U.S.A.
REPORTING AND DISCUSSION OF RESULTS
During the performance of Task I the objectives for Task II were
modified as a result of information generated during the first
continuous gasifier run under Task I. Consequently, the reporting
and discussion of results of Tasks I and II is set out in Section
VI through IX in chronological order, so that the sequential
logic of changes to objectives, equipment and techniques can be
readily followed.
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SECTION IV
GENERAL
EXPERIMENTAL EQUIPMENT
The experimental equipment used in this study consists of two
batch reactor units and the continuous CAFB pilot plant. These
units have been described (1) previously in detail. For the
current work, the batch units remain essentially unchanged.
However several modifications were made to the pilot plant on
the basis of experience gained in the first three runs.
BATCH UNITS
Each batch unit contains a reactor, air and fuel systems, flare
for product gas disposal, and gas sampling and analysis system
as shown in the flow plan, Figure 1.
FIGURE 1 Batch Unit Flow Plan
Sançle Gas
Pump
Sample
flon*
Sample Gas
to
Analysars
Air Blower
CAFB
Botch Reactor
From Heated
Air Meter
1’
Feed
Fuel Injector Air
Propane for start-up
10.
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A reactor is illustrated in Figure 2.
FIGURE 2 Batch Unit Reactor
The reactor is of refractory lined carbon steel construction. The
lower section, which contains the fluid bed, is 7 inches i.d. by
33 inches high. The upper section is expanded to reduce gas
velocity and is non symmetrical to permit internal cyclones to
drain externally.
The plenum beneath the gas distributor is also refractory lined to
serve as a combustion chamber for propane—air mixtures used during
unit start up. The distributor is a “top hat’ shape cast refractory
design. The central raised section, five inches in diameter,
contains 16 horizontal holes around its circumference.
I SwlIigL .
led $awiIsP,let.
IsID,i & —
Diat$eeter PI&t.
“ Ii ’
11.
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For the current work, a rapper was installed on one cyclone in
each reactor to prevent fine particles sticking to the cyclone
walls. The pneumatic activator of the rapper is located outside
the reactor and drives a striker rod through a gland to tap on
the cyclone wall.
The gas analysis equipment used in this study is the same as used
in Phase I, and is listed in Table 1.
Table 1
Batch Unit Gas
Analysis Equipment
Type Manufacturer Model Response Range
SO 2 Infra-red Maihak Unor 6 ContinuouS 0-1,000 ppm
so 2 infra-red Maihak Unor 6 0-20% by vol.
so 2 Conductiznetric Wostoff 0-1000 ppm
Co 2 infra-red Maihak Unor 6 0-20% by vol.
Co Infra-red Maihak Unor 6 0-20% by vol.
Parainagnetic Servomex OA 137 0-25% by vol.
During fully combusting conditions and also regeneration of
suiphide, the gas from the bed is monitored directly for CO
co, 0 and SO and the appropriate information on combusti n
effic &ncy, su’phur removal efficiency or sulphur release deduced.
During gasification, a portion of the product gas is burned in
a sample flame located just above the reactor and the combustion
products analysed for SO , 0 , CO and CO 2 . The desuiphurising
efficiency of the gasifi r i calculated from this analysis of the
fully combusted gas.BecauSe of the wide range of sulphur
compounds present in the gasifier product itself, this is the
only practical method for measuring gasifier sulphur removal
efficiency.
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CONTINUOUS PILOT PLANT
Process Flow Plan
Figure 3 is a process flow plan of the continuous pilot plant.
The heart of the system is the gasifier—regenerator unit cast
of refractory concrete contained in an internally insulated steel
shell. The product gas of the gasifier fires a 10 rnhllion/
Btu/hr pressurised water boiler. The hot water is heat exchanged
with a secondary water circuit which loses its heat through a
forced convection cooling tower. The rest of the system consists
of the necessary blowers, pumps and instruments to operate the
gasifier, regenerator, burner and solids circulating system.
FIGURE 3 CAFB Pilot Plant Flow Plan
The gasifier itself sits within a pit to permit alignment of the
gasifier outlet duct with the burner inlet. Fuel pumps, flow
meters, and start up burner controls are mounted on a mechanical
equipment console. Electrical instrumentation and manometers are
mounted on a separate control console. Gasifier blowers are
located in a separate blower house outside the main building, and
the cooling tower is mounted on the roof.
13.
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The gasifier and regenerator reactors are cavities in a single
refractory concrete block. The block contains other cavities
which make up the gasifier outlet cyclones, the gas transfer
ducts, and the transfer lines through which solids circulate
between gasifier and regenerator. The gasifier cavity is
rectangular in cross section, tapering from 17.5 x 37 inches at
the distributor level to 19.5 x 39 inches at the 21 inch level.
The upper portion has. parallel sides. The regenerator tapers
from 7” diameter at the bottom to 8” diameter 22” above dist-
ributor datum and remains parallel thereafter.
Modifications to Pilot Plant
A number of changes were made to the pilot plant between the end
of Run 3 and the beginning of the current work phase. These
modifications were intended to improve both working conditions
and to overcome operating deficiencies encountered in the earlier
runs.
Additional modifications made following Run 4 are discussed in a
later section of this report dealing with preparations for Run 5.
Those changes discussed here were performed in the period between
December 1971 and July, 1972.
Jontrol Cubicle -
The area including the control panel, analyser bank, and mechanical
console, (pumps, rotameters, etc) was walled off from the operating
area containing the gasifier, boiler, and heat exchangers. This
control cubicle was supplied with its own ventilation and heating
system to provide a dust free area both to simplify maintenance
of delicate equipment and to reduce dust exposure of personnel.
Dust Extractor -
A central dust extractor was installed in the main operating area
with flexible ducts leading to areas of possible dust emission
such as sample points, cyclone fines drains, and solids draw offs.
Limestone Feeder —
To avoid temperature upsets which had been experienced when lime-
stone feed was interrupted to fill the original stone feed hopper,
and to provide a constant measure of stone feed rate, a new stone
feed system was designed and installed. In the new system lime-
stone is fed to the gasifier through a gravity feed line from a
pressurised weigh hopper. A vibrator is used to control the rate
of stone addition. The feed line enters the gasifier through its
side at a point below the bed level. A flow of air through the
feed line keeps this line purged of gasifier product and helps
move the stone into the bed. A ground level hopper is charged
from bags of stone. A pneumatic transfer line moves the stone
from the ground level hopper to an upper hopper from which stone
is periodically dropped into the weigh hopper. The upper hopper
acts as an air lock to avoid depressurising the weigh hopper.
Readout from the weigh hopper is continuously recorded on a chart
from which both hopper content and rate of feed can be determined.
14.
-------
Regenerator Drain Control -
In the original control system, gasifier bed level was controlled
by automatic adjustment of fresh limestone feed rate. Rate of
solids draw—off from the regenerator was determined by manual
adjustment of a vibrator on the regenerator drain line
Experience showed that regenerator solids drain flow was not
dependable when throttled to a very low rate. Under low flow
conditions the drain line would frequently block with solids.
Also, when stone addition rate was controlled by gasifier bed
level it was not possible to set stone feed to a predetermined
rate. With the new limestone feeder it became possible to set
stone feed rate at any desired level. A modification was there-
fore made to transfer control of gasifier bed level to the rate
of regenerator solids draw off. The new regenerator drain
system incorporated an air operated plug valve in the drain line,
an adjustable cycle timer which controlled both the length of
valve opening time and the frequency at which the valve opened,
and a pressure switch operated by the depth of bed in the gasifier.
When gasifier bed depth was below the set point, this pressure
switch prevented the regenerator drain valve opening. When
gasifier depth exceeded the set point, operation of the
regenerator drain valve was under control of the cycle timer. With
the cycle timer system, solids flowed from the regenerator in short,
sharp bursts, which kept the drain line clear. Nitrogen injection
into the drain line above the valve helped keep the solids free
when the valve was closed.
Flue Gas Recycle -
A new blower was installed to boost pressure of the flue gas
recycle stream returned to the gasifier air blowers. A more
efficient cyclone was installed in the line to improve dust removal
from the recycle flue gas.
Re nerator Pressure Control -
Good control of solids circulation in the gasifier-regenerator
circuit requires that the pressure difference between the vessels
be closely controlled. In the first three pilot plant runs this
control was achieved by manual adjustment of a butterfly valve
in the regenerator outlet line. This procedure was inconvenient
and failed to provide completely satisfactory control. To
improve matters, an automatic pneumatic instrument was installed
to adjust valve position in response to the signal from a
differential pressure measuring cell.
Regenerator Gas Outlet -
The regenerator gas outlet line was repositioned to enter the flue
stack down stream of the external cyclone. This was done to remove
the hazard of SO 2 release when solids were removed from the cyclone
drain.
Air Distributors -
oEh the gasifier and the regenerator air distributors were
modified. The original regenerator air distributor was of the same
design as used in the batch units, a cast refractory “top hat” shape
with horizontal holes. After each of the first three pilot plant
runs, a lump of agglomerated solids was found in the regenerator
on the distributor. The shape of this lump indicated that it had
originated in the defluidised region immediately above the centre
of the “top hat’ distributor. Also the refractory distributor
cracked around its edge in each of the first three runs.
15.
-------
The refractory construction which led to the “top hat” shape is
needed in the batch units which employ propane firing beneath
the distributor for start up. However conditions in the pilot
plant regenerator are much less severe at the distributor, and
stainless steel construction is possible. Therefore a stainless
steel distributor was designed, tested and installed before
Run 4. This distribitor contains 5 stainless steel nozzles,
each with 4 horizontal 1/8 inch i.d. holes. Hole spacing in the
nozzles was selected to equalise air rate per unit area in the
regions inside and outside of the nozzle circle.
Nozzles of the gasifier distributor were modified to reduce the
attriting energy of the nozzle air jets. A shroud ring with
offset, larger diameter holes, was placed around each air nozzle.
Pressure drop through the inner holes of each nozzle still
determines overall pressure drop and assures uniform air
distribution. The inside surface of the shroud absorbs the jet
energy however and permits a more gentle entry of the air into
the bed.
yclone Outlet Tubes -
The central gas outlet dip tubes from the gasifier cyclones had
been a trouble area during the first three runs. These tubes
become coated with coke during gasification and are subjected to
extreme conditions when the coke is burned off. The original
gas outlets were simple stainless tubes which burned severely
during Run 1. The tubes were omxnitted completely during Run 2
with consequent high dust loss rates.
Double wall stainless steel tubes used in Run 3 had provision
for steam cooling during decoking operations. These tubes
survived the run with slight damage. However, their thickness
impaired cyclone operation.
Test specimens of various materials had been exposed to cyclone
tube conditions during Runs 2 and 3, and self bonded silicon
carbide was found to survive intact. New cyclone tubes of this
material were therefore obtained and installed before Run 4.
Gasifier Outlet Duct -
In early runs, coke and lime deposits formed in the gasifier
outlet ducts at points where the gas underwent sharp changes
in direction. Major deposits formed at the junctions between the
vertical and horizontal sections. After Run 3 the refractory
lining of the bifurcated gas duct was recast to provide a
smoothly rounded curve from the vertical to horizontal sections.
Gas Burner -
The gas flame reached nearly to the end of the primary fire tube
of the boiler during the first set of runs. An attempt to shorten
the flame by increasing the proportion of air added to the gas at
the burner entrance was limited by increasing temperature within
the burner throat. The burner was therefore redesigned after
Run 3 to add a third air inlet point to increase mixing and
turbulence at the gas outlet. Figure 4 shows the revised gas
burner.
16.
-------
FIGURE 4 Revised Main Burner
Dampers and pitot tubes on the three air ducts leading to the
burner permit individual adjustment of the rate to each point.
A new air blower was installed to replace the 1200 CFM blower
supplied with the boiler. The new blower will permit a higher
firing rate and can overcome the added back pressure of dust
removal equipment added to the flue gas system prior to Run 3.
Boiler Probe -
An air cooled probe was installed at the rear end of the primary
fire tube to simulate a boiler superheater tube and test for the
nature and extent of deposits or corrosion formed on such a
tube. The probe is 1.5 inches L.d. by 39½ inches extension
inside the boiler wall. Air for probe cooling is supplied by a
diesel driven auxiliary compressor. A temperature controller
which senses probe metal temperature by a thermocouple adjusts
air rate to the probe. Tube temperature profile is recorded
continuously by other thermocouples mounted along its length.
c’J
Pnmary Air
17.
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SECTION V
OIL AND LIMESTONE
OIL
The oils used in the present work have been heavy fuel oils from
Venezuelan crudes obtained from Axauay refinery of Creole
Petroleum Co. One supply, obtained in drums direct from Amuay
has been used in batch unit tests. A second supply from bulk
storage in the U.K. has been used in pilot plant work. An
additional supply of very heavy residue has been obtained for
batch unit studies which have not yet begun. This oil, also from
Amuay, is vacuum pipe still bottoms produced when ordinary
atmospheric pipe still bottoms is further distilled to give a
vacuum gas oil which can be hydrofined to give a reduced sulphur
fuel oil. Chemical and physical inspection results of these
three oils are listed in Table 2.
Table 2
Properties of
CAFB Test Fuel Oils
Amuay Amuay Pilot Plant Tests Vacuum
Batch Unit 1/12/71 14/3/73 26/3/73 Bottoms
Tests
Property
Specific Gravity 0.957 0.955 0.960 0.960 1.015
Kinematic Viscosity
CS at 140 deg. F 201 221 — 254
210 deg. F 41.4 39.3 40.3 45.6 3180
280 deg. F 321
350 deg. F — — 71
Carbon % by wt. 85.9 85.6 85.8 85.3 85.7
Hydrogen 11.3 11.4 11.6 11.3 10.3
Sulphur 2.35 2.48 2.42 2.43 2.95
Nitrogen 0.35 0.26 0.09 0.35 0.63
Conradson Carbon ‘ 11.6 10.9 11.1 10.8 17.4
Asphaltenes 7.1 4.8 5.3 6.0 6.9
Vanadium ppm 366 345 300 315 530
Nickel 43 40 65 41 6
Sodium 36 35 37 38 9
Iron 3 4 — 3 108
18.
-------
LIMESTONE
Pilot plant runs have employed U.S. limestone BCR 1691 and a U.K.
limestone from Denbighshire. Batch unit work has employed these
two stones as well as U.S. stone BCR 1359. Two additional “.S.
stones have been received. They are Tymochtee Dolomite and
a limestone supplied by New England Electric Systems which we have
designated NEES (1) . This was selected because of its proximity
to the proposed demonstration unit. No tests have been made
on these two stones as yet. Chemical inspections of all the
stones are listed in Table 3.
Table 3
Chemical Properties of Test Limestones
Limestone Composition
Stone: BCR 1691 BCR 1359 Denbighshire Tyrnoch- NEES
tee (1)
Dolomite
Component
CaO % by wt. 45.6 54.1 55.2 31.1 55.0
MgO 3.35 0.60 0.30 20.9 0.6
Si0 2 13.65 0.75 0.68 3.1 0.8
Fe 2 0 3 I’ 0.35 0.09 0.10 0.4 0.09
A1 2 0 3 2.80 0.31 0.25 1.13 0.3
CO 2 35.7 44.0 43.4 43.6 43.2
S (Total) 0.44 0.12 0.05 0.13 0.03
Vanadium ppm 41 50 26 < 25 <20
Sodium 219 < 20 59 175 250
Nickel 40 30 21 10 50
19.
-------
SECTION VI
PILOT PLANT OPERATION - RUN 4
GENERAL
The experimental plan for Run 4 called for use of limestone
8CR 1691 at a series of pilot plant conditions to test correlations
based on batch unit studies. Limestone replacement rate, gasifier
bed depth, and qasifier bed temperature were the major variables
to be examined. A brief test of the effect of regenerator excess
oxygen content was also included in the experimental plan.
A major goal of the study was to find if increasing gasifier bed
depth would have the beneficial effect that batch studies
had indicated to be possible. As events developed, properties
of limestone BCR 1691 prevented accomplishment of these test
goals and focussed attention on development of means to start up
and operate with a stone which produces a great deal of dust
under certain conditions.
The dust forming characteristics of this stone under full
combustion conditions were found to be much worse than had been
encountered with Denbighshire stone. This high rate of fines
production caused a number of operating problems, extended the
start up period, and eventually caused termination of the run
after nine hours of gasification. Results of the post run
inspection of the unit appear in Appendix B. Results and problem
areas encountered in the run are described here.
GASIFICATION RESULTS AND DISCUSSION
Table 4 summarises gasification conditions and results obtained.
Initially a high lime replacement rate of 2.1 mole CaS was employed
to build gasifier bed level. A slight reduction to 1.7 mole CaS
was used during the final 4 hours. Sulphur removal efficiency of
nearly 98% at the higher rate declined to about 93% when stone
rate was reduced. However the gasification period was too short
to consider these results to represent lined out conditions.
Table 5 lists the distribution of particle sizes in the solids
from gasifier and regenerator beds and in solids recovered from the
boiler fire tube and regenerator cyclone during gasification.
20.
-------
Table 4
Gasification Suinmxnary - Run4
Day 1-lour
Temperature deg. C
Gasifier Regenerator
Superficial
Air Rate
ft/sec
Fluidised
ed Depth
in.
Lime Replacement
Mol CaO/Mol S
Air/Fuel Sulphur Removal
% Stoich %
1
2030
870
—
3.7
19.5
2.1
23.1
1
2130
882
1035
3.8
20.0
2.1
24.0
t
1
2230
872
1040
3.7
21.9
2.1
23.6
1
2330
870
1080
4.3
23.3
2.1
24.3
95.7
2
0030
875
1100
4.1
23.6
2.1
24.1
97.9
2
0130
881
1110
4.1
23.8
1.7
24.1
93.3
2
0230
875
1015
4.0
22.7
1.7
23.7
93.3
2
0330
872
1068
4.1
22.2
1.7
24.0
87.3
2
0430
872
1060
4.1
23.0
1.7
24.2
92.6
-------
Table 5
Particle Size Analysis Run 4
Sample Location Gasifier Regenerator Solids from Boiler Regenerator
Bed Bed Fire Tube Cyclone Fines
____________ ___________________ ( gasification )
Particle Size,
wt % in Fraction
1400 Micron + 39.8 21.6 19.3 .27 *
1400 — 1180 11.5 8.8 6.3 )
1180 — 850 22.2 18.0 13.0 .27 *
850 — 600 15.3 13.9 13.0
600 — 355 11.0 14.7 20.5 .27 *
355 — 250 0.5 5.8 7.1 .27 *
250 — 150 0.2 12.7 5.1 10.6
150 — 106 0.1 0.8 2.8 11.4
106 — Dust 0.5 3.7 13.0 76.9
Bulk Density q/cc 1.13 1.08 .85 1.03
* These particles had a white appearance,
distinctly different from that of the
finer particles.
The elutriation effect of the gasifier bed in removing particles
smaller than the 355—600 micron fraction is apparent in Table 5.
We would expect particles smaller than about 500 microns to be
entrained at Run 4 test conditions. It is evident that little of
the entrained material was returned to the gasifier by the cyclone.
The presence of a wide spectrum of particle sizes in solids from
the boiler fire tube also indicates poor cyclone performance.
However the gasifier cyclone which drained back to the regenerator
evidently was operating as there was an appreciable fraction of
150—250 micron solids in the regenerator bed.
The bulk density of the bed solids was higher than has been
observed in earlier studies. Batch unit tests with 9CR 1691 had
given settled bed densities of about 0.83 g/cc compared with values
over 1.0 observed here. A change in density of the fluidised bed
had also been noted during the start up period of Run 4. This
density increased from about 0.8 to nearly 1.1 during the start up.
It is possible that a selective loss of lower density particles
contributed to this increase in bed density.
The chemical analyses of bed samples listed in Table 6 show that
silica content of the beds, and indeed all solids samples, increased
over those of the raw limestories. This change indicates that
minerals other than Si0 2 were preferentially lost from the system,
probably as very small particles.
22
-------
Fable 6
ChemicaI lsisofLimaSamP e Run4
Gasifier
Units (After Shutdown)
Regenerator
(During
Gasification)
Boiler
(After
Shutdown)
Regenerator Cyclone Fines
(During (During
Conthustiori) Gasification)
CaO
wt%
53.9
58.8
57.0
68.1
59.9
?4g0
“
3.95
4.7
4.35
4.1
4.2
A 1 2 0 3
Si0 2
Fe
“
“
‘
2.35
22.1
0.73
3.0
24.1
0.87
2.85
18.6
1.02
4.1
21.2
0.97
3.6
21.6
087
Na
“
0.06
0.07
0.07
0.06
0.08
V
“
0.13
0.15
0.11
0.04
0.38
CO 2
S (total)
“
“
1.84
4.85
0.73
2.09
0.36
1.56
0.21
0.74
0.12
5.51
S as sulphate
“
4.39
2.06
1.28
0.72
2.58
Loss on Ignition
“
—
—
7.46
0.31
Gain on Ignition
u
0.21
0.16
—
—
4.09
Si0 2 /CaO
Mg O/CaO
0.41
0.074
0.41
0.080
0.33
0.076
0.31
0.060
0.36
0.070
A 1 2 0 3 /CaO
0.043
0.051
0.050
0.060
0.060
S10 2 /CaO in Original
MgO/CaO
Limestone
I I
Al 0 /Cao
23
0.27
0.07 4
0 II
0.046
-------
The difference between gasifier and regenerator bed sulphur contents
was 2.8% on stone indicating a good level of regeneration. A
very high fraction, 99%, of the regenerator sulphur appeared as
sulphate. This represents a considerably higher degree of
suiphide oxidation than achieved in earlier runs and may indicate
some oxidation of the sample during its collection. The
regenerator cyclone fines show a slightly higher sulphur content
than the gasifier bed sample. They also show a high content of
sulphide which indicates that the fines passed through the regener-
ator without undergoing much reaction.
START UP PROBLEMS
Nature of Start Up Problems
During start up of the continuous pilot plant in Run 4, problems
were encountered in the following areas:-
(a) Blockage in solids transfer line
(b) Plugging in regenerator gas outlet system
(c) Dust emissions to boiler from gasifier
(d) Dust in flue gas recycle stream
(e) Dust emissions to atmosphere
(f) Regenerator Agglomerates
All of the problems were related to differences in the characteristics
of stone BCR 1691 from those of the Denl? hshire stone used in the
continuOus unit during Phase I studies . The major differences
are lower fusion temperature, the cause of problems (a) and (f) above
and production of a higher proportion of very fine dust in a
fluidised bed under fully cornbusting conditions, the cause of
problems (b) through (e). The dust produced from BCR 1691 is more
difficult to retain in collection equipment than that originating
from Denbighshire stone. It also clings to surfaces of pipes,
cyclones, control valves etc , and is difficult to dislodge without
application of direct mechanical force. It does not drain from
hoppers, or even vertical pipes, without continuous rapping.
Transfer Li Block
Heatup of the pilot plant for Run 4 started on July 28. Stone
addition was begun the afternoon of August 1, and by early on
Aug 2, a hot fluidised lime bed was established under kerosene
combustion conditions. However, efforts to establish good solids
circulation through gasifier and regenerator were unsuccessful.
Attempts to rod out the transfer lines did not improve circulation
very much. The unit therefore was shutdown on August 4, allowed
to cool and opened for inspection on August 8. The blockage was
found to consist of a fused mass of lime particles which obstructed
most of the mixing pocket in the regenerator to gasifier (R to G)
transfer line. The transfer line from this mixing pocket to the
gasifier also contained a quantity of material with the appearance
of foamed slag.
24.
-------
Reconstruction of the startup procedure indicated that the
obstruction was caused by limestone particles and fines entering
the R to G transfer line during the initial stages of stone
addition while the’pocket and transfer line were heated by direct
gas flame. The geometry of the system is now such that stone
enters the bed from the stone feeder at a point directly opposite
the R to G line. The start up burner is between the stone feed
point and the R to G line. During start up, gasifier pressure
had been maintained well above regenerator pressure to drive
hot gas into the regenerator to raise its temperature. It is
evident now that flame had actually entered the transfer line
along with stone. The silica content of the BCR 1691 stone
lowers its melting point enough to allow fusion under these
conditions, producing, in effect, flame spraying of fused stone
directly into the R to G transfer line.
The solution to this problem is to adjust pressure balance during
the early stages of stone addition to avoid overheating the
R to G transfer system. This method was adopted for the second
start up. As added precautions, a thermocouple was installed in
the transfer pocket, and low silica Denbighshire stone was added
initially to form a bed deep enough to cover the transfer line.
The experience gained during the second bed addition on 15 August
indicates that flow of hot gas and stone into the R to G line
can be prevented by adjusting the pressure balance and that no
trouble would have been experienced with the 1691 stone alone. The
thermocouple in the transfer pocket is a valuable guide to
temperature and flow conditions at that point, and its continued
use is recommended. No blockage was encountered during the
second stone addition, and once good fluidisation and combustion
were established, high circulation rates between beds were
easily obtained. Some difficulty was encountered in establishing
initial fluidisation with 300 to 3200 micron stone. Indeed each
startup had encountered some trouble during the initial period of
stone addition because of the high heat load required for stone
calcination and the high gas velocity required to fluidise
uncalcined stone. In this startup, good fluidisation and operating
bed temperature were achieved after addition of some precalcined
stone removed from the unit in Run 3. Use of calcined lime is
recommended for startup in future runs.
9 enerator Gas Outlet System
During the startup period of Run 4, dust in the regenerator off
gas stream continually blocked the regenerator cyclones and the gas
exit line and control valve down stream of the cyclone. It was
possible to keep the cyclone functioning only by continuously
rapping on its body Without this rapping, solids failed to drain,
quickly filled the cyclone interior, and went overhead to form
restrictions further down stream. It was found that the overhead
line remained relatively clean in straight sections and smooth
bends but plugged at sharp bends and fittings.
The character of regenerator fines changed when gasification began
on August 20. Almost immediately the regenerator cyclone became
free draining and operative without rapping.
The colour of the fines also changed to a darker hue. Microscopic
examination showed the dust from the combustion period to have a
high proportion of very fine particles.
2S
-------
fl ’; ? -
1 % - S. # F -. ‘-
r ,
- .
— * .. .
t t t.. -\
t . tt2 t .
Fig. 5A Photonicrograph of Regenerator Fines
Under Combusting Conditions.
____ ca
: 4ç ‘ç a __ p
___ 7 ;.
. i es r ¶4 CF 2flZ ’ -
: € :g 1
- I •.-
I , S , 4 i ___ .e 4r -
Sfl
-. ..e Ef4jf4 r
Fig. SB Photomicrograph of Regenerator Fines
Under Gasifying Conditions.
2€.
-------
Figure 5 shows photornicrographs of regenerator solids obtained
under com.busting conditions (5A) andgasifying conditions (5B).
The larger particle size and reduced agglomeration tendency of the
gasifying samples is apparent.
Nevertheless, the regenerator outlet control valve eventually
plugged after nine hours of gasification and the run was terminated.
Inspection of the outlet system revealed no accumulation of fines
except at the valve itself. The solids forming the plug appeared
to be more like those formed during combustion than during
gasification, and it is possible that they were a remnant of the
pregasification period which had dislodged from the transfer line
and moved down stream to the valve.
Dust Emissions to Boiler
A large quantity of dust passed from the gasifier into the boiler.
Much was retained within the boiler, particularly at the back end
of the main fire tube, where the flue gases change direction
abruptly through 180 degrees, but some passed through to be
caught in the external cyclone or to escape from the stack. There
are indications that fines losses decreased during gasification,
but the period was too short to confirm this observation.
Unlike previous operations where fines also entered the boiler,
this time we were unsuccessful in withdrawing solids from the
drain points at the boiler end. The solids failed to drain
because of their steep angle of repose.
A contributing factor to high losses from the gasifier was the
greater bed pressure drop used in the current run. In earlier
runs, a maximum bed pressure drop of 19 inches water gauge (w.g.)
was used. In the current run this was increased to 26 in w.g.
This increase in pressure drop together with a reduced density
of fluidised cyclone fines (due to lower average particle size)
caused level of solids in the cyclone drain line to reach the
cyclone itself and decrease cyclone efficiency.
Dust in Flue Gas_Rec 1e
Two stages of cyclones in the flue gas recycle system failed to
remove fines to a degree which would assure long term cleanliness
of the gasifier distributor. No pressure drop increase in the
distributor was observed during the short test period, but fines
observed in the flue gas sample line filter (downstream of the
cyclones) indicated that a problem eventually would have occurred.
External Dust Emissions
The external settling chamber and cyclone which proved adequate
for final flue gas clean up in Run 3 was unable to cope with the
fine dust produced under combustion conditions in the current run.
In part, poor performance of the external cyclone was also caused
by the sticky nature of the fines. The interior of the cyclone
was quickly coated with a layer of solids which impaired its
efficiency.
27.
-------
Regenerator Agglomerates
Analysis of Run 4 temperature records and inspection of samples
retrieved from the regenerator revealed an additional problem with
BCR 1691 stone. During the early portion of the Run 4 gasification
period, an upset in pressure balance interrupted solids circulation
and allowed a brief temperature excursion. The temperature in the
lower portion of the bed reached 1130 deg. C. After this upset,
the temperatures in the upper and lower regions diverged with the
lower temperature logging about 80 deg. C. below the upper one.
This condition indicates poor fluidisation. When the regenerator
was opened for inspection a nwnber of agglomerates were found.
We believe that these lumps formed during the brief high temperature
period in spite of bed fluidisation. In previous runs with
Denbighshire stone there had been temperatures of over 1130 deg. C.
without encountering similar losses in fluidisation. This
difference in behaviour is attributed to the lower fusion point of
the lower purity 8CR 1691 stone.
28.
-------
SECTION VII
BATCH UNIT STUDIES
GENERAL
In the original programme of work, batch unit studies were to
determine the suitability of additional fuel-limestone combinations
for CAFB applications. However since Run 4 operations revealed
a serious problem with BCR 1691 stone which had not appeared
in earlier batch unit tests, the batch programme was revised to
include an investigation of dust forming tendencies under a
variety of operating conditions.
In earlier batch work there had been a comparison of dust losses
between stones during gasification-regeneration cycles. No
such comparison had been made under fully combusting conditions.
[ n the normal batch unit test procedure there had been little
exposure of the solids to combustion conditions in the absence
of sulphur except during calcinatiori. Consequently the conditions
employed during Run 4 start up produced an entirely unexpected
result in that the BCR 1691 stone formed copious quantities of a
dust with a very sticky nature.
The batch unit test programme was revised to accornodate the
following objectives.
• Determine if continuous unit conditions which produced
large quantities of sticky dust could be duplicated in batch
units.
c Compare dust producing tendencies of Denbighshire and BCR 1691
stones under different conditions.
o Provide a quantitative measurement of dust production to be
expected under start up and operating conditions with
Denbighshire, BCR 1691, BCR 1359, and two additional stones
to be provided by New England Electric System (NEES).
• Measure sulphur absorption performance of BCR 1359 and the
two NEES stones.
o Conduct tests of the feasibility of operating CAFB with vacuum
pipe still bottoms.
BATCH UNIT PROCEDURES
Several test procedures were employed to measure sulphur absorption
and dust producing characteristics of the stones. These included
fresh bed tests in which a new batch of calcined lime was used for
each test, cyclic gasification tests in which a single batch of
lime was cycled between gasification and regeneration conditions,
and kerosene and fuel oil combustion tests in which the appropriate
fuel was burned in the fluid bed with excess air to control
temperature.
29.
-------
Startup_and Calcination
The bed is calcined by combustion of propane below the distributor
and by direct injection of kerosene into the bed. First of all,
the bed space temperature is raised to 950 deg. C. by gas combustion
below the distributor. Then, 4000g of limestone is added. This
is heated to 750 deg. C. using gas before switching to direct
kerosene injection. Due to the strongly endothermic nature of
the calcination reaction, the temperature remains in the region
of 800 deg. C. until C , evolution ceases. When the temperature
rises to 950 deg. C, inaicating that calcination of the original
charge is complete, 2500g limestone is added and calcined. Further
batches of this size are added and the procedure repeated until
the target bed depth is reached.
Combustion Test
For a test with kerosene combustion, the injection of kerosene
is continued with the bed past the point of complete calcination.
For fuel oil combustion, the fuel supply is simply switched to
heavy fuel oil from its heated supply drum. The fuel rate is set
to give a bed temperature slightly in excess of the target, and
fine adjustment is carried out with the aid of a cooling coil.
Relevant data on gas analysis and unit behaviour are recorded
and appropriate bed and cyclone samples taken. Combustion is
continued for the required length of time.
Gasification Test
To achieve gasification, the fuel supply is changed from kerosene
to fuel oil when calcination is complete, and oil rate is increased
to obtain an air/fuel ratio of about 25% of stoichiometric.
The sample flame burner and external flare are lit and the gas
analysers connected. Relevant data are recorded and bed samples
and cyclone samples taken at prescribed intervals. Gasification
is continued for the requisite length of time.
Regeneration
Regeneration of the suiphided stone from a gasification cycle is
performed by stopping the fuel supply and continuing the flow of
air. Oxidation of carbon and CaS in the bed raises temperature
to the regeneration level. Relevant temperature and analytical
data are collected, and a bed sample is taken when regeneration
is complete.
Shut Down
At the end of a run, the bed temperature is allowed to drop to
700 deg. C before the bed is removed through the drain point just
above the distributor. Draining the unit is much easier when the
bed is hot since the solids flow better under these conditions.
When all the bed has been removed, all ancillary equipment is
shut off.
30.
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Tes ts
Cyclic tests are the nearest simulation to continuous gasifier
operation that can be obtained in batch units. The same charge
of lime is subjected to repeated cycles of sulphur absorption
and regeneration. After each regeneration a portion of lime is
removed and replaced by an equivalent amount of fresh limestone.
The limestone calcines to lime during the early part of the next
gasification cycle. Without replacement, the activity of the lime
bed gradually declines. With replacement, the activity falls
initially, but in a few cycles lines out at an equilibrium level
which is influenced by the rate of replacement.
Enough cycles are performed at each set of operating conditions
to establish the lined—out sulphur removal efficiency for those
conditions.
Cyclic tests use the same calcination and start—up procedures
as the fresh bed tests. However, after the initial start, a
series of gasification and regeneration cycles follow each other.
Sampling and gas analysis procedures also are the same as used in
fresh bed tests.
When a regeneration cycle is complete, the fluidisation cools
the bed very rapidly. Since solid sample withdrawal is very
difficult in the absence of fluidisation, it is not possible to
draw a regenerated bed lime sample before temperature drops
below the level needed for the next gasification cycle. Therefore
the following procedure was adopted.
(1) When regeneration was complete as indicated by end of SO 2
emission and fall of bed temperature, fluidising air was
stopped.
(2) Replacement limestone was added.
(3) Fluidising air was resumed and when temperature reached the
desired point, oil feed was resumed.
For the purpose of stone comparison, the target conditions in each
test were the same. These conditions are listed below:-
Air/fuel Ratio (% of stoichiometri-c) 25
Gasificiation Temperature (deg. C.) 870
Make-up rate (wt CaO/wt S) 1.6
Bed Depth (inches w.g.) 15
Gas Velocity (ft/sec) 6
Potential Sulphur Differential (% wt) 2
Limestone Particle Size (microns) 600-3175
BATCH UNIT PROGRAMME
The batch unit test programme for Phase II is summarised in Table 7.
The design of this programme with regard to fines production rates
and properties was based very much on our experience in continuous
pilot plant run 4. There, fines produced from stone BCE 1691 under
fully combustirig conditions did not drain freely from cyclones,
whereas those produced under gasifying conditions did.
31.
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Table 7
Test Programe for CAFB Batch Units
Cond it i one
Calcinatton and prolonged combustion
at 870 deg. C
Same at 1050 deg. C
Calcination and combustion at
870 deg. C
Calcination and gasification at
870 deg. C
Gasification - regeneration cycles
ProLonged combustion using cycled
bed from Test. 1 -E
Gasification regeneration
cycles
Entire prograntrne same as in Test 1
through step F
Ca1cin ttion and prolonged combustion
at 870 dog. C
Calcination and prolonged gasification
at 870 deg. C
Gasification - regeneration cycles
Calcination and prolonged combustion
at 870 deg. C
Same
Gasification - regeneration cycles
Calcirtation and prolonged combustion
at 870 deg. C
Same
Test
1-A
1-B
1-C
1-0
1-E
1-F
1—G
2-A
through
ut 2 - F’
3-A
3-B
3-C
4-A
4-B
4-C
5-A
5—B
Stone Fuel
8CR 1691 Kerosene
Axnuay 2.5% S
Fuel oil
Kerosene
Aznuay 3% S
Vacuum Resid.
Denbighshire Kerosene and
Nnuay 2.5% S
F.0.
8CR 1359 Kerosene
Aniuay 2.5% 5
fuel oil
NEES stone Kerosene
Antuay 2.5% S
fuel oil
NEES stone Kerosene
2
Axnuay 2.5% S
fuel oil
Objective
Measure fines loss rate and properties of
dust produced with BCE 1691 stone
S ante
Sante
Same
Measure lined out sulphur removal effici-
ency and fines loss rate
Measure fines loss rate during kerosene
combustion in conditioned lime bed
Measure lined Out sulphur removal
efficiency and fines loss rate with very
heavy fuel
Measure fines loss rate under different
sets of conditions with oenbighshire
stone
Measure fines loss rate and dust
properties with BCR 1359
S ante
Measure lined out sulphur removal
efficiency and fines loss rate
Measure fines loss rate and properties of
dust produced with NEES stone 1
Same
Measure lined out sulphur removal
efficiency and fines loss rate
Measure fines loss rate and properties of
dust produced with 14EES stone 2
S ante
Measure lined out sulphur removal
efficiency and fines loss rate.
5-C
Gasification - regeneration cycles
-------
It was suspected that the higher resistance to flow of the
combustion fines was due to their containing a much higher
proportion of very fine particles. Why less of the very fine
particles should be produced during Continuous gasification and
regeneration was not clear. It was considered possible that the
presence of sulphur on the stone, the higher regenerator bed
temperature, an ageing effect or a combination of these could
be the answer. Any of these changes could have altered the
particle surfaces in such a way as to make them less susceptible
to decrepitation.
Work on this programme has now been completed to Test 3C except
for Test 1G with Amuay vacuum bottoms.
EXPERIMENTAL RESULTS
The results to date permit a comparison of sulphur removal
efficiencies, fines production rates, and fines properties from
three limestones, BCE 1691, Denbighshire, and BCR 1359.
Sulphur Removal Efficiency
Sulphur removal efficiencies (SEE) for the three stones were compared
by cycle tests. Adverse operating conditions were chosen so that
SEE would be less than 100%. When successive cycles of gasification
and regeneration are repeated, SEE, under adverse operating conditions,
falls for several cycles and then lines out. Comparison values of
SEE are measured at the lined out level. Detailed results of
the tests are listed in Appendix A. Results are summarised in
Table 8.
SEE’s are similar for the three stones. However, although the
target conditions were the same for each test, actual conditions
did vary slightly as shown in the table. Consequently, for
comparison, lined Out SEE was calculated for each of the test
conditions for each stone from the equation derived for 9CR 1691
From the ratio of measured to calculated SEE it appears that the
Deribighshire stone is the most active and 9CR 1691 the least
active. BCR 1359 is intermediate in activity but appears to be
closer to 1691 than Denbighshire.
Fines Production Rate
The measurement of fines production rate was based on bed losses.
Table 9 suinmarises the bed loss results for the three lixnestones
under a variety of CAFB conditions. Detailed results are given
in Appendix A.
With 9CR 1691, highest bed loss rate occurred during kerosene
combustion with a fresh bed at 870 deg. C. Loss rates during fuel
oil combustion and gasification where sulphur was being absorbed
by the stone were lower, those during gasification where sulphur
absorption was more rapid being less than during combustion.
33.
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Table 8
Sun n y of Batch Unit SRE Results
Residence P.S.D. 7 ’ Make—up Rate SRE SRE * Ratio of
Limestone Time CaO/S (Measured) (Calculated) SRE (Meas’d)
(sec) (wt %) wt. Mole to SRE (cal)
BCR 1691 0.20 1.96 1.59 .91 74.5 78.2 0.95
w
• BCR 1359 0.15 1.87 1.61 .92 76.0 76.0 1.0
Denbighshire 0.17 2.18 1.49 .85 76.0 64.9 1.17
* Calculated from equation for predicting SRE’s for BCR 1691.
7’ Projected Sulphur Differential
-------
Table 9
Summary of Batch Unit Fines Loss
Test Test Condition Limestone Temperature Total Solids Loss fron Bed, grams Average Loss Rate over
___________ _______ deg. C 1 hour 2 hours 3 hours 4 hours - 4 hours (g/ mini
1-A Kerosene Combustion 8CR 1691 870 2900 4220 5000 5430 22.6
1—B Kerosene Combustion ru 1050 555 960 1220 1360 5.7
1—C Fuel 0i1Coxthustion 870 1110 1910 2690 3250 13.5
1-0 Fuel Oil Gasification 870 680 1190 1560 1860 7.8
1—E asjfication-Regeneration Cycles 850 - 1050 . 19.8/4.5 *
1-F Kerosene Combustion Su lphided
Aged Bed 870 760 1300 1700 2040 8.5
2—A Kerosene Combustion Denbighshire 870 5 00 760 930 1020 4.3
2-B Kerosene Combustion 1050 520 680 750 790 3.3
2-C Fuel Oil Combustion 870 1160 1600 1920 2160 9.0
2-0 Fuel Oil Gasification 870 1460 1930 2160 2330 9.7
2-K asification—Regeneratlofl Cycles 850 — 1050 6.6/1.4 *
2-F Kerosene Combustion Suiphided
Aged Bed 870 48 96 140 188 0.8
3-A Kerosene Combustion 8CR 1359 870 200 340 440 510 2.1
3-B Fuel Oil Gasification tI 870 480 630 750 810 3.6
3—C Gasification—Regeneration Cycles 850 - 1050 6.4/1.9 *
* First value is loss rate during 1st cycle, second is lined out loss rate after 5 cyclel.
-------
The sulphided and aged stone from the cycle tests also showed a
relatively low loss rate during kerosene combustion at 870 deg. C.
The lowest loss rate measured by tests of the fresh bed type,
however, was obtained during kerosene combustion at 1050 deg. C,
the temperature level used in regeneration. This was only slightly
above the lined out rate for the cycle tests in which the rate fell
from 19.8 g/min in the first cycle to a stable level of 4.5 g/min
after 5 cycles. We can conclude, therefore that raising bed
temperature or introducing sulphur into the bed decreases BCR 1691
loss rate. This explains the reduction in losses observed when
gasification was commenced in Run 4. Sulphur had been adsorbed
by the bed which was also being subjected to the high temperatures
of regeneration. Prior to this, no sulphur had been introduced
to the bed since kerosene combustion had been employed, and the
regenerator temperature was low since no reaction was taking place.
A bed ageing effect is also indicated. Hourly losses during fresh
bed tests decreased as the tests proceeded whilst loss rate during
cycle tests had fallen from 19.8 g/min to 4.5 g/min by 5 cycles.
It is important to note, however, that ageing during fresh bed tests
is distorted to some extent due to the fact that bed depth was
decreasing during each test. That bed depth has an important
effect on loss rate was established in the Phase I work (l)
The relative importance of all variables effecting bed loss rate
of 3CR 1691 is summarised in the equation below. This was derived
from further analysis of the fresh bed test results.
223.31 X D 2 7
A 0.44 X (T—750) 1.80 S 0.41
L = Loss Rate (g/min)
D = Bed Depth (inches)
A Bed Age (hours)
T = Bed Temperature (deg. C) (750 deg. C is taken as
CaCO decomposition temperature)
S = Bed u1phur Content including inherent sulphur
(% by weight)
This equation shows the approximately square relationship between
losses and bed depth as observed previously. The loss rate of
3.1 g/min calculated from the above equation for cycle test
conditions is in fair agreement with the measured rate of 4.5 g/rnin.
With Denbighshire stone fresh bed tests showed a decrease in loss
rate with increased temperature and an increase when fuel was
used instead of kerosene. The stable rate during cycle tests was
lower than for fresh bed tests. The lowest rate, however was
measured with aged,sulphided stone under kerosene combustion.
In light of our experience with the two previous stones, only
kerosene combustion, fuel oil gasification and cycle tests were
studied with BCR 1359. The gasification fresh bed test gave a
higher loss rate and cycle tests a lower rate than the fresh bed
test with sulphur free kerosene combustion.
36.
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Table 10
Summa LBatch_Unit Loss Rates
Loss Rate (g/inin) *
Conditions Kerosene** KerOsene** Fuel 0i1 Fuel Oil** Kerosene Combustion Gasification
Combustion Combustion Combustion Gasification on Sülphided Aged Bed Regeneration
870 deg.C 1070 deg.C 870 deg.C 870 deg.C 870 deg. C Cycles
(. )
Limestone
BCR 1691 22.6 5.7 13.5 7.8 8.5 4.5
8CR 1359 2.1 3.6 1.9
Denbighshire 4.3 3.3 9.0 9.7 0.8 1.4
* At all conditions except gasification/regeneration cycles, loss rate has been calculated over the
first four hours of the test. The cycle loss rate is the stable loss rate.
** Fresh Bed Tests
-------
The bed ageing effect which decreased losses with BCR 1691 was
also evident with Denbighshire and BCR 1359. For example, within
5 cycles, loss rate dropped from 6.6g/mirt to l.4g/rnin with Denbigh-
shire and from 6.4 g/min to 1.9 g/xnin with BCR 1359.
These results reveal several interesting comparisons between the
stones. These are summarised in Table 10.
Neither of the higher purity stones gave the very high loss rate
found with BCR 1691 during fluid bed combustion with kerosene at
870 deg. C. Both Denbighshire stone and BCR 1691 gave lower loss
rates during kerosene combustion when temperature was increased
from 870 deg. C to 1050 deg. C.
Although the temperature effect was less dramatic for Denbighshire,
its loss rate remained below that of BCR 1691. No high temperature
test was made with BCR 1359.
All three stones exhibited a decreasing loss rate with age during
the initial test periods. The change became less significant at
long exposure times. This age effect may be due to a strengthening
of particles by a sintering process, to elimination of particles of
lower initial strength, or to a combination of these factors. The
significant decrease in loss rate at the higher temperature appears
to be a consequence of the more severe sintering which would be
expected under those conditions.
The effect of changing from kerosene to fuel oil differed between
the low purity and high purity stones. Whereas with BCR 1691,
fresh bed loss rate decreased when fuel oil replaced kerosene
combustion and decreased again on going to fuel oil gasification,
opposite directional results were found with Denbighshire stone.
Results from the shorter test programme on BCR 1359 indicate that
its behaviour is similar to that of Denbighshire. In spite of its
high loss rate during fresh bed gasification, the Denbighshire stone
gave the lowest rate of the three during gasification - regeneration
cycles.
We believe that loss rate differences between kerosene and fuel
oil operation are due to sulphur in the oil. However, the mechanism
of the sulphur effect must be complex to increase losses with the
pure stones whilst decreasing losses with the lower purity BCR 1691.
In addition to the results already discussed, bed losses were also
measured during calcination and are summarised in Table 11
38.
-------
Table 11
Summary of Batch Unit Calcination Losses
Stone Low Sulphur Fuel
Kerosene and High Sulphur Fuel
propane * 2.3% S Oil *
BCR 1691 18 15
Denbighshire 16 24
3CR 1359 6
* Losses as % of calcined stone
The low calcination loss rate from stone 3CR 1359 makes it
particularly attractive. The effects of sulphur on loss rate in
calcination of Denbiqhshire and 3CR 1691 stones are in the same
direction as observed in the fresh bed tests with these stones.
Nature of Fines
Information on fines properties was gathered from material
collected in the cyclone and deposited in pipes downstream of the
cyclone.
Throughout the tests the cyclone operated satisfactorily with the
aid of a mechanical rapper, and regular samples of fines were
obtained. Table 12 contains the results of a microscopic examination
of these fines together with cyclone efficiencies. With respect
to physical appearance, the fines have been separated into five
groups. Examples from each group are shown in Figures 6 to 10.
The only fines represented by Figure 6 are those collected during
kerosene combustion at 8.0 deg. C in 3CR 1691. Most of the
particles appear to be less than 5O)i. The fine particles seem to
be adhering to each other to form loose agglomerates and to the
surface of the few larger particles that are around. This stickiness
was also observed during kerosene combustion in Run 4. Figure 7
shows the type of particles obtained when BCR 1691 was subjected
to kerosene combustion at 1050 deg. C, to fuel oil combustion at
870 deg. C, and to gasification-regeneration cycles followed by
kerosene combustion at 870 deg. C. There appears to be some larger
particles in this group and there is less evidence of the stickiness.
In Figure 8, the type of fines collected in all instances where
gasification or gasification/regeneration cycles were carried out
are illustrated. Due to the presence of carbon, it is difficult
to estimate the particle size range present. However, it would
appear that the great majority of particles are less than 50,u.
The fines collected under combustion conditions with Denbighshire
and 3CR 1359 stones are represented in Figure 9.
39
-------
.
, -;
Figure 6 Cyclone Fines - BCR 1691 (Kerosene Combustion
870 deg. C.)
Figure 7 Cyclone Fines - 3CR 1691 (Kerosene Combustion
1050 deg. C. Fuel Oil Combustion 870 deg. C.,
Kerosene Combustion 870 deg. C on aged bed).
40.
-------
Figure 9 Cyclone Fines - BCR 1359 and DenbighShire
(All combUstiOn conditions).
Figure 8 Cyclone Fines - BCR 1691, BCR 1359 and
DertbighShire (Gasification - Regeneration CycL
41.
-------
Figure 10 Cyclone Pines - Denbighshire
(Kerosene Combustion 870 deg. C. on conditioned bed)
42.
-------
Table 12
Nature
Test Condition ______
Kero combustion 870 deg.C 8CR 1691
Kero combustion 1050 deg.C
Fuel oil combustion 870 deg.C
Fuel oil gasification
870 dog. C
Gasification/Regeneration
Cycles
Kero combos tion-suiphided
aged bed 870 deg. C
Kero combustion 870 deg.C
Kero combustion 1050 deg,C
Fuel oil combustion
870 deg. C
Fuel oil gasification
870 deg. C
Gasification/Regeneration
cy ci es
Kero combustion - suiphided
aged bed 810 deg. C
Kero combustion 870 deg. C
Fuel oil gasification
870 dog. C
Gasi fication/Regeneration
Cycles
Limes tone nce
Loose agglomeration of
particles less than 50AJ
Mainly discrete particles
less than 50,u
Mainly discrete particles
less than 50 ,u
Mixture of carbon & lime
Size
99% 50/u
90% 50 Al
90% <50A1
Denbighshire
Mixture of carbon & lime -
Mainly discrete particles 90% .c 5 O .iu
less than 50 u
Discrete mixture of particles —l000,&i
up to 1000 l
-1 000AI
-1000 Al
Mixture of carbon & lime
Mixture of carbon & lime
one_Efficiency
80
100
32.6
49.1
83.9
59.5
79.4
62.0
49.0
56.6
100
44.4
86.3
83.0
100
Mainly particles greater 90% ‘5041
than 50 si
8CR 1359 Discrete mixture of particles —1000 .i
up to 1000 o
Mixture of carbon & lime —
Mixture of carbon & lime -
-------
In this case there is a discrete mixture of particles with a
maximum size around l000. u. Finally, Figure 10 illustrates the
type of particles collected during kerosene combustion at 870 deg. C
of the suiphided, aged, Denbighshire bed. Here, there is a discrete
mixture of particles which are mainly in the size range 50-1000 p.
These results indicate an approximate correlation between fines
product rate and particle size of the fines collected by the
cyclone. The large cyclone particles were collected at the
lowest production rate i.e. during kerosene combustion of suiphided,
aged, Denbighshire stone. Also it would appear that kerosene
combustion at 870 deg. C in BCR 1691 which gave the highest fines
production rate resulted in the smallest particles being collected
in the cyclone.
It is not possible, however, to extend this argument and correlate
fines production rate with particle size of total fines produced
because of differences in cyclone efficiency. One of the lowest
cyclone efficiencies was recorded for the large cyclone particle
sizes and one of the highest for the smallest cyclone particle size.
If particle size of cyclone material was a true reflection of
particle size of fines produced by the bed, then one would have
expcted cyclone efficiencies to have been highest with the largest
particle size. Why cyclone efficiency should act in this way is
indeed puzzling. However, it does mean that this information of
fines size and appearance cannot be used with respect to fines
production mechanisms.
Further tests on the fines collected showed that those collected
during gasification flowed best and equally well for all three
stones. Under all other conditions those from BCR 1359 and
Denbighshire flowed more freely than those from 8CR 1691. The
superior flow characteristics of fines from gasification are
attributed to their high carbon content which could be as much
as 45% by weight. The superior flow characteristics of fines
from Denbighshire and 8CR 1359 stones in relation to BCR 1691
under all the other CAFB conditions are attributed to the presence
of fewer of the very fine particles. Although these differences
were encountered in the flow properties of the fines, we never
encountered batch unit fines with the severe stickiness of pilot
plant fines made during Run 4 startup with BCR 1691.
Batch unit fines and bed samples from 3CR 1691 were analysed for
calcium and silicon. The results showed that the batch unit
conditions also caused the preferential loss of calcium from the
bed observed in pilot plant Run 4. It appeared to begin during
calcination and continue through the tests. As in the pilot plant,
calcium lost from the bed did not appear in recovered fines but
was lost from the system. The extent of calcium loss in the batch
unit tests was not as great as found in the pilot plant where
the SiO /Cao ratio increased to 0.41 compared with an initial
value o 0.27. In batch unit cycle tests, the Si0 2 /CaO ratio
lined out at approximately 0.33.
44.
-------
During each test unit pressures were monitored to determine if
any blockages were occurring downstream of the bed and after each
test exit gas lines were dismantled and examined. Only one
blockage of any significance was ever encountered. This occurred
with BCR 1691 under kerosene combustion conditions at 870 deg. C
and was traced to a period of cyclone rapper malfunction. This
emphasises the importance of fines concentration in the gas
stream in relation to blockages and shows that even with BCR 1691 ’s
high rate of sticky fines production, proper draining of cyclones
with the aid of rappers where necessary prevents blockages.
45.
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SECTION VIII
RUN 5 PILOT PLANT MODIFICATIONS
MODIFICATIONS AFTER RUN 4 AND BATCH STUDIES
Between Runs 4 and 5 a number of modifications were made to the
pilot plant to permit improved operations under conditions like
those encountered in Run 4. The experience of Run 4 and data
gained in the batch unit programme revealed that operation is
more difficult with limestone BCR 1691 than with Denbighshire
stone. The batch work indicated that BCR 1359, another high
purity stone, should behave more like Denbighshire Stone.
However there are many locations where high purity limestone will
be considerably more expensive than lower purity stones available
locally. It therefore was desirable to assess more fully the
consequences of operating with a lower purity stone. Although
8CR 1691 is not necessarily typical of all low purity stone,
it has deficiencies which, if overcome, would assure that the
CAFB gasifier could operate with stones of a wide quality range.
Five major changes were made to assist operations with the dusty
and more easily agglomerated stone.
• External cyclone drainage
• Non obstructing regenerator pressure control system
• Flue gas recycle scrubber
• Regenerator overtemperature quench
• Flue gas stack scrubber
The revised pilot plant flow plan in Figure 11 shows these
modifications. Other minor changes were made to the unit to
improve pilot flame stability and to assist in diagnosing boiler
performance.
Cyclone External Drain System
Pressure balance calculations on the gasifier cyclone return
system indicated that there would be insufficient height of leg
available to return fines to the gasifier through the internal
passages if bed depth were increased to the levels desired for
high sulphur recovery with low replacement rates of BCR 1691 stone.
This problem increases in severity when fouling increases the
pressure drop across the cyclone inlet. By using external cyclone
drains, the pressure at the cyclone drains could be made independent
of the gasifier bed pressure. It was not sufficient however just
to drain the cyclones externally. With deep beds the rate of
entrainment into the cylones could be high and stone losses severe
unless the coarse fraction were returned to the gasifier. Therefore
an external system to both drain the cyclones and return the fines
was needed.
46.
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Cooling Tower
v t f
IPJ
Fuel Oil H
Kerosene
Elutriator
Stack
Air
Air
Flue Gas
Recycle
- —--
Figure 11
CAFB Revised Pilot Plant Flow Plan
-------
After consideration of several designs, a system was selected
which met the constraints of available space, pressure, and gas
consumption. Details of the system appear in Appendix C. In
summary the system for each cyclone consists of a conical
bottom pot receiver mounted beneath each cyclone drain, a
butterfly valve to isolate the pot from the cyclone when the
pot is being emptied and a Warren Springs Laboratory pulsed
flow powder pump to transfer solids from the conical pot to an
overhead receiver, common to both cyclones. An elutriator
removes the very fine fraction from the cyclone solids, and a
pneumatic injector returns the larger size fraction to the
gasifier bed. Nitrogen is the operating gas for the transfer
system. Most of the time the butterfly valve beneath the
cyclone remains open draining solids to the conical pot. At
timed intervals the butterfly valve closes and the pot is pumped
out to the overhead receiver.
Regenerator Pressure Control
To avoid a repetition of the regenerator off gas line blockage
which terminated Run 4, the pressure control valve was removed
from the outlet line. This valve had become blocked by fine
solids in the gas stream during that run. In order to vary
regenerator back pressure, a new blower was installed to inject
air into the outlet line just downstream of the cylone outlet.
Air flow from this blower is regulated by a control loop which
senses the difference between gasifier and regenerator pressure
and adjusts a valve in the air line to achieve the desired
pressure difference.
Flue Gas Recycle Scrubber
Run 4 demonstrated that simple cyclones were unable to provide
sufficient cleaning of the recycle flue gas stream to prevent
gradual blockage of the gasifier air distributor nozzles. A
venturi scrubber system was designed to provide greater cleanup.
The scrubber was designed to handle 200 CFM of gas at a pressure
drop of 14w.g. Water is sprayed into the gas at the throat
of a venturi. A knockout vessel at the venturi outlet removes
the water and entrained dust. The venturi was placed on the
suction side of the recycle blower to protect the blower from
dust, and a recycle line was provided to permit a high gas
circulation rate through the venturi even at low rates of flue gas
flow to the gasifier.
Regenerator Quench
The circulation of fresh solids from the gasifier to the regenerator
controls regenerator temperature. Upsets in the pressure balance
between gasifier and regenerator or temporary obstructions in
one of the solids transfer lines can sometimes interrupt this
solids circulation and allow regenerator temperature to increase.
48.
-------
If regenerator temperature gets too high there is danger of
sintering the lime particles and forming agglomerates. An
emergency quench system was installed to prevent this occurrence.
The lower regenerator bed thermocouple was connected to a
controller which admits a flow of nitrogen to the intake of the
regenerator air blower when bed temperature reaches the alarm
point. The alarm was set to operate at 1100 deg. C. Nitrogen
fed to the blower dilutes the regenerator air supply and reduces
the rate of oxidation to prevent over temperature. The circuit
is fitted with a manual switch so that the process operator can
inject nitrogen at will in the event of other forms of upset.
Stack Top Gas Scrubber
To avoid particulate emissions to the atmosphere during periods of
high lime losses from the gasifier cyclones, a final stage of water
scrubbing was added to the pilot plant flue gas stack. Experience
in Run 4 had indicated that the flue gas cyclone was not completely
effective In recovering lime fines produced from BCR 1691 under
combustion conditions.
The new scrubber consists of a section of ductwork shaped like
an inverted “U” mounted on top of the stack. The down leg
directs the gases Into the top of a funnel shaped receiver which
causes another reversal of gas direction upward to the atmosphere.
Water is sprayed into the down leg and collected by the funnel.
This water which picks up limedust by passage through the flue gas
is conducted to a ground level settling vessel. Overflow from this
vessel is circulated back to the scrubber nozzle by a centrifugal
pump. The system is designed to circulate water to the scrubber
at a rate of approximately 100 gallons/mm.
49.
-------
SECTION I)
PILOT PLANT OPERATION - RUN 5
GENERAL
Objectives of Run 5 were to measure sulphur removal efficiency
at different bed levels and lime replacement rates, to test
feasibility of desulphurising gasification at temperatures above
900 deg. C approaching adiabatic conditions, and to determine
effectiveness of pilot plant modifications in solving problems
met in Run 4.
The test programme was to start up with precalcined lime from
Denbighshire stone, operate for a test period with Denbighshire
stone, and then switch to BCR 1691 stone. Provision was allowed
for returning to Denbighshire stone if the BCR 1691 proved
inoperable.
In actual fact, Denbighshire stone was used during the first two
days and the final week of operation. CR 1691 was used for the
rest of the run.
Table 13 sununarises the various periods of operation during the
run. For purposes of computer identification of data, each run
hour is designated by a decimal number with its whole part
signifying run day and its fractional part the hour. February 6
was run day 1.
For example, 4.1630, represents 4.30 p.m. on February 9. The same
system is employed on the abscissa representing time in the data
graphs of Appendix D. These graphs show the variation with time
of major operating variables during the run. Tables of detailed
operating data also appear in Appendix D.
OPERATIONS SUMMARY
Operability of the pilot plant was much better in Run 5 than in
any previous test. Prolonged test periods were achieved with
lined out conditions. Some initial difficulties were experienced
with some of the new equipment, and some of the problems encountered
in earlier runs reappeared but in less severe form.
Heat up began on 2 February 1973 and proceeded smoothly without
interruption. Precalcined lime from Denbighshire stone was used
for the initial bed, and no difficulty was experienced in
achieving fluidisation. During the first 109 hours of continuous
gasification, four data periods were obtained, two with Denbighshire
stone and two with 8CR 1691. This first gasification period was
terminated when pressure drop through the boiler became excessive.
The increase in pressure drop was caused by a build—up of deposits
at the entrance to the first pass of small boiler tubes. The
gasifier was suiphated and decoked by controlled oxidation, and
the boiler tubes were cleaned.
50.
-------
Table 13
Summary of aun S OperafingA Periods
Number of Test Gasification
Periods Hours
Feb
Feb
Dates
6 — Feb
9 - Feb
9
11
Run D y
1. 2000—4.0500
4.0500—6.1000
Limestone Feed
2
2
)
)
)
)
109
Cause of Termif ation
Denbighshire
8CR 1691
pressure drop through
boiler.
Change to BCR 1691
caused no interruption
of gasification.
Feb
13 - Feb
14
8.1920—9.0630
8CR 1691
0
11
Regenerator DefluidisatiOfl
Feb
16 — Feb
18
11.1215—13.1000
BCR 1691
4
45
Fuel Injector Failure
Feb
21
- Feb
22
16.0600—17.2200
8CR 1691
2
40
Regenerator DefluidisatiOn
Feb
25
- Feb
27
20.1800—22.1900
Denbighshire
4
49
Cyclone Inlet Pressure Drop
Feb
28 — March 3
23.1300—26.1900
Denblghshire
5
•
78
Voluntary - End of Run
-------
Resumption of gasification was delayed by malfunction of the fines
return system and by behaviour which suggested an obstruction in
the regenerator. However removal of the regenerator distributor
revealed no blockage. The trouble with the fines system was
caused by chunks of coke dislodged during decoking.
Gasification resumed on 13 February but was interrupted 11 hours
later by regenerator defluidisation and a regenerator hot spot of
1150 deg. C. Removal of the regenerator distributor revealed some
agglomerated solids in the hot spot region. Cause of the
defluidisation is suspected to be partial gas by—passing and lack
of fines in the regenerator.
The next gasification period began on February 16 and lasted 45 hours.
Five data periods were obtained with BCR 1691 stone. This period
was interrupted by failure of one of the fuel injectors. The tip
of the injector was destroyed causing poor local fuel distribution
and local defluidisation. After suiphation and decoking, the
other injectors were inspected and found to be intact. Since the
damaged injector was nearest the start—up burner, it is possible
that it had been weakened by over—temperature during the heat up
period. The fuel injector was replaced, the boiler tubes were
cleaned, and gasification resumed on 21 February. This gasification
period lasted 40 hours with 8CR 1691 stone and provided two
data sets.
This period also was terminated because of partial defluidisation in
the regenerator.
The regenerator superficial gas velocity, based on air and nitrogen
inputs, was over 5 ft/sec during this period, more than enough
for good fluidisation. The cause of the loss of fluidisation was
again suspected to be gas by—passing, but this hypothesis was not
confirmed. The difficulty was aggravated by malfunction of one of
the regenerator blower-stages which limited the available air
capacity. Following this interruption, the regenerator distributor
and interior were again inspected.
The regenerator interior contained only a thin collar of fused
material just above the distributor, not enough to explain the
regenerator difficulties. The refractory lining of the distributor
was rebuilt, and cracks in the lower portion of the regenerator
were patched with fire clay before reassembly.
The next gasification period consisted of 49 hours with DenbighShire
stone. Fuel rate was increased to test gasifier operation at
900 deg. C. with lower air/fuel ratio and reduced flue gas recycle.
Pressure drop through the boiler prevented increasing fuel rate
enough for completely adiabatic conditions but smooth operation
was maintained at over 460 lb/hr oil rate and 19 to 20% of
stoichiometric air—fuel ratio. Two test periods were obtained at
the high fuel rate and two more at an oil rate of about 410 lb/hr.
52.
-------
No difficulties were experienced with regenerator defluidisation
during this period. Gasification was interrupted for decoking
due to rising pressure drop through the gasifier cyclones. The
boiler was not cleaned during this down period. Gasification
was resumed for a final 78 hours of gasification with Denbighshire
stone. Six more data sets were obtained in this final period
after which the unit was shutdown. Mechanical operation was smooth
in the final period. The external gasifier fines return system
functioned well, except that cyclone efficiency declined, and a
reducing amount of fines was caught for return.
At the end of gasification, the unit was shut down without
suiphation or decoking in order to permit examination of the
interior with deposits in place.
EQUIPMENT PERFORMANCE
Equipment performance was greatly improved over that experienced
in previous runs, particularly that of Run 4. A number of the
new features operated well, but some continued to be troublesome
throughout the run.
Stone Feeder
The stone feed system which used a vibrator in a pressurised shell
to feed from a weighed hopper proved reliable throughout the run.
Stone feed rates were usually quite steady and easily measured.
Regenerator Drain Valve
The gasifier bed level control system which used a pressure switch
in the gasifier to activate a drain valve in the regenerator proved
to be reliable and to give good control of gasifier bed depth.
Regenerator Pressure Control
No blockages were encountered in the regenerator off gas line
during Run 5. This line and its control system had plugged
continually during Run 4 start up. In Run 5 the line remained
clear and showed no sign of pressure build up. The system used
controlled introduction of excess air into the outlet line down-
stream of the cyclone and avoided restrictions in this line. It
was not possible to operate the system in automatic mode due to
the large pressure pulses introduced by the solids circulating
system, but manual control of the pneumatic valve position proved
satisfactory for control of the pressure difference between
regenerator and gasifier. Pressure difference was regulated to
within one or two inches water gauge. Normally the regenerator
pressure was adjusted to be 3 to 5 in. w.g. below gasifier gas
space pressure although higher differences were sometimes used.
Cleanliness of the regenerator gas line was also aided by
continuous use of a pneumatic rapper on the regenerator gas cyclone.
This rapper ensured drainage of solids from the cyclone walls. Also,
the conditions that produced the very sticky fines, kerosene
combustion in a bed of 8CR 1691 stone, were avoided as much as
possible.
53.
-------
Flue Gas Recycle Scrubber
The venturi scrubber on the flue gas recycle stream removed a great
deal of lime fines from the gas, but was not completely effective.
Some particles passed the scrubber, and some fouling of the
recycle gas line, control valve, and gasifier distributor was
encountered. The rate of gasifier distributor pressure rise in
Run 5 was much less than in prior runs. In the initial stages
of Run 5 there was frequent plugging of the inlet of the venturi
throat itself with lime deposits. Increasing the gas flow
through the venturi to the maximum rate available (estimated at
200 CFM) by using maximum recirculation eliminated plugging at
this point.
The drain line from the water separator occasionally blocked and
required cleaning.
On at least three occasions blockage of this discharge caused
water carryover to the gasifier plenum itself with a consequent
sharp decrease in gasifier temperature.
Regenerator Over Temperature Protection
The new regenerator emergency quench system dilutes the inlet air
with nitrogen when regenerator temperature reaches the set point
of 1100 deg. C. This system proved to be quite valuable and
avoided excess temperature several times when malfunction of the
solids circulation system reduced lime flow rate through the
regenerator. In only one case did regenerator temperature
seriously exceed 1100 deg. C, and that was due to emptying of the
quench N 2 supply bottle before normal conditions were restored.
This quench system is believed to be responsible for avoiding
regenerator blockages by agglomerated solids which occurred in
previous runs.
Boiler Pilot Flame
The new boiler pilot burner and its gas and air system were quite
effective in providing a stable and reliable pilot. No difficulty
was met in lighting the pilot over a wide range of conditions nor
in keeping it lit.
Stack Top Washer
The water spray scrubber installed to prevent discharge of dust
to the atmosphere was operated during part of the run when it
appeared that some dust was passing the external flue cyclone.
When operated the scrubber appeared to be effective in avoiding
dust emissions. Because of the diffuse upward discharge of gas
from the system, it was not possible to obtain a quantitative
measure of the actual dust content.
Corrosion of the water recycle piping in this scrubber system was
severe because of SO from the regenerator which was remixed with
the boiler gas in th stack.
54.
-------
Cyclone Fines Return System
The gasifier cyclone fines return system operated well during much
of the run in spite of several deficiencies. Operation was
trouble-free for most of the first 109 hours of gasification and
for the final 127 hours. During other periods there were a number
of upsets. Two problems were encountered in the first period:—
(1) Dust worked its way back to the pneumatic control system
through a pressure measurement line and caused stoppage.
This problem did not recur after installation of a fine
filter and additional N 2 bleed in the pressure line.
(2) Residual pressure remained in the conical dust receiver
after a transfer of solids when the butterfly valve to the
cyclone drain opened to begin refilling, this pressure
caused a surge of gas back up the cyclone leg and upset the
cyclone operation. The result was a burst of fines into the
boiler after each transfer operation. These puffs were
evident in the peaks observed in boiler SO emission. This
difficulty was removed by installation of delay device
which allowed pressure to discharge down stream from the
conical receiver before the valve to the cyclone could open.
Most of the problems met during the mid run period were caused by
chips, flakes, and chunks which found their way into the cyclones
and transfer system after each temporary shutdown and decoking
operation. It was necessary to disconnect vessels and lines on
several occasions to remove these flakes and chunks. In other
cases these solids prevented good operation of the butterfly
valves. When the butterfly valves failed to seat properly
before a transfer, N 2 gas again blew back through the cyclone
and sent dust to the boiler.
Installation of a chunk trap in the elutriator drain during the
run improved operation a great deal. Installation of additional
chunk traps in the conical vessels will be a further aid.
Efficiency of the cyclones themselves deteriorated during the run.
In the initial period efficiency was fairly high, and only a
small amount of very fine solids entered the boiler. In later
stages of the run the efficiency deteriorated, and a considerable
quantity of quite coarse solids entered the boiler.
Inspection of the cyclones at the end of the run revealed them to
be nearly completely choked with a mixed deposit of lime and
carbon in the annular space betceen walls and gas outlet tube.
The steel liners which had been installed before Run 5 were
severely burned and distorted. The silicon carbide gas outlet
tubes were strong, smooth, and intact. It is evident that
demonstration of an effective way to maintain cyclone efficiency
must remain an important problem area of this work.
55.
-------
In view of performance of the process, it is clear now that recycle
of cyclones fines to the gasifier without a means to achieve
their regeneration is not desirable. The fines, with their large
surface area, pick up a considerable load of sulphur and make a
number of cycles through the gasifier and cyclones without
entering the regenerator. If any of these fines escape the
cyclone, they enter the boiler and cause loss of sulphur removal
efficiency. A means of providing preferential regeneration of
the fines is a desirable process feature.
Regenerator Operation
DefluidisatiOn of the regenerator bed occurred twice during
gasification in Run 5 and once under combustion conditions. The
cause of this behaviour has not been established, but gas by—passing
in some manner is suspected. The effects of such by—passing would
be aggravated by lack of fines in the regenerator solids which
would increase minimum fluidisation velocity. Two of the ways
in which by passing could occur are leakage of gas through the
solids circulation passages and leakage through cracks. It is
possible for air to enter a crack in the refractory near the
bottom of the bed, travel upward through the crack, and return
to the vessel higher up. A small crack was observed in the
regenerator wall and patched during the run, but it did not
appear large enough to account for the troubles observed.
The fact that the tendency to defluidise became more severe with
time during the gasification periods involved suggests that it
may have been related to another time dependent factor such as
the increase in gas space pressure which took place as gasifier
outlet passages gradually fouled. Leakage of air back into the
gasifier—to—regenerator solids transfer line and up the unused
portion of the cyclone fines return leg could follow such a
course. Such leakage would have to pass into the cyclone past
the steel sleeve insert which had been dropped into the
cyclone leg as a seal. However, distortion of the steel sleeve
by heat following gasifier decoking is a distinct possibility.
On the other hand, such a loss of air near the regenerator bottom
does not accord with the apparent low SO 2 concentration measured
in the regenerator gas. Indeed, sulphur material balance
considerations imply that the gas flow through the regenerator
was higher than that supplied by the regenerator air blower. It
is probable that further tests will be needed to establish the
cause of this unusual regenerator behaviour.
Boiler De sits
Boiler deposits found within the boiler were of two types:
(1) Loose accumulations of dust or coarser particles in the
soot trap areas at the boiler ends.
(2) Agglomerated deposits formed from very fine particles
which build up at the inlets to the first pass of
small fire tubes.
The loose accumulations of particles do not appear to present a
long term problem. They would be subject to easy removal by
normal soot blowing techniques.
56.
-------
The agglomerated fines represent a potential problem area which
requires additional study to define its severity in large scale
equipment. Certainly it is an inconvenience in our pilot plant
equipment. However it must be stressed that the fire tube
boiler used for our pilot plant tests is in no way typical of a
water tube power generation boiler and the problems we have
experienced may be typical only of the particular boiler we are
using. From the point of view of deposit buildup the pilot
plant boiler is far from ideal. Deposits are very local in nature;
being found only at the inlets to the first set of water cooled
fire tubes, where the gases change direction by 180 deg. in a
downward direction.
They did not form after the first few inches of tube length nor
were they found to any significant degree on a test probe inserted
radially into the gas stream at the end of the main fire tube as
a simulation of a superheater tube in a water tube boiler. These
deposits evidently form from fine particles which are in a sticky
state following their passage through the flame. The growth of
the deposits was faster during the first 109 hours of Run 5
gasification than during the final 127 hours. Whether this
difference was due to the presence of BCR 1691 stone in the
first period or to the eroding effect of a higher concentration
of coarse particles during the final period remains to be
established. It is possible that deliberate injection of a small
amount of coarse stone into the boiler could prevent deposit
formation. No deposits were found at the tube inlets during
the 111 hours of gasification in Run 2 during which a high
concentration of solids passed through the boiler.
PROCESS PERFORMANCE
Table 14 lists values of operating conditions and results for
the various test periods of Run 5. Each value is the average
for four hours operation. In most test periods a set of solids
samples was collected for analysis.
Sulphur Removal
The degree of sulphur removal in the pilot plant is calculated
from the measured SO and CO 2 contents of the boiler flue gas
compared with sulphu and carbon contents of the boiler fuel.
The carbon content of the pilot burner propane is considered
in this calculation, as is the CO 2 released by calcination of
limestone makeuc.
During Run 5, sulphur removal efficiency (SRE) varied from
60 to 99%. Appendix figure D-l shows the hourly levels of SHE
along with other important variables. Figure 12 shows the
effect of lime replacemeflt.rate on SHE for individual test
periods. It is apparent that lime replacement has a major
effect on SHE.
57.
-------
Table 34
rat tog ! !!! A2 i
Day Time
2.2130
3.0530
3.1530
3. 2 130
5.1030
6.0730
Ui 12.0830
12.1530
12, 1930
13.0330
13.0630
17. 1230
17. 1730
21.0930
21.1830
22.0730
22.1830
24.0730
25.0530
25. 1530
26.0530
26.1130
26.1830
Temperature dog. C
Gasifier Regenerator
883 1047
886 1063
884 1061
894 1066
8S2 1053
856 1055
877 1050
870 1036
871 1055
868 1031
876 1024
863 1052
861 1047
903 1062
889 1065
878 1070
873 1069
862 1057
878 1060
875 1060
880 1060
866 1060
872 1060
Lime Replacement Gasifier Depth
14o1 Ca/Mci S in w.g.
.62 19.2
.64 20.6
.48 21.5
.62 18.0
1.76 18.0
1.19 17.6
.81 21.1
1.14 22.3
1.07 21.8
1.01 18.4
.87 18.2
.90 25
1.55 25.3
1.16 22.0
1.40 22.7
.96 23.0
1.03 24.0
2.01 24.6
1.48 25.3
1.36 25.6
1.04 25.0
1.32 24.9
1.41 25.4
Fuel Rate
lb/hr
398
396
395
404
400
383
398
396
395
392
390
399
401
469
469
43.3
409
411
405
404
405
410
409
Sulphur Removal
77.5
72.8
74.8
60.2
95.0
94.5
79.1
84.6
85.5
83.4
87.2
84.0
76.9
93.3
89. 1
91.2
85.0
99.3
93.7
90.3
90.8
92.0
86.1
Regenerator Stone
Conc. Vol% % of S Led
4.1 32.7 Denbighehire
4.4 43.3
3.0 28.0
2.6 23.1
5.1 44.7 bCR 1691
4.6 36.9
4.1 46.3
3.7 43.9
4.1 4a.5
4.0 47.0
3.9 48.5
4.7 46.1
3.8 38.4
3.8 32.8 Denbighahire
4•4 35.8
e.o 33.4
3.7 34.6
3.7 34.4
3.6 3S.4
3.6 34.9
3.2 32.3
3.4 34.5
3.4 33.3
-------
100
90 —
>-
U
z
80 -
U
U-
La
L i i
-J
4 70 —
>
0
U i
60 - 0 SYMBOL RUN LIM STONE
• 3 DENBIGHSHIRE
-J
4 BCRI69I
50— V 5 BCRI69I
0 5 DEN8IGHSHIRE
40 I 1
0 2 3 4
LIME RELACEMENT RATE, Mol Co/Mol S
• I
.• • .. •
.
V.
0
V
.
V
Figure 12 CAFB Pilot Plant Sulphur Removal Efficiency
-------
The effects of other variables are less clearly defined. In
particular, increasing gasifier bed depth did not produce the
expected improvement in SRE over results obtained in earlier
runs with shallower beds.
It appeared in this run that the beneficial eftect of increased
bed depth was offset by increasing fines loss rate. The loss
of fines hurts sulphur removal efficiency in two ways. Firstly
it removes the high surface area fraction of the lime which has
greatest potential for sulphur pickup, and secondly, when sulphur
laden fines enter the boiler they partly regenerate to contribute
SO 2 to the flue gas.
The effect of bed depth was also obscured by the fact that
deterioration in cyclone performance made it difficult to
operate with deep beds at very low lime replacement rates.
The test at 900 deg. C gasifier temperature on day 21 produced
sulphur removal efficiency over 90% and demonstrated the
feasibility of operating at this temperature with a low air/fuel
ratio. The fuel rate in this test was 469 pounds per hour, the
highest yet used in the pilot plant. Pressure drop in the boiler
prevented increasing fuel rate still further to test completely
adiabatic gasifier operation without flue gas recycle.
Metals Retention
Comparing the metals content of spent lime with the metals content
of fresh limestone and fuel oil, it is possible to estimate the
degree of metal retention by the solids. Figures 13, 14 and 15
compare the retained weights of vanadium, sodium, and nickel with
quantities of these metals fed during various operating periods
of the pilot plant.
Vanadium retention is essentially complete, in agreement with
predictions based on batch unit tests. Sodium retention was
36%, somewhat higher than the 20% level obtained in batch tests.
Nickel retention, which was not studied before, averaged 75%.
There appeared to be no significant difference between metals
pick up efficiencies with Denbighshire and BCR 1691 stones.
However because of differences in stone loss rates, there is a
difference in absolute metals retention levels in the unit. With
Denbighshire stone there was practically no lime loss from the
system. That lime which escaped the gasifier cyclones was caught
either in the boiler or by the external flue gas cyclones. However
with BCR 1691, losses amounted to as high as 40% of the lime
replacement rate and of course any associated metals were lost as
well. While not affecting the ability of the CAFB gasifier to
remove vanadium as a source of high temperature corrosion of
boiler superheater tubes, the loss of metals on lime particles
would be a pollution factor which could be reduced even further
by increasing the efficiency of the particulate removal portion
of the system.
60
-------
VANADIUM RETENTION (RUN 5 )
SYMBOL
4
V
x
S
0
+
A
GAS1FICATIO
PERIOD
3
4
5
6
TIME
PERIOD
1.1730 — 40000
40000— 6.0830
11.1230— 13.0630
16.0630—17.2130
20.1830 — 22.1930
23.0730— 26.1830
2
LIMESTONE
FEED
DENBIGHSHIRE
8CR 1691
8CR 1691
8CR 1691
DE NB UGH $ H IRE
DENBIGHSH IRE
FUEL
VANADIUM
0
CONTENT 300 ppm
H
0
1 4 )
z
I—
14)
a:
>
8
*
6
4
+
V
V
2
0
10,000 20,000 30,000 40,000
FUEL CONSUMED (Ib)
50,000
Figure 13 Vanadium Retention
(Run 5)
-------
SODIUM RETENTION (RUN 5 )
SYM9QL GASIFICATION
PERIOD
3
4
5
V
0
+
I5
1 0
TIME
PERIOD
11730 — 40000
40000- 6.0830
111230 — 13.0830
6.0630—17.2130
20.1830 — 22.1930
LIMESTONE
FEED
DEN8IG HSHIRE
9CR 1691
8CR 1691
8CR 1691
DENSIGH SHIRE
D ENBIGHSH iRE
6
rUEL
23.0130— 261830
SODIUM
CONTENT 37ppm
a’
I’ -)
.0
a
w
z
I—
w
0
z
05
x
0
10,000 20,000 30,000 4OL O0
FUEL CONSUMED (Ib)
50,000
Figure 14 Sodium Retention (Run 5)
-------
NICKEL RETENTION (RUN p
10,000 20,000 30,000 40,000
FUEL CONSUMED (Ib )
L )
w
z
4
I-
w
z
50,000
Figure 15 Nickel Retention (Run 5J
-------
Solids Losses
To obtain a more comprehensive picture of solids losses during
Run 5, both the amount of material emitted by the gasifier into
boiler and the amount escaping the external cyclone into the
stack have been computed.
Results are suxrsnarised in Table 15. The water scrubber on the
stack top was able to catch the worst of the material which
escaped the cyclone, but since the quantity recovered by the
scrubber was not measured, it is included with the stack losses.
Table 15
Summary of Solids Loss - RunS
Time Limestone Make—up Gasifier Loss Rate Bed
Feed Rate Fluid Bed Depth (lb/hr) Si0 2 /
(lb/hr) (inches) CaO
Day.Hour Gasifier Stack Ratio
3.1530 Denbighshire 7.3 27.5 0 0 0.006
3.2030 7.3 22.5 5.3 2.7 0.006
5.1030 BCR 1691 60.0 30.0 29.9 23.3 0.203
6.0730 27.0 26.7 18.3 15.3 0.236
12.1730 26.0 31.5 31.0 22.2 0.224
13.0330 26.0 27.6 21.5 14.8 0.217
17.1130 21.3 35.7 18.4 9.7 0.199
17.1730 43.0 37.1 33.2 25.7 0.185
21.1830 DenbighShire 22.0 30.7 18.9 12.0 0.053
22.0630 22.0 28.7 17.2 10.3 0.053
22.1730 17.0 30.0 13.2 6.3 0.038
25.0530 25.0 35.7 16.2 0.1 0.023
25.1430 25.0 36.4 17.0 4.3 0.016
26.0430 17.5 35.7 13.2 1.0
26.1730 17.5 39.2 12.0 0
64
-------
Taking gasifier losses first, it is apparent that these varied
considerably during the run. Previously, in batch units it had
been found that gasifier loss rate was dependant on make—up rate
and bed depth. (1) In Run 5, however, the situation was more
complicated since cyclone performance was known to have deteriorated
sometime during the run. This was evident from the after—run
inspection which showed the cyclones to be in poor condition.
Statistical analysis of gasifier loss rates showed an inconsistancy
between the first and second data points (3.1530 and 3.2030).
Also all later variations for both stones could be explained by
changes in bed depth and make-up rate in a single correlation.
From this analysis, the following was deduced. Firstly, cyclone
performance deteriorated between the 3.1530 and 3.2030 data
points and did not alter appreciably thereafter. The reason for
the deterioration is not as yet clear. Secondly, gasifier loss
rates at constant cyclone performance were shown to depend on
make-up rate and bed depth. Whilst the effect of make—up rate
was similar to that observed in batch units, bed depth appeared
to be less significant. Thirdly, gasifier loss rate was
independent of limestone type in this instance. This would not
always be the case. Here, it would appear that the cut-off point
for the reduced performance cyclones and the attrition patterns
for the two limestones is combined to cause this phenomenon.
For stack losses, a very different picture emerged when these results
were examined in detail. The two limestones behaved differently.
Under all conditions examined, stack losses were small when the
bed was composed mainly of Denbighshire limestone. With
predominantly BCR 1691 (5.1030 to 17.1730) , however, they were
appreciable. They were also higher when of the order of 20% of
the bed as estimated from Si0 2 /CaO ratios was BCR 1691 (21.1830 to
22.0630) . Statistical analysis also showed that stack losses from
a 5CR 1691 bed correlated with make-up rate and bed depth.
Since the performance of the stack cyclone did not change during
the run, these variations in stack losses can only be explained
if it is accepted that BCR 1691 produces a fraction of material
of much smaller particle size than Denbighshire. This has been
indicated from Run 4 and batch test data.
Since the stack cyclone was designed to have the same efficiency
as the gasifier cyclones, the results from Run 5 indicate that a
gasifier following the same principles for solids handling as the
pilot plant and with gasifier cyclones fully operational would
give negligible gasifier losses with a limestone of Denbighshire
type and gasifier losses of the order of stack losses with a
limestone of BCR 1691 type.
65
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Bed Homogeneity
Both Denbighshire and 3CR 1691 limestones were tested in Run 5.
A measure of bed homogeneity with respect to limestone type was
obtained by analysing bed samples for silica and calcium oxide
and comparing the ratio of these two compounds. The silica
to calcium oxide ratio for the high purity Denbighshire stone
is much lower at 0.006 than BCR 1691 at 0.28. Results
are summarised in Table 16.
For the first two data times, the system was completely homogeneous
since only Denbighshire stone had been added. During the subsequent
BCR 1691 test period, the bed always had a Si02/CaO ratio below
that of the raw stone. Also no persistent increase in the ratio
was observed.
During the final Denbighshire test period, a reduction in Si02/
CaO ratio took place and the ratio of the raw Denbighshire stone
was approached. Material from the stack cyclone gave simi.lar
results to that from the regenerator cyclone throughout. During
the 3CR 1691 test period, the Si02/CaO ratio in these fines was
generally higher than that for the raw stone.
The significance of the Si/Ca ratios during the 3CR 1691 tests is
somewhat obscured by the fraction of Denbighshire stone which
remained in the bed following changeover to BCR 1691.
We believe that the low Si02/CaO ratios in the bed during BCR 1691
tests was due to residual Denbighshire stone. The absence of any
increase in ratio over the period can be attributed to the addition
of Denbighshire stone during burnout and maintenance periods
between 6.0730 and 12.1630 and also 13.0600 and 17.1130.
We discount the possibility of a preferential loss of silica from
the beds being the cause of the low results even though higher
Si0 2 /CaO ratios were observed in cyclone fines. The reason being
that beds composed only of BCR 1691 in batch tests and Run 4 showed
the opposite effect in that silica was concentrated in the bed.
In those instances the fines which were trapped also showed a
higher ratio than the raw stone indicating that the material lost
completely from the system was rich in calcium. Assuming the
same to have happened here with the BCR 1691 fraction of the bed,
then the higher Si0 2 /CaO ratio in the fines can be explained by
BCR 1691 being lost preferentially.
After the final change to Denbighshire feed, the Si0 2 /CaO ratios
indicated that some 3CR 1691 was present in the bed to the end
of Run 5, albeit in ever decreasing amounts.
66.
-------
Table 16
Silica/Calcium Oxide Ratios (Run 5 )
Time Limestone Gasifier Bed Regenerator Stack Cyclone Regenerator Cyclone
Day Hour Feed Si0 2 /CaO S 10 2 /CaO Si0 2 /CaO Si0 2 /CaO
3.1530 Denbighshire 0.006 0.006 0.006 0.006
3.2030 0.006 0.006 0.006 0.006
5.1045 BCR 1691 0.203 0.215 0.283 0.305
6.0730 0.236 0.218 0.261 0.300
12.1630 0.228 0.258 0.301 0.283
12.1800 0.224 0.225 0.316 0.286
13.0400 0.217 0.313
13.0600 0.242 0.225 0.312 0.304
17.1130 0.199 0.106 0.303 0.266
17.1800 0.185 0.214 0.301 0.273
21.0700 Denbighshire 0.056 - 0.009
21.1800 0.053 0.015 —
22.0715 0.053 0.007
22.1745 “ 0.038 0.010
25.0530- 0.023 0.020 0.006 0.006
25.1430 0.016 — —
-------
Particle Size of Solids
The particle size distribution of solids in the reactor beds
depends on size of the feed, particle attrition, and effectiveness
of the cyclones in returning fines. Sieve analysis of a number of
gasifier and regenerator samples are presented in Appendix D.
Histograms of the average stone feed and two sets of gasifier and
regenerator bed samples are given in Figure 16.
The two sets of bed samples illustrated represent extremes of the
samples taken. Performance of the fines return system was poor at
13.0400, and fines were being lost from the unit. The cyclones
and fines return system were operating relatively well at 21.1600
as shown by the larger fraction of small particles.
These figures show that the fraction of particles in the 250 to
1400 micron size range was increased in the unit at the expense
of both larger and smaller sizes. The fraction below 250 microns
in the gasifier was lost altogether while the quantity of material
above 1400 microns was reduced.
Poor performance of the fines return system, as at 13.0400, causes
significant loss of particles as large as the 600-850 micron range.
Regenerator and gasifier beds were quite similar in size
distribution with the regenerator showing a slightly higher fraction
of fines.
Figure 17 shows the variation, during the run, of the fraction of
bed in the size range below 600 microns.
This figure indicates a rapid deterioration of fines return
effectiveness during the early part of the run. It indicates
that fines return was restored during decoking before day 21
startup but again declined toward the end of run.
A similar pattern is shown in Figure 18 where average particle
size of the bed is plotted against run time. The particle size
here is calculated from sieve analysis, using the relationship:
w
W/d
which gives a surface area mean particle size.
68
-------
r C
—o
— 100/v
Li
z
4
Li
NJ
U)
2
— 100/0
>-
C-)
z
Li
0
Li
Figure 16
Size Distribution of Fluid Beds
(Run 5)
0
1000 2000
PARTICLE SIZE, MICRONS
3000
69
-------
35
I 1 I I
30 — GASIFIER 0
REGENERATOR X
U)
z
025
a:
o I
o
0
X
z 0
\ \
4
a:
w 15- 00
-J
\X
-J
x
(I )
10
I-
‘I
0 &
I
5
0 I 1 H 1 i I1 1 II I
0 5 10 5 20 25 30
RUN DAY
Figure 17 Gasifier and Regenerator Fines
below 600 microns (Run 5)
x
70
-------
I I I I
0
XO
x
GASIFIER
RE GE NERATOR
I l I H 1 I I 1 L 1
0 5 (0 (5 20 25
RUN DAY
Figure 18 Average Size of Bed Particles (Run 5)
0-
x
00
U)
2
0
U
U)
0
-J
0
(I)
w
0
w
NJ
C,)
w
-J
(-3
I-
w
4
U i
>
4
Ui
U
4
D
C /)
1400 —
1200 -
(000 -
800-
600-
400-
200-
0
x
x
x
30
71
-------
Nitrogen Oxides
The concentrations of nitrogen oxides measured in the boiler flue
gas during Run 5 are compared in Table 17 with values measured
in previous tests. All samples were taken from the boiler flue
gas by means of gas sample bags and analysed off line in the
laboratory. A chemiluminescence method was used for Run 5
samples.
The ASTM D16o8 phenol-disulphonic acid method was used for the
other samples. Laboratory cross checks have indicated that the
two methods are in agreement.
Run 5 results agree generally with earlier gasification test results.
All of the gasification results have lower NO concentrations than
the tests with the oil burner. This improvein nt during CAFB
gasification is probably due to a combination of the effects of
two stage combustion and the use of flue gas recycle. It is
possible that nitrogen compounds in the fuel are converted to a
harmless form during gasification, and that flue gas recycle
reduces maximum flame temperature which reduces equilibrium NO
concentration. x
Even with the original oil burner operation the NO concentrations
were low when compared with concentrations in the flue gas of large
power station boilers. This difference in overall level is believed
to be caused by the close proximity of the flame to the large water
cooled surface in the fire tube boiler used in the pilot plant.
Therefore, although the reduction in NO level caused by CAFB
gasification is believed to be realisti , the absolute low level
achieved probably would not be reached in large power generation
boilers where flame temperatures are much higher.
Thermal Behaviour
Most of the heat released by partial combustion of fuel in the
gasifier is retained as sensible heat in the gas going to the burner.
The heat release in the gasifier has been estimated from thermal
equations for the masses and heat contents of the various streams
entering and leaving the gasifier. Table 18 lists results for
Run 5 test conditions. The equations used are explained in Appendix E.
Heat losses from the pilot plant gasifier bed, based on reasonable
values of thermal conductivity and heat transfer coefficients,
amount to about 1% of the heat released in the gasifier. Depending
on lime replacement rate, about 2 to 4% of the heat released goes
to calcine stone and raise the lime to gasifier temperature. At
the air/fuel ratio employed, which did not deviate much from 20%
of stoichiometric, the heat release per pound of fuel was estimated
to be approximately 3100 BTU/lb corresponding to 155,000 BTU/mole
of oxygen.
72
-------
Table 17
Nitrogen Oxides in CAFB Boiler Flue Gas
Operating Mode Sample Date-Time Method NO p Flue Gas 0 _ Vol% Oil Rate GPH
Oil Burner-low fire May 1971 ASTM D1608 256 3
II It it II 249 0
Oil Burner—High fire 280
it II 266
Gasification Run 1 August 1971 179 —
II It It II 155 —
II II II II 172 —
It II TI I I 200 — —
Run 3 1 Dec 1971 15:55 126 2.9 — 3 39.4
16:00 120 11 II
16:05 130 1
3 Dec 1971 0950 163 2 4 — 2 3 42.2
09 : 5 5 173 I I
II 1005 1 1 169
° ° “ “ 10: 10 163 0 II
° 6 Dec 1971 1345 237 1 0 35.0
II II It l3 5O II 186
II It 14 00 ° 18 1
14:05 (a) 166 II I I
Rut 5 9 Feb 1973 15:00 Chemiluminescence 101, lOs, 110 3.’) ,13.1
17 Feb 1973 0000 180 3 7 44 0
22 Feb 1973 16:10 185 2.9 43.5
(a) Thermo Electron Corporation N0 Analyser Model 10 (a)
-------
Table 18
Heat Release in CAFB Gasifier
Run 5
Air/Fuel
Run.Time % of
Heat Lost
% of
Calcination arid
Stone Heating
Heat From
Regenerator
Heat
Heat
R lease
Stoichiometric
Heat Release
%
Heat
of
Release
%
Heat
of
Release
Release
B U 4 Hr 6
BTU/lb oil
2.2130
19.8
.73
1.49
4.10
3.0530
20.4
.78
1.58
5.33
1.23
3102.
151459.
3.1530
20.6
.71
1.16
5.14
1.26
3197.
3.2130
20.3
.66
1.49
4.64
1.29
3195.
5.1030
21.3
.83
4.42
4.88
1.31
3276.
152768.
.0730
20.4
.77
3.02
4.31
1.24
3236.
149531.
-
12.1530
12.1930
20.6
20.6
.94
.97
3.06
2.93
5.66
7.57
1.23
1.19
3105.
3025.
145776.
13.0330
20.7
.80
2.60
5.71
1.26
3215.
154366.
12.0830
20.6
.86
2.15
5.39
1.26
3165.
152362.
17.1230
21.3
1.09
2.38
5.31
1.25
3128.
145641.
17.1730
21.9
1.04
4.03
5.56
1.28
3186.
144305.
21.0930
19.7
.80
3.05
4.96
1.39
2967.
21.1830
19.0
.82
3.67
4.07
1.40
29/7.
155911.
22.0730
19.6
.86
2.43
5.02
1.27
3066.
154131.
22.1830
19.8
.88
2.58
5.14
1.26
3078.
24.0730
19.2
.97
4.92
5.30
1.29
3145.
162932.
25.0530
19.2
1.06
3.66
5.89
1.27
3130.
162062.
25.1530
19.].
1.07
3.40
5.62
1.25
3084.
160495.
26.0530
18.9
1.07
2.63
6.06
1.25
3048.
160315.
161531.
26.1830
18.7
1.16
3.57
5.74
1.25
3046.
13.0630
20.8
.80
2.22
5.77
1.26
3225.
166185.
26.1130
18.1
1.10
3.37
5.98
1.24
-------
Figure 19 shows the observed variation of fuel heat release with
air/fuel ratio, and compares the measured values with a line
calculated for a release of 155,000 BTU/mole of oxygen. Values
from all the pilot plant runs to date fall along this line. This
value of 155,000 BTU/mole 0 agrees well with heat release
calculated from the fractio of carbon and hydrogen oxidised arid the
Co/CO ratio formed in the gasifier. Details of this calculation
also appear in Appendix E.
-J
0
-J
L i i
cc
—I
Figure 19
Heat Release vs. Air/fuel Ratio
During Gasification
AIR/FUEL RATIO, % OF STOICHIOMETRIC
75
-------
Product Gas Composition
Four samples of the gasifier product vapour were collected during
the run and analysed by gas chromotograph. This analysis gives
dry gas composition on a water and liquid hydrocarbon free basis.
Results are listed in Table 19.
Table 19
Product Gas Composition
Run 5
Sample Time 22.1030 22.1745 26.0400 26.1800
Composition, Vol %
(Air Free Basis)
N 2 61.7 62.5 63.4 64.7
Co 2 10.82 10.89 10.88 10.17
Co 8.36 7.52 9.36 8.97
112 6.86 7.75 6.42 5.80
CH 4 6.75 6.29 6.09 6.02
C 2 H 4 5.36 4.83 3.81 4.38
C 3 H 6 .11
76
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A material balance calculation on the product gas composition
and unit feed rates permits an estimate of the quantity and
composition of that portion of the product vapour which is missed
by the gas chromatograph because of condensation in lines or in
the sample container.
In this calculation nitrogen and oxygen balances are forced to
100% and the amount of hydrogen and carbon not accounted for is
assumed to make up the liquid fraction. The flue gas recycle
stream was assumed to contain the same H 2 0/C0 ratio as the
boiler products of combustion. Table 20 summ rises results of
this calculation. Details appear in Appendix TableF—l.
The results are fairly consistent in indicating the fractions
of carbon and hydrogen oxidised and the quantity of carbon which
goes either to coke or heavy hydrocarbons. The results on hydro-
gen disappearance show greater variability however.
This variability is probably due to the method of calculation
which depends on finding small differences between relatively
large numbers.
In particular, the hydrogen/carbon ratios of .25 and .23 calculated
for the heavy components on day 22 appear to be unreasonably low.
It is unlikely that the true H/C ratios could be much less than
1.0. These results show the desirability of obtaining accurate
samples of the total gasifier product, including light and heavy
components. However, collection of such a sample is quite
difficult in practice, and will require development of a suitable
quench and recycle system to collect the liquid fraction without
plugging.
Re nerator
Performance of the regenerator during Run 5 appeared to be less
satisfactory than in earlier tests and was somewhat inconsistent.
It was disappointing in that sulphur concentration in the off gas
and selectivity of calcium sulphide oxidation to calcium oxide
plus SO appeared to be much lower than the levels which earlier
runs ha shown to be possible. Results were inconsistent in that
the apparent sulphur production rate of the regenerator could
account for only about half the sulphur being absorbed in the
gasifier.
Furthermore, SO release based on gas analysis did not agree with
SO based on so ids analysis. Obviously this matter requires
additional study to locate the cause of the discrepancy.
The run data shown in Appendix Figure D—2 are the gas analysis based
values of SO 2 concentration and selectivity.
77
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Table 20
Su ary of Gasifier Cc spoment Distributions
(Calculated from Flows and Product Gas Composition)
Sample Time 22.1030 22.1745 26.0400 26.1800
Oxygen In, % of Total
With Air 67.7 68.4 66.9 66.4
With Flue Gas Components 23.7 23.6 23.0 23.6
From Solids Reactions 8.6 7.9 10.0 10.0
Oxygen Out, % of Total
As Sulphate on lime 1.7 1.6 1.3 1.5
As Carbon Oxides 83.7 81.6 84.4 77.9
As H 2 0 (by Difference) 14.5 16.7 14.4 20.5
Hydrogen In, % of Total
In Oil Feed 94.1 94.2 94.3 94.3
In Flue Gas H 2 O 5.9 5.8 5.7 5.7
Hydrogen Out, % of Total
As Dry Gas Components 82.2 79.6 67.1 49.8
As H 2 0 13.8 15.8 13.6 18.7
As Heavy Components (By Difference) 4.0 4.6 19.3 31.4
Carbon To Gasifier, % of Total
In Oil Feed 93.2 93.5 93.5 93.1
In Flue Gas, C oxides 5.8 5.6 5.5 5.5
In Stone 1.0 0.9 1.0 1.3
Carbon From Gasifier, S of Total
As Oxides from Gasifier 40.6 38.7 41.5 37.1
As CO 2 from Regenerator 1.0 0.4 0.2 0.3
As Hydrocarbon in Dry Gas 37.0 33.5 28.1 28.6
As heavy Components (By Difference) 21.5 27.4 30.3 34.0
C oxidised, S of Feed 35.8 34.0 37.0 31.9
H oxidesed, S of Feed 14.6 15.8 14.4 19.9
C in Heavy Components, S of Feed 23.9 30.0 33.1 37.4
H in Heavy Components, S of Feed 4.3 4.8 20.4 33.3
CO/CO 2 in fresh oxides 1.102 .97 1.22 1.36
H/C in Heavy Components .29 .26 .99 1.43
Air/Fuel Ratio, of Stoichiometric 19.5 19.8 19.3 18.3
‘T8
-------
Table 21
Summary of Regenerator Performance
Run 5
Regenerator Sulphur Output
By Gas By Solids Gasifier
Analysis Analysis SRE
% of Fed lb/Hr % of Fed ________
30.8 5.55 59.4 72.8
39.7 5.44 56.6 60.5
44.3 7.83 81.1 93.8
38.9 5.83 63.1 96.6
36.5 7.41 77.1 84.6
40.7 6.68 70.2 86.1
47.8 5.06 53.4 87.9
45.2 5.13 54.8 88.2
51.7 9.60 99.5 80.5
26.9 5.77 50.9 92.9
40.7 6.20 54.9 87.3
34.1 4.83 49.6 90.9
40.9 5.94 59.9 85.2
37.4 5.16 52.3 95.7
27.3 5.80 59.1 90.6
31.8 4.07 40.2 88.9
39.2 5.28 53.0 90.8
34.0 5.82 58.9 85.4
Time
S in Solids
wt.%
Gasifier Regenerator
Selectivity
% CaS to CaO
Gas Solids
Analysis y j s
L!
3.1530
2.89
1.82
27.8
50.5
2.87
3.2230
2.72
1.81
37.6
52.9
3.82
5.1030
3.66
1.59
40.9
70.5
4.28
6.0730
3.72
1.64
40.5
61.0
3.60
12.1630
2.88
1.85
30.4
59.2
3.51
12.1830
2.80
1.80
32.8
53.5
3.88
.
D
13.0430
13.0630
2.76
2.53
2.06
1.83
37.2
34.2
40.7
40.7
4.53
4.23
17.1130
3.11
1.00
44.0
78.]-
4.99
21.0730
2.69
1.78
28.2
50.0
3.05
21.1830
2.85
1.75
50.0
64.3
4.59
22.0730
2.90
2.21
32.5
46.0
3.32
22.1730
2.91
1.91
40.0
56.2
4.06
25.0530
2.74
1.88
32.0
43.7
3.66
25.1430
2.52
1.72
26.8
53.3
2.68
26.0430
2.74
2.17
27.5
34.5
3.16
26.1030
2.70
1.96
34.1
44.8
3.90
-------
Table 21 compares regenerator sulphur emission figures based on
gas analysis with values calculated from solids analysis.
Solids compositions were those of the gasifier and regenerator
beds. It is apparent that sulphur production values based on the
solids analysis are much higher than the gas analysis based
figures. Similarly, the calculated values for oxidation
selectivity are much higher when based on solids analysis. The
solids based selectivities for this run are in good agreement
with the gas analysis based selectivities of Run 3 (Figure 38
of Reference 1)
In Run 5 the level of sulphur in the bed was relatively low
compared with the sulphur content of the cylcone fi ies. With
the external cyclone fines return system these fines had little
chance to be regenerated. In Run 3 however, fines from one
cyclone drained to the regenerator. This arrangement must have
decreased the sulphur content of the circulating fines and
reduced the effect of fines loss to the boiler on sulphur removal
efficiency.
SULPHUR BALANCE
Use of the gas analysis figures for regenerator sulphur output
leads to very low sulphur material balances for the unit.
Assuming that measured values of sulphur removal efficiency are
correct, it is not possible to account for the missing sulphur
by assuming that it left with the lime purge stream of fines
losses. A fault in the regenerator off gas analyser could
explain the discrepancy, but checks and calibrations made on
the instrument during the run indicated that it was functioning
properly. Similarly, a larger gas flow through the regenerator
than measured would cause a low estimate of sulphur production.
Such a large error in gas measurement is unlikely because air,
to the regenerator is measured both by orifice and gas meter,
nitrogen to the solids transfer system is metered, and nitrogen
to instrument bleeds is negligibly small. The only other
possibility is a major leakage from gasifier to regenerator,
and again this is believed unlikely. To help solve this mystery,
additional analyses and measurements will be made in future runs.
Gas flow rate out of the regenerator will be checked by a helium
tracer method, and a gas chromatograph will be used to check
regenerator gas composition for SO 2 and other sulphur compounds.
We also plan to modify the boiler flue gas sampling system to
reduce further the possibility of losing sulphur in the sample
lines and filters. It is possible that SO 2 is absorbed on lime
dust which enters the sample line, and such absorption would
produce an optimistic estimate of sulphur removal efficiency.
Such errors are believed to be small however, as the lines are
80
-------
frequently cleaned, and spot checks with Draeger tubes (direct
reading SO 2 colour change tubes) made directly through the
boiler door sample point agree with the continuous reading
instrument.
81
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SECTION X
SCOPING OF ENGINEERING EFFORT
GENERAL
As part of this project, Esso Engineering, Florhain Park,
New Jersey, USA, was requested to scope the engineering effort
which might be required to carry CAFB from its present stage
of development through the construction, startup, and testing
of a large scale demonstration unit. A 100 MW scale unit was
assumed as a basis. This scoping study is summarised here.
Detailed results have been supplied to the Environmental
Protection Agency in a separate memorandum.
SUMMARY
The total development of CAFB through a 100+ MW demonstration
test period is expected to take about 6-¼ years and require
$3,320,000 in engineering effort. Of this total, $570,000 is for
developmental engineering and pilot plant guidance, the
remaining $2,750,000 is associated with the demonstration
project. Optimistically, the total development might be
accomplished in 4—¾ years with $2,520,000 of engineering effort.
Conversely, greater costs and times could be experienced.
The approximately 1 MW pilot plant at Abingdon can be scaled to
100+ MW without an intermediate pilot plant but with some risk.
To reduce the risk, large scale mock-up studies of the fluidi-
sation system and special engineering development of critical
equipment should be carried out. This scope includes the
engineering manpower for these studies but does not include
money to build any large scale test rigs. These are assumed to
be included in the laboratory programme.
DISCUSSION
The approach used in preparing this study was to assess the state
of development and the complexity of the process and then determine
a reasonable schedule to carry the development to completion
consistent with the criteria of reasonable risk. Previous
experience in process and project developments similar to this
was used as a guide. An estimate was then made of the various
types of engineering activities required during the development
and the extent of the effort required for each activity. This
estimate excludes engineers who are directly associated on a
full—time basis with research activities or the operation of
pilot plants or demonstration plants.
The total programme has been divided into four activities, and a
“most probably” and an “optimistic” development schedule has been
estimated for each activity. Of course, there is also the
possibility of a longer and more costly development programme if
developmental problems prove extensive or if significant modif i-
cations are required to the demonstration plant due to start up
difficulties.
82
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The four activities in the development programme are:-
(1) Process Development
(2) Process Design
(3) Detailed Engineering, Procurement, Erection
(4) Startup and Test
The engineering effort is suinmarised in Table 22. The costs
shown there are grouped to distinguish the engineering guidance
costs during pilot plant development work from the engineering
costs associated with design, erection, startup, and test of
the large demonstration project. The project work is divided
into two categories; basic engineering and the prime contractor’s
effort. Basic engineering includes the basic process design,
owner’s interest protection and follow-up during the detailed
engineering and erection stage, and the engineering associated
with start-up and testing. The prime contractor’s effort
involves the engineering required for mechanical design and
erection of the demonstration plant.
In this schedule it was assumed that a client for the demonstration
plant would be obtained near the end of the small scale develop-
ment phase at which point a site would be selected. Certainly
the earlier the client is located, the sooner it will be possible
to direct both engineering and pilot plant activities towards a
specific project with improved chances for shortening the time
and cost of the development work.
The schedule and cost of the development activity is a major
uncertainty in this type of effort. The schedule depends to a
large extent on the degree of effort expended and the degree of
risk which might be considered acceptable when starting the
demonstration plant design. The estimate of 18 months for
additional small scale development assumes minimum future pilot
plant problems, very little process optimisation, and a higher
risk in proceeding with the demonstration plant than if more
extensive (experimentation and engineering) development work were
undertaken.
83
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TABLE 22
CAFB DEVELOPMENT PROGR1 MME
SUMNA RY OF ENGINEERING EFFORT( 1
Most Likely Opti1t iStiC
$M (2) DateS 2Y Dates
DEVELOPMENT 570 1/73—7175 360 1/73—7/74
PROJECT
• Basic Engineering
— Process Design 420 10/74-10/75 270 1/74-8/74
— Owners Interests 580 10/75—7/77 390 6/74-9/75
— Startup and Test 550 7/77—4/79 500 9/75—4/77
1,550 1,160
• Contractors Design
and Erection
1,200 10/75—7/77 1,000 6/74—9/75
3,320
(1) Cost of Engineering work only - costs of experimefltatiofl#
pilot plant work, construction, and demonstration plant
operation are excluded.
(2) M Thousands.
84
-------
SECTION XI
REFERENCES
1. Study of Chemically Active Fluid Bed Gasifier for Reduction
of Sulphur Oxide Emissions. Final Report, OAP Contract
CPA 70—46, Esso Research Centre, Abingdon, Berkshire,
June, 1972.
2. Curran, G.P., Fink, C.E. and E. Gorin.
Phase II Bench—Scale Research on CSG Process, R&D Report
No. 16. Report to Office of Coal Research, Contract No.
14-01-0001-415, Consolidation Coal Co. July 1st, 1969.
85
-------
SECTION XII
INVENTIONS
1. UK 50014/72 Moss, Craig, Taylor and Tisdall
Preventing agglomeration during regeneration of sulphides
by passing stone from the gasifier into a region of the
regenerator which is separated by a layer of fluidised
stone from the regenerator distributor.
2. UK 24739/72 G. Moss
Production of a highly suiphated lime from CAFE regenerator
off—gas, to avoid the need for reduction of SO 2 to sulphur
or production of sulphuric acid.
3. UK 29513/72 Moss and Taylor
Reduction of attrition in fluidised beds by a two stage
nozzle, the first stage being a high pressure drop orifice,
and the second stage providing dissipation of kinetic energy
and a non—attriting gas velocity into the fluid bed.
86
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SECTION XIII
GLOSSARY
TERMINOLOGY
Sulphur Differential - The difference in total sulphur
level on the fluid bed between
start and finish of a batch
gasification run, or between the
inlet and outlet streams from
gasifier to regenerator in a
continuous unit.
Superficial Velocity — The velocity of the fluidising
gases (air plus flue gas recycle,
but excluding gas and vapour
formed from the fuel) in the
empty gasifier or regenerator bed,
at the temperature of the bed.
Fluidised Bed Depth - (Pressure drop from above distri-
bution to gas space above the
bed) +Bed density
Lime Replacement - Fresh limestone added to the
gasifier, expressed as weights
of CaO in the limestone added
over a given period per unit
weight of sulphur in the fuel
gasified during the same period.
Alternatively expressed as a
ratio of moles CaO added per mole
S in the fuel gasified.
Sulphur Removal Efficiency(SRE)— ( 1 SO 2 _observed in flue gas xlOO%
S02 if none absorbed
Calciriation - Removal of CO 2 from limestone by
heating above approximately
750 deg. C.
Adiabatic Gasification - Operation at low air to fuel ratios
in the gasifier such that heat
released by partial oxidation of
the fuel just serves to maintain
the gasifier temperature at the
required level.(Air supply about
14% of that needed to fully
combust the fuel)
Combustion — Operation at high excess air
levels during combustion in the
gasifier such that the gasifier
temperature just remains at the
required level (Air supply about
400% of that needed to fully
comhust the fuel).
87
-------
Megawatt (MW) — Used in this report only for
electrical power generation rate.
Thermal energy rates are all
expressed in Btu/hr.
SYMBOLS USED IN TEXT
A Bed age in hours (batch tests)
D Bed Depth (inches)
d particle size, microns
day Surface area mean particle size, microns
I. Loss rate from bed, g/min
S Bed Sulphur Content (total), weight %
T Bed Temperature, deg. C.
W Weight of fraction in sieve analysis
,u Micron (10-6 metre.)
88
-------
SECTION XIV
APPENDICES
P age
APPENDIX A - Data from Batch Unit Studies 90
APPENDIX B — Cyclone External Drain System 99
APPENDIX C — Inspection of Pilot Plant After 109
Run S
APPENDIX D - Data from Run 5 124
APPENDIX E - Heat Balance Equations 185
APPENDIX F - Gasifier Product Composition 197
-------
APPENDIX A
t ta From Batch Unit Studies
90
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
Table A-4
Kerosene Combustion (Tests 1-A, 2-A, 3-A)
Stone: BCR 1691 Deflbi hah1re BCR 1 59
Time Bed Bed Cyclone Bed Bed Bed Cyclone Bed Bed Bed Cyclone Bed
Depth Density Dust Temp. Depth Density Dust Temp. Depth Density Dust Temp.
(hi) (inches) (lb/SCF) (g) (‘C) (inches) (lb/ 3CF) (g) (‘C) (inches) (lb/SC! ?) (g) (‘C)
0 23.) 46.8 - 875 20.7 46.8 - 70 22.7 46.8 - 850
18.4 49.4 170 900
19.1 49.4 280 870 21.5 49.4 90 842
1 15.9 49.4 920 880 19.6 46.3 135 865 20.4 52.0 70 860
15.3 46.8 1005 855
19.3 46.8 125 870 20.9 19.4 6o 860
2 14.4 48.1 390 890 18.4 49.4 80 870 20.0 52.0 50 865
2 13.2 48.1 420 88o 18.1 43 J 4 8 885 20.0 52.0 87 865
3 12.3 48.1 290 893 17.8 49.4 35 890 19.4 52.0 55 872
11.9 48.1 280 900 18.0 49.4 40 880 19.3 55.0 40 870
17.3 51.0 o 860 19.2 53.6 30 880
- - - 19.4 52.0 30 88o
5 - - 19.9 51.0 27 860
- - - - - 20.4 49.9 23 860
6 - - - 19.6 51.1 35 86°
-------
Table A—S
Ker sene Combustion (Tests 1-B, 2—B)
Stone: BCR 1691 Denhi hshire
Time Bed Bed Cyclone fled fled Bed Cyclone Bed
Depth Density Dust Temp. Depth Density Dust Temp.
(hr) (inches) (lb/SCF) (g) (°C) (inches) (lb/SCF) (g) (°c)
0 21.8 45.8 0 1020 19.8 52.0 - 1010
20.4 45.8 317 1080 17.6 56.2 270 1015
1 21.8 44.2 55 1065 17.3 57.2 90 1080
20.0 46.8 235 1078 15.5 62.4 27 1050
2 19.3 46.8 380 1080 - - - -
18.3 47.8 1)0 1070 16.5 57.2 47 io8o
3 18.6 47.8 80 1058 15.4 61.4 18 1080
18,2 47.8 75 1054 15.5 62.4 20 1060
4 18.5 46.8 75 1078 15.3 62.4 14 1050
17.7 49.4 60 1090 15.0 63.4 12 1050
5 17. 49.9 50 107 15.2 62.4 11 1050
5 - 17.2 51.0 50 1065 15.2 62.4 10 1060
6 17.0 49.9 30 1065 15.2 62.4 6 1080
-------
Table A—6
Fuel 011 Combustion (Tests l -C, 2—C)
Stone: BCR 1691 Denb hshire
Time Bed Bed Cyclone Bed Fuel Bed Bed Cyclone Bed Fuel
Depth Density Dust. Temp. Rate Depth Density Dust Temp. Rate
(hr) (inches) (lb/SCF) (g) (°c) (g/min) (inches) (lb/sCF) (g) (°c) (g/min)
0 24.5 41.6 - 870 - 22.8 46.8 - 850 -
23.0 41.6 180 810 34 21.1 46.8 160 850 27
1 20.4 44.2 155 870 47 19.7 49.4 260 850 27
l c 17.8 49.4 113 875 44 19.1 49.4 310 868 27
2 18.0 46.8 170 870 49 18.2 52.0 100 868 23
2 17.6 46.8 70 870 49 17.6 75 872 27
3 17.6 46.8 70 880 2 47 17.1 53.6 60 878 25
16.5 49.4 57 860 47 17.6 52.0 85 878 25
4 16.3 49.4 23 8 o 47 16.9 53.0 68 965 25
16.5 249.4 35 840 40 17.2 52.0 62 890 25
5 16.4 49.9 35 840 38 15.9 57.2 45 850 23
15.7 52.0 850 38 15.6 57.2 67 860 23
6
-------
Table 6-7
Fuel 011 GasI Flea) ton (Tests 1—B, 2—B, 3—13)
3tone tyR 1691 Denb t hohire 8CR 1)59
Time Bed Bed r’yelore Bed Fuel l Ied Rod 5yclor e Bed Fuel Bed Bed Cyclone Bed Fuel
Depth Densify Dest Temp. (late Depth Density Dust Temp. Rate Depth Density Dust Temp. Bate
(he) (inches) (lh/GCF) ( ) 570) (e,/mtn) (inches) (Ib/COF) ( ) (57) ( /mln) (Inches) (lh/SCF) (g) ( 0) ( /rnln)
0 ‘1.6 57.4 - (368 - 22.7 46.8 - 900 — 22.0 46.8 — o -
2250 46.5 40 060 170 19.1 51.0 190 850 178 20.) 57.11 180 835 17)
1 l9. 4 52.4) 545 860 178 17.0 .1 290 360 204 20.3 49.14 127 860 iSo
1 m 19.6 51.1) 135 8 O 170 16.2 ,7.2 100 850 178 i9.k 52.0 90 85) 192
2 oo.4 si.o 5 860 1)7 16.? 57.2 60 818 208 19.1 54.6 90 868 204
19.7 49.4 10 855 200 16.7 57.2 150 862 206 19.4 52.0 60 860 204
3 18.8 51.0 100 855 2011 16.2 57.2 40 862 211 19.2 52.0 67 868 214
18.8 52.0 70 855 211 15.6 57.2 75 850 221 18.) 511.6 62 870 221
4 19.11 52.0 70 850 259 15.9 57.2 3) 040 242 iR.6 54.6 60 860 221
19.11 52.0 86 862 221 16.2 57.2 60 850 238 18.) 55.6 60 860 242
5 18.0 54.6 75 870 2)1 15.7 59.8 50 860 224 19.5 57.2 50 860 2)4
18.0 54.1 70 870 185 15.4 59.8 60 865 221 18.7 59.8 52 860 2)4
6 17.8 54.1 55 86 216 15.7 59.8 35 870 212 18.7 59.8 — 860 234
-------
Table .4-8
Kerosene Combustion on Cycled Bed (Tests 1-F, 2—0)
Stone: BCR 1691 Denbighshire
Time Bed Bed Cyclone Bed Bed Bed Cyclone Bed
Depth Density Dust Temp. Depth Density Dust Temp.
(hr) (inches) (lb/SCF) (g) (°C) (inches) (lb/SCF) ( ) (°c)
0 22.7 46.8 - 900 18.7 59.8 - 850
- 20.0 52.0 240 880 18.0 62.4 18 830
1 19.4 52.0 175 815 17.8 62.4 10 843
18.8 52.0 170 830 17.6 62.4 12 853
2 18.2 52.0 280 600 17.5 63.4 8 860
2- 17.6 52.0 100 800 17.7 62.4 12 870
3 17.6 52.0 40 850 18.0 62.4 10 872
17.3 51.0 230 620 18.0 62.4 10 862
4 17.3 51.0 100 F70 17.5 64. 8 865
16.4 52.0 40 850 16.9 65.0 5 68
5 16.2 52.0 40 850 17.4 65.0 8 875
5 j 15.8 52.0 100 850 17.0 65.0 7 880
6 15.9 49.4 140 870 16.4 65.0 7 880
-------
APPENDIX B
CPFB Cyclone External Drain System
99
-------
CAFB Cyclone Fines Beturn System
A major modification to the CAFB pilot plant made between runs
and 5 was installation of a system for externally draining the
main gasifier cyclone legs and returning the coarse fraction of
recovered solids to the gasifier. The system is shown in figure B..J .
Beneath each cyclone drain line is a conical recei ver which can
be isolated from its cyclone by a butterfly valve A. A line leads from
the bottom of each conical vessel to a common fines receiver mounted
above the gasifier. A pulser controlling valve D injects bursts of N
into this transfer line at a “N 2 knife” location just beneath conical 2
vessel. Another supply of nitrogen enters each conical vessel through
valve C.
Each transfer line is isolated from the fines receiver by a ball
valve E. The fines receiver drains into the side of an elutriator
vessel I luidised by nitrogen end fitted with a slug breaker near its
top. Gas from the top of the elutriator, together with gas from the
top of the fines receiver goes to a filter vessel in which fines are
retained. The coarse solids fraction from the bottom of the elutriatOr
drops through a water cooled heat exchanger to a bottom outlet from which
it falls into a pick—up line through which it is air injected back into
the gaslfler. A valve in the verti l line just above the air pickup is
controlled by a differential pressure switch to prevent flow reversal
from the gasifier toward the elutriator.
The operating sequence of valves A, C, D, and E is regulated from a
control panel in the control room. There is a separate control panel and
set of valves for each of the two gasifier cyclones.
The controllers operate pneumatically wlthflUidic elements
performing all logic and time delay functions. The sequence of operations
of the controller is as follows:
1. Valve A is open, valves C, D and E are closed. The vessel is filling
with solids from the cyclone. A small bleed of N 2 into the body of
valve P. purges the cyclone drain leg of gasifier product gas.
2. When a timer reaches its end point, a signal closes valve A. After
a short delay to allow A to shut completely, valve C opens to admit
N to the vessel to aerate and pressurise the solids in it. When
t e pressure reached a suitable level, say 5 psi, valve D begins
opening and closing in rapid sequence to inject pulses of N 2
into the transfer line. Simultaneously, valve E opens to deliver
material to the fines receiver. Material forced from the vessel
is broken into short slugs by ti-a action of the N 0 pu .ses. These
slugs travel with reduced pressure drop and with ‘ess gas consumption
than would be required for conventional pneuthatiC transport.
100
-------
When the vessel is empty, the pressure within falls rapidly
because there is no longer a resistance imposed by the solids.
The pressure sensor detects this fall and closes valves C, D and E.
Valve A then opens and a new filling period begins. We expect
that fill times will be of the order of 15 minutes and emptying
times about 1 minute or less. An interlock between the two
parallel systems prevents both operatIng at the same time.
Solids drop from the fines receiver into the elutriator vessel
where they are contacted with nitrogen at about 3 ft/sec.
Fine solids are carried overhead by this gas while coarser mater..al
falls into the downoomer which also is fluidised by N 2 . This
downcomer is water cooled to reduce solids temperature enough to
prevent rapid reaction with air. Fluidising N 0 for the downeomer
is injected about one foot above the bottom outlet. Nitrogen from this
injector passes both upward to fluidise the downeomer and downward
to seal against baokflow of air. A restriction at the bottom of
the downoomer limits the flow rate of solids to prevent choking the
air pickup. If the height of th4 fluid bed in the downcomer falls
too low to provide an adequate seal against air backflow, the pressure
detector senses the reduction and closes the bottom valve. The
valve opens again when the pressure difference is restored.
A detailed description of the operation of the fluidic controller
is given after Fig. B-l.
101
-------
nt
FIG. B-I CYCLONE DRAIN SYSTEM
Fines
Gasifier
Product
N2
N 2
102
-------
Description of’ Pneumatic Circuit
There are two parallel cyclone drain systems which empty to a
common e]ut,rlator. Fpch drain and transfer system is controlled by
ltw own pneumatic control panel. The two systems are interlocked to
prevent simultaneous transfer by both. -The controllers are based on
“Fluid Log” pneumatic elements made by Lang Pneumatic Ltd., Telford.
Shropshire. A diagram of the pneumatic”circuit appears in Figure B-2
Table I compares the numbering system used in the circuit diagram
with nimbers on the physical panel.
Most of the logic elements in the circuit are 5 port valves with
either pilot inlets at each end or a spring at one end and pilot pressure
at the other.
The symbols used in this memo, for these elements are as follows:
• If an element is energised left, it means that the left pressure
is higher than right. In this case, the element lines are as follows:
Pressure Higher a_I :I j_. / ‘
3 1 c
• If an element is energised Right, the opposite applies:
/2 H 1 - ‘
/ Pressure higher
5 f(
The numbers refer to the manner in which the element parts are numbered.
The symbol used at point 2 on the above example is a restrictive
orifice. Point ) is open to atmosphere with no restriction. Point 5 is plugged.
Element No.) in the diagram is a four port valve, air operated,
spring loaded. With no air signal, it takes the configuration;
2*
When air energized, it reverses to:
fry- fl ( >ç si&)
103
-------
Elements ( end ) are two port valves, spring loaded. They are
normally closed, that is, no flow passes through them unless pressure is
applied to thair pilot sides.
The oscillator consists of four elements and a regulator which
cause pulses of air to be applied to the pilot of element 9 when
a signal is applied to the oscillator. Details of this circuit are not
illustrated here.
The figure illustrates the configuration of the valves when
operating air at 80 pslg has been applied to all points marked “ “, but
the switch is in the “off” position.
At this point there is no pressure on either pilot side of element
12. The position shown was maintained from the previous operation by a “detent”
feature which prevents movement of the element when pressure is off. The
element is normally in this position except for the brief period during which
tank pressure is building before the outlet valve has opened.
The sequence of operations is listed in Table II and described briefly
below.
Pressure is supplied up to the switch from element 12 through element
7, but is blocked to the pilot of 10 and inlet valve by the switch. Thus the
inlet is closed while the switch is off.
When switch is turned on, pressure is applied to to open the inlet
valve and to reverse 10 to pressurised the delay accumulator and
start accumulator exhausting through an orifice. The exhausting of R-2
supplies the main time delay for solids to drain from the cyclone to the
transfer tank. This time is expected to be of the order of 10-15 minutes.
Because of the pressure applied through to the pilot of 2, reverses.
At this point there is still no pressure on either side of but it remains
held in position by a “detent” spring. The pilot of is pressurised by the
air pressure to valve A, and It reverses. Element 5 remains held by “detent”
with no pressure on either side. —
When the pressure in za accumulator drops below that supplied by
“Delay-2” regulator, element reverses. This starts R-l exhausting and
pressurises the pilot of l4which reverses. Pressure is removed from pilot
which reverses to close the inlet valve.
Pressure is released from # pilot which reverses in readiness to supply
pressure to the left pilot of 5.
When R-l pressure falls below that of Delay 1 regulator, reverses.
This delay was provided to give inlet valve about 5. .seconds to close fully.
Pressure now is supplied through 2to the left pilot of 12 which reverses.
Pressure from 12 then goes through to reverse 5. Pressure from also
reverses to pi ”x” pressure on the pilot of 7. The reversal of 5 applies
pressure from 11 to the pilot of 6 which opens £ admit nitrogen to The transfer
tank aerator. Element supplies this signal to operate its aerator so long
as the second system is not operating the other tank aerator. If the other
system is operating, the signal to open valve6is delayed until transfer
from the other system is completed. Likewise, a signal from this system
prevents the second from starting if the first is aerating. The purpose of
this interlock is to prevent simultaneous high nitrogen demand from both
transfer units.
104
-------
While the transfer tank is building pressure, the outlet valve
remains closed until tank pressure exceeds “f pressure, say 5 psig. When
this happens, element 7 reverses to apply a si al to the oscillator, to
element 8, and to the ght side of 12. Element removes the si al
from right of 5 and from 10. The oscillator starts pulsing N 2 through
the knife via 9. Element 8 reverses to open the outlet valve. Solids
should now start transferring from the transfer tank to the elutriator. The
element 12 reverses to change the si a1 to which reverses to apply
“w” pressure to the right pilot of .
When transfer of solids is completecj, tank pressure falls quickly.
When it drops below pressure tt fl element Lreverses to restore conditions
to the starting point. Signal is removed from the oscillator, from 8, and from
the right side of 12. The oscillator stops pulsing the knife through 9 and
element 8 reverses to close the outlet valve. Pressure is applied th ugh
7 to5 which reverses to stop the N 2 flow to the aerator, and signal is
pplied to 10 and to 3 which reverses to open the inlet valve. Reversal
of 10 starts the long time delay exhausting of R-2 and repressurises R-l
to verse element 2. The transfer tank is now again receiving solIds from
the cyclone.
105
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Table B-i
Composition of Element Numbering System
Number on Fig.l Number on Panel Catalogue No
1 17 PLV 52/16F/B
2 18 PLV 52/16F/B
3 - (on front)
9 PLV 52/15H7B
5 11 PLV 52uSD/B
6 - (in pit)
7 13 PLV 52/16F/B
8 15 PLV 52/15H/B
9 - (in pit)
10 16 PLV 52/15H/B
11 mt PLV 52/1511/B
12 7 PLV 52/16D/B
13 12 PLV 52/16F/B
i behind panel PLV 52/15H/B
lcY6
-------
TableB _ 2
Se uence o CofltrOllel’. ( perat ions
Supply On Turn on Switth R.2 < Delay 2 P H-i P < Delay 1 P Tank R > )C Powder Empties Repeats
8 mm 10 sec Tank P < W Cycle
Switch ORF (1) ON ON ON ON ON
Valve 1 L L (1) * H () ) L L L
2 H (3) * L - L (1) * H H (3) L
3 L (2) * H (3) L L L (2) * H
14 (7) * L short (14) H - short H R * L ahort delay
delay delay
5 H H H 3)*L L (2) R
6 H H H (4 )*L L (3) *R
7 H R H H (i) L (1) H
8 H H H H (2-) L (2) H
9 H H H H (2) Pulsing (2) * B
10 L (2) B H (3) L L (2) * B
11 H - H H P1 B B
12 H R R (2) L (2) H H
13 L L L ) H ) L L
114 L L (2) * H (5) * L L L
P1-i P 0 80 80 -Exhausting> —> 0 0 80
R—2 P 80 Exhausting > < Delay 2 —> 80 80 Exhausting > Delay 2 P
W iPSI 1 1 1 1
X 5P5 1 5 5 5 5
inlet Closed - Open * Close Closed Closed * Open
Outlet Closed Closed Closed Closed Open Closed
Aerator OfT Off Off On On * Off
Knife Off Off Off Off On Off
Tank P 0 0 0 Rising < X Bready 0
-------
Inlet Valvs
Open Close
To Other
Panet
mt.
Aerator
on Tank Knife Outlet Valve
4— F;
Close
N 2
0
Of f
On—Off switch
FIG. 8-2 CIRCUIT OF CYCLONE FINES PUMP CONTROLLER
-------
APPENDIX C
Inspection of CAFE Pilot Plant After Run 5
lOgr
-------
APP DIX C
Examination of the CAFB Gasifier and Auxiliary Equipment
Gasifier and Regenerator Unit
Gasifier Concrete
The walls of the gasifier were generally blackened with carbon and
in the upper regions around the junction between flue lid and the walls
the carbon was up to *8 thick. There were patches immediately above the
cyclone inlet ducts where the carbon had burnt away suggesting that an
air leak has occurred after the shut down. In the lower areas, the wall
was glazed with a hard thin tenacious carbon deposit.
The cracks in the concrete hot faces which have been present from
the first firing of the unit showed their customary fine black deposit
of carbon about 1” wide in the upper areas of the wall. The lower areas
were clear, because of the splashing action of the bed material. There was
no sign of any deterioration in the concrete from this test run.
There were areas at the junctions between the distributor and the
walls where a deposit of fine material had agglomerated together to form
a small coving between the distributor and the wall, most likely caused
by an area of poor fluidisation due to a blocked or partially blocked
distributor nozzle.
Gasifier Penetrations
The thermocouples, fuel injectors and pressure tappings were in good
order throughout although the left hand fuel injector had been replaced
during the run.
The thermocouple in the lid had a considerable growth of carbon around
its protri ing end whilst those in the bed area were generally clean, apart
from the one at the lowest point in the bed close to the fines return pipe.
There was some agglomeration of lime and carbon bridging between this
thermocouple, the distributor and the gasifier wall. This condition probably
arose from poor local fluidisation due to an obstructed nozzle in the
distributor together with the introduction of the fines from the return
system into this poorly fluidised zone. This thermocouple did show a
consistently low reading during the latter part of’ the run which was probably
caused by this local static material.
110
-------
The fuel injectors were layered with carbon ontheir protruding sections
and the i.njector at the right hand side bad a hollow cap of carbon and lime
enclosing its end.
The two air injection tubes which pass through the lid to direct air
into the cycloxa inlets were both heavily scaled and burnt away at their
protruding tips.
The stainless steel tubes for the stone feed and fines return were
both in good order.
Cyclones
The cyclone bodiea and their inlet sections which had both been lined
with type )lO stainless steel to improve their surface finish and hence
performance were heavily obstructed with a mixture of carbon and lime.
The inlet ducts are illustrated in figure C-l and figure C-2 which
show the black deposits around the inlets which gradually become lighter
further into the cyclone.
The white area on the gasifier wall immediately above the cyclone
inlets can only be explained by some areas of carbon burning out after the
unit was shut down possibly due to some local movements which may have broken
any seals existhg between the underside of the lid and the gasifier wall.
Figure C-l also shows the white irregular deposits on the cyclone
outlet tube which can be seen hanging down inside the cyclone body.
Figure C- shows a view into the right hand cyclone after removal of
the out. let tube shown in figure C-4. The cyclone upper body was obstructed
around its total circumference leaving the hole in the centre formed by the
outlet tube. The material was laid down in a very irregular manner consisting
of layers and folds of fine hard material with a tortuous gas path amongst
this deposit. The deposit becomes less pronounced towards the bottom of the
cyclone compared to the upper section, but still filled a considerable volume
of the lower section. The stainless steel liner was badly scaled and distorted,
in some areas it was completely burnt away due to the high surface temperatures
when carbon was burnt off at the various burn out operations during the run.
In the upper sections of the cyclone the liner was soft and eaaily broken
into flakes but nearer the bottom of the cone the lining was generally
stronger and some quite large pieces of steel were removed intact. The steel
lining tube which sealed off the cyclone drain passage to the gasifier to
regenerator transfer line was heavily scaled with a small area burnt away
on the top retaining collar.
111
-------
The deposit at the intersection of the rectangular inlet with the
cylindrical cyclone body was layered with white fine material separated by
thinner blacker layers, suggesting that the thicker white layers are laid
down during some of the oombuzting periods. This is confirmed by the hatch
unit tests which showed that combustingconditions give rise to the most
severe material deposition.
The silicon carbide outlet tube of the right hand cyclone shown
in figure C- was removed with very little material adhering to the outer
surface. Inside the tube there was an overall thin layer of’ carbon arid lime
about 1/16” thick deposited around the bore and in some areas near the
bottom of the tube there were thicker irregular deposits up to 5/8” thick.
These thicker layers could be removed fairly easily but the thinner layers
were very firmly bonded to the tube.
The left hand cyclone was also severely blocked in the inlet section
and in the body of the cyclone although it was not so severe as the right
hand cyclone (figure c-5). The cyclone outlet tube was coated on its outside
with a thick black flakeydeposit (figure C- 6 )but the deposits did not completely
bridge across the gaspassage between the outlet tube and the cyclone wall.
The bore of the tube was coated with a layer of’ carbon and line about
1/16” thick deposited fairly evenly over the surface of’ the tube.
The stainless steel cyclone liner was severly scaled and locally
distorted or burnt away in many locations near the top of the cyclone.
Towards the lower end the lining was almost intact although still heavily
scaled.
The stainless steel tube sealing off the drain to the regenerator to
gasifier transfer line was heavily scaled and around the upper outside surface
there were crystalline deposits of carbon which could have come up the transfer
line from the gas if ler bed. The upper retaining collar around the tube had
burnt completely away.
Gas ifier Distributor
The distributor was generally in good conditiori apart from some
obstruction in the holes in the nozzles. The obstructions arose from two
sources - one was from lime particles which entered from the gasifier bed
and the other deposits have come from fine rust scale carried through from
the flue gas recycle line which was carrying a saturated gas from the water
scrubber.
The distributor nozzle design has three staggered holes in series with
each other) the first smaller one to provide the pressure drop needed in a fluid
bed distributor and the third larger one provides a low exit velocity to
minimise damage to the stone in the bed. The middle hole forms a plenum
between the inlet and outlet holes.
112
-------
Generally, the outer holes were obstructed more with lime particles
and the inner holes with deposits of rusty coloured material. There were
considerable problems encountered in removing the distributor because it
had become wedged with limestone and heavy fuel oil when one of the fuel
injectors failed during the run. Some of the mechanical force used to
remove the distributor may have dislodged some of the material found in
the nozzles but examination did show that 2 of the 192 outlets were obstructed
completely, 18 of the inlet holes were completely blocked, 109 partly
obstructed and the remaining 65 holes were clear.
The stainless steel material used for the nozzles was in excellent
condition and showed no sign of deterioration.
Gasifier Lid
The lid was heavily coated with carbon on its lower face and the
protruding thermocouple had a considerable deposit of carbon around it.
The refractory concrete was in good condition but the vermiculite and calcium
silicate insulating slab was cracked in a number of places.
Gasifier ed Material
The gasifier bed was slumped without suiphation at shut down and
figure C-7 shows the typical bed material after half the material had been
removed. There are two interesting features in this picture, one of them
being the random area of completely white material which proved to be an
area of fine particles free of carbon obviously formed after the bed was
slumped which must then have had some oxygen in-leakage to burn off the
carbon. This area of carbon free material existed in about the middle
third of the bed depth. Another interesting and typical feature is the
blackened area at the right hand side which corresponds to No.1 fuel injector
location. After further bed removal this fuel injector was found to be
encased with a hollow sphere about “ diameter of carbon and lime particles
which would have restricted the throw of the injector.
The material was generally free from any large agglomerates apart from
a few lumps of carbon and fine lime particles up to l ” across. There were
some areas where there bad been static zones in the f luid bed around the
perphery of the distributor particularly near the fines return pipe from
the elutriator shown in the right hand side of fig. C-8.
Regenerator
The regenerator material was free from agglomerates and the bore
of the regenerator was generally clean apart from a few small deposits
at the top outlet section and around the joint between the distributor
and the wall. The refractory did not show any deterioration.
113
-------
The distributor was in good condition with the nozzles clear and un-
damaged. The refractory lip around the distributor had cracked away during
the run and the repair that was made had withstood th i remainder of the run
without any deterioration.
The transfer passages to and from the regenerator were free of any
agglomerates and the refractory concrete around the transfer sections was in
good condition.
Ductwork and Burner
Gasifier Outlet and Burner
The bifurcated duct had a uniform deposit of carbon on its Inner
surface about 1/16” thick and at the junction between the two ducts there
were thicker deposits where the two gas streams converged. The thermocouple
in this area was heavily coated with a carbon and lime layer about “ thick.
The premix section upstream of the burner had a 1/16” carbon layer on
the stainless steel section and the 1/8” dia. stainless steel thermocouple
in the burner throat was burnt away.
The main burner was in good shape with a layer of carbon about 1/16”
thick deposited on the internal gas ductwork. The outlet ring of the burner
had a local build up shainly of lime on its outer face shown in fig C-9
which was deposited from the turbulence in the gas streams arising from the
stainless steel taffle plates located to shield the pilot flame from the main
combustion air.
The pilot burner whie h had been very successfully modified to provide
a forced gas and air premix system prior to the flame retention nozzle was
in good condition apart from some local lime deposits around the nozzle end
which protruded into the burner quarl in the boiler.
The stainless steel deflector plates in the boiler burner quarl were
heavily deposited with lime.
Regenerator Outlet
The outlet pipe from the regenerator top gas space was coated with long
thin light purple deposits of material laid along in the direction of flow
and projecting out from the wall of the duct (figure C-JO). In one area there
was one projection which extended out almost 1” from the wall but generally
they were “ or less and only about 1/8” thick at the furthest tip. They
were held fairly firmly to the wall and appeared to have passed through a molten
phase having a fairly smooth outer face. Further downstream the pipe was
coated with a much more uniform deposit about 1/8” thick which was composed of
fine particles rather than the liquid type of deposition In the hotter
upstream section of the pipe (Figure C-11 .
114
-------
Downstream of the dust extraction cyclone the pipe was clean apart
from a very thin white fine deposit.
Boiler and Stack
Boiler
The back end of the boiler had been cleaned periodically during the
test run and figure C-12 shows the condition after shut down. There was
some coarser material laying in the bottom of the boiler fire tube with
some finer material deposited around the lower sides of the fire tube mainly
on the left band side when viewed from the rear, Indicating that there
was some swirl in the flame at the burner.
The entries to the first pass of fire tubes were deposited with rings of
fine material although none of the tubes was totally blocked. Some of the
deposits were smooth and rounded whilst others were spikey, shown clearly
in figure C-13.
The return tube passes were generally clear apart from a few tubes which
had some deposits at their ends. Some material had been deposited out of
the gas stream and collected in the boiler space at the end of this second
tube pass.
The refractory on the boiler rear door showed some signs of pitting
but there has been a gradual deterioration not just associated with this
particular run.
Boiler Probe
The boiler probe acquired some local light brown lumpy deposits which
tended to be more concentrated around the root” of tube figure C-lk, in front of
the entrance to the first tube pass. There was no indication of any deterioration
of the tube which was cooled during the run to approximately 600°c.
Stack and Cyclone
There was a quantity of lime at the base of chimney which was otherwise clean
and the external cyclone and knock out pot was also clear of any obstruction.
Both these vessels were continually rapped during the run and this prevented any
significant build up of material.
115
-------
Flue Gas Scrubber and Recycle Line
Scrubber
There were considerable operational problems with the scrubber initially
due to the wet fine slurry which as discharged and the deposits which built up
at the scrubber entry. The problems were eased by running the scrubber with
some recycle so reducing the dust burden concentration and increasing the
velocity of the gas and hence efficiency of operation.
Examination afterwards showed that the knock out vessel had some hard
fine deposits on its wall opposite the gas entry when the material would first
impinge on the wall. The scrubber and its entry pipe was clear.
Recycle Line
The scrubber has cleaned up the solids content of the gas very well but
introduced some problems due to the cold saturated gas leaving the scrubber
which contained enough very fine particles to seize up the cycle line control
butterfly valve and the control valve to the burner air valve which was in a
static leg and probably contained a considerable quantity of condensation.
116
-------
:J — p . 1 1 -2b - -
;tf i ec ’ - a A1 1 ‘& - - . - .
zr -
-
sts
t : I \ ;zt
4FF- - ‘*at - : C
a
• _ _ ; ?- • - -
- -, .• — ‘•1J *
• f —.
I
Figure c-i R.II. cyclone Inlet
• ffure c-2 L.H. cyclone Inlet
117
-------
g re C-4 R.H. Cyclone Outlet Tube
118
Outlet Tube Removed
:f ’ : :*
-------
C-5 L.H. C cicne, Outlet Tube 1 emoved
Figure c-6 L.H. Cyclone Outlet Tube
-------
Figure C-7 Gasifier
Bed Half Removed
L
ure C- 8 Gasifier Bed
- - __
—
120
-------
• . •• • 1
121
-------
Figure C -il Regener&tor Outlet Pipe, Downstream End
122
ts_ S3ç
I
C-i2 Boiler Back End After Shut Down
-------
Figure C-13 Fire Tube First
Pass Inlets
Boiler Probe
lI_
Figure C-li
123
-------
Table D-1
D-2
D-.3
D-k
D-5
D-6
D-7
D-8
D-9
D-lO
D-ll
D- 12
D-l3
D-lk
D-15
Figure D—l
APPENDDC 1)
Data for Run 5
Pa
125
130
135
139
143
150
157
164
172
173
174
175
176
177
178
179
182
Denbighahire Stone CAFE Gas compositions
Denbighshire Stone CAFB Pilot Plant Performance
BCR 1691 Stone CAFE Gas Compositions
5CR 1691 Stone CAFE Pilot Plant Performance
Temperature. Fuel and Stone Rates
Gas Rates
CAFE Pressures
Computed Solids Circulation
Composition of CAPE Solids (Sulphur)
Composition of CAPE Solids (Carbon)
Composition of CAFE Solids ( Egnition at 900°C)
Lime Metals Content
Calcium Oxide and Silica Contents of Bed
Sieve Analyses o± Gasifier Bed
Sieve Analyses of Regenerator Bed
Operation Log - Gasifier Temperature, Bed Depth,
Lime Replacement Ratio and SRE.
D-2 Operation Log - Regenerator Temperature, Gas
Velocity, Selectivity and Gas
802 Concentration
124
-------
1 of 5
Table D-l
Run 5 Denbigh hire St one
CPFR GAS COMPOSITIONS
02
FLUF GAS
C02
S02
02
REGENERATOR
C02
GAS
SOP
GASIFIER
VOL 02
INLET C-AS
VOL T C02
DAY.HOUP
PPM
ANAL
CALC
ANAL
CALC
2.013(1
2.11
14.4
209
(1.
0.5
2.7
16.5
16.3
2.7
3.7
2.11230
3.0
14.4
227
0.
11.6
3.6
1.5
16.3
2.6
3.7
2.0330
3.3
13.9
327
0.
0.6
4.9
15.0
14.9
3.4
4.5
2.0430
2.7
13.0
254
0.
11.5
3.9
15.0
15.1
3.3
2.0530
2.11
1.3.5
54
2.
11.9
4.9
15.0
15.2
3.2
4.3
2.9630
3.5
12.7
97
(1,
0.5
4.3
15.0
15.6
3.1
3.9
2.0730
2.5
13.5
427
0.
11.7
3.6
1S.
15.4
3.11
4.1
2.0930
•1
14.
754
0.
11.6
2.5
15.5
16.0
2.9
3.2
2.1030
3.0
14.4
745
(1.
0.6
3.6
16.0
14.9
2.4
4.9
2.11311
0.11
15.6
545
2.
1.7
2.7
14.0
1.1.4
3.6
7 .
2.1230
1.5
15.3
512
9.40
11.2
1.6
14.0
13.6
3.5
5.2
2.1330
3.11
14.4
563
0.
0.6
4.1
14.5
14.5
3.3
5.2
2.1439
7.9
15.0
4911
2.
1.3
5.2
1•0
14.2
3.5
5.6
2.1530
2.11
14.7
472
.
1.3
I 6
1 .5
14.
3.3
5.1
2.1630
2.1
‘ 1°
0.
1.1
4.6
14.5
l4
3.1
5.11
.1730
2.4
14.4
345
0.
1.0
4.3
14.5
14.6
3.1.
4.9
2.1 30
2.2
14.7
363
0.
11.11
5.2
14.0
14.6
2.9
5.11
2.1930
2.0
14.7
336
0.
11.2
5.2
14.
14.6
2.9
5.11
2.2030
9.5
14.1
236
0.
0.6
2.1
15.0
15.
2.2
4.3
p.2130
2.0
14.7
336
ç1•
1.5
3.9
14.0
14.4
3.3
5.1
2.2230
2.3
1 .b
412
0.
1.7
4.6
1.4.5
14.9
3.11
4.7
9.2330
2.4
1 .4
412
0.
1.0
4.6
15.0
15.3
2.2
4.4
3.003(1
2.11
14.4
41°
11.
1.P
3.2
15.0
15.2
2.2
4.3
3.0130
11 ,
3.023p
2.6
14.4
41
0.
1.11
5.4
15.0
15.6
2.7
4.2
3.0330
‘ .6
14.1
4110
2.
1.3
4.3
15.0
15.7
2.6
4.1
3.04311
2.3
14.’
4011
0,
1.9
4.1
15.5
15.6
2.6
4.7
3.0530
2.3
14.1
409
0.05
1.11
3.6
15.5
1.5.7
2.5
4.11
3 . 2632
2.1
14.7
3111
0.05
1.1
3.6
15.5
15.9
2.5
4.11
1.0730
7.1
14.4
4011
0.
1.4
.1
15.5
15.9
2.5
3.9
3.0113 ?
9.7
14.7
312
0 ,
1.3
4.1
16.0
15.2
1 .5
4.’?
3.0932
2.0
14.4
299
0.
1.7
4.3
16.2
16.2
2.4
3.11
3.1030
1.
1h.
29(1
0.
1.1
3.11
15.5
15.1
2.11
4.4
3.1130
1.
14.4
145
0 ,
0.9
2.1
15.2
14.9
2.0
4,9
1.17311
2.11
14.4
154
11 ,
11.2
3.2
15.0
15.1
7.9
‘ .5
2.1330
2.11
14.7
31°
1.00
11.2
11.9
14.5
14.2
2.9
4.11
3.1430
2.5
13.
436
2.
1.7
4.6
14.5
14.7
3.1
4.7
3.15311
1.5
14.4
4011
11.
0.11
3.2
15.5
14.3
3.1
4,9
3.1.6311
2.11
1.4.4
367
0.
0.4
7.9
1.5.5
1.4.7
2.9
4.9
3.1730
2.11
14.4
47?
0.
1.0
3.6
15.7
14.5
3.11
4.9
125
-------
2 of’ 5
Table D-l
Run 5 Denbighsbire Stone
CAFP GAS COMPOSITIONS
FLUE GAS
02 C02 S02
DAY.HOUR 2 2 PPM
REGENERATOR GAS GASIFIEP INLET GAS
02 COP S02 VOL 2 02 VOL 2 COP
2 2 2 ANAL CALC ANAL CALC
3 • 1830
3 • 19 30
3.2 030
3 • 21 30
3.2230
3.2330
4.0030
4.0130
4. p230
P.R 13.5 472 0.
2.8 13.5 563 0.
2.2 14.4 618 0.
2.1 14.7 600 91.
1.7 14. 581 91.
2. 1 601 91.
2.1 0910 91.
2.3 14.4 418 0.
2.6 14.4 410 91.
4.91330
4 • 0430
20 • 2830
20 .21 30
20.22391
20.2330
21 .0030
2 1 • 0 1 30
21 .0230
21.0330
21 .91430
21 .0530
21.91630
21 .0730
21.0830
21 .09 30
21 • 1030
21 • 11 30
21.1230
21 • 1330
21 • 1430
21 • 1 530
21.1630
21 .1730
21.1030
21 • 1930
21 .2030
21.2130
21 .230
21.2330
22.0030
1.6
1.6
0.8
( 7 1 .6
1.6
0.2
1.1
1.91
1.3
1.9
1.3
1.3
2.0
2.6
1.1
1.5
1.3
0.8
1 .9
2.6
1.8
P. O
3.7
1 .9
1 .9
2.3
4.7
1.3
3.3
1.7
6.7
4.3
4.6
4.9
2.0
3.5
4.5
3.1
4.5
4.1 15.5 15.1
4.1 15.5 15.1
1.1 15.5 14.8
1.1 15.2 14.9
4.3 15.2 14.7
1.4 15.1 14.0
3.9 15.2 14.8
4.6 15.0 1.5.5
5.4 15.0 15.6
4.3 15.0 15.7
4.1 15.0 15.6
2.5 18.0 17.8
1.6 17.5 17.1
3.2 17.7 17.0
3.8 18.0 17.3
3.0 19.0 10.5
3.9 19.0 10.5
2.9 19.3 18.5
3.2 10.7 18.1
3.9 10.8 18.1
4.6 18.8 18.1
3.9 19.0 10.2
3.2 19.0 1.1
4.3 19.0 18.1
3.9 19.0 10.0
3.6 19.0 10.0
4.1 19.0 1.0
4.3 19.0 17.9
4.3 19.0 10.0
4.3 19.0 17.9
3.6 19.2 17.9
4.9 19.2 17.9
3.9 19.5 10.0
5.1 18.0 17.2
5.2 20.5 10.0
1.9 17.0 )5.0
3.4 15.8 14.0
3.2 15.7 14.7
3.6 16.4 15.5
3.2 16.5 15.6
2.6 14.1 4910
2.3 )4.4 400
2.0 15.0 72
1.6 15.6 54
1.4 15.6. 7?
2.91 15.3 45
2.0 15.0 27
P. O 15.0 18
2.91 15.0 18
1.0 15.0 45
1.8 15.3 9
2.0 15.3 9
2.? 15.3 1910
2.3 15.0 109
2.2 15.0 109
2.0 15.0 109
2.9 ’ 14.7 181
1.5 15.3 36
1.2 15.3 54
2.91 14.4 136
1.5 15.3 127
I.? 15.3 81
1.91 15.0 36
1.8 )4.7 145
1.2 15.1 2091
1.8 15.3 200
2.91 14.7 163
2.0 14.7 210
1.7 15.0 245
1.7 15.0 154
2.91 15.0 209
71.
0.
91 • 20
(71 • 20
91 • 20
0 • 20
0.20
91 • 2 ( 7 1
91 • 40
91 • P R
171 • 20
91 • 20
0 • 20
91 • 391
0.20
91 • 20
0.20
0.10
0.10
0.10
0.10
0.10
0.10
1 7 1.10
0.10
(71.10
0 • 20
0 • 20
(71 • 2 ( 71
0.
3.91 4.4
3.91 4.4
2.9 4.8
2.9 4.0
2.8 4.7
2.9
2.9
7.8 4.7
2.7 4.2
2.6 4.1
2.6 4.1
1.6 2.5
2.2 3.2
2.1 3.2
).0 3.0
1.3 2.91
1.2 2.0
1.1 1.9
1.4 2.2
1.3 2.3
1.3 2.3
1.3 2.3
1.3 2.3
1.3 2.3
1.3 2.3
1.2 2.3
1.3 2.4
1.3 2.4
1.2 2.2
1.2 2.4
1.2 2.4
1.2 2.3
1.2 2.3
1.8 2.9
0.2 2.4
2.6 4.91
3.3 4.7
3.3 4.0
2.9 4.3
2.9 4.3
126
-------
of 5
Table D—l
Run 5 Denbighahire Stone
CAFA GAS COMPOSITIONS
FLUE GAS
0 C02
DAY.UOUR S S
RNEFATO8
OP COO
GAS
S02
G SIFIEF
5 02
INLET
VOL
GAS
5 COP
S
S
S
CALC
ANAL
CALC
502
VOL
PPM
ANAL
22.0130
2.1
15.0
100
2.50
2.6
.1
16.7
19.6
2.2
4.3
22.2230
2.2
14.7
91
0.30
2.6
3.
, .o
15.6
p.o
4.2
2°.fl332
2.0
15.1
280
0.32
5.7
.3
2.2
.3
22.0 /i30
2.1
15.8
145
0.20
4.5
3.6
16.2
15.6
2.2
4.3
22.2530
.2
15.0
145
0.20
2.6
4.3
l6.
15.6
2.0
4.3
22.0438
2.6
l5 .
127
2.22
3.1
4.6
16.9
15.7
2.7
4.3
22.2730
2.5
14.7
136
0.20
2.9
3.4
17.0
16.2
2.5
4•0
22.0232
2.5
14.7
129
0.
5.2
/ .6
16.5
15.6
).2
4.3
22.2930
2.5
14.7
136
2.12
1.6
3.6
16.5
15.7
2.2
4.2
22.1032
2.0
15.2
163
2.10
1.7
2.9
16.5
15.6
2.9
4.2
22.1130
2.2
15.0
191
0.
2.6
3.6
16.5
15.6
2.8
4.2
22.l 30
2.5
14.4
121
2.
5.2
3.9
16.5
15.7
2.6
4.1
22.1330
2.0
14.4
309
0.10
3.1
I 9
16.5
19.6
2.7
4.)
22.1430
2.5
14.7
302
.12
4.3
4.3
16.5
15.7
2.7
4.2
22.1530
2.5
14.7
281
2.12
3.3
4.3
16.5
15.7
2.4
i .2
22.1630
2.0
14.7
190
0.10
3.3
3.2
16.5
15.6
2.4
4.2
22.1730
2.5
)4.4
210
0.10
3.8
4.3
16.5
15.7
2.6
4.1
22.1030
.0
14.5
227
0.10
3.3
2.9
16.5
15.6
2.6
4.)
23.1730
2.0
14.7
31%
2.30
1.2
1.6
16.1
1.2
3.1
5.4
23.1032
7.5
14.7
154
2.20
2.0
2.9
16.0
14.4
3.1
5.2
23.1932
2.5
14.7
11%
0.12
2.2
3.9
16.0
14.4
3.1
5.2
73.2232
2.5
14.7
0
0.20
1.1
I.)
16.0
14.6
1.1
5.1
23.2132
2.5
15.2
112
0.50
2.2
2.5
16.0
14. /i
2.2
5.4
23.2230
2.3
14.7
112
0.40
2.0
3.2
1.0
14.4
2.1
5.1
23.2330
9.8
14.7
112
16.0
14.6
2.9
5.1
24.0032
2.6
15.
11
0.10
).5
4.3
)4.0
l5.
2.9
4.9
24.2130
2 .6
14.4
11
9.10
1.7
3.4
16.5
1S.
9.9
4.7
24.0930
2.3
14.4
10
0.12
1.9
3.2
16.5
15.1
2.2
4.5
24.0330
2.5
14.4
15
0.
2.5
3.9
16.5
1.5.2
2.0
4.5
?4.0430
15.0
15
2
2.5
3.6
16.5
15.0
2.2
4.7
°4.0532
2.2
15.0
9
0.
3.5
3.1
16.5
15.4
2.3
4.4
24.0632
2.2
15.2
9
..
2.2
3.4
16.5
15./i
2.2
4.4
24.2732
°.0
15.2
9
2.
3.3
4.6
16.7
15.0
7.6
L .7
24.2239
2.0
15.2
9
3.2
4.3
17.0
15.9
2.4
4.2
24.0932
9.3
15.2
9
0.
1.3
2.9
)7.0
)5.2
2.4
4.2
24.1232
4.0
14.7
9
0.20
1.4
2.9
17.0
16.3
2.4
4.2
74.1)32
1.2
14.7
9
0.22
2.5
3.9
)A.2
15.5
2.4
4.2
94.1238
2.2
14.7
9
0.28
2.2
3.6
17.0
15.1
2.4
4.5
24.1130
2.1
14.7
9
0.12
3.1
4.3
16.1
15.2
9.2
4.6
24.)430
2.1.
14.7
54
0.
2.6
3.9
16.0
14.9
2.9
4.7
127
-------
l ç f 5
Table D-l
Run 5 1)erthiglshire Stone
CAFB GAS COMPOSITIONS
FLUE GAS
02 C02 SOP
DAY.HOUR % PPM
REGENERA T0R GAS
02 COP SOP
1 1
C-ASI F! ER INLET GAS
VOL 1 02 VOL 1 C02
ANAL CALC ANAL CALC
24.1530
2.0
15.0
54
0.
2.8
3.9
16.0
14.9
2.0
24.1630
1.0
15.8
36
0. 8
2.5
3.6
16.0
14.0
2.8
4.0
24.1730
2.0
15.0
110
8.20
2.5
3.6
16.0
15.1
2.0
4.6
24.1030
1.0
15.0
0
0.
2.6
3.9
16.1
15.1
3.8
4.6
24.1930
1.2
15.3
0
0.
2.5
3.6
16.0
14.6
3.9
4.9
24.2030
1.5
15.0
0
8.
2.0
3.2
16.0
14.7
3•9
4.8
74.2130
1.6
15.0
0
0.
1.7
2.9
16.0
14.8
3.8
4.0
24.2230
0.
2.2
3.9
16.0
14.2
3.8
24.2330
2.5
14.4
5a
0.
1.7
2.5
16.0
15.0
3.0
4.6
25.0038
2.1
14.7
110
0.
1.6
2.5
16.2
14.9
3.8
4.7
25.0130
2.0
15.0
01
0.
1.7
1.8
16.3
14.9
3.9
4.0
25.0230
2.0
14.7
150
0.
1.4
2.9
16.3
14.9
3.0
4.7
25.0330
2.3
14.4
100
0.
2.5
3.9
1.3
15.0
3.7
4.6
25.0430
2.3
14.4
72
0.
2.2
3.8
16.4
15.0
3.7
4.6
25.0530
2.4
14.4
63
0.
2.1
3.8
16.4
15.0
3.7
4.6
25.0630
2.4
14.4
100
0.
1.7
3.9
16.4
15.8
3.7
4.6
25.8730
3.0
13.0
76
0.
2.5
3.9
16.6
15.2
3.5
4.4
25.0830
2.0
14.7
136
0.
1.7
3.4
16.3
14.9
3.7
4.7
25.0930
2.1
14.4
63
0.
1.9
3.2
16.3
15.0
3.7
4.6
25.1030
3.1
13.5
0
0.
2.2
3.6
16.6
15.5
3.5
4.2
25.1130
2.0
13.0
10
0.
2.3
4.3
16.7
15.2
3.5
4.4
25.1230
.7
13.0
288
0.
2.2
3.9
16.7
15.0
3.5
.4
25.1330
2.1
14.4
100
0.
3.2
4.6
16.3
15.2
3.7
4.4
25.1430
2.9
1 ”.l
136
0.
2.3
3.1
16.5
15.4
3.5
4.3
25.1530
2.8
14.1
136
0.
2.5
2.9
16.4
15.4
3.6
4.3
25.1630
2.7
1 .I
172
0.
2.5
3.4
16.3
15.4
3.6
4.3
95.1730
2.1
14.4
136
0.
2.2
3.9
16.3
15.2
3.7
4.4
25.1830
2.6
13.0
136
0.
.3
3.9
16.2
15.1
3.8
4.4
25.1930
2.2
14.4
190
0.
2.3
3.9
16.2
15.2
3.8
4.4
25.2030
.1
14.4
163
8.
2.6
3.6
16.2
15.2
3.0
4.4
25.2130
2.1
14.4
145
0.
3.1
3.9
16.3
15.2
3.8
4.4
25.2230
2.0
1 ”.4
120
0.
2.4
3.4
16.4
15.1
3.7
4.4
25.2330
p.c
14.4
136
0.
2.7
3.2
16.1
14.7
3.9
4.8
26.0030
1.9
14.7
130
0.
3.3
4.3
15.9
14.6
4.1
4.9
26.0130
1.9
14.4
120
0.
3.0
3.6
16.0
14.9
4.0
4. ,
26.0230
2.0
14.1
122
0.
1.7
2.7
16.0
14.0
4.0
4.6
26.0330
9.2
14.1
140
0.
1.7
3.2
16.7
15.6
3.5
4.1
26.0430
1.0
14.4
163
0.
1.9
3.2
16.5
15.1
3.6
4.4
26.0530
2.0
14.4
1 7
0.
3.3
3.9
16.5
15.1
3.5
4.4
26.0630
9.0
14.1
15
0.
3.0
4.3
16.3
15.1
3.6
4.3
128
-------
5 of 5
Table D-1
Run 5 Denbighahire Stone
CAFO GAS COMPOSITIONS
FLtIE GAS
09 COP
DAY.ROnJR 7 7
RE GENERA TO S
02 C02
7 7
RASI El ER
VOL 7 02
ANAL CALC
15.0 14.
26.0730
26.11030
26 • 09311
26. 130
26. 1130
26. 1233
26. 1 330
26.14311
26.1532
2 6. 6311
26.1732
26.10311
1.6
1.7
2.0
2.0
2.11
2.0
2.0
2.11
2.0
2. 1 7
1.11
2.11
INLET GAS
VOL 7 C02
ANAL CALC
4.1 4.6
S02
PPM
145
1119
136
136
1 (111
1 36
136
1119
163
2011
21111
2° 7
14.4
14.1
14.4
14.4
1 /j.4
14.1
14.1
14.4
14.1
I 4.4
1 Ia. 4
14.1
Op S
702
7
2.1
3.2
3.4
3.9
3.2
4.3
3. A
3.2
3.2
3.4
3.6
0.
0.
0.
0.
0.
2.
0.
11.
0.
11 • 20
( 20
2.
2.11
P .R
2.3
2.5
3.7
4.3
2.5
.5
2.6
2. S
3.1
2.5
16.1
1.1
15.5
15.0
16.11
I 5.
I A • (1
1 A. 3
16.3
I 6.°
16.3
15.11
14.5
14.5
14.11
1.4.9
5.2
15.3
15.3
15.3
15.4
3.9
3.9
4. 1
L i. I
3.9
3.11
3.9
3.6
3.7
3.6
3.6
4.5
4.5
Li. 9
4.6
4.5
Li. L i
L . 3
4.3
4.3
4.1
129
-------
1 of 5
Table D-2
Run 5 Denblghshire ftone
CAFO PILOT PLANT PERFORMANCE
SULPHUR GAS 6-BED AIR/ CA/S REGENERATOR
REMOVAL RATE DEPTH FUEL CAS SULPHUR OUTPUT
DAY.HO1’R PCT FT/SEC U i. STOIC MDL TO CAO LB/HR 7 OF FF0
2.0130 05.0 4.90 27.1 23.? 2.47 22.9 2.33 23.0
2.9230 04.6 4.00 2R.6 23.4 2.25 30.2 3.29 33.7
2.0330 75.9 4.96 27.1 22.5 0.75 42.2 3.56 37.9
2.0430 02.0 5.05 26.4 23.1 9. 33.6 2.99 31.0
2.9530 9.1 5.01 23.1 23.2 0.19 41.0 4.16 44.4
2.9639 97.9 4.93 24.7 23. 0.19 36.4 3.20 34.1
2.9730 69.1 4.99 24.0 23.7 0.14 31.7 2.47 26.4
2.0939 5.97 24.7 22.3 9.27 23.0 1.04 19.4
2.3039 49.5 5.23 23.1 22.0 0.35 30.7 2.95 29.4
2.1130 65.9 5.09 27.1 20.4 0.43 24.4 2.37 23.6
2.3230 66.9 4.91 23.7 20.5 9.39 14.4 1.44 14.6
2.1330 61.7 4.91 23.7 21.0 0.46 34.5 3.56 36.4
2.1430 60.0 4.09 23.7 21.1 0.53 45.3 4.74 49.3
2.3530 60.6 4.74 27.1 20.0 0.55 39.7 4.12 42.0
9.1630 71.6 4.79 25.3 20.7 0.50 39.4 4.04 43.0
2.1730 76.5 4.64 26.0 20.7 0.56 36.5 3.70 30.2
2.1030 75.0 4.26 23.7 19.3 0.66 41j.R 2.17 43.4
2.1930 77.6 4.26 23.7 19.3 0.66 44.0 4.17 43.2
2.2030 7.7 4.29 25.0 20.2 0.59 10.0 1.66 17.2
2.2130 77.6 4.52 95.3 20.2 0.59 34.6 2.59 26.0
2.2230 71.6 4.37 22.9 19.3 0.59 40.7 4.30 42.5
2.2330 71.6 4.29 22.9 20.3 0.66 30.9 2.37 45.1
3.0030 71.6 4.29 24.4 23.0 0.76 32.7 3.37 36.7
3.0330 4.21 25.0 19.0 0.60
3.9230 71.6 4.22 25.0 20.0 0.65 45.0 5.09 52.1
3.0330 72.2 4.39 25.6 20.9 0.72 36.0 4.03 43.0
3.0430 72.0 4.21 24.7 20.5 0.64 36.6 4.00 42.2
3.0539 71.6 4.10 26.2 20.9 0.59 30.7 3.46 35.0
3.9639 74.6 4.20 24.7 20.2 0.52 30.9 3.47 35.6
3.0730 7.0 4.20 74.7 20.3 0.62 35.6 3.95 49.0
3.0030 70.0 4.23 24.7 20.2 0.62 35.3 3.92 40. 3
3.0939 00.2 4.20 26.9 21.3 0.67 37.6 4.99 42. ’
3.1030 73.4 4.33 25.3 20.0 0.50 32.6 3.3w 35.1
3.1130 76.5 4.50 24.7 20.9 0.43 37.9 1.02 10.
3.1230 75.9 4.59 25.3 21.2 9.23 27.0 2.05 29.9
3.1330 70.0 4.39 21.0 19.9 0.46 0.0 0.01 0.3
3.3430 9.1 .4R 21.5 20.4 0.54 40.6 4.16 43.2
3.1539 79.0 4.40 27.5 20.7 0.53 27.0 2.07 39.0
3.1630 75.3 4.44 26.7 20.9 0.40 32.0 3.47 37.3
3.3730 67.0 4.53 23.7 21.9 0.61 31.7 2.60 39.2
130
-------
2 of 5
Table D-2
Run 5 Denbighshire Stone
CAF PiLOT PLANT PERFORMANCE
SULPHUR
GAS
G-000
AIR/
CA/S
RECENERATOR
REMOVAL
RATE
DEPTH
FUEL
CAS
SULPHUR OUTPUT
DAY.HOi R
PCT
FT/SEC
IN.
.
STOIC
MOL
TO CAO
LP/HR Z
OF FED
3.1030
66.0
4.51
25.7
19.0
0.66
36.1
3.73
36.7
3.1930
59.2
/i.53
22.5
22.2
0.63
36.5
3.40
30.3
3.2030
50.2
4.45
22.5
10.3
0.57
9.0
1.03
9.6
3.2130
60.1
4.45
72.5
20.4
0.69
12.2
0.75
7.0
3.2230
60.5
4.51
22.5
20.4
0.43
37.6
3.02
39.7
3.7330
4.43
20.0
20.3
0.
11.6
1.27
13.2
4. 003
4.47
p2.5
21.3
0.
29.7
3.46
37.6
4.0130
71.6
4.16
2.5
19.0
0.
39.14
4.12
42.1
4.0730
71.6
4.01
22.5
20.0
0.
45.7
5. 4
52.6
4.0330
7.2
4.17
1.9
20.9
3.
36.7
4.05
43.4
4.0430
72.0
4.25
21.9
20.5
0.
36.5
4.04
47.6
20.2030
95.3
4 .zil
“6.6
22.1
2.70
23.2
2.55
74.6
20.2130
96.6
4.36
26.7
20.0
1.53
14.4
1.65
15.5
2 0.2230
95.5
i.30
29.3
20.0
1.71
29.1
3.35
3 . 4
23.7330
97.!
4.48
26.7
‘0.6
1.71
33.2
3.06
34.5
21.0030
90.2
4.30
99.3
20.0
1.62
30.7
3.75
33.5
21.0130
90.9
4.29
27.3
22.1
1.60
32.0
3.66
314.2
21.0230
92.2
4.36
30.3
20.9
1.57
24.7
2.75
24.3
21.2330
97.1
4.27
30.0
19.7
1.16
26.2
3.14
7.1
21.3 30
99.4
4.25
29.3
20.2
1.34
33.4
3.70
33.1
21.0530
99./
4.28
27.5
20.0
1.34
42.3
4.47
39.2
21.0630
93.6
4.72
29.3
20.1
1.10
33.3
3.73
32.9
21.0730
92.9
4.09
29.3
19.4
1.26
22.2
3.05
6.9
21.0033
92.9
4.09
27.5
19.0
1.28
40.0
.1l
35.4
21.0933
92.9
8.39
29.3
19.0
1.09
33.4
3.70
33.2
21.1030
07.9
4.09
30.0
19.4
1.21
30.0
3.30
29.1
21.1130
97.7
4.03
7.5
19.1
1.26
35.5
3.72
33.4
21.1230
9.5
4.33
30.0
10.9
1.29
42.4
4.06
35.2
21.1330
90.0
4. 6
27.5
19.2
1.43
34.0
3.92
34.5
21.1430
91.9
4.05
27.5
19.1
1.35
39.0
4.00
35.14
21.1530
9.0
1 .0t
30.7
10.9
1.45
33.3
3.29
20.7
p1.1630
97.6
4.02
30.3
10.0
1.81
55.2
4.75
81.4
21.1731’
93.4
4.33
30.3
19.0
1.45
30.3
3.72
32.5
21.103:
07 3
4.32
30.7
10.0
1.144
53•fl
.59
43.7
21. 1930
07.3
4.00
22.1
19.0
1.54
52.2
4.27
42.2
21.2030
09.1
4.06
20.7
10.8
1.36
35.0
3.65
34. A
21.2130
25.0
4.34
27.5
10.9
0.74
32.1
3.31
33.5
21.2233
08.3
4.02
29.3
10.4
0.98
32.4
3.22
31.9
21.”330
09.9
4.35
24.9
19.2
1.26
32.7
3.47
38.8
22.3333
06.4
4.08
29.3
19.5
1.54
32.1
3.05
33.7
131
-------
Table D-2
Run 5 ] nbighshire Stone
CAFO PILOT PLANT PERFORMANCE
3 01 5
SULPHUR
GAS
0-RED
AIR/
CA/S
REGENERATOR
REMOVAL
RATE
DEPTH
FUEL
% CAS
SULPHUR OUTPUT
DAY.HOUR
PCI
FT/SEC
IN.
%
STOIC
MOL
TO CAO
LP/HR
OF FED
22.0130
93.5
4.04
20.7
19.3
1.45
36.9
3.90
30.7
P2.0230
94.6
4.06
30.0
19.7
1.00
30.2
3.20
32.4
22.0330
87.1
4.04
30.7
19.5
1.18
45.6
4.15
41.7
22.0430
90.5
4.04
28.7
19.1
0.94
30.2
3.64
35.7
22.0530
90.5
4.05
20.7
19.0
0.93
37.6
4.06
39.6
22.0630
91.7
3.90
20.7
19.6
1.04
41.4
.31
44.0
22.0732
93.9
4.01
32.7
20.0
1.04
32.5
3.32
34.!
22.0030
92.0
4.26
20.7
19.1
1.12
46.5
4.45
43.5
92.0930
90.9
4.22
20.7
19.6
1.15
30.1
3.26
32.6
22.1030
09.3
4.03
30.7
19.5
1.02
24.5
2.54
25.5
22.1130
00.1
4.02
29.4
19.6
0.93
31.7
3.31
33.4
22.1230
07.7
4.04
29.4
19.7
1.07
40.6
3.26
32.0
22.1330
79.0
4.02
30.0
19.0
1.06
44.3
4.55
44.7
22.1430
20.0
4.04
30.0
19.0
0.93
41.3
4.01
40.5
92.1530
01.2
4.04
30.0
19.9
0.96
39.0
4.01
40.5
22.1630
07.3
4.01
30.0
19.7
1.16
30.0
3.01
30.4
22.1710
05.2
4.05
30.0
19.0
0.96
40.0
4.06
40.9
20.1030
04.7
4.03
30.0
19.7
1.13
26.0
2.61
26.4
23.1730
70.0
4.01
96.4
17.0
2.37
13.5
1.49
15.7
93 . IS3fl
89.7
4.07
20.6
10.
2.37
24.9
2.74
20.0
23.1930
97.1
4.07
30.0
10.0
2.43
33.7
3.70
39.5
23.2030
100.0
4. 1
30.7
19.7
2.23
9.5
1.03
10.0
93.2130
92.3
3.04
31.4
17.0
2.21
22.0
p.30
23.9
23.2230
92.1
3.99
32.9
10.7
2.32
20.3
2.99
31.1
93.2330
92.1
4.01
32.9
18.9
2.33
24.0030
99.2
4.10
31.4
20.0
2.31
35.3
3.86
40.1
24.2130
99.2
4.20
32.9
19.0
2.29
20.9
3.18
32.6
94.0230
99.3
4.10
36.9
19.5
2.30
33.3
3.61
36.0
94.0338
90.9
4.11
34.3
19.4
2.17
34.3
3.57
36.0
24.0430
99.0
4.30
34.6
19.2
2.06
31.4
3.30
33.3
4.0530
99.4
4.09
34.3
19.5
2.14
20.5
9.09
29.0
24.0630
99.4
4.28
34.3
19.5
7.02
29.4
0.39
32.1
24.0730
99.4
3.93
34.3
10.1
2.04
41.6
4.37
43.3
24. 0 032
99.4
4.00
34.3
19.4
2.50
38.0
3.92
39.4
24.0930
99.4
4.30
35.2
19.9
2.49
23.0
2.54
25.5
24.3030
9.4
4.17
36.4
21.0
2.29
24.2
2.58
26.1
24.1130
99.4
4.00
34.0
19.7
2.32
34.6
3.67
37.7
74.1230
99.4
4.21
35.7
10.7
1.93
32.3
3.37
34.1
24.3332
99.4
4.15
34.2
19.2
0.67
30.5
4.20
41.1
24.1430
96.4
4.28
33.3
10.9
0.96
34.4
3.76
30.0
132
-------
Table D-2
Run 5 Denbi hsh1re Stone
CAF’B PILOT PLANT PERFORMANCE
k of’ 5
SULPHUR
GAS
C-OFTD
AIR/
CA/S
REGENERATOR
REMOVAL
RATE
DEPTH
FUEL
i GAS
SULPHUR
OUTPUT
DAY.HOUJR
PCT
FT/SEC
IN.
2
STOiC
MOL
TO GAO
LA/HR 2 OF’ FED
24.1530
96.4
4.00
3•7
18.8
1.22
34.0
3.74
37.6
24.1630
97.6
4.01
32.7
18.8
1.29
32.0
3.45
34.9
24.1730
9.3
4.20
34.3
19.7
1.51
3 1.7
3.40
34.5
24.1838
100.0
3.99
35.7
18.7
1.47
34.5
3.70
37.3
24 . 1930
100.0
4.04
35.7
12.6
1.68
31.4
3.35
34.0
24.2030
100.0
4.06
35.7
18.7
1.55
27.7
3.12
31.6
2 .Pl30
100.0
4.09
35.7
18.9
1.52
24.3
7P
27.7
24.2230
4.0.9
‘37.1
1.2
1.62
33.5
3.27
30.2
24.7330
96.3
4.1!
35.7
19.2
1.46
1.0
2.37
24.!
25.0030
92.1
4.06
35.7
19.0
1.62
20.2
2.55
25.9
25.0130
94.7
4.10
35.7
19.0
1.75
15.7
1.76
17.9
25.0230
90.0
4.08
35.7
19.0
1.62
23.9
2.73
27.2
25.0330
93.2
4.30
36.4
19.1
1.62
34.!
3.84
39.0
25.0430
95.1
4.10
3A.4
19.2
1.30
30.2
3.67
37.5
25.0530
95.7
4.09
35.7
19.2
1.46
32.0
3.66
37 . 4
25.0630
93.2
4.10
‘35.7
19.2
1.59
32.7
3.22
39.2
25.0730
94.6
4.12
35.7
19.6
1.37
34.2
3.86
39.
25.0530
90.9
4.09
35.7
39.7
1.60
28.6
3.31
34.1
25.0930
95.7
3. 0
36.4
18.2
1.47
27.4
3.13
32.1
25.1030
100.0
4.02
34.3
19.4
1.59
30.7
3.51
35.2
25.1130
92.7
3.92
34.3
18.5
1.27
36.5
4.16
42.7
25.1230
25.9
3.25
34.7
18.2
1.53
33.6
3.77
38.6
25.1330
93.2
4.01
36.4
19.1
1.47
41.3
4.44
45.5
25.1430
90.6
4.01
36.4
19.3
1.36
26.2
2.68
27.3
25.1538
90.6
4.00
36.4
19.4
1.21
25.3
2.76
22.2
75.1630
82.0
4.02
43.3
19.5
1.4!
29.9
3.29
33.9
25.3730
90.2
3.99
37.1
1.7
0.96
33.5
3.83
32.3
25.1038
90.4
4.09
36.4
19.1
1.06
33.2
3.81
32.5
25.1930
27.1
4.00
32.5
10.2
1.02
33.9
.71
37.3
25.2030
88.9
2.01
32.5
18.8
1.15
31.6
3.38
34.1
25.2130
90.1
4.12
37.7
19.2
0.99
35.5
3.72
37.3
25.2230
91.8
4.11
37.3
19.2
1.15
29.2
3.20
32.1
25.2330
90.5
3.9/i
37.3
12.0
1.16
28.0
3.00
30.3
26.0230
91.3
3.97
35.7
17.9
1.16
38.7
4.21
48.5
26.0130
93.2
4.06
35.7
12.2
1.23
32.2
3.43
34.5
26.2230
91.5
3.98
35.7
18.4
1.12
2.7
2.56
25.7
26.0330
90.3
4.00
35.7
19.3
1.03
27.2
3.16
31.9
26.2430
80.9
4.03
35.7
10.7
1.09
27.5
3.16
31.9
26.0530
91.4
4.03
35.7
18.8
1.03
35.7
3.92
39.5
26.0630
09.3
/i.O3
35.7
1.2
0.87
37.0
4.72
42.5
133
-------
5 of 5
Table D-2
Run 5 Denbighahire Stone
CI’FB PILOT PLANT PERFORMANCE
SULPHUR
CAS
C-RED
AIR/
CA/S
RECENJEPATOR
REMOVAL
RATE
DEPTH
FUEL
CAS
SULPHUR OUTPUT
D4Y.HOUR
PCT
FT/SEC
IN.
T
STOIC
MOL
TO CAO
LP/HR %
OF FED
26. 0730
90.1
4. 4
35.7
18.7
1.30
17.9
1.99
PP.1
26. 0030
92.4
3.95
35.7
10.3
i . i
20.9
3.21
32.3
26.0930
90.7
4.00
35.7
18.0
1.22
29.6
3.33
33.6
26.1030
90.0
3.81
30.5
17.
1.37
34.1
3.90
39.2
26.1130
93.2
3.00
37.7
17.5
1.41
30.4
3.27
33.0
26.1230
92.6
3.05
30.5
17.9
1.44
40.9
.33
43.6
26.1332
90.6
3.00
37.7
17.9
1.35
31.3
3.56
35.9
26.1430
92.6
4.22
30.5
20.5
1.44
20.3
3.10
32.1
26.1530
80.7
3.93
39.2
18.8
1.67
30.1
3.32
33.6
26.1630
06.4
3.93
39.2
18.7
1.30
28.7
.13
31.5
26.1730
06.4
3.04
39.2
10.3
1.35
31.4
3.20
33.2
26.1030
04.3
3.09
30.5
10.0
1.30
31.3
3.44
34. ’
134
-------
lofk
Table D .-3
Run 5 R 1691 Stone
CAFO 0 S COMPOSITIONS
FLUE GAS REGENERATOR GAS GASIFIER iNLET GAS
02 COP 502 02 CD? SO? VOL T 02 VOL COP
DAY.HOUR S PPM S S ANAL CALC ANAL CALC
.oc in 2.1 R01 2. 3.9 15.2 l1 . 7.9
1.9132 2.3 / 1 2. 1.1 1.6 15.0 15.5 2.° 1.2
z ,.423( 2.6 l .14 /-i70 0. 1 0 5.1 15.0 15 2.7 .2
1.2334 2.0 11.1 00 0. 1.3 / .3 15.0 15.7 2.0 1.1
1,.41 ,j31 2.3 11 . h O 4 0. 1.9 ‘-.1 15.0 75.4 7.0 1.7
1. ’5 ’34 2.5 13. ’ 130 1.50 1.2 10.2 11.2 3.1 5.1
1.0630 3.0 11.4 372 0. 7.0, 2.7 70.0 75.5 2.4
1.Ø73Ø 3.0 14.1 230 0. 1.9 5.2 10.4 15.9 2.9 1.0
1.0410 1.0 73.7 1 ,90 . 1.1 4.9 10.5 15 . 2.0 1.1
4.0910 3.5 13.5 2. 0.° 1.3 76.5 15.7 2.1 1.1
4.7030 3.0 7h.L 310 0• 1 • 3 5.2 15.5 75.1 2.9 1.7
4.1110 2.5 13.4 0. 0.4 0.0 11.0 14.5 3.0, 1.9
1.1230 2.0 12.4 510 2.02 ‘.3 .9 10.2 11.7 2.0 h.P
3.0 1/ .hi 451 0.2 15.0 71.5 3.1 5.2
1.1 410 2.5 12.2 59 7. 1.3 5.0 15.0 1 4.9 2.1 1.0
4.7532 2.5 13. 31,5 0. 1, 75.4 11.4 2.9 4.0
4.7032 1.0 12.2 151 .. 2 2.9 15.0 15.2 2.4 1.5
4.1710 1.0 12.7 930 (i 4.2 2.9 75.5 15. 2.9 4.2
4.1430 3.2 72.2 1 . 1.2 5.6 15.5 15.1 2.9 1.4
1.7932 2.0 12.4 720 2 1.1 5.2 15.5 15.2 2.0 1.5
1.2030 3.5 12.9 . 5.2 15.0 15.1 1.0
‘.2130 3.2 13. . 1.1 0. 15.5 15.0 2.9 1.7
4.2230 2.5 14.1 0• 1.0 5.6 10. 15.1 2.0 4.0
4.°3 ’ 3P 1.5 13.0 163 ‘‘. 0.5 5.1 15.0 11.5 3.3 5.1
5.0032 2.5 12.0 154 0. 0.0 6.0 17.0 15.1 2.1 1.4
5.0710 5.0 13.0 71,5 . 2.2 79.0 15.0 2.2 1.14
5.2230 3.0 1/4./4 127 . 7.7 7.0 7•0 7/4.9 .9 1.9
5.0332 3.0 11,.1 12 7. 0.6 1.1 10.7 11.9 2.9 1.9
5.0130 3.0 74./i 730 . (1.0, 1.6 16.0 11.0 9 . 0 1.9
5.2530 3.0 14.1 163 0. 7.1 3.9 16.2 15.1 2.2 1.7
5.0634 3.0 73.4 90 ( . 1.0 4.9 10.0 15.1 7.9 1.5
5.0732 2.5 73.2 ‘ 5 0. 4.0 1.0 10 . 0 15.2 9•2 4.4
5.2232 2.” 74.4 7’) 0. 1.0 1.9 76. 15.1 ‘ ) . /.7
5. ” 2 2.7 1 .1 07 7. 2.2 5.0 16.2 15.1 2.9 4.6
5.173 3.2 74.1 99 . 7 • 7 5.2 10.2 75.1 2.’) 4.7
5.1132 2.4 1 .L’ 1 ‘. 2.1 5.1 70.5 15. 9.0 .3
5.1932 9.5 74.7 797 2• 2.4 5.9 10.5 15.7 2.4 1.7
5.1430 2.2 11 .4 763 2. 1.3 5.0 76.5 75.2 2.6 1.5
5.1530 2.9 11.7 100 2. 4./, 5.6 37.0 15.3 2.0 1.5
5.1630 2.5 7 .7 72 2. 4.7 5.3 17.7 75 . 2.0 /4.2
135
-------
2 of
Table D-)
Run 5 B 1691 Stone
CAFR GAS COMPOSITIONS
FLUE GAS REGENERATOR GAS GASIFIER INLET GAS
02 C02 S02 02 COP S02 VOL % 02 VOL 2 C02
DAY.HOUR 2 2 PPM 2 % 2 ANAL CALC ANAL CALC
6.0239 ! 2.5 14.4 336 0. 1.7 3.9 15.0 14.1 3.4 5.4
6.0339! 1.5 15.0 245 0. 2.5 1.3 15.0 13.7 3.5 5.1
6.0430 1.5 15.9! 172 0. 1.7 3.2 15.5 13.9 3.5 5.4
6.0530 3.0 34.4 90 0. 2.0 4.6 16.0 14.5 2.9 5.9
6.0130 2.0 15.0 27 0. 1.7 5.6 16.9> 14.1 3.1 5.1
6.0739 ’ 1.5 15.6 54 0. 2.2 5.1 16.0 14.9 2.9 4.9
6.0039’ 2.9! 35.0 27 9’. 3.1 11.0 1 .3 2.9 5.?
8.7230 3.0 13.0 354 0.20 0.0 o.o 15.7 14.9 .o 4.6
8.2330 3.9’ 14.1 354 0.29 1 0.5 0.6 35.7 14. 3.0 4.0
9.0039’ 7.5 14.4 37 0 1. 93.4 0.9 > 15.3 15.1 4.0 4.6
9.0130 3.2 14.1 39’Q 9’. 0.0 1.6 15.0 15.3 3.1 4.5
9.0230 2.7 14.1 11>3 0. 0.6 1.1 35.0 I5.3 3.1 4.4
9.0339’ 3.0 14.1 330 9>. 0.9 91.0 11.91 5.5 3.5 4.1
9.0 39’ 3.0 14.4 172 0. 1.0 9.9 36.2 15.5 3. 4.4
9.0539’ 2.? pa . 272 0. 0.9 3.1 16.0 35.5 3.5 4.7
11.1430 3.5 13.5 354 9>. 0.0 2.5 16.0 14.7 4.4 2.0
11.1530 3.5 13.0 236 0. 1.0 1.1 16.0 15.5 4.91 4.’>
11.1630 4.91 33.5 727 0. 1.9> 3.2 16.5 15.6 4.0 4.3
11.3730 3.0 13.0 97 0. 0.6 2.1 16.0 15.6 4.0 4.4
31.1030 3.5 13.7 90 0. 0.7 4.3 16.5 15.4 3.0 4.2
11.1930 3.5 13.2 54 2. 91.6 3.2 17.0 15.9 3.6 3.0
11.2030 4.0 13.2 109 9>. 1.91 4.1> 17.0 11.3 3.5 3.1
11.2130 4.0 13.2 254 0. 1.0 3.6 16.5 16.3 3.3 3.7
11.2230 4.91 13.5 29>0 0. 0.0 4.6 16.5 35.7 4.91 4.?
13.2330 3.5 13.5 336 0. 1.91 4.9 36.5 35.5 3.9 4.2
32.0030 4.91 13.0 31>3 0. 3.5 5.1 16. 15.7 2.0 4.?
12.0230 7.0 10.7 200 0. 91.1. 3.9 16.5 16.9 2.0 3.1
P.0330 4.0 13.2 27? 0. 0.0 4.9 16.5 16.1 2.6 3.
12.0430 3.5 14.1 254 0. 1.3 4.9 16.5 15.9 2.0 4.1
12.0530 3.5 14.1 245 0. 1.1 3.4 16.5 15.9 2.0 4.1
12.0630 3.0 14.1 3991 0. 3.3 4.3 16.0 15.1 9.9 4.6
32.0730 4.01 33.9 290 0. 1.2 4.3 16.0 15.0 3.1 4.7
17.01030 4.01 33.7 272 0. 91.0 4.3 1.0 15.0 3.0 4.7
12.0930 4.2 13.5 236 0. 3.1 3.0 16.7 15.9 9•Q 4.3
32.3030 3.7 13.0 236 91. 1.3 3.6 j7.0 3.0 2.6 4.91
12.3230 3.5 13.0 254 0. 1.6 2.9 17.0 36.0 2.6 4.91
1 .I330 4.0 13.0 209 0. 1.3 4.1 37.0 36.3 2.4 4.0
12.1430 3.5 I3.R 209 9!. 3.4 3.2 17.0 15.9 2.4 2.0
12.1530 3.3 12.1 227 0. 1.6 4.2 17.0 16.3 2.4 3.9
12.1630 3.6 13.0 710 0. 1.9 3.2 37.0 16.2 2.4 3.0
136
-------
Table D-3
Run 5 BCR 1691 Stone
COFO GA COMPO IT1 ON
) of 1
02 ( 02 02
2OY .Hfl7 ’ 0
000FNESATOO C,AS
32 C02 S02
0
OASIFIFP INLET S
VOL 0 09 VOL 0 Cfl
ANAL (TALC ANAL CALI
1 0 .17Y
.1
17.0
10
.
1.9
i.A
17.2
10.9
2.4
3.0
1°.1 ’ 1”
7.7
1L.1
2 ”fl
( .
1.3
4..
17.2
10.3
2.4
3. ’
12.1912
7.0
12.0
2
3.
1.7
4.7
17.0
10.3
2.4
7.0
12.2079
4.11
J.2
227
0.
J .3
ii.3
J7.7
10.1
2 .z,
7•k
1”.2l2
3.’’
17.
277
0.
1.9
.A
16.7
15.3
2.7
4.3
19.2230
7.7
13.5
309
‘ .
1.0
z.3
16.7
15.2
0.7
L.A
12.93311
3.2
1 .l
379
.
1.9
3.0
15.7
14.1
7.2
5.7
1 .7 ’17
4.0
14.1
270
1’.
.2
2.9
15.7
14.°
7.2
5.3
1 .713
7.4
14.1
202
.
1.0
3.0
15.”
•1
5.3
11.11912
‘ .1
12.’
90
‘ ‘.
1.4
1.4
16.0
14.7
3.0
5.i?
11.02 fl
7 7
1 . A
262
‘ ‘.
l.A
7.o
15.0
14.0
3.1
5.0
11.0”19
Ji.2
11. ’
102
9.
2•Q
IA.?
1’ .7
2.9
4.°
1’.fl 3fl
7 ()
12.0
170
7.
1.’
L ..2
16.2
14.0
2.9
4.9
13.7630
2.9
12.5
I6k
‘.
1.5
A
jS.°
14.0
3.”
5.1
IA.”712
4.7
17.’-
151
‘.
3.f
o -.c’
14.0
5• ’ -,
16. - ’ 1’
14.4
7”
‘ ‘.
.2
1.1
91.5
l5.
1.3
/1.0
I A • ‘ (‘
I •
1 i • 1
145
P.
1 •
. ‘ • ‘
I A • C
1 Li • ‘)
3 • 2
L •
11.1 ‘2 ’1
7.7
14. /i
37 ’2
0.
1.
3.9
1 .5
1 L.
3.0
4•)
16.1170
2.1
14.1.
lA S
.
.0
‘.
15 . A
10.1’
3.5
/.4
1c -.123
3.4
14.1
437
1.
0./i
2.0
10.9
15.1
3.1
4.7
10.1272
3.4
14.1
‘.
‘ .0
L.A
14.5
35.2
3.3
1.4
1A.1 .30
I-i . - ’
12.4
397
‘ .
7.0
.9
16.
10.0
9.0
Z .
IA.1 /I ’
4.4
1 .0
230
1’•
0.6
4.7
16.4
10.2
0 2
7.4
I -.l7
2.’
! .“
3/S
-?•
.A
0.1
1A.
lA. ’
2.7
4.2
10.1 . 2
. ‘
I’-.”
2 ’
‘.
1.11
5.3
17.fl
JS.7
2.”
/ ‘)
1A.1 211
1.4
1/.4
0’
1.
1.0
1.3
17.7
15.0
2.0
10.7039
3 -i
2.4
27
1.
2.0
5.2
17.’’
I 5.4
2.0
4•(
10.0132
7.4
14.1
1.9
1.4
37.0
10.0
2.(•
‘.1
1A.0 39
.5
14.4
1R
-‘.
1.7
2.9
17.
IS.7
2.6
1.1
16.9110
2.0
17.4
1 ’
.
1.6
7.9
7.)
15.0
9.4
7.9
17.4’ -’
7•3
10.4
07
9.
1.1
2.6
17.9
1A. ’
2.0
7.’)
l7.’12
“.
1.1
L.A
17.0
90 ’)
2.4
“.1
lf.’ 7’
.
1.0
1.2
17.1’
°1.
2.4
“.1
17. ’ ’
17.
12.0
2.4
9.0
)7 . 4L7 0
‘.
1.1
L•2
17.’
10.4
2./.
r’ ( .
17.’5 ’
0.
1 .
1.2
17.”
19.4
9./i
2.5
17.2 12
‘.
1.0
7•9
17 . ’
12.?
9./ 1
-‘.
17.’7 2 ”
‘.
1.1
4.2
17.”
°1.
9•7
17. ’ 7
“.
17. ’
90.’’
2.2
0.1
17.’ )7
2.’’
13.4
107
‘.
3.7
L .1
17.”
10.”
2./i
2.P
137
-------
4 o ’ 4
Table D-3
Run 5 1691 Stone
Cc FP GAS COMPOSITiONS
FLUF GAS REGENE ’ATOR GAS GASIFIER
OP COP SOP 02 C02 502 VOL 02
DAY.140 1’R PP1 ANAL CALC
INLET GAS
VOL C02
ANAL CALC
17.l03f
0
14.1
227
0.
1.9
4.3
17.0
1 .0
2.4
3.9
17.1130
3.0
14.1
201
0.
1.6
5.1
17.0
J5.9
2.4
4.1
17.1930
3.9
14.1
272
‘.
1.7
4.9
17.5
16.5
2.2
3.5
17.1330
2 .
l4.
236
2.
1.6
4.6
17.9
16.4
2.4
3.7
7.1430
9.7
14.1
499
9.
2.6
4.6
17.2
16.3
2.4
3.6
17.1530
3.0
14.1
1 1
2.
1.6
4.3
17.9
16.4
.2
3.6
17.1630
2.0
1 .4
299
9.
2. /i
2.5
17.0
16.2
.4
3.
17.1730
.2
p4.4
390
..
1.6
3.
17.5
36.6
.3
3.5
17.I 30
P.
14.4
72
! .
3.1
4.9
17.5
16.5
2. 1
3.5
17.1930
9.5
227
?.
1.6
4.3
17.5
16.3
2.
3.6
17.2930
2.7
127
.
3.1
4.1
17.7
16.5
.1
1.5
136
-------
1 of 4
Table D-4
Run 5 B R 1691 Stone
COFR PILOT PL INT PF2FO ANCE
. !JLPH’’F 1 $ fl-PED J / PFrFNFPATO P
RF ’1OVr L RATE DEPTH FTJEL CA% LPHID O’JTP2T
flAY. OHE PCT FT/ FC IN. 7.cTOIC VCL TO CI’O LF./HP 7 OF FF2
0 . 7 22.5 7J.3 r , , 29.7 3.0 ,
L.913 71. .1A 92.5 19.2 . ‘ 39.1 1 .12 02.1
1 .f230 7l.c /i.I 92.5 20. 2. 05.7 5.1Z 52.c ’.
0.23 9 72.) 1l . 7 1.° 9( .9 0. 7A.7 / .75 03. / 1
4 .9012 72 . i ’.05 O .Q 9$.5 r. ‘ 3 .5 /1.00
/ 1.’5 ’ 3P 90.’ 90•7 19.0 9.79 fl.9 2 . 1
/2.0c 3p 7/.7 0.10 20.1 29.3 .(0 07,9 .7 2 .7
/.0731’ Q7 ( 0.21 ‘75.1 91’.0 ( .97 / 5.1) 5.22 53.7
4.2111 ” (7. 711.7 2’.c< 1. ’7 z .79 /49.2
iS 2.5 2r’.O f 52 10. u. 1’1’ 02.1
/.1’31 7 .’i o.fl ’ 0.5 V.7 (1. /15.’ Ij.Q5
/1.’! 9”.5 19.S . o°. 5.5 57.1’
/:.l , ’ A7.7 /J• 74 ) 71. 29.2 1.2/2 12•9 .92 20.7
0.1239 ( .i o./ ° 25.7 2fl. 1.71
1.1 /.)0 ( 0. 01 . 7 ‘1 • S I • 7 ’ • 0 /j.70 02
1’l.15 ’ ” 75 . ‘ .S 95.7 1.S 1.77 72.9 /j. 1 ’.7 00.5
11.1 (30 71.9 5” 05•7 2°.o 1.71 2.7 7.20 7/.3
/i.I’llP c 1 . ) / .u ’ 25.7 22.2 1.(’( 04 ,9 0.59
/2.1 32 25.7 21. 9.1: o7.r ‘ 11’
/1.1419 21 ./l 25.7 21 . 2.11 4 .1 0.77
0.51’ 911.7 21.7 9.11 /2.5 /i.79
Li. 5J 95.7 2l. 11 1.A2 51’.L 5.0 57.11
‘1. 5L 25.7 0’7•7 • 0 07.9 5.01 52.7
1’i .7 0.5! 27.7 7l. 1.77 li 1’ . 1’ 0.A0 49.9
5 .1 ’1’lfl <./2 ,7Q 27,7 9fl. 1 . ‘“‘ /19.9 5.59 55.0
5. 1 1’ 27.2 0.7 97.7 21./
S. 23:’ ° .7 0.15 27.7 91.2 1. 1’ 1/ .2 l. V.!
5.7129 94.11 /4.211 27.7 21.1 1.70 °.2 9.93 7,0
S.9 1 ’ 21’•7 0.’I 07.7 19.A 1.10 l’-’.l L.1 01.7
11•’ S- )fl c711• 4 / 1.0/4 ‘32,9 91.2 1.72 1(’ Q 3.11 ‘/1.1
.1’ 31’ o / 4./ IL 99.5 21.7 1.72 ‘37. ‘3.7/
-, ,“72() OA.A /4.0/2 ‘ l 21.A 1.7S IA.2 7.27
11.:’23 ’ 1’S.! / 1./I/I 77.7 p1.2 1.711 /11’.r / I . 25
5.92’ 7.7 / 1• 2 I1’.V 21.” 9 ,94 /17 2 L. 9 59./.
.1 30 91.0 49.’ _1.’ 1.k ’ 0(1.2 0.24 40. ’
‘-.117 ” 1’/J./4 /4.1’ i1’.’ 22.5 1.7’ 5. 1’ /1.i 09.1’
5.1’ 11i 1 •s / .r’ 72.0 19.7 1.10 01.5 0.50 07.9
11.102,7 ‘2.Q 4.21 29.2 72.5 (24.22 /15.3 0.75 50 4.5
.15’1’ 07,7 1 4 .9A 2 .0 21.9 1.2! 07.5 o . c5 49.7
11.1( 12 95.4 4.19 2 .9 211.2 “• 2 77. ’ 7 3
139
-------
2 of 4
Table D.,4
Run B 1691 Stone
C.4FR PILOT PLA JT PERFORMA JCF
MOL
SULPHUR
AS
2-REt
AIR/
C /S
R E 4FR6TOR
REMOVPL
RPTE
I)FPTH
FUEL
Z C6S
SULPHUR OUTPUT
DAY.HOUR
PCI
FT/SEC
I? $.
ZSTOIC
TO C O
LP HR
OF FF.D
6.2P3 ’l
17.1
4.32
2.1
22.6
7.91
33.)
3.16
32.3
6.2330
23.9
4.37
25.4
20.1
1.39
40.7
3.59
39.0
6.2437
2 .1
4.90
25.4
19.9
1.32
29.4
2.75
29.5
6.7510
9.2
4.25
26.1
27.7
1.3
39.2
3.27
35.4
6.0630
92.2
4.04
27.7
27.1
1.13
46.9
4.05
43 7
6.0730
96.6
4.02
25.7
20.2
1.10
o.s
•
6.2737
97.p
3.70
27.7
12.9
1.20
53.5
4.72
53.3
7.2230
74.9
4.64
21.2
22.0
1.53
2.73 7
75.4
4.54
24.7
21.6
1.56
5.1
( .59
6 .
9.2737
77.7
4.54
p5.3
22.)
2.00
9.2137
72.6
4.35
26.2
20.9
9.21
13.1
).64
17.2
9.0210
75.7
2.41
26.7
9 1. 1
l.Rl
13.6
1.77
1 .7
9.0337
91.7
.nS
9R.6
21.6
‘.A?
0.)
0.01
2.1
9.2432
72.3
4.49
25.2
21.7
1.12
24.2
1.4
16.°
9.7530
R ).5
2.47
26.7
71.3
0.91
29.7
4.27
46.9
11.1h1
74.3
4.57
9 1.7
p1.6
2.39
71.4
9 %
OR.9
11.1510
23.3
4.32
22.5
20.3
7.17
12.7
1.24
12.5
11.1632
73.6
4.37
27.9
21.5
2.96
29.2
3.55
37.3
11.1737
73.9
4.3
22.1
p1.4
7.96
1 .3
2.23
23.4
) ).)732
93.3
4.1%
10.1
91.4
1.02
30.6
4.29
51.5
11.1930
96.7
4.43
31.4
p2.2
1.24
29.7
3.66
30.5
11.223?
92.2
4.37
3).4
22.2
1.07
42.0
5.26
55.2
1l.° 137
71.2
4.39
31.4
22.3
1.21
33.)
4.25
4.2
11.2232
74.9
4.36
30.7
21.3
1.12
40.0
4.93
51.1
11.2337
75.7
4.12
32.7
2 ).)
1.25
43.2
2.27
52.i
19.0032
7 .3
4.42
32.7
2 1.4
1.20
52.3
6.72
62.9
19.223?
R ).9
4.11
30.2
22. !
1.’ 3
36.2
4.24
42.9
12.0330
79.9
4.36
31.4
7 !.!
0.14
S.7
5.07
52.7
12.2430
p2.4
4.35
30.2
20.3
7.61
45.3
5.93
52.2
12.053?
73.2
4.36
76.9
9) .fl
2.72
33.9
3.93
39.9
1 .7630
72.9
4.37
30.0
22.1
7.90
42.7
4.75
4 •
)2.0730
7 .6
2.30
39.0
20.4
2.66
39.3
5.12
53.7
1 .0237
79.9
4.99
32.0
20.9
2.95
39.2
4.76
51.2
12.3932
72.9
4.J9
31.2
p ).4
I.34
36.9
4.43
47.2
12.1030
21.1
4.12
31.
90,7
1.29
34.2
2.09
42.9
12.1732
R2.(
4. )2
99.3
90.7
1.15
99.3
3.40
35.6
12.1330
75.7
4.12
3).4
20.7
1.12
37.7
4.66
47.6
12.1432
75.9
4.)9
11.4
22.6
1.13
32.6
3.73
40.2
12.1510
74.2
4.05
32.9
70.5
1.27
22.3
4.92
51.2
12.1630
04.6
4.04
31.4
20.5
1.92
30.4
3.5)
36.5
140
-------
3 of k
Table D-4
Run 5 BCR 1691 Stone
CAFF4 PILOT PLANT PERFOR ANr,E
RE GEN E PA TOE
2 CA2 . 5 (LPHIIR OUTFIT
TO CAO LP/HE 7. 00 FED
142.3
32.0
3 / 4 • 14
36. L
32 I
97.3
3 / 4.6
35.1
1 / 4 • 2
41.7
• 2
17.2
1 40 •
34.2
5.23
3 • 08
/4.141
14. Li
3.72
14.0 1
‘4 • 1 4 0
/ 4. 07
. 27
• 149
44 • 10
‘4.53
/ 4.97
/ 4.23
SVLPH!JE
GAS
G-PED
Al R/
CA/S
REMOVAL
RATE
DEPTH
•FIJEL
04Y.HOIJE
PCT
FT/SEC
IN.
STOIC
MOL
12.1732
21./4
4.03
31.
.0.5
0.93
12.1230
2 .1
4.25
31.14
20.7
1.07
12.1932
05 .
14.2/4
30.0
20.7
1.15
19.2117
03.2
4.12
20.6
22.9
1.0 / 4
19.7132
p2.7
“.27
‘ i3
22.6
1.13
l .?°10
77.7
1.142
27.7
21’.!
).P / 4
12.213(
79.2
/4./ i /i
27.7
22.9
(14.95
13.1’fl32
21 .6
/4.140
p’4 .9
70 .9
I3. !32
0 1.7
4.30
27.4
20.5
I . ! !
13.2210
0 .3
14.11
27.6
20.7
0.914
11.fl33
P .1.3
/ 4./i’
27.6
20.5
14•93
13.VIi IO
07.9
/i. / 5
9/4.9
90.2
0.95
13 . 5 r
H9.9
/4.15
2/4.9
20.7
(14.77
13.0630
1 •2
4•147
97.7
7 (14 9
Ø•97
I 2i32
/47.2
14.26
9 .7
22.7
J .90
1/4.”032
67.P
/4.149
9L.Ø
2.0
1.56
11.1
1.11
1.0912
AQ.7
4.5C
A.2
21.5
1.53
74.6
1.146
27.9
21.5
1.36
11.7
1.59
16.1130
76.5
14.47
90.6
21.5
1.25
99.3
1.0/4
14.123
70.9
/ 4./i(.
39.2
21.7
1.69
90.3
3.33
16.1317
71.0
/4.5/4
304.2
22.6
1.77
32./i
/i.142
16.1 YI0
76.3
/4.147
3i .f,
93.1
1.70
7.3
1.714
I’ , .I/4lfl
75.7
4.22
12.9
2. .
1. 4
35.7
3.0/4
16.1730
7/4.5
14.31’
19.9
22.5
1.90
/42.1
14.73
16.103”
79.’
4 .149
31.4
27, / 4
9.12
1414.0
5.22
I .1 3 ’
95.7
/4.22
36.9
“2.!
2.12
‘32.9
14.32
1 . ”’l
92.1
/4.13
/4, P
21.7
.07
/45.4
5.30
1 / 4.2132
90.0
/4.19
99.9
1.96
13.6
1.32
1A.’ 3fl
90.0
/.1
30.5
72.0
1.90
36.9
3.26
92.7
/ 4.99
32.5
21.3
0.93
30.3
1.20
17.2230
90.1
/4.97
IA.9
71.7
9.92
1 / 4.
1.914
17. ” 130
‘.31
30.5
26.0
(44•09
141.1
1.6 / 4
17.2212
/ i.32
36.9
27.9
0•99
37.6
14.52
J7. 0730
14.32
36.9
96.14
1.20
17.2/410
/4.39
36.9
24.9
114.77
32.0
14.53
17.2532
/ .31i
36.9
77.5
1.20
35.2
6.99
17.0 / 43
14.3!
39.5
76.4
1.00
36.1
4.(143
l7 . 47 1
‘i.ll
37.7
27.9
1.11
30.2
4.36
1 7 .r 4 0- 1 42
4.314
37.7
26.6
(44•97
17. ’930
10.4
4.3t
30.5
21.6
2.41
37.5
14.10
52 • S
40 • 7
14/4 • 2
46.2
49 • 3
45.7
47 • 0
145.3
53.7
116. 1
147 • 0
52.6
145.2
11.4
37. 2
12 • 7
15.14
147.2
73.1
.4
51.
56 • 1’
• /4
50 • S
13.5
•0 • / 4
34.2
/40.”
/40 • 9
/47. 1
/4/i •2
76. ‘
1 • 7
/45.14
141
-------
4of 1
Table D-4
Ruu 5 B I 1691 Stone
CAFB PILOT PLANT PERFORMANCE
SULPHUR GAS G-RED AIR/ CA/S REGENERATOR
REMOVAL RATE DEPTH FUEL 2 CAS SULPHUR OUTPUT
OAY.HOUR PCT FT/SEC IN. 2 STOIC MOL TO CAO LP/HR 2 OF FED
17.1 3 R4.2 4.P7 3R.5 21..’? 1. 5 37.R 4.1.4 43. 1 .
17.1L3 Ri’I.S 4.1.4 35.7 2V .7 1.P2 .99 51.7
)7. !23 RI.! 4.12 35.7 PI. I . Ia .7 4.4 46.2
1?.1331 R3.9 4. S 35.7 2c .9 1.19 42. 4.72
17.1 3O 71.7 4.32 35.7 21.5 1.31 42.5 4.45 44.7
7.)53(1. 73.5 4.3 35.7 22.5 1 .R 37.2 4.2R 43.6
17. 163R 86 . 4.17 35.7 21.4 1.R4 2 .R 2.46 25.6
17.1.732 73.4 4.74 37 . l 22.3 1.33 32.9 3.79 39.5
17.IR3R R1.5 4.17 37.! 21.9 1.17 46.5 4.9R
l7.t93 R4.A 4.15 37.1 21.3 I. 9 37.% 4.25 44.2
17.P03 91.4 37.1 21. .9R 42.5 8.35 45.3
142
-------
1 of 7
Table D-5
Run 5 Temperatures, Fuel and Stone Rates
DAY.TIME
TEMPERATURE,
DEC.
C ’
RATE.
LB/HR
GASIFIER
REGEN
RECYCLE
OIL
STONE
2.013 0
933.
1067.
72.
404.5
76.0
2.91230
923.
1078.
72.
403.6
63.0
2.91330
939.
1069.
75.
300.0
22.0
2.91430
920.
1060.
76.
300.0
0.
2.0530
9491.
1030.
76.
307.1
5.5
•91630
920.
1053.
76.
307.1
5.5
2.0730
933.
1050.
77.
4.0
2.0930
9491.
1060.
75.
392.6
0.0
2.1030
952.
1055.
75.
414.5
11.0
2.1130
002.
191391.
75.
414.5
13.0
2.j230
080.
1060.
72.
4917.2
12.0
2.1330
085.
1050.
75.
403.6
1 . 0
2. 1 430
000.
1050.
75.
399.0
16.0
2.1530
082.
1050.
00.
405.4
17.0
2.1630
095.
1050.
00.
407.2
10.0
2. 1730
092.
1050.
00.
200.0
17.0
2.1030
000.
1040.
75.
397.2
20.0
2.1930
0091.
1040.
75.
397.2
20.0
2.2030
890.
1070.
80.
399.0
18.0
2.2 )30
882.
1023.
05.
399.0
10.0
2.2230
873.
1040.
95.
398.1
10.0
2.2330
880.
10 8.
00.
399.9
20.0
3.0030
090.
1052.
79.
379.0
22.0
3.0130
080.
70.
404.5
21.0
3.0230
082.
1061.
70.
403.6
20.0
3.0330
004.
1062.
00.
305.3
21.0
3.0430
009.
1065.
79.
391.7
19.0
3.0530
090.
1063.
00.
399.9
10.0
3.0630
000.
1057.
75.
402.6
16.0
3.0730
89 1.
1058.
75.
399.9
19.0
3.0030
099.
1060.
75.
404.7
19.0
3.0930
097.
1055.
70.
392.6
20.0
3.1030
091.
1041.
70.
397.2
15.0
3.1130
091.
1055.
70.
399.0
13.0
3.1230
092.
1055.
70.
393.5
13.0
3.1330
002.
1060.
791.
402.6
14.0
3. )430
002.
1070.
70.
390.1
16.5
3.1530
081.
1061.
70.
304.3
15.5
3.1630
096.
19165.
60.
302.3
14.0
3.1730
096.
1061.
70.
366.0
17.0
3.1030
0991.
1062.
70.
420.9
21.0
3.1930
9091.
1065.
75.
375.2
10.0
3.2030
095.
1060.
70.
439.2
19.0
3.2)30
092.
1070.
00.
399.0
21.0
3.2230
9912.
1065.
05.
390.1
13.0
3.2330
092.
1067.
03.
397.7
0.
143
-------
2 of T
Table D-
Run Temperaturea, Fuel and Stone Rates
DAY.TIME TEMPERATURE’ DEG C, R4TE L8/HR
GASIFIER REGEN RECYCLE OIL STONE
4.0030 901. 1062. 03. 379.0 0.
4.0130 876. 1062. 82. 404.5 0.
4.0230 831. 1055. 82. 403.6 0.
4.0330 888. 1060. 02. 385.3 0.
4.0430 910. 1065. 03. 391.7 0.
4.0530 850. 1067. 55. 399.9 101.0
4.0630 050. 1068. 65. 402.6 136.0
4.0730 853. 1066. 105. 399.9 110.0
4.0830 859. 1067. 80. 0 1.7 57.0
4.0930 070. 1050. 70. 392.6 19.0
4.1030 870. 1050. 78. 397.2 0.
4.1130 868. 1070. 78. 399.0 0.
4.1230 o75. 1020. 78. 393.5 49.0
4.1330 870. 1050. 00. 396.2 63.0
4.1430 861. 1050. 80. 392.6 63.0
4.1530 862. 1055. 00. 396.2 65.0
4.1630 852. 1049. 80. 390.0 62.0
4.1130 850. 1050. 80. 395.3 61.0
4.1830 060. 1050. 80. 394.4 70.0
4.1930 062. 1058. 00. 397.2 85.0
4.2030 860. 1052. 80. 399.0 78.0
4.2130 862. 1051. 86. 393.5 59.0
4.2230 865. 1053. 85. 393.5 60.0
4.2330 873. 1052. 60. 384.3 63.0
5.0030 880. 1055. 75. 413.6 73.0
5.0130 870. 1055. 80. 399.0 45.0
5.0230 867. 1059. 80. 314.3 63.0
5.0330 848. 1050. 02. 307.1 64.0
5.0430 858. 1058. 80. 416.4 67.0
5.0530 858. 1052. 85. 400.8 64.0
5.0630 858. 1058. 82. 400.0 63.0
5.0130 859. 1058. 80. 399.0 50.0
5.0830 858. 1052. 80. 400.0 65.0
5.0930 845. 3050. 80. 400.8 84.0
5.1030 845. 1050. 80. 399.0 67.5
5.1130 851. 1050. 80. 399.0 64.0
5.1230 848. 1049. 80. 398.1 42.0
5.1430 870. 1055. 80. 308.0 30.0
5.1530 072. l OSS. 80. 386.2 29.0
5.1630 888. 1054. 80. 386.2 35.0
6.0230 865. 1049. 85. 300.7 32.0
6.0330 862. 1050. 83. 380.1 49.0
6.0430 852. 1054. 80. 384.3 47.0
6.0530 868. 1050. 85. 301.6 46.0
6.0630 852. 1060. 83. 382.5 40.0
6.0730 850. 1055. 83. 381.6 39.0
6.0830 850. 1065. 82. 367.9 41.0
144
-------
3 of 7
Table D- 5
Run 5 Teinj,eratures, Puel and stone Bates
DAY.TIME
TEMPERATURES DEG C s
RATE ’
CASIF’IER
RE(EN -
RECYCLE
OIL
STONE
.2230
995.
1005.
70.
301.6
54.0
9.2330
299 ’.
1035.
19’.
300.7
55.9’
9.0030
97?.
106?.
79’.
303.4
71.0
9.0130
970.
19’a7.
79’.
393.5
74.0
9.0239’
075.
1042.
7Q .
392.6
66.0
9.0339’
072.
1 0 R.
7 .
393.5
59.0
9.o 39’
990.
1060.
72.
391.7
43.0
9.0530
92 ?.
1055.
72.
393.5
3 .0
11.1439’
07 1.
1069’.
oQ ’.
392.6
07.9’
11.1539’
069’.
1069’.
80.
1411.0
03.0
fl.163Q
060.
1067.
09’.
393.5
35.9’
11.1739’
966.
1 064.
00.
393.5
35.9’
11.1039’
979.
1069
74.
392.6
30.0
11.1930
992.
10 60.
o o.
92.e ,
11.2030
092.
9’75.
99’.
393.5
39.9’
11.2130
992.
1067.
00.
391.7
44 .9’
11.9230
000.
1060.
05.
398.1
fl.233 9 ’
867.
1060.
00.
390.1
12.00391
069.
1065.
03.
407.2
53.0
12.0230
875.
1065.
0?.
409.1
39.9’
2.G 33 0
001.
1065.
90.
407.2
20.0
IP. 0 30
909’.
)9’5P .
o 9 ’.
7?.053V
001.
1050.
00.
407.?
31.0
12.0630
079.
1050.
00.
407.2
3 1 1.0
17.0730
075.
1050.
75.
193.5
12.0039’
979.
1050.
75.
304.3
34.0
40.0
17.0939’
R 6 0.
l0 4 .
7S.
397.1
12.1030
065.
oo.
394.4
45.0
12.1230
075.
1030.
80.
395.3
12.1330
0 6 0.
19’4i .
00.
396.2
41.5
12.1430
Q 9.
1035.
90.
396.2
47.0
12.1530
079’.
1042.
75.
395.3
12.1630
060.
1071.
03.
397.2
12.1739’
065.
10 ’49.
02.
395.3
34.0
12.1030
R76.
1071 •
05.
393.5
12.1930
079.
1022.
5.
394.4
j2.9 030
074.
1030.
05.
3914.4
12.2130
071.
1012.
91.
394.4
12.2230
079.
1075.
80.
403.6
12.2330
075.
1020.
75.
390.1
42.0
13.0030
965.
1020.
75.
392.6
40.0
13.0130
062.
1025.
75.
309.9
34.0
13.0230
270.
1060.
75.
i9t.7
13.0330
875.
1020.
80.
392.6
34.5
13.0430
000.
1029’.
V 0.
391.7
20.8
13.0530
RRQ ).
1035.
75.
390.0
33.0
13.0630
070.
1020.
75.
303.1
145
-------
k of 7
Table D-5
Run 5 Temperatures, Fuel and stone Rates
DAY.TIME
TEMPERATURE’
REGEN
RECYCLE
OIL
STONE
GASIFIER
80.
387.1
39.0
16.0730
868.
990.
387.1
56.0
16.0830
875.
1070.
388.9
55.8
16.0930
87%.
1068.
90.
388.0
49.0
16.1030
86%.
1070.
80.
387.1
45.0
16.1130
881.
1029.
79.
388.9
61.0
16.1230
866.
1050.
72.
384.3
63.0
16.1330
869.
1051.
7?.
380.7
60.0
16.1530
870.
1060.
71.
381.6
65.0
16.1630
861.
1060.
7*.
381.6
70.0
16.1730
865.
1060.
381.6
77.0
16.1830
872.
1061.
70.
381.6
75.8
16.1930
865.
1060.
70.
379.8
73.8
16.2030
859.
1055.
70.
402.6
73.0
16.2*30
855.
975.
72.
402.6
74.8
16.2230
848.
1012.
72.
396.2
34.0
16.2330
862.
1050.
72.
397.2
17.0030
861.
1061.
7 1.
399.0
33.0
17.0130
870.
1052.
70.
397.2
34.0
17.0230
87?.
1039•
70.
399.0
40.0
17.0330
872.
1031.
70.
418.2
30.0
17.0430
872.
1038.
62.
378.9
45.0
17.0530
876.
1039.
70.
399.0
40.0
17.0630
870.
1020.
70.
397.2
41.0
17.0730
869.
1059.
69.
399.0
36.0
*7.0830
876.
1052.
70.
399.0
15.3
17.0930
868.
1053.
70.
397.2
46.0
17.1030
861.
1052.
70.
399.0
37.6
17.1130
860.
1051.
70.
400.0
37.0
17.1230
861.
1050.
70.
399.0
44.0
17.1330
862.
104%.
70.
411.8
50.0
17.1430
862.
1059.
70.
396.2
69.0
17.1530
865.
1041.
70.
398.1
69.0
17.1630
858.
1041.
70.
397.2
49.0
17.1730
860.
1049.
70.
397.2
43.0
)7.)830
85%.
1050.
65.
397.2
40.0
17.1930
869.
1051.
6..
397.2
36.0
17.2030
862.
1050.
65.
428.3
88.0
20.2030
91%.
1012.
68.
438.3
51.0
20.2130
900.
1012.
70.
454.0
59.0
90.2230
888.
i 55.
146
-------
5 of 7
Table D-5
Run 5 Temperatures, Fuel and Stone Rates
D Y.TIME
TEMPERITUREs
bEG
C
RATEs
LB/HR
GASIFIER
REGEN
RECYCLE
OIL
STONE
20.2330
890.
1052.
70.
462.1
60.0
21.0030
915.
1055.
65.
462.1
57.0
21.0130
915.
1053.
65.
434.7
53.0
21.0230
915.
1060.
60.
467.6
56.0
21.0330
910.
1050.
5.
477.7
42.0
21.0430
905.
1060.
65.
471.3
48.0
21.0530
902.
1060.
65.
471.3
40.8
21.0630
914.
1060.
65.
469.4
42.0
21.0730
900.
1060.
65.
468.5
45.0
21.0830
890.
1060.
62.
477.7
45.0
21.0930
901.
1067.
65.
459.4
38.0
21 . 1 030
900.
1064.
65.
460.5
43.0
21.1130
899.
1061.
65.
467.6
45.0
21.1230
900.
1062.
66.
468.5
46.0
21.1330
903.
1065.
3.
460.5
51.0
21.1430
21.1530
901.
089.
1068.
1064.
62.
64.
h67.6
/472.2
48.0
52.0
21.1630
895.
1063.
63.
474.0
51.0
2 1. 1730
898.
1066.
6 .
472.2
52.0
21.1830
p75.
1066.
6.
465.8
51.0
21.1930
893.
1069.
60.
470.4
55.0
21.2030
892.
1069.
70.
435.6
35.0
21.2130
860.
1069.
70.
409.1
23.0
21.2230
856.
1068.
70.
417.3
30.0
21.2330
870.
1071.
70.
416.4
40.0
22.0030
871.
1068.
70.
410.9
48.0
22.0130
871.
1070.
70.
4)6.
46.0
22.0230
878.
1068.
70.
408.1
31.0
22.0330
872.
1069.
70.
411.8
37.0
22.0430
874.
1070.
70.
420.9
30.0
22.0530
875.
1069.
70.
423.7
30.0
22.0630
875.
1070.
70.
405.4
32.0
22.0730
888.
1070.
70.
402.6
32.0
22.0830
875.
1070.
70.
4 2.B
36.0
22.0930
072.
1071.
77.
412.7
36.0
22.1030
072.
1069.
72.
410.9
32.0
22.1130
871.
1064.
72.
410.0
29.0
22.1230
875.
1065.
72.
410.9
32.0
22.1330
873.
1069.
72.
420.0
34.0
22.1430
875.
1070.
72.
409.1
29.0
22.1530
872.
1069.
71.
409.1
30.0
22.1630
071.
1069.
72.
409.1
36.0
22.1730
877.
1071.
71.
410.9
30.0
22.1830
873.
1068.
70.
408.1
35.0
147
-------
6 of 7
Table D 5
Run 5 Temperatures, Fuel and Stone Rates
DAY • TIME TEMPERATURES
GASIFIER
REGEN
RECYCLE
OIL
STONE
23.1730
091.
1053.
75.
393. 5
71.0
23.1030
069.
1055.
75.
404.5
73.0
23.1930
868.
1052.
75.
395.3
73.0
23.2030
869.
1061.
74.
394.4
67.0
23.2130
860.
1061.
75.
67.0
23.2230
850.
1060.
75.
396.2
70.0
23.2330
866.
1060.
75.
395.3
70.0
24.0030
862.
1060.
75.
398.1
70.0
24.0130
069.
1060.
75.
402.6
70.0
24.0230
862.
1060.
75.
405.4
71.0
24.0330
865.
1060.
75.
409.1
66.0
24.0430
062.
1055.
75.
409.1
64.0
24.0530
060.
1054.
74.
410.9
67.0
24.0630
861.
1059.
75.
410.0
63.0
24.0730
865.
1061.
74.
412.7
64.0
24.0830
878.
10 (1.
72.
410.9
70.0
24.0930
873.
1061.
72.
411.0
78.0
24.1030
073.
1061.
70.
409.1
69.0
24.1130
005.
1060.
72.
401.7
71.0
24.1230
079.
1061.
71.
409.1
24.1330
875.
1061.
72.
410.0
21.0
24.1430
870.
1060.
70.
409.1
30.0
24.1530
870.
1060.
70.
410.9
30.0
24.1630
869.
1060.
75.
409.1
40.0
24.1730
869.
1059.
72.
408.1
47.0
24.1030
871.
1060.
75.
410.0
46.0
24.1930
865.
1060.
75.
407.2
24e2030
870.
1061.
75.
407.2
48.0
24.2130
876.
1060.
71.
406.3
24.2230
875.
1060.
70.
406.3
52.0
24.2330
883.
1060.
70.
406.3
45.0
25.0030
870.
1060.
70.
407.2
52.0
25.0130
878.
1062.
70.
406.3
54.0
25.0230
874.
1060.
70.
406.3
50.0
25.0330
880.
1060.
70.
406.3
52.0
25.0430
880.
1060.
70.
403.6
40.0
25.0530
880.
1060.
70.
404.5
45.0
25.0630
880.
1060.
70.
404.5
25.0730
25.0030
880.
878.
1060.
1060.
70.
70.
401.7
401.7
49.0
148
-------
7 of 7
DAY • TI ME
Table -
Run 5 Temperatures, Fuel and Stone Rates
T EM P E RAT UREa
25.0930
25.1030
25.1130
25.1230
25.1330
25.1430
25.1530
25.1630
25.1730
25.1830
25.1930
25.2030
25.2130
25.2230
25.2330
26.0030
26.0130
26.0230
26.0330
26.0430
26.0530
26.0630
26.0730
26.0830
26.0930
26.1030
26.1130
26.1230
26.1338
26.1430
26.1530
26.1630
26.1730
26.1830
GASIFIER
800.
883.
890.
870.
078.
878.
873.
870.
871.
881.
379.
883.
802.
081.
876.
07 !.
872.
870.
878.
880.
085.
806.
866.
870.
872.
062.
860.
064.
868.
865.
860.
870.
877.
072.
REGEN
1060.
1061.
1060.
1060.
1059.
1060.
1060.
1068.
1060.
1061.
1061.
1060.
1060.
1060.
1058.
1060.
1060.
1060.
1068.
1060.
1060.
1060.
1060.
1060.
1060.
1060.
1060.
1060.
1068.
1062.
1860.
160.
1062.
1859.
DEG C.
RECYCLE
75.
75.
70.
73.
75.
75.
75.
75.
75.
75.
74.
74.
70..
70.
70.
70.
70.
70.
70.
70.
70.
70.
70.
78.
70.
70.
78.
70.
74.
74.
73.
74.
74.
72.
RATE’
OIL
403.6
405.4
402.6
403.6
402.6
405.4
403.6
400.8
412.7
09.1
410.9
410.0
411.0
410.9
409.1
409.1
410.0
418.9
409.1
410.0
410.8
418.0
410.0
410.0
409.1
411.8
418.0
418.9
409.1
410.0
409.1
410.9
408.1
408.1
LB/HP
STONE
45.0
49.0
39.0
47.8
45.0
42.0
37.0
43.0
30.0
33.0
32.0
36.0
31.0
36.0
36.0
36.0
32.0
35.0
32.0
34.0
32.0
27.0
43.0
44.0
38.0
43.0
44.0
45.0
4.0
45 .
52.8
43.0
42.0
43.0
149
-------
1 of 7
Table D-6
Run 5 Gas Bates
DAY.TIME
GAS RATESP
SCFM
REGEN
VELOCITY
GASIFIER
PILOT
NITROGEN
FT/SEC
AIR
FLUE GAS
PROPANE
2.1
4.69
2.013 8
274.
89.
2.5
18.6
2.1
4.96
2.0238
27 .
89.
2.5
3.81
2.0338
247.
117.
2.4
14.7
2.1
4.02
2.0438
258.
111.
2.4
15.8
4.34
2.0530
258.
111.
2.5
17.5
3.92
2.0638
258.
186.
2.5
15.5
.1
3.59
2.8730
265.
l OS.
2.5
2.1
3.95
2.8938
281.
93.
2.5
15.5
2.1
4.32
2.3038
268 .
123.
2.5
17.3
2.1
4.58
2.3338
249.
205.
2.5
18.7
2.1
4. 1
2.1230
244.
335.
2.5
19.2
2.1
4.57
2.1330
249.
329.
2.5
2.1
4.67
2.1430
239.
329.
2.5
2.1
4.67
2.1530
244.
1.?.
2.5
18.7
2.1
4.59
2.1630
244.
122.
2.5
I.4
2.1
4.54
2.3730
239.
136.
2.5
18.2
2.1
4. 12
2.1830
223.
105.
2.5
16.2
2.1
4.12
2.1930
223.
105.
2.5
16.2
2.3
4.37
2.2030
233.
94.
2.5
17.1
2.1
3.50
2.2130
234.
116.
2.5
13.9
2.1
4.63
2.2238
223.
99.
2.5
18.5
2.1
4.99
2.2330
234.
96.
2.5
20.1
2.1
4.72
3.0030
234.
94.
2.5
2.1
3.0130
234.
98.
2.5
2.1
4.91
3.0230
234.
90.
2.5
19.5
2.1
4.97
3.0330
234.
88.
2.5
19.8
2.1
5.16
3.0430
234.
88.
2.5
20.6
2.1
5. 15
3.0530
234.
85.
2.5
2.1
5.14
3.0638
239.
82.
2.5
20.6
2.1
5.08
3.0730
239.
82.
2.5
28.4
2.1
5.06
3.0830
239.
.
2.5
2.1
4.94
3.0938
244.
R I.
2.5
2.3
4.70
3.1830
234.
95.
2.5
2.1
4.74
3.1130
245.
186.
2.5
2.1
4.67
3.1230
244.
102.
2.5
2.1
5.30
3.1330
234.
105.
2.5
2.1
4.77
3.3430
233.
112.
2.5
2.1
.73
3.3538
233.
312.
p.5
2.1
4.72
3.1630
234.
108.
2.4
18.7
2.1
3.94
3.1730
234.
112.
2.5
15.4
2.1
4.79
3.1830
239.
107.
2.3
19.1
17.6
2.1
4.46
3.1938
239.
106.
2.5
19.9
2.1
4.99
3.2030
234.
107.
2.5
14.0
2.1
3.65
3.2130
237.
104.
2.5
18.7
2.1
4.71
3.2230
3.2330
239.
235.
104.
185.
2.5
2.5
20.3
2.1
5.07
150
-------
Table 13-6
Run 5 Gas Rates
2 of 7
DAY.TIME
GAS RATESP
SCFM
REGEN
G SIFIER
PILOT
REGENERATOR
VELOCITY
41R
FLUE G S
PROPANE
AIR
NITROGEN
FT/SEC
4.003 0
235.
l OS.
2.5
19.1
2.1
4.0130
233.
90.
2.5
1 .9
2. )
4.71
4. 0230
234.
90.
2.5
19.7
2.1
4.9)
4. 33Ø
234.
07.
2.5
20.0
2.1
4.90
4.0430
234.
07.
2.5
2 0.s
.l
5.16
4.0530
219.
1)5.
2.5
20.8
2. )
5.17
4.0630
235.
96.
2.5
20.0
2. )
5.17
4.0730
235.
05 .
.5
20.6
2 .)
5.10
4.0 03 1 ’ )
735.
100.
2.5
20.5
2.1
5. O R
4.0930
234.
94.
2.5
20.0
2.1
2.89
4.1030
223.
99.
2.5
19.2
2.1
4.69
4.1130
223.
110.
2.5
19.2
2.1
2.76
4.1230
2 3.
110.
2.5
.9
. )
4.54
2. )33 0
933.
121.
2.5
18.9
2.1
4.65
4.1430
237.
117.
2.5
17.3
2.1
4.20
4.1530
.45.
114.
2.5
)0.7
.1
4.62
4.1630
243.
117.
2.5
17.4
2.1
4.30
4.1730
243.
112.
2.5
10.9
2.1
4.62
4.1030
242.
112.
2.5
19.1
2.1
4.70
4.1930
242.
111.
2.5
16.0
2.1
4.21
4.2030
243.
115.
2.5
19.2
2.1
4.69
4.7)30
242.
1)6.
2.6
12.9
2.1
4.65
4.2730
243.
116.
2.6
3•5
2.1
2.57
4.2330
222.
132.
2.6
19. 1
2.1
4.71
5.0030
233.
1 1 1.
2.6
19.5
2.1
4.79
5.0)30
233.
ii ).
2.6
19.3
1.0
4.53
5.0230
233.
I I !.
2.6
19.6
2.1
4.03
5.0330
233.
lIt.
2.6
19.1
1.0
4.47
5.0430
233.
I I I .
2.6
19.)
2.)
4.72
5.0530
243.
110.
2.6
19.0
1.0
4.45
5. 0630
243.
111.
2.6
19.0
0.2
4.25
5.0730
2 43.
111.
2.6
19.0
1.0
4.26
5.0930
243.
111.
2.6
19.1
1.0
4.47
5.0930
24 .
111.
2.5
19.0
1.2
4.45
5.1030
242.
111.
2.5
19.1
0.2
4.24
5.1130
237.
94.
2.5
19.0
1.0
4.46
5.1230
226.
99.
2.6
19.2
1.0
4.50
5.1430
233.
99.
2.6
18.6
1.0
4.37
5.1530
5.1630
238.
233.
97.
93.
2.6
2.6
18.3
16.9
1.0
0.0
4.32
3.78
6.0230
6.0330
6.0430
6.0530
6.0630
6.0730
6.0030
224.
224.
224.
223.
223.
234.
202.
122.
122.
116.
116.
105.
99.
99.
2.6
2.6
2.6
2.6
2.6
2.6
2.6
17.0
17.3
17.7
15.5
15.6
16.5
16.9
1.0
.1
2.1
1.2
1.0
0.0
1.0
4.14
4.25
4.35
3.63
3.67
3.62
3.97
151
-------
3of7
Table D-6
Run 5 Gas Rates
DAY.TIME GAS RATES. SCFM REcEN
GASIFIER PILOT REGENERATOR VELOCITY
AIR FLUE GAS ROPANE AIR NITROGEN FT/SEC
8.2230 239. 112. 2.5 79.2 1.0 5.00
0.2330 234. 312. 2.5 22.1 2.1 5.39
9.0030 245. 106. 2.5 23.4 2.1 5.01
9.0130 234. 303. 2.5 3.6 1.0 5.54
9.0230 39. 100. 2.5 24.5 2.3 5.96
9.0330 244. 100. 2.5 24.5 2.1 6.01
9.0 430 245. 100. 2.5 26.7 2.1 6.56
9.0530 244. 94. .5 27.3 2.3 A. 67
13.3430 730. 117. 2.7 21.1 Ii.1 5.71
13.3530 230. 300. 2.7 2 .1 . 3 5.92
1l.16i 730. 99. 2.5 22.6 .3 5.02
11.1730 230. 99. 2.5 22.6 3.1
11.1030 23g. 100. 2.5 22.2 2.3 6.03
11.1930 249. 9.. 2.5 22.2 4 .3 5.99
31.2030 249. 06. 9.5 22.2 4.3 6.00
11.9130 249. 06. 2.5 21.3 4.1 6.21
31.2230 239. 90. 2.5 91.0 3.1 5.65
11.2330 239. 99. 2.5 19.2 3.1 5.31
32.0030 46. 99. 2.5 21.2 3.1 5.53
12.0230 224. 93. 2.5 90.9 4.3 5.65
12.0330 244. 93. 2.5 19.1 4.3 5.74
12.0430 224. 93. 2.5 23.4 4.1 6.15
32.0530 244 . 93. 2.5 21.3 5.3 5.91
12.0630 234. 105. 2.5 21.4 .3 5.72
19.0730 223• 111. 2.5 23.5 4.1 6.39
32.0830 223. 11 1. 2.5 73.6 4.3 5.76
12.0930 234. 94. 2.5 21.0 5.1 6.00
12.3030 934. OR. 2.5 22.2 1.1 5.87
32.3230 934. 07. 2.5 22.1 5.1 6.01
12.1330 734. 7. 2.5 22.0 4.1 5.02
12.1430 233. OR. 2.5 72.0 5.1 6.02
32.3530 933. 8. 2.5 22.4 3.1 5.69
12.1630 234. 01. 2.5 29.0 3.1 5.7
12.3730 231. 01. P. S p2.0 3.1 5.A
12.3030 933. 81. 2.5 73.6 0.0 4.95
39.3030 233. R I. 2.5 Pp.1 2.1 5.33
19.2030 234. 07. 2.5 22 . ! 2.1 5.33
32.2130 233. 99. 2.5 22.1 2.1 5.26
12.2230 233. 110. 9.5 73.3 3.1 5.90
3. 2330 223. 323. 2.5 97.7 4.1 5.06
33.0030 223. 12. 2.5 22.0 2.3 5.26
13.0130 223. 122. 2.5 2.0 4.1 5.72
33.0230 273. 122. 2.5 2.2 4.3 5.93
33.0330 223. 122. 3.6 22.7 4.3 5.05
13.0430 223. 122. 2.5 22.2 2.1 5.79
13.0530 23. 32. 2.5 22.6 4.1 5.91
33.0630 223. 13. 2.5 23.0 4.) 5.93
152
-------
k of 7
Table D-6
Run 5 Gas Rates
DAY.TIME
GAS RATES
SCEM
REGEN
GASIFIER
PILOT
REGENERATOR
VELOCITY
AIR
FLUE GAS
PROPANE
AIR
NITROGEN
FT/SEC
16.0730
241.
140.
2.6
19.6
4.1
5.06
l6. 30
230.
110.
2.6
19.2
4.?
5.29
16.0930
238.
110.
2.6
19.2
2.1
4.02
16.1030
230.
110.
2.6
10.0
2.1
4.74
16.1130
238.
107.
2.6
19.0
3.1
4.06
16.1230
238.
111.
2.6
21.0
3. ?
5.4?
16.1330
249.
106.
2.5
20.3
2.1
5.02
16.1530
249.
100.
2.5
20.1
5.01
16.1630
243.
o.
2.5
19.2
2.1
4.00
16.1730
249.
00.
2.5
19.5
2.1
4.80
16.1830
249.
94.
2.5
20.5
3.1
5.33
16.1930
243.
00.
2.5
20.3
3.1
5.27
16.2030
230.
88.
2.5
20.4
3.1
5.27
16.2130
243.
00.
2.6
19.5
3.1
4.76
16.2230
244.
00.
2.6
19.5
3.1
4.90
16.2330
244.
88.
2.6
20.4
3.1
5.25
17.0030
249.
09.
2.5
20. 6
4.1
5.56
17.0130
250.
98.
2.5
20.3
3.1
5.23
17.0230
249.
00.
2.6
21.5
3.1
5.45
17.0330
249.
0 0.
2.5
20.5
2.1
4.97
17.0430
249.
08.
2.5
21.5
3.1
5.45
17.0530
249.
99.
2.5
36.4
2.1
17.0630
249.
89.
2.5
20.6
3.1
5.17
17.0730
29.
oo.
2.5
20.6
3.1
5.32
17. 0030
249.
80.
2.5
20.6
3.1
5.29
17.0930
249.
00.
2.5
20.7
2.1
5.10
17.1 30
249.
00.
2.6
20.4
2.1
5.03
17.1130
239.
89.
7.5
20.5
2.1
5.05
17.1?30
249.
76.
2.6
20.1
1.0
4.74
17.1330
243.
7..
2.5
20.5
3.1
5.26
17.1430
259.
82.
7.5
20.!
2.1
4.99
17.1530
259.
82.
2.5
20.7
2.1
5.05
17.1630
240.
02.
.5
21.7
2.1
5.77
17.1730
259.
76.
2.5
21.5
2.1
5.26
17.1830
254.
76.
2.5
70.9
2.1
5.13
17.1930
248.
76.
2.5
21.1
2.1
5.1k
17.2030
254.
76.
2.5
21.3
2.1
5.22
20.2030
293.
52.
2.4
21.0
3.1
5.17
20.2130
273.
63.
2.4
22.3
2.1
5.24
20.2230
272.
63.
2.4
23.1
1.0
5.35
153
-------
5of 7
Table D-6
Run 5 Gas Rates
DAY.TIME GAS RATESs SCFM REGEN
GASIFIER PILOT REGENERATOR VELOCITY
AIR FLUE GAS PROPANE AIR NITROGEN FT/SEC
20.2330 283. 63. 2.4 22.8 1.0 5.28
21.0030 288. 40. 2.4 2 .5 1.0 5.22
21.0130 288. 40. 2.5 20.9 1.0 4.04
21.0230 293. 40. 2.5 21.7 1.0 5.04
21.0330 283. 46. 2.5 22.0 1.0 5.11
21.0430 282. 46. 2.5 21.5 1.0 4.99
21.0530 282. 46. 2.5 21.5 1.0 4.99
21.0630 283. 46. 2.5 21.2 1.0 4.93
21.0730 272. 46. 2.4 21.1 1.0 4.89
21.0830 272. 46. 2.4 21.0 1.0 4.89
21.0930 272. 46. 2.5 21.0 1.0 4.09
21.1030 27 0. 46. 2.6 20.5 1.0 4.78
P.1.1130 268. 46. 2.6 20.5 1.0 4.76
21.1230 267. 46. 2.6 20.5 1.0 4.78
21.1330 268. 46. 2.6 20.6 1.0 4.00
21.1430 268. 46. 2.6 20.6 1.0 4.81
21.1530 269. 46. 2.6 20.5 1.0 4.76
21.1630 268. 46. 2.6 20.2 1.0 4.70
21.1730 268. 46. 2.6 20.5 1.0 4 .7
21.1830 063. 58. 2.6 19.2 1.0 4.48
21.1930 268. 46. 2.6 19.7 1.0 4.50
21.2030 236. 81. 2.6 20.5 1.0 4.80
21.2130 226. 98. 2.6 21.4 1.0 4.99
21.2230 226. 98. 2.6 21.7 1.0 5.07
21.2330 235. 87. 2.6 21.4 1.0 5.01
22.0030 235. 86. 2.6 20.6 1.0 4.81
22.0130 p35. 86. 2.5 20.9 1.0 4.91
22.0230 236. 86. 2.5 20.9 0.0 4.67
22.0330 236. 86. 2.6 20.8 1.0 4.85
22.0430 236. 86. 2.6 20.8 2.1 5.09
22.0530 236. 86. 2.6 21.0 1.0 4.91
22.0630 230. 86. 2.5 20.5 1.0 4.80
22.0730 236. 80. 2.5 20.4 1.0 4.77
22.0830 235. 88. 2.5 20.8 1.0 4.85
22.0930 235. 86. 2.6 20.4 1.0 4.76
22.1030 236. 86. 2.6 19.9 1. 4.66
2.1130 236. 86. 2.6 20.5 1.0 4.77
22.1230 236. 86. 2.6 17.7 1.0 4.15
22.1330 235. 86. 2.6 20.1 1.0 4.70
20.1430 236. _86. 2.6 20..4 1.0 4.76
22.1530 237. 86. 2.6 20.6 1.0 4.80
22.1630 237. 86. 2.6 20.5 1.0 4.78
22.1730 237. 86. 2.6 20.8 1.0 4.85
22.1830 236. 86. 2.6 20.1 1.0 4.68
154
-------
6 of 7
Table D-6
Run 5 Gas Rates
DAY.TIME
GAS RATES.
SCFM
REGEN
GASIFIER
PILOT
REGENERATOR
VELOCIT’T’
AiR
FLUE GAS
PROPANE
AIR
NITROGEN
FT/SEC
23.1730
203.
108.
2.7
21.3
1.0
4.97
23.1830
214.
108.
2.7
21.5
1.0
5.02
23.1930
213.
108.
2.7
21.4
1.0
5.00
23.2030
224.
l O B.
3.0
21.2
1.0
4.97
23.2130
203.
103.
2.6
20.9
1.0
4.91
23.2230
213.
105.
2.6
20.6
1.0
4.83
23.2330
214.
105.
2.6
20.0
1.0
4.70
24.0030
230.
103.
2.6
20.2
1.0
4.74
24.0130
230.
103.
2.6
20.9
1.0
4.89
24.0230
230.
97.
2.6
20.5
1.0
4.80
24.0330
230.
97.
2.6
20.1
1.0
4.71
24.0430
30.
97.
2.6
20.4
1.0
4.77
24.0530
235.
91.
2.6
20.9
1.0
4.87
24.0630
235.
91.
2.6
20.9
1.0
4.88
24.0730
219.
9!.
2.6
20.6
1.0
4 .8
24.0830
235.
80.
2.6
20.?
1.0
4.73
24.0930
239.
6.
2.6
20.0
1.0
4.60
24.1030
244.
86.
2.6
20.3
1.0
i .73
24.1130
234.
86.
2.6
20.7
1.0
4.83
24.1230
224.
9..
2.7
20.8
1.0
4.07
24.1330
230.
97.
2.7
21.0
1.0
4.92
24.1430
225.
97.
2.6
21.2
1.0
4.96
24.1530
225.
97.
2.6
21.1
1.0
4.93
24.1630
225.
97.
2.6
21.4
1.0
5.00
24.1730
236.
97.
2.6
21 .1
1.0
4.92
24.1830
225.
91.
2.6
20.9
1.0
4.08
24.1930
225.
97.
2.7
20.8
1.0
4.05
24.2030
25.
97.
2.7
21.6
1.0
5.05
24.2130
225.
97.
2.7
21.4
1.0
5.00
24.2230
225.
97.
2.6
2 1.7
1.0
5.06
24.2330
225.
97.
2.6
21.7
1.0
5.05
25.0030
225.
97.
2.6
23.5
1.0
5.45
25.0130
225.
97.
2.5
21.9
1.0
5.10
25.0230
225.
97.
2.5
21.6
1.0
5.03
25.0330
225.
97.
2.4
21.7
1.0
5.Ø
25.0430
225.
97.
2.5
21.8
1.0
5.07
25.0530
225.
97.
2.5
21.0
1.0
5.07
25.0630
225.
97.
2.5
21.8
1.0
5.00
25.0730
275.
97.
2.5
21.9
1.0
5.09
25.0830
225.
97.
2.4
21.8
1.0
5.00
155
-------
7of 7
Table D-6
Run 5 (las Rates
DAY.TIME GAS RATES. SCFM REGEN
GAS! Fl ER PILOT REGENERATOR VELOCITY
AIR FLUE GAS PROPANE AIR NITROGEN FT/SEC
25.0930 215. 91. 2.5 21.8 1.0 5.07
25.1030 225. 91. 2.5 21.9 1.0 5.10
25.1130 215. 92. 2.5 21.7 1.0 5.05
25.1230 2)4. 91. 2.5 21.4 1.0 4.99
25.1330 225. 91. 2.5 21.2 1.0 4.93
25.1430 226. 91. 2.5 19.6 1.0 4.57
25.1530 226. 91. 2.5 21.6 1.0 5.02
25.1630 225. 91. 2.5 21.5 1.0 5.01
25.1730 225. 91. 2.5 21.8 1.0 5.07
25.1830 226. 97. 2.5 21.6 1.0 5.02
25.1930 225. 91. 2.5 21.0 1.0 4.89
25.2030 225. 91. 2.5 20.9 1.0 4.86
25.2130 231. 94. 2.5 20.8 1.0 4.85
25.2230 231. 94. 2.5 20.9 1.0 4.85
25.2330 215. 97. 2.5 20.6 1.0 4.79
26.0030 214. 97. 0.5 20.6 1.0 4.00
26.0130 225. 97. 2.5 21.1 1.0 4.91
26.0230 220. 97. 2.5 21.6 1.0 5.02
26.0330 231. 86. 2.5 22.0 1.0 5.10
26.0430 225. 92. 2.5 22.0 1.0 5.10
26.0530 226. 92. 2.5 22.0 1.0 5.10
26.0630 226. 92. 2.5 21.9 1.0 5.07
26.0130 226. 97. 2.5 21.9 1.0 5.08
26.0830 220. 95. 2.5 22.1 1.0 5.12
26.0930 225. 94. 2.5 21.8 1.0 5.05
26.1030 209. 97. 2.5 22.1 1.0 5.12
26.1130 209. 97. 2.5 2.3 1.0 5.16
26.1230 2 14. 95. 2.5 22.1 1.0 5.12
26.1330 214. 91. 2.5 22.1 1.0 5.12
26.1430 247. 91. 2.5 22.0 1.0 5.10
26.1530 225. 89. 2.5 21.7 1.0 5.02
26.1630 26. 89. 2.5 21.6 1.0 4.99
26.1130 220. 85. 2.5 21.2 1.0 4.93
26.1830 225. 86. 2.5 21.3 1.0 4.94
156
-------
1 of 7
2.0130
2 • 0230
2.0330
2.0430
2.0530
2.0630
2.0730
2 • 0930
2. 1030
2.1130
2.1230
2.1330
2.1430
2.1530
2.1630
2.1730
2 • 1830
2.1930
2.2030
2 • 2130
*2230
2.2330
3.0030
3 • 0 1 30
3 • 0230
3 • 0330
3.0430
3.0530
3 • 0630
3 • 0730
3.0830
3 • 0930
3.1030
3.1130
3.1230
3.j 330
3. 1430
3 • 1 530
3 • 1 630
3.1730
3. 1830
3.1930
3 • 2030
3.2130
3.2230
3.2330
Table D-7
Run 5 CAFB P:ressures
DAY.TIME GASIEJER
GAS
SPACE
16.0
15.0
16.0
17.0
17.5
17.5
I 7. 5
18.0
18.5
16.0
16.0
16.0
14.5
16.0
16.0
16.0
14.0
I 4.0
14.5
15.5
13.5
13.5
14.5
14.0
14.0
14.5
14.0
14.0
14.0
15.0
15.0
14.5
JA.5
15.0
15.5
15.5
15.0
16.0
16.5
16.5
16.5
16.5
16.6
1 7.0
16.5
PRESSURES.
DI STRI H
D.P.
14.0
14.0
15.0
16.0
16.0
16.0
16.0
16.0
16.0
16.0
16.0
16.0
15.5
15.5
15.0
I 5.
12.5
12.5
1.5
14.0
12.0
12.5
1.5
14.0
14.0
14.5
14.0
12.0
12.0
12.0
12.0
12.0
13.0
14.0
12.0
14.0
14.0
13.0
13.0
14.0
14.0
13.0
13.5
12.0
13.5
13.5
IN. W.G.
BED
D.P.
19.0
20.0
19.0
18.5
18.5
18.5
18.0
18.5
18.5
19.0
19.0
19.0
19.0
19.0
19.0
19.5
19.0
19.0
20.0
19.0
19.5
19 • 5
19.5
20.0
20.0
20 • 5
21.0
21 .0
21.0
21.0
21.0
21.5
21.5
21 .0
21.5
21 .0
2! • 5
2.0
20.0
19.0
18.0
18.0
18.0
18.0
18.0
18.0
RE EN
SO. CR. BED
D.P.
0.70 19.0
0.70 23.0
0.70 25.0
0.70 27.0
0.80 28.0
0.75 30.0
0.75 32.0
0.75 32.0
0.80 33.0
0.70 32.0
0.80 78.0
0.80 33.5
0.00 34.0
0.70 28.0
0.75 28.0
O •75 2 .0
0.80 28.0
28.0
0.80 20.0
0.75 20.0
0.85 20.0
0.85 20.0
0.80 0.0
0.80 20.0
0.80 20.0
0.80 20.0
0.85 20.0
0.80 20.0
2.85 20.5
0.85 23.0
0.85 23.0
0.80 22.0
0.85 21.0
0.85 .5.0
0.85 27.0
1.00 24.0
1.00 76.0
0.80 27.0
0.75 6*0
0.80 2.0
0.70 26.0
0.80 27.0
0.80 28.0
0.80 27.0
0.80 27.0
0.90 77.0
157
-------
2o1’ 7
Table D-7
Run 5 CAFB Pressures
DAY.TIME GASIFIER PRESSURES. IN. W.G. REGEN
GAS DISTRIB BED SF. GR. BED
SPACE D.P. D.P. D.P.
4.0030 16.8 33.5 - 38.0 0.80 27.0
4.0130 16.3 13.91 18.0 91.80 28.0
4.0230 16.2 13.0 18.0 0.80 27.0
4.0330 17.7 13.5 17.5 0.80 27.5
4.0430 18.0 13.5 17.5 0.891 27.5
4.0530 18.0 13.5 17.0 0.75 28.0
4.0630 19.0 13.5 17.0 0.75 27.0
4.91730 19.9 13.5 19.0 0.75 27.0
4.0830 19.5 13.5 19.3 0.75 27.0
4.0930 22.0 15.0 18.0 0.8 9 1 18.0
4.1030 22.0 15.0 18.0 0.80 10.0
4.1130 22.0 15.0 19.0 0.8 9 1 18.0
4.1230 20.5 15.0 17.5 0.80 18.0
4.1330 20.5 16.0 10.0 0.70 18.0
4.1430 21.0 16.0 18.0 0.70 19.91
4.1530 20.5 16.0 18.0 0.70 10.0
4.1630 21.0 16.0 18.0 0.70 19.0
4.1730 21.0 16.0 18.0 0.70 10.0
4.1830 21.0 16.0 18.0 0.70 18.0
4.1930 21.0 17.0 18.0 0.70 18.0
4.2030 21.0 17.0 18.0 0.70 18.0
4.2130 21.5 17.0 18.0 0.70 19.0
4.2230 20.8 17.0 18.0 0.70 20.0
4.2330 20.5 17.5 10.0 0.65 21.0
5.0030 20.5 15.5 18.0 0.65 293.0
5.0130 20.5 15.5 10.0 0.65 20.0
5.912391 20.5 15.5 10.0 0.65 21.0
5.0330 20.0 15.0 18.0 0.65 21.0
5.0430 20.5 16.0 18.0 0.65 20.0
5.0530 20.7 19.0 18.0 0.60 19.0
5.0630 20.7 17.0 17.7 0.60 22.0
5.0730 21.5 17.5 18.0 0.60 21.0
5.0830 21.0 16.5 18.0 0.60 22.0
5.0930 20.5 16.0 18.0 91.691 21.0
5.1030 20.5 16.0 18.0 0.60 22.0
5. 1130 20.0 17.0 1.0 0.60 23.0
5.1230 19.5 18.0 18.0 0.60 23.0
5.3430 20.5 15.5 17.5 0.60 23.0
5.1530 20.5 15.5 17.5 0.65 22.0
5.1630 21.0 15.5 17.5 0.65 22.0
6.0230 25.2 14.5 15.5 0.70 20.0
6.0330 24.0 15.0 16.5 0.65 22.0
6.0430 25.5 14.5 17.0 0.67 21.0
6.0530 25.0 14.5 17.5 0.67 23.0
6.0630 26.0 14.5 18.0 0.65 20.0
6.0730 28.0 14.0 18.0 0.70 23.0
6.0930 27.0 11.5 18.0 0.65 24.0
158
-------
3of 7
Table D-7
Run 5 CAFB Pressures
DAY.TIME G4SIFIER PRESSURES. IN. ! 1 J.D. REGEN
GAS DISTRIB BED SP. DR. RED
SPACE D.P! - D.P. D•P•
0.223 0 13.0 16.9 1 10.5 0.85
0.2330 13.5 17.0
9.0232 13.2 17.0 19.0 0.75
9.0130 13.0 17.0 19.5 2.75
9.0230 13.0 17.2 20.0 2.75
9.2330 13.0 17.0 20.2 0.72
9. 43O 13.0 17.0 22.0 0.80
9.2530 13.0 17.0 22.0 2.75
11.1430 15.5 18.0 0.65 17.5
11.1530 15.5 10.5 0.65 19.0
11.1630 15.0 19.5 0.70 19.5
11.1730 15.0 19.7 0.72 19.0
11.1030 15.5 21.5 0.70 23.0
11.1930 15.0 22.0 p.70 27.0
11.2030 16.0 22.0 0.70 23.0
11.2130 15.5 22.0 0.70 22.0
11.2230 16.0 21.5 0.70 20.0
11.2330 16.0 21.5 0.70 20.0
12.0030 16.0 2 1.0 21.5 p.70 24.0
12.0230 16.5 21.0 21.0 0.70 22.2
12.0330 16.5 21.0 2.0 0.70
12. 430 16.5 21.0 21 .0 0.70 0.0
12.0530 16.0 21.0 2 1.5 0.00 22.2
12.0630 16.0 20.0 21.0 0.70 20.0
12.0730 16.0 21.0 21.0 0.70 21.0
)2.0R30 16.5 21.0 21.0 0.70 22.0
12.0930 16.0 20.5 22.0 0.70 22.0
12.1030 15.0 10.5 22.0 0.70 22.0
12.1230 15.5 17.5 22.0 0.75 24.0
12.1330 15.0 10.5 22.0 0.70 24.0
12.1430 15.0 10.0 22.0 0.70 23.0
12.1532 15.0 19.0 23.0 0.70 75.0
1.1630 15.0 19.0 22.0 2.70 25.0
1.1732 15.2 10.? 22.0 0.70 24.0
12.1030 15.5 19.5 22.0 0.70 25.0
12.1932 15.5 20.0 21.0 0.70 24.0
12.2030 16.0 20.0 20.0 0.70 22.0
12.2130 16.0 21.5 3p. 5 0.72 23.0
12.2230 16.7 23.0 10.0 0.65 23.0
12.2330 16.7 23.0 10.0 0.65 19.0
13.0030 16.7 23.2 18.0 0.67 24.5
13.0130 16.7 22.5 10.5 0.67 21.5
13.0230 17.0 22.5 10.5 0.67 23.0
13.0330 17.2 22.7 10.5 0.67 22.0
13.2430 17.0 23.0 10.0 0.67 22.5
13.0532 17.0 22.5 10.0 0.67 2.0
13.2630 17.0 22.7 10.2 2.67 22.5
159
-------
4of 7
16 • 0730
6.0830
16.0930
16. 1030
16.1130
16. 1230
16 • I 330
16. 1530
16.1630
16. 1730
16.1830
16.1930
16.2030
16.2I 30
16.2230
I 6 • 2330
I 1.0030
1 7 • 01 30
I 7.0230
I 7.0330
17.0430
1 7 • 0530
17 • 0630
17.0730
17.0030
17.0930
17.1030
17 • 1130
17.1230
17 • 1 330
17.1430
17 • 1 530
17.1630
17. 1730
1 7 • 1830
1 7 • 1930
Ii .2030
20.29 30
70.2 130
20.2230
Table D-7
Run 5 C&FB Pressures
DAY.TLME GASIFIER
GAS
SPACE
17.0
16.0
16.0
16.0
16.0
16.0
16.0
16.0
16.0
16.0
16.0
17.0
17.0
17.0
17.0
17.0
17.5
17.5
17.5
17.5
17.5
18.0
‘7.5
17.5
17.5
17.5
17.5
17.5
17.5
17.5
‘7.5
17.5
17.5
17.5
17.5
1 1.5
11.5
22.0
22 • 0
2. • S
P RE S SUB E S
DI STRI B
U-P.
20 • S
19.5
20 • 0
20 • S
22 • 0
20 S
20 • 0
20.0
19.0
19.0
19.5
19.0
20.0
20 • 0
20.0
19.0
20.0
20 • 0
20.0
19 • S
20.0
20 • 0
19 • S
19.5
19.5
2 0.0
19.5
19.5
19.0
19.0
19.0
19.0
19.0
18.0
18.5
18.5
1% • 5
19.4
20.5
19.0
IN. .G.
BED
0-P.
11.0
19 . 5
19.5
20 • 0
21.0
21.0
22.5
23.0
23.0
2.0
94.0
24.5
34.0
25.0
25.0
24.0
25.0
24.0
24.0
24.0
24.0
25.0
24.5
24.5
25.0
25.0
25.0
25.0
25.0
25.0
25.0
25.0
26.0
26.0
26.0
76.0
19.2
20.0
20 • S
RE (3 EN
SP. CR. BED
D.P.
0.75 26.0
0.75 1.0
0.75 15.0
0.70 20.0
0.70 10.0
0.70 90.0
0.70 21.0
0.65 22.5
0.70 22.0
0.70 23.0
0.70 21.0
0.65 23.0
0.60 20.5
0.65 23.0
0.65 23.0
0.65 21.0
0.65 20.0
0.65 21.0
0.65 22.0
0.65 23.0
0.65 23.0
0.65 23.0
0.65 22.0
0.65 23.0
0.65 22.0
0.65 24.0
0.65 26.0
0.70 25.0
0.70 24.0
0.70 24.0
0.70 25.0
0.70 21.0
( 1.70 23.5
0.70 23.5
0.70 21.0
0.7(1 23.0
24.0
0.73 23.5
0.75 25.0
0.70 25.0
160
-------
5 of 7
20.2330
21 .0030
21 *01 30
21 .0230
21.0330
21 .0430
21 .0530
21 .063Q
21.0730
21.0830
21.0930
21 • 1030
21 • ii 30
2) • 1230
Pt . I 330
Pt • 1830
21 • I 530
21 • 1630
21 • 1730
21 • 1030
.1 • 19 30
21.2030
21.2130
21 .2230
21.2330
22.0030
22.0130
22.0230
22.0330
22.0430
22.0530
22.0630
22.0730
22.0830
22 • 0930
22. 1030
22. I 130
22. 130
2. 1330
22. 1430
22.1530
2• 1638
22. I 738
22. 1030
Table D-7
Bun 5 CAPB Pressures
DAY.TIME (3ASIFIER
GAS
SPACE
23.0
23.5
24.0
24.0
24.5
24.5
24.5
25.0
25.5
25.0
26.0
25 • 5
25.5
25.0
25.5
25.5
26.0
26.5
27.0
27.5
20.0
24.5
25.0
24.5
24.5
25.0
25.0
25 • S
25.5
26.0
26 • 0
26.0
27.0
26.5
26. 5
26.0
26.5
26.5
26.5
26.5
26. 5
26.5
27.0
27.5
PRE 5SURES
D I STRI B
D.P. -
19.0
10.0
17.5
17. 5
18.0
18.0
18.0
18.0
17.5
17.0
16.5
17.0
17.0
16.5
16. 5
16.5
16.0
16.5
16.5
17.0
16.0
17.5
19.0
19.0
19.0
19.0
19.0
19.0
19.0
19.0
19.0
19.0
19.0
19.5
‘9.5
19.5
19.5
19.5
19.0
19.0
19.0
19.0
19.0
19.0
IN. W.G.
BED
D.P.
20.0
20 • S
PB • 5
21.0
21.0
22.0
22 • B
22.0
22.0
22.0
22.0
22 • S
27.0
72 • 5
22 • B
22.0
23.0
22 • 5
22 • 5
23.0
22.5
2 1.5
2.0
22 • B
21.5
22.0
23 • B
2.5
23.0
23.0
23.0
23.0
23.0
23.0
23.0
23.0
23.5
23.5
24.0
24.0
28.0
24.0
24.0
24.8
REGEN
SP. GR. BED
D.P.
0.75 24.0
0.70 24.0
0.75 26.0
0.70 26.0
0.70 27.0
0.75 26.0
0.80 27.0
0.75 27.0
0.75 27.0
0.80 27.0
0.75 27.0
0.75 27.0
0.80 26.0
0.75 25.0
0.00 26.0
0.00 23.0
0.75 23.0
p.75 75.0
0.75 29.0
0.75 24.0
0.00 22.0
0.75 25.0
0.20 23.0
0.75 23.8
0.00 24.0
0.75 27.0
0.00 25.0
0.75 2 5.0
0.75 25.8
0.80 24.0
0.00 25.0
0.80 25.0
0.75 25.0
0.08 25.0
0.80 25.0
0.75 26.0
0.00 26.0
0.00 26.0
0.00 23.0
0.00 24.0
0.80 05.0
0.80 25.0
0.00 25.0
0.08 22.0
161
-------
6 of 7
Table D-7
Run 5 C&FB Pressures
DAY-TIME GASIFIER PRESSURES. IN. W.G. REGEN
GAS DISTRIB BED SP. CR. BED
SPACE D.P. - D.P. D.P.
23.1730 21.0 24.0 18.5 0.70 20.0
23.1830 21.5 24.0 20.0 0.70 22.0
P3.1930 21.0 24.0 21 .0 0.70 23.0
P3.2030 21.0 24.0 21.5 0.70 23.0
23.2130 20.5 23.5 22.0 0.70 22.0
23.2230 21.0 24.0 23.0 0.70 22.0
23.2330 21.5 24.5 23.0 0.70 24.0
24.0030 2 1.5 25.0 22.0 0.70 2.0
24.0130 22.0 25.0 23.0 0.70 24.0
24.0230 22.0 25.0 24.0 0.65 24.0
24.0330 22.5 25.0 24.0 0.70 24.0
24.0430 22.0 25.0 24.2 0.70 24.0
24.0530 22.0 24.5 24.0 0.70 24.0
24.0630 2.5 24.5 24.0 0.70 24.0
24.0730 22.0 24.5 24.0 0.70 24.0
24.0830 23.0 25.5 24.0 0.70 24.0
24.0930 94.0 25.0 24.5 0.70 24.0
24.1030 24.0 25.0 25.5 0.70 24.0
24.1130 23.0 25.0 25.5 0.75 24.0
P4.1230 22.5 24.5 25.0 0.70 24.0
24.1330 2 .0 25.5 25.5 0.75 24.0
24.1430 22.5 25.5 25.0 0.75 24.0
24.1530 22.0 25.5 24.5 0.75 24.0
24. 1630 22.0 25.0 24.5 0.75 24.0
24.1730 22.5 25.0 24.0 0.70 24.0
24.1830 22.5 25.0 25.0 0.70 24.0
24. 1930 23.0 26.0 25.0 0.70 24.0
P4.2 30 23.0 27.0 25.0 0.70 24.0
24.2130 23.0 26.8 25.0 0.70 84•0
24.2230 23.0 27.0 26.0 0.70 26.0
24.2330 23.5 27.0 25.0 0.70 27.0
25.0030 23.5 27.0 25.0 0.70 27.0
25.0130 23.0 27.0 25.0 0.70 27.0
25.0230 3.S 27.0 25.0 0.70 27.5
25.0330 23.5 27.0 25.5 0.70 28.0
25.0430 23.7 27.0 25.5 0.70 28.0
25.0530 24.0 27.0 25.0 0.70 28.0
25.0630 23.5 27.0 25.0 0.70 28.0
25. 0730 23.5 27.0 25.0 0.70 28.0
25.0 830 24.0 27.0 25.0 0.70 2R.0
162
-------
7 of 7
Table D-1
Run 5 CAFB Pressures
DAY.T!ME
GASIFIER
PRESSIiRES
IN. W.G.
REGEN
GAS
DISTRIB
BED
SP. GR.
BED
SPACE
D.P.
D.P.
D.P .
25.0930
24.0
27.0
25.5
0.70
22.0
25. 1030
24.0
27.0
24.0
0.70
26.0
25.1130
23.5
27.0
74.0
0.70
26.0
25.1230
23.5
27.0
26.0
0.75
25.0
25.1330
23.5
27.0
25.5
0.70
25.2
25.1430
24.0
27.0
25.5
2.70
24.0
25.1530
23.5
27.0
25.5
0.70
24.0
25.1630
23.0
27.0
26.0
0.60
26.2
25.1730
23.0
27.0
26.0
0.70
26.0
75.1030
24.5
29.0
25.5
0.70
20.0
25.1930
25.0
.9.0
25.0
0.65
25.0
25.2030
5.5
29.0
25.0
0.65
25.0
25.2130
25.0
29.5
24.5
0.65
25.0
25.2230
25.0
29.0
25.0
0.67
24.0
25.2330
25.0
29.2
25.
0.67
25.0
26.0030
25.0
29.0
25.0
0.70
25.0
26.0130
25.0
29.0
.5.2
0 . 70
25.0
26.0230
25.2
79.0
25.0
0.70
24.0
26.0330
25.1
29.2
25.0
0.70
2L . 2
26.0430
25.5
20.0
25.0
0.70
23.0
26.0530
25.5
72.5
25.0
0.70
23.0
26.0630
25.5
70.5
25.2
0.70
2 /4.0
26.0730
25.5
20.5
25.0
0.72
25.0
26.0830
25.5
20.5
25.0
0.70
25.0
26.0930
26.0
20.5
25.0
0.70
26.0
26.1032
26.2
22.5
25.0
0.65
26.5
26.1130
26.2
22.5
24.5
0.65
26.5
26.1230
26.0
22.2
25.0
2.65
26.5
26.1330
26.0
20.0
94.5
0.65
27.0
26.1430
26.0
22.0
25.0
0.65
22.0
26.1530
26.0
28.0
25.5
2.65
23.0
26. 1630
27.0
22.0
25.5
0.65
25.2
26.1730
26.5
22.0
25.5
0.65
6.0
26.1030
26.5
22.0
25.0
0.65
26.0
163
-------
1 of 8
Table D-8
Computed Solids Circulation
CAFB Rwi 5
SOLIDS CIRCULATIONS LB/HR
TOTAL SOLIDS
CIRCULATING TO FROM
DAY.HOUR CAO REGEN REGEN
2.0130 913 936 946
2.0230 796 821 827
2.0330 658 678 678
2.0430 876 897 901
2.0530 1103 1126 1126
2.0630 845 865 868
2.0730 744 762 766
2.0930 842 862 870
2.1030 1073 3095 1101
2.1130 830 853 861
2.1230 705 728 741
2.1330 705 728 732
2.1430 684 709 707
2.1530 680 704 705
2.1630 726 749 75)
2.1730 717 741 743
2.1830 574 596 594
2.1930 574 596 594
2.2030 626 647 657
2.2130 606 624 626
2.2230 670 694 694
2.2330 743 769 771
3.0030 734 758 763
3.0130
3.0230 644 670 668
3.0330 682 708 710
3.0430 702 727 730
3.0530 756 782 788
3.0630 775 801 807
3.0730 754 780 783
3.0830 780 806 809
3.0930 761 786 788
3.1030 810 834 838
3.1130 781 804 816
3.1230 743 767 774
3.1330 744 767 784
3.1430 585 609 610
3.1530 674 698 705
3.1630 664 688 693
3.1730 573 592 597
164
-------
2of 8
Table D—8
Computed Solids Circulation
CkFB Run 5
SOLIDS CIRCULPTION, LB HR
TOTAL SOLIDS
CIRCULATING TO FROM
DAY.HOUR CAO REGEN REGEN
3.1830 674 698 700
3.1930 638 660 663
3.2030 792 816 832
3.2)30 521 537 549
3.2230 688 732 714
3.2330 810 835 851
4.0030 813 839 847
4.0130 610 634 635
4.0230 524 550 548
4.0330 717 742 745
4.0430 807 833 836
4.0530 440 459 475
4.0630 611 636 646
4.0730 556 582 580
4.0830 590 616 616
4.0930 722 748 752
4.1030 699 724 725
4.1130 555 591 578
4.1230 747 769 775
4.3330
4. 1430 527 550 547
4.1530 607 631 633
4.1630 568 591 596
4.1730 660 684 694
4.3830 592 618 615
4.1930 491 513 511
4.2030 604 630 629
4.2130 583 606 602
4.2230 580 604 602
4.2330 659 684 685
5.0030 660 687 693
5.01 30
5.0230 693 716 730
5.0330 687 710 726
5.0430 593 618 60
5.0530 6S9 684 689
5.0630 631 656 658
5.0730 620 645 64 %
5.0830 613 637 638
5.0930 555 579 576
165
-------
3 of 8
Table D-8
2 puted Solids Circulation
CAFB Run
SOLIDS CIRCULATION. LB/HR
TOTAL SOLIDS
CIRCULATING TO FROM
DAY.HOUR C O REGEN REGEN
5.1030 601 625 625
5.1)30 574 598 596
5.1230 608 634 634
5.1430 615 640 638
5.1530 634 658 658
5.1630 682 7 0 706
6.0238 622 644 648
6.0330 530 552 552
6.0430 5 5 567 572
6.0530 519 539 540
6. 9 1630 439 459 457
6.917391 5919 5391 530
6.0830 427 448 442
8.2230 1569 1594 1616
8.2330 1109 1135 1155
9.0030 903 930 954
9.0130 1003 1031 1049
9.0230 1049 1079 1097
9.0330 1027 1056 1080
9.0430 997 1030 1042
9.0530 1050 1084 1093
11.1430 696 722 133
11.1530 722 748 765
11.1630 700 729 736
11.1730 755 783 797
11.1830 683 713 715
11.1930 725 754 761
11.2030 679 708 708
11.2130 796 825 831
11.2230 687 715 716
11.2330 549 574 573
12.0030 594 622 6)7
12.0230 643 670 674
12.0330 563 588 586
12.0430 781 812 809
12.0530 723 750 755
12.0630 720 747 748
12.0730 789 820 821
12.0830 715 743 745
12.0930 12) 748 751
166
-------
of 8
Table D-8
Computed Solids Circulation
FB Run 5
SOLiDS CIRCULATION, L8/Ht
TOTAL SOLIDS
CIRCULATING TO FROM
DAY.HOUR CAO REGEN REGEN
12.1030 755 783 787
12.1230 852 879 886
12.1330 751 779 78!
12.1430 773 80! 806
12. 1530 777 805 806
12.1630 648 674 681
12.1730 693 721 720
12.1830 735 763 768
12.1930 950 978 981
12.2030 911 939 942
12.2130 1028 1055 1061 ,
12.2230 702 732 735
12.2330 942 970 973
13.0030 908 935 938
13.0130 811 839 842
13.0230 657 686 686
13.0338 954 982 986
13.0430 973 1001 1004
13.0538 2.60 82.9 890
13.0630 949 978 983
16.0730 1182 1205 1225
16.0930 ‘708 730 749
16.1030 574 598 682
16.1130 825 842. 855
16.1230 739 766 774
16.1330 699 725 727
16.1538 693 718 7.6
16.1630 604 629 632
16.3730 611 636 636
16.3838 624 651 649
16.1930 620 646 648
16.2030 603 630 628
16.2130 1143 1166 1179
16.2230 733 758 760
16.2330 665 690 696
17.0030 598 62.4 629
17.0130 657 62.3 684
17.0230 797 2.24 827
17.0330
17.0430 797 824 827
17.0530 1479 1525 1531
167
-------
5of 8
Table D-8
Computed Solids Circulation
C&FB Run
SOLIDS CIRCULATIONa LR/ )fR
TOTAL SOLIDS
CIRCULATING TO FROM
DAY.HOUR CAO REGEN REGEN
17.0630 951 877 880
l1. ?30 643 669 671
17.0830
17.0930 684 710 712
17.1030 648 674 676
17.1130 639 666 665
17.1230 665 691 691
11.1330 641 667 667
17.1430 588 6)3 612
17.1530 734 760 762
17.1630 824 851 863
17.1730 738 745 750
17.1830 609 635 631
17.1930 716 742 745
17.2030 654 679 679
20.2030 1477 1503 1513
20.2 )30 1411 1438 1453
20.2230 868 895 902
20.2330 877 904 908
21.0030 1057 1085 1092
21.0130 972 998 1003
21 .0230 980 1006 1016
21.0330 1002 1029 1038
23.0430 873 898 902
23.0530 808 834 835
21.0630 917 943 9 8
21.0730 835 860 867
21.0830 746 770 770
21.0930 787 813 917
21.1030 787 8)2 818
2 3 . 3130 779 804 807
21.1230 698 721 720
21.1330 807 833 837
21.1430 718 742 743
21.1530 748 774 779
21.1630 5 18 600 592
21.1730 693 717 717
23.3830 530 552 547
21.1930 577 600 594
21.2030 697 722 725
168
-------
6 of 8
Table D-8
Computed Solida Circulation
CAFB Run 5
SOLIDS CIRCULATION, LB/HR
TOTAL SOLIDS
CIRCULATING TO FROM
DAY. 1-(OUR CAD RECEN REGEN
21.2130 606 630 634
21.2230 586 611 614
21.2330
22.0030 607 630 633
22.0)30 611 636 630
22.0230 699 724 730
22.0330 538 560 557
22.0430 563 586 507
22.0530 6 44 670 671
22.0630 605 630 6.9
22.0730 676 700 104
22.0830 565 589 585
22.0930 653 678 604
22.1030 657 6 1 690
22.1130 662 687 691
22.1230 505 525 524
22 . 1330 54 600 606
22.1430 58 ) 604 603
22.1530 600 632 633
22.1630 626 650 655
22.1730 609 633 633
22.1830 632 655 66 )
169
-------
7 of 8
Table D—8
Computed Solids Circulation
CAFB Run 5
SOLIDS CIRCuLATION. LB/HR
TOTAL SOLIDS
CIRCULATING TO FROM
DAY.HOUR CAO REGEN REGEN
23.1730 9)9 944 960
23.1830 754 780 789
23.1930 737 764 768
23.2030 775 800 8)7
23.2130 663 688 697
23.2230 639 664 671
23.2330
24.0030 645 670 674
24.0130 708 733 740
24.0230 654 679 684
24.0330 641 665 669
24.0430 671 696 701
24.0530 669 693 699
24.0630 678 703 710
24.0730 614 639 638
24.0830 663 688 689
24.0930 716 741 750
24.1030 711 735 744
24.1)30 7 3 748 751
24.1230 698 723 727
24.1330 670 695 696
24.1430 693 718 722
24.1530 683 708 711
24.1630 682 708 713
24.1730 689 714 719
24.1830 684 709 713
24.1930 672 697 702
24.2030 738 764 772
24.2130 776 802 812
24.2230 742 769 773
24.2330 826 852 864
25.0030 842 871 883
25.0130 815 841 855
25.0230 782 808 818
25.033g 755 78) 785
25.0430 771 798 803
25.0530 774 800 805
25.0630 780 807 812
25.0730 758 785 788
25.0830 786 813 820
1170
-------
8 of 8
Table D-8
ComDuted Solids Circulation
CARB Run 5
SOLIDS CIRCuLATION, LB/HR
TOTAL SOLIDS
CIRCULATING TO FROM
DAY.HO IJR GAO REGEN REGEF’3
25.0930 794 82 1 829
25.1030 789 8)5 821
25.1130 793 820 823
2 5 . 1230 7)2 738 742
25. 1 330 690 7)6 716
P5.1 3Q’ 691 7)5 722
25.1530 748 773 782
25.1630 751 777 783
25.1730 728 755 759
25.1830 754 780
25.1930 794 750 753
25.2030 741 766 771
25.2130 710 735 730
25.2230 743 7( 774
25.2330 718 743 7.9
26. 0 030 648 673 674
26.013 0 697 722 726
26.0230 764 790 800
26.0330 798 825 833
26.0430 837 864 072
26.0530 7 I 787 789
26.0630 76) 787 7.9
26.0710 766 79) 804
26.0830 737 763 770
26.0930 743 770 776
96.1030 699 726 730
26.1130 683 708 71 /i
26.1230 646 671 671
26.1330 730 757 762
26.1430 715 741 748
26.1530 715 741 747
96.1630 716 741 748
26.1730 97 722 726
26.1830 722 748 753
171
-------
Table 0-9
Compoiltion of CAFR SoJ.lde Run 5 (3u1 ur )
nm.
Day. Hour
sampling Position:
Ossttiar
Upper
Gasirier Lower
Regenerator
Roller Stack Cyclone
ElutriatorCoerse Elutriator Finea
Rege. erator Cyclone
I d
-3
I ’ . )
7.1530
3.2200
5.10 445
6.0730
12.1600
12.1800
13.0400
13.0600
17.1130
17.1800
21.0700
21.1800
22.0715
22.1745
25.0530
25.1400
26.0400
26.1000
26.1800
Total S
% Ut
2.89
.66
2.88
2.80
2.76
2.53
3.11
3.98
2.69
2.85
2.90
2.91
2.74
2.52
2.74
2.70
2.87
u as
Sulphate
% Ut
0.22
0.37
0.43.
0.31
0.31
0.26
0.21
0.22
0.28
0.19
0.48
0.48
0.24
0.42
0.29
0.35
0.38
TO aL S
% Ut
2.64
2.72
.48
3.72
2.95
7.29
2.55
2.56
3.29
3.29
2.64
2.97
2.82
2.92
2.48
2.52
2.87
2.88
2.79
$u phate
0.23
0.38
0.142
0.28
0.39
0.38
0.79
0.43
0.36
0.21
0.43
0.19
0.47
0.38
0.53
0.45
0.40
0.87
0.29
Total
% wt
1.82
1.81
1.59
1.84
1.85
1.80
2.06
1.83
1 • 00
1.78
1.75
2.21
1.91.
1.88
1.72
2.17
1.96
1.96
S 58
Su , ate
1.27
1.19
1.29
1.48
1.12
1.18
1.33
1.28
0.80
1.19
0.80
1.29
1.26
1.35
1 • 12
1. 37
1.26
1.27
Total 5
% Ut
5.15
4.99
3.34
3.38
3.26
3.70
3.69
3.50
2.88
4.03
4.27
4.52
5.16
4.62
3.85
4.08
14 • 10
4.30
4.90
o as
Su4pI atte
1.88
3.20
1.74
2.11
2.58
2.93
2.44
1.96
2.56
1.41
1.18
1.37
1.71
1.66
2.24
1.80
2.31
1.47
2.19
S
% wt
6.48
8.04
1.96
1.74
3.26
2.96
1.68
1.75
1.44
1.56
2.94
3.59
7.53
3.36
3.08
3.18
3.57
7.5*
‘.37
Su te
2.62
Ut
4.26
0.32
1.65
8.39
0.17
1.53
-
-
1.57
5.24
0.71
1.73.
4.56
0.25
1.53.
3.72
0.23
1.38
5.13
0.28
1.34
3.57
0.25
1.24
4.40
0.27
1.26
4.92
0.28
1.94
6.12
0.24
1.79
2.09
7.28
7.45
0.21
0.48
2.26
7.09
0.44
2.014
7.86
0.26
230
6.21
0.81
2.31
6.94
0.27
2.22
2.19
6.68
7.46
0.32
0.36
Total S
7.85
7.17
4.97
2.34
4.41
4.66
4.03
5.12
14.81
5.82
5.57
6.70
630
6.89
6.94
4.9*
2.75
6.47
S Se
Su p te
3.85
0.25
0.32
0.146
0.30
0.29
0.37
0.25
0.29
0.35
0.36
0.71
0.66
0.20
0.55
0.32
0.36
0.39
Total S
% Ut
8.11
4.59
4.69
4.78
14.59
2.14
4.21
4.24
5.11
5.19
5.82
6.21
6.62
5.75
5.79
5.97
5.88
6.10
5 as
S a
7.148
3. 1 40
3.47
2.91
1.26
2.71
7.07
7.17
3.27
2.99
14.25
44.55
3.83
3.68
4.00
3.83
4 • 11
-------
Table D-l0
Composition of CAFE Solids Run 5 (Carbon )
Time
Da cr. Hour
Gasifier
Upper
Gasifier
Lower
Regenerator Boiler
Stack
Cyclone
Elutriator
Coarse
Elutriator
Fines
Regenerator
Cyclone
3.15:50
3.2200
5.10145
6.c 1730
12.1600
12.1800
17.01400
13.0600
17.1130
17.1800
21. 1Y700
21.1800
22.0715
22.1745
25.05)0
25.11400
26. 0400
26. iooo
26.1800
Carbon
wt %
0
0.02
0.19
0.114
0.12
0
0.13
0.28
0
0.13
0
0.08
0
0.03
0
0.03
0
Carbon
wt %
0
0
0
0.09
0.04
0.06
0 .06
0.11
0.02
0.13
0
0.14
0
0
0
0
0.014
0.07
0.04
Carbon
wt %
0
0.02
0
0
0.04
0
0
0
0
0.50
0
0
0
0
0
0
0.06
0
0
(;artDon
0.32
0. (17
0.12
0.17
0.114
0
0.05
0.19
0
0.07
0.43
0.56
0.13
0.29
0
0
0
0.19
0.12
Carbon
wt %
8.8)
9.09
0.27
0.65
0.50
0.52
0.45
0.55
0.15
0.24
3.21
2.90
1.61
0.65
0.41
0.50
1.66
0
0.89
Carbon
wt %
1.1l 1
9.147
1.7)
0.92
0.50
1 • 20
0.38
0.76
0.55
2.35
3.91
2.82
2.67
1.149
1.26
1.99
1.73
2.37
Carbon
wt %
19.97
26.05
1.51
3.18
2.49
2.71
2.91
0.87
1.79
8.54
9.19
5.56
5.78
2.35
4.51
5.85
5.91
5.114
Carbon
wt %
2.81
0
0.09
0.146
0.08
0.1 1-5
0.11
0.02
0
o.cyr
0.20
0.05
0.04
0
0.19
0.07
0
-------
Table D-ll
of CAPH Solids Run 5 (Ignition at9OQ°CJ
Sampling Position:
Boiler
Stack
Cyclone
m.utriator
Coarse_-
Elutriator
Fines_______
Regenerator
Cyclone
Time LoSe uarn Loss Gain Loss a1n Loss Gain Loss Gain
Day, Hour wt% wt wt j wt wt
3.1530 - 3.39 7.91 - 0 0 11.5 1 4 - - -
3.2200 1.03 - 20.146 - - 5,53 21.2 - - 0.80
5.10 1 45 - 0.147 15.62 - - - - 1 4.69 - 0.86
6.0730 0.24 - 7.56 - - 5.67 - 5.45 - 1.24
12.1600 0.10 - 4.77 - - 6.78 - 4.14 - 1.13
12.1800 - 0.147 7.20 - - 5.97 - - - 1.65
13.0400 - 1.40 3.91 - - 6.55 - 3.90 - 1. ’Y7
13.0600 - 1.90 1.75 - - 5.314 - 2.82 - 1. 146
17.1130 1.18 - 2.22 - - 6.52 - 7.27’ - 1.18
17.1800 - 3.(Y7’ 2.89 - - 5.97 - 6.10 - 1.87
21.0700 - 2.53 5.32 - - 7.22 0.39 - - 1.33
21.1800 - 1.68 14.12 - - 5.13 0.88 - - 2.26
22.0715 - 1.6 14 3.19 - - 8.00 - 5.70 - 1.35
22.17145 - 0.75 3.57 - - 7.97 - 3.50 - 1.146
25.0530 — 1.32 3.30 - - 10.82 - 6.11 - 1.52
25.1400 - 2.31 3.46 - - 8.39 - 6.72 - 1.53
26.01400 - 1. 1l 9 5.2 14 - - 8.79 - 14.47 1.40
26.1000 - 2.18 4.95 — - 7.78 - 3.114 1. 1 414
-------
Table D-12
Lime Metal-s Content (Th&n 5 1
Time Samp ling Vanadiu$i Sodiug Niokel
Day-, Hour Position
3.1530 Gasifier Lower 5000 259 477
Gasifier Upper 4800 296 454
Regenerator 8200 375 451
7.2200 Gasifier Lower 4900 215 849
Regenerator 6000 392 653
5. 1045 Gasifier Lower 3200 49k 454
Gasifier Upper 4000 522 434
Regenerator 3400 580 440
6.(YT30 Gasifier Lower 4400 530 552
Regenerator 5000 541 622
12.1600 GasifieD Lower 2600 369 3 76
(lasifier Upper - 2500 361 767
Regenerator 2000 330 701
12. 1800 Gasifier Lower 2800 300 413
Gasifier Upper 2700 317 337
Regenerator 2800 317 351
13.0400 Gasifier Lower 3100 466 326
Gasifier Upper 3200 469 360
Regenerator 4200 510 470
17.0600 Gasifier Lower 3100 425 349
Gasifier Upper 3100 492 354
Regenerator 3600 487 363
17.1130 Gasifier Lower 1700 330 325
Gasifier Upper 2000 415 280
Regenerator 1300 300 265
174800 Gasifier Lower 2800 415 320
Gasifler Upper 2300 430 320
Regenerator 3000 505 360
21.(7700 Gazifier Lower 2200 205 255
Gasifier Upper 1900 185 266
RegeneratOr 2500 195 315
21.1800 Oasifier Lower 2900 285 365
Gasifier Upper 3300 215 420
Regenerator 3600 255 406
22. (7715 -Gasifier Lower 4200 295 365
Gasifier Upper 4700 295 405
RegeneratOr 5300 280 470
22. 3.745 Gasifler Lower 6300 275 585
Gasifier Upper 5700 320 495
Regenerator 5300 280 480
25.0530 Gasifler’ Lower 5100 280 500
Gasifier Upper 5000 245 470
Regenerator 5700 235 625
25.1800 Gastfier Lower 5000 745 465
Gasifier Upper 5800 230 495
Regenerator 5400 245 475
26.0800 GasifieD Lower 63.00 180 580
GasifieD Upper 6300 155 500
Regenerator 6500 240 620
26.1800 Gasifier Lower 5700 195 561
GasifieD Upper 5800 185 520
RegeneratOr 5600 245 580
175
-------
Table D-13
Calcium 0,Ude and Silica Contents of Bed during Run
Sampling Position:
Time -
Day, Hour
Gasifier
UD ?r
CaO 810
Wt % Wt
(}aeifler
L Mer
CaO S W
Wt % Wt
Regenerator
CaO S10
Wt % Wt
Boiler
CeO 810
Wt % Wt
Stack
Cyclone
CeO Sb
‘dt % wt
Elutriator
Pines
CeO sip
Wt % Wt
Elutriator
Coarse
cao Sb -
Wt $ Wt
— —
Regenerator
Cyclone
cao— Sb
Wt % Wt
—
5.W45
76.7
15.14
73.6
15.0
75.9
16.3
147.7
11.2
714.6
21.1
63.8
22.4
-
-
62.7
19.1
6.cq’ o
-
-
74.5
1.7.6
75.4
16.4
66.14
1.5.5
67.5
16.6
64.3
19.4
61.1
20.6
61.3
18.4
12.1630
75.1
17.2
75.3
17.1
72.8
18.8
69.6
17.6
60.8
18.3
63.5
20.2
70.1
18,5
64.3
18.2
12.1800
714.5
17.2
76.7
16.5
79.5
17.9
67.8
17.7
59.5
18.8
..
-
68.4
18.6
65.8
18.8
13.0400
73.6
16.1.
75.5
16.2
-
-
66.0
18.2
63.2
19.8
62.5
20.8
66.0
20.1
-
-
13.0600
76.4
18.4
75.6
18.4
75.5
17.0
69.6
3.9.4
614.8
20.2
63.2
20.8
74.2
19.0
65.1
19.8
17.1130
714.5
1 .6.1
79.7
14.5
87.2
9.2
71.0
17.9
67.8
19.3
67.6
17.4
72.6
16.2
65.6
17.4
17.1800
75,8
15.)
79.4
13.3
72.8
15.6
70.2
1.6.)
614.7
19.5
66.3
18.1
7],.2
15.9
66.7
18.2
21.0700
86.7
5.0
89.7
14.8
-
-
-
-
-
-
82.5
1.0
-
-
84.7
0.8
21.1800
87.8
14.7
84.7
14.5
- .
-
-
80.8
1.2
85.0
0.8
-
-
-
-
22.0715
87.4
14,2
85.8
14.9
-
-
-
-
86.8
0.6
-
-
-
-
-
-
22.1745
87.8
3.3
87.1
3.14
-
-
-
-
78.8
0.8
-
-
-
-
-
-
25.0530
-
-
96.3
2.3
96.3
1.9
96.1
1.5
89.6
0.5
94.4
0.8
95.0
1.0
87.2
0.5
25.1470
-
-
97,1
1.6
-
-
-
—
-
—
-
—
-
—
-
-
-------
Table D-14
of Qasifier Bed Run 5
f
Sample No
Gasifier
Bed
+2800
wt.%
2800
1400
ii.
Wt.%
1400
1180
Wt.%
1880
850
w.
wt.%
850
600
wt.%
600
250
i
wt.
250
150
Wt.%
lO0 i
Wt .2
Less
Than
100
Total
Day/Hr.
51225
2.53
6.33
22.15
33.55
22.87
12.61
0
0
0
100
5.1045
51233
0.44
41.28
13.02
19.87
14.13
ii.O4
0.22
0
0
100
6.0730
51236
1.37
45.21
15.07
20.55
10.96
6.85
0
0
0
100
12.1500,
51245
1.36
45.92 ‘
13.27
20.41
12.93
6.11
0
0
0
100
12.1800
50001
50029
1.70
3.53
48.02
49.17
13.35
13.07
19.60
17.63
11.65
10.17
5.68
6.22
0
0.21
0
0
0
0
100
100
13.0400
17.l8 0
50047
0.28
34.27
11.89
18.88
15.66
19.02
0
0
0
100
18.1100
50048
0.22
28.40
12.12
20.99
16.95
21.32
0
0
0
100
i.i6oo
50055
0.68
41.91
11.78
18.08
12.74
14.66
0.14
0
0
100
21.1800
50061
0.24
38.59
12.86
18.20
13.59
16.28
0.24
0
0
100
22.0730
50073
0.72
41.17
11.89
18.38
13.96
14.08
0
0
0
100
22.1745
50114
0.66
42.90
13.20
19.64
12.87
10.40
0.17
0.17
0
100
25.0530
50123
0.64
44.16
13.25
19.87
13.56
8.52
0
0
0
100
25.1400
50133
0.63
47.79
14.11
18.95
12.21
6.32
0
0
0
100
26.0400
50149
0.95
1 4.62
14.24
19.94
17.29
6.96
0
0
0
100
26.1000
Sieve Analyses
-------
-J
Table D-15
Sieve Analyees o e enerator Bedjfiuri5 )
Sample No
Regen Bed
Run 5
+2800
lb
Wt.%
2800
1400
Wt.%
11400
1180
Wt.%
1180
850
w
850
600
600
250
250
150
.%
150
100
-100
Total
Wt., ,
b
ay 1 r.
51197/73
O.6
39.30
10.22
18.53
16.29
15.02
0
0
0
100
2.1730
51198/73
0.33
31.146
8.914
20.2
18.514
20.5)
0
0
0
100
2.19)0
51232/73
1.15
141.67
12.93
22.141
14.66
7.18
0
0
0
100
6.0730
51239/73
O.6Ji
47.92
13.142
18.85
12.14
7.03
0
0
0
100
12.1600
50006/73
2.1l4
145.99
11.2)
18.72
13.36
8.56
0
0
0
100
13.0600
501214/7)
0.46
1iO.2 )
13.79
20.2)
1L#.25
10.80
0.23
0
0
100
1.5.11400
50025/73
500)6/73
1.1414
0.143
38.li.6
17.93
10.58
7.34
16.83
17.06
12.50 20.19
18.79 27.65
0 0
2.16 1.73
0 100
6.91 100
17.11)0
17.1800
50044/7)
0.1ii#
34.37
9.31
22.62
20.40
12.142
0.244
0
0
100
21.0700
50051/73
0.21.
27.66
10.614
17.98
16.49
25.85
1.06
0.11
0
100
21.1800
50059/73
0.14)
32.8)
11.59
18.67
15.45
20.82
0.21
0
0
100
22.07)0
50071/73
0.36
28 8 )
10.85
18.714
16.94
24.114
0.18
0
0
100
22.17145
50012/73
0.31
35.614
12.75
20.28
15.67 114.90
0.15 0.15
0.15
100
25.0530
50131/73
0.57
40.92
13.58
20.84
114.15
9.75
0.19
0
0
100
26.0400
50150/7)
0.74
41.98
13.09
19.26
16.79
7.9
0.25
0
0
100
26.1000
-------
GASIFIER DEPTH
G AS I ER I EMf ER I EM PER ATURE
2—1200 2400 3-1200 2400 4- 1200 2400 5—1200
RUN DAY-TIME
2400 6-1200
D I. SHEET.) .
SULPHUR REMOVAL EFFICIENCY
-j
4
>
z
—ju-
I —
z
CI)
w.
L u 0
-J
0
-J
2
LIME REPLACEMENT RATIO
DENBIGHSHIRE STONE
0
25
20
15
oIz
Lu—
‘900
Lu
I .-
8CR 1691 STONE
850
800
-------
IF . ,’
T\U RE FDCI CY
LINE REPLACEMENT RATIO
8CR 1691 STONE
GASIFIER DEPTH
GASIF1ER TEMPERATURE
I I I I
-i -—I
Jv ’
4 8CR 1691 STONE
__ - t
8-2400 9-1200 II -2400 12-1200 2400 6-1200 2400 17-1200
90 —
80 —
70—
60—
50 —
2—
C . )
z
U
C)
L i.
L i-
U)
w2
(Jo
V
I
STONE-’
4—- 8CR 1691
0
W25
20
wz
IS
U) u900
4.
(Do.
850
I -
I \
RUN DAY—TIME
DI, SHEEI2 .
-------
tOO
9O
80
z
70
jtL .
uJ 6 O
50
I-
z
() 0
a.
0 ,
H
o -J
H
0
25
c 2O
w -
a. 15
U-
“900
40
850
800
SULPHUR REMOVAL EFFICIENCY
LIME REPLACEMENT RATIO
DENBIGHSHIRE STONE
GAS F(ER TEMPERATURE
2400 2 1-1200 2400 22-1200 2400 24-1200 2400 25-1200 2400 26-1200
RUN DAY-TIME D.ISHEET.3 .
-------
( 1
0
In>
0 —
6
100
5
U :
wWIIOO
1050
,O00 I I I
REGENERATOR GAS $02 CONCENTRATION
REGENERATOR SELECTIVITY CoS TO CoQ
REGENERATOR SUPERFICIAL GAS VELOCITY
RFGfl RATOR TEMPERATURE
2-Iaoo
2400 3-1200 2400 4-1200
RUN DAY—TIME
2400 5-1200 2400 6- 2OO
-------
z
0
0>
00
uJ
c , ) - 3
U
U
U)
U-
I- .
I.- - ’
LgJ
U)
z 4
tAJ 0
U i
lI-J
00
U)>
z
0-
I. -, --
c d
U i , .-
zz
WU
U
a,
8-2400 9-1200 11-2400 12-1200 2400
RUN DAY—TIME
16-1200 2400 17-1200
D,2.SHEET.Z
-------
00
(I , >
z
(I,
0
U
I-
“S
-J
(I)
I .- .
2400 21-1200 2400 22-1200 23-2400 24-1200 2400 25-1200 2400 26-1200
RUN DAY —TIME
D 2.SHEET.3
-------
APPENDIX E
Gas Ifier Heat Balance
185
-------
Appendbc E
Gas ifler Heat Balance
The heat released in the gasifieD under partial oxidation conditions
depends on the air fuel ratio and on the relative extents of the various
competing reactions within the bed.
As a first step in development of a eathematical model to predict gasifier
operating conditions, a series of heat and material balance equations have
been prograimsed to compute the heat release rate in the gas ifier from measured
fl i rates and temperatures. A second set of equations is used to estimate the
heat release rate from heats of reaction and from estimates of the relative
amounts of hydrogen and carbon oxidised and the CO/C02 ratio produced.
List of Variable Names
Variable Meaning Units
F lb,4 r
A Sew
D inches
3 microns
3’l’NRT 1b44r
It Deg.C
TR Deg.C
V ft/sec
C T wt.%
H T wt.%
SIVT wt.%
CP OPCT wt.%
O2INFG Vol.%
3T IR lb 1 Br
O SC 4
CO2FO Vol.%
IN Deg.F
O2 D lb Mole/Hr
RATIO Dimensionless
i oe Dimensionless
TO I)eg.C
HCCO2 Btu/lb mole C
HOCO BttVlb mole C
i . H2i i 2O Btu/lb mole He
lb Mole 09Jl00 lb oil
a percent
Deg.P
IX !R Btvltr
Fuel hate
Air rate
Gasifier bed depth
Average bed particle size
Fresh stone feed rate
Gas ifier bed temp
Regenerator bed temp
Superficial gas velocity
Carbon in fuel oil
Hydrogen In fuel oil
Sulphur in fuel oil
CaO in stone
Oxygen In flue gas
Lime circulation rate
Flue gas inlet rate
Carbon dioxide in flue gas
Temperature of inlet plenum
Oxygen feed to gas ifier
CO/C02 ratio produced in gasifier bed
combustion
Fraction of C02 in CO and C02 produced
in gasifier bed combustion
Gasifter temperature
Heat of Combustion for C to C02
Heat of combustion for C to CO
Heat of Combustion for 112 to ikO
Oxygen needed for stoichiometric
combustion of fuel oil
Oxygen fed as percent of stoichiometric
Regenerator temperature
Heat transferred from regenerator to
gasifier by circulating lime
186
-------
Meaning
Units
Btu/Hr
B’tu/Hr-CR I
Dimensionless
Dtu/Hr
Bti r
Heat required to raise temperature of Btu/Hr
oil from inlet to gasifier temperature
Heat required to crack oil to H2 + C
Heat required to raise inlet air from
plenum temperature to 1600 Deg.P
Heat required to raise inlet air from BtW’Hr
plenum temperature to gas iuier temp-
erature
Ratio of CaO to S fed
Heat required to calcine limestone
Heat required to raise temperature of
C02 from limestone to gasifier temp-
erature
Heat required to raise nonvolatile part BtWHr
of limestone to gasifier temperature
Heat lost from gas if ier bed through walls BtWJ3r
Net heat input to gasifier except for BtiVHr
reaction and flue gas contributions
Constant in equation for percent of
carbon oxidised
Portion of fuel carbon oxidised
Amount of fuel carbon oxidised
Heat released by oxidising carbon to
C02
Heat released by oxidising carbon to CO
Heat released by oxidising hydrogen
to H20
Heat required to raise flue gas from
plenum to gasifier temperature
Heat released by combustion in gasifier
BTUBRN x io-6
Heat released per lb oil by combustion
in gasifier
Heat released per mole oxygen by
combustion in gasifier
Amount of fuel hydrogen oxidised in
gasifier
Portion of fuel carbon oxidised Percent
Variable
BTUOIL
BTUCT*(
BrUAIR
C
UCAL
BTIJCO2
BTUSTN
R JWS
B FXD
AA
C OXDP
RTUCC O2
BTUCCO
HPUH2 O
RTUFG
BTUBI*
HTREL
RPULB
HOXD
HOXLt’I
H OXDP
CO2M
COM
02H20
HOXLM2
H OXDP2
HDFL
Dimensionless
Percent
lb Mole/100 lb fuel
BtW100 lb fuel
Btu/l00 lb fuel
BttVlOO lb fuel
Bti /Br
(Bt Ir) x 10-6
Btu/lb oil
Btu7lb mole
lb Mole/laO lb fuel
H0XE* and HOXDP are based on heat
balance calculation
Amount of fuel carbon oxidised to C02 lb Mole/100 lb fuel
Amount of fuel carbon oxidised to CO lb Mole/lOO lb fuel
Amount of oxygen converted to H20 lb Mole/lOO lb fuel
Amount of fuel hydrogen oxidised lb Mole/lOO lb fuel
Portion of fuel hydrogen oxidised Percent
HOXL*12 and HOXDP2 are based on material
balance calculation
Heat release per mole oxygen by Btu/lb mole
combustion in gasifier, calculated by
stoichiometry and reaction heats
187
-------
Thermal Equations
The thermal equations are linearized forms which adjust for departure
from tabulated values at 1600 Deg.F.
HCCO2 = 169790 + .4)4 *
IICCO 48525 + 7375 * (Ta. - 1600)
HH2H2O 106941 + 1.327 * (Ta -l600)
The air heat equations are based on published enthalpies at 100°F arid
1600°F and on specific heats (Cp) at these levels.
(60 mm lbNole ) -(H + C 4 ,100 TAiN-loo))
31fl Enthalpy change to heat air from inS temp. to 1600°F.
The term 60 converts from C ?4 to lb Mole4ir
.1579 * (15057 -(3825 + 698 *(TAIN_100.)))
3 UATh A *( + ( * * (P0-1600))
1’tJAIR A t(WA + .0439 * (pa - 1600) )
The enthalpy change to heat the flue gas is based on a flue gas of
asstzmed composition:
2.44% 02; 75% N ; 10.1% I1 0; 12.5% CO 2
Weighted average values of enthalpies and specific heats were combined
to obtain the final equation.
BTUP 0 * ( * (H 1 - H)O0 + pi6 00 * (Tal600) + Cpyio (300-PAIN))
ffl9JF } a * .1579 * (10719 + 8.902 Ta - 14243 - 7.423*TAIN + 2227)
1’U FG a * (1.24056 * Ta - 1.172* TArN - 204.79)
Heat of limestone calcination is taken at 1410 Btu/1b of CaO
H1 JCAL = STh ’ * * .01 1410
AL rNm ’ * CAOPcT * 14.10
188
-------
Heat to raise temperature of C02 liberated in calcinatiori.
H UC02 * CAOPCT * .01 *
(Hi6oo + C i 6 (Ta- 1600))
HPUCO2 STNR 1 2 * CAOPCT * . * + 17.36 (TG-1600))
* CAOPCT * (7.203 + .00239 * (TG-1600))
Heat to raise temperature of nonvolatile portion of stone.
ruST1 (STNWf) (Fraction Nonvolatile) * (Hi9j + Cp 1 6 00 * (TG-1600))
TU$TN $TNRP * (1 1 4 * CAOPCT * (362.9 + .28k * (To-1600))
56 100,
BTTJSTg STNRT * (1- .00786 * cA0PCT) * (3629 + .28k * (TG-1600))
Heat supplied to oil.
Heat uptake by the oil is made up of sensible heat, latent heat,
and heat of cracking.
Sensible beat and latent heat are included in a linear equation based
on the oil enthalpy at 1200 °F and the specific heat of the oil between 1200°F
and gasifier temperature.
ffl’U OIL * (780. + 0.8k *
The heat of cra cking Is assumed to be 600 BtW’lb
- F * 600
The heat loss from the gasifier bed is estimated from an anaLysis
which indicated that the effective product of thermal conductivity and
bed area is equal to 0.2k2 x D. The heat loss equation Is therefore:
ruLos . o.2uI 2*D* (Ta. 7Q)
Material Balance EquatIOnS
Oxygen fed.
O2FED 0.21 * 60 * 60 * O2INFG
A+ G*
O2FED = .0332 * A + .00158 * G * O2INFG
189
-------
CO/CO2 Ratio
The ratio of CO to C02 produced by partial combustion of oil in the fluid
bed is a complex function of themodynamics. reaction kinetics, arid contacting
in the bed.
Analytical results indicate that the values obtained are much lower than
equilibrium for the temperatures encountered.
It is believed that considerable C02 forum by reaction of carbon on lime
in the highly axidising region near the air inlet nozzles and that insufficient
contacting time is available for this C02 to reach equilibrium with CO and
carbon In the upper portion of the bed.
The equation used here is an empirical equation relating CO/C02 ratio to
temperature, based on gas analysis data obtained in the batch CAFE reactors
during the phase I study and reported in Appendix D of reference (1). The
equation is;
RATIO 2.91 E- * EXP(9.76 E-3 *
Fraction of Carbon 0x11&sed
In the current analysis it is assumed that components of the flue gas,
recycled for temperature control, do not react in the gasifier. The validity
of this assumption remains to be verified. With this assumption, fixing the
C0/G02 ratio and the fraction of feed carbon oxidised also fixes the fraction
of feed hydrogen oxidised since the oxygen fed must appear as CO, C02 or
H20.
The fraction of carbon oxidised is based on an emptrical equation first
derived from batch unit studies but modified to Improve match with the heat
results of the continuous pilot plant runs. This equation is:
COXDP =EXP(-AA) “ (R .9 2) * ( v Pl3, )
A value of 8. i. for (AA) has given the best match between measured and
calculated heat release rates.
The moles of carbon oxidised per 100 lb of fuel is the product of the
fuel carbon content and fraction oxidIsed.
co a*i .! * c r * 0XDP
12 100
coxi = 8.333 E- * CPCT * COXD ’
190
-------
Calculated Heat Release of Fuel Oil
The calculated heat release per 100 pounds of fuel is given, for each
element, by the moles of that element oxidised multiplied by the heat of
combustion of that element.
BruCCo2 HCCO2 C0XT A * FRCO2
ZIUCC O = HCCO * COXI* * (1 - Ff 1002)
BrUH2 O = HH2H2O * HOXDM
The total heat release is the sum of that for the elements.
“Measured” Heat Release
The “measured” heat release is the difference between the sensible heat
inputs to the gas ifler and the heat required to raise the products to gasifier
temperature.
BTUBRN = TUFG ÷ BTULOS + HTUCAL ÷ BTUOIL + BTUCHK + TUAIB + BTUSTN + BTUCO2
- H UCIR
Hydrogen Oxidised
The amount of hydrogen oxidised can be estimated from the heat balance
or from the oxygen balance.
Derivation of the heat balance equation for HOXW is as follows:
Heat release per mole 02 Heat of Combustion of elements
02 needed to burn elements
HOXD = BTUCCO2 + BTU000 + BflJH2O
* (HOXL 4 + COXDM * (]. + FRCO2)))
Also, from the measured heat release in the gasifier,
HOXD = ml ’UBHN/ 02FED
$ibstituting, HPUH2O = HOXE 4 * 1ffl2B20
gives,
HOXD = H UCCO2 + BTUCCO + HH2H2O * HOXDM
* (uoxDM + COXDM * +FRC0 ) )
191
-------
Rearranging and solving for HOX t4 gives:
nomi (Br1 co2 + m’jcco - * (1. + FJ O2) )*uoxD )/ ( .5*fl HIi2W 3)
Material Balance bydropen oxidised
Since all oxygen fed to the gas ifler is assumed to make CO , CC 2 , or
R20, fixing the amoumts of CO and C02 by empirical expressions also fixes the
quantity of H20.
CO2M COXL 4 *
COM COXDM - CO2M
02R20 O2PED * 100/F CO2M - COM/2
HOXD 02H20 * 2
Calculated Heat Release Per Mole Op
A purely calculated heat relea8e per mole oxygen can be computed for
comparison with the “measured” value. For this purpose a value of BTUH2O is
calculated using the material balance value of hydrogen oxidised.
B’ruH2o HOXI)42 *
then,
HIZL = .01 * * (BIUCCO2 + BrLXC0 + BTUH2O )/02FED
These heat and material balance equations have been used in several
computer programa to analyse experimental data and to predict new operating
conditions. Table E-1 list the Fortran statements of programe JHTROC which
calculates tk gasiuler heat release rrom experimental operating conditions
and compares the value with the predicted value. Results of calculations on
CAFB data appear in Table E-2.
192
-------
Table E-l Fortran Listing of
CAFB Heat Balance Programme
JHTHOC
)OOC CALCULATE HEAT OF COMBUSTION FROM CAFB CONDITIONS
hOC PROGRAM NAME JHTHOC
120 PRINT 600
130 DIMENSION P(25)
140 STiLE JHTFZI
150 50 CONTINUE
160 READ (I) (PCI), 1=).17)
170 IF CENOFILE 1) 10. R0
180 10 CONTINUE
190C DEFiNE VARIABLE VALUES
200 100 F = Pit) ‘FUEL RATEs LB/HR
210 A P12)
220 D = p1.3)
230 S P14) ‘AVERAGE PARTICLE SIZE, MICRONS
240 STNRT P15)
250 TC = P(6) ‘GASIFIER TEMP DEG C
260 TR = P17) ‘REGENERATOR TEMP DEC C
270 V P18) CJAS VELOCITY (INITIAL ASSUMPTION)
2g0 CPCT = p19) CARBON TN FUEL. WT. PCI
290 HPCT = P110) ‘HYDROGEN IN FUELS WI. PCI
300 SPCT = plfl) ‘SULFUR IN FUEL’ WI. PCT
310 CAOPCT = PC 12) CAO IN STONE’ WI. PCT
320 0119F6 = P 113) ‘02 IN FLUE GAS* VOL PCI
330 STNCIR = p 114) ‘LIME CIRCULATION RATE, LB/HR
340 G = P115)
350 CO2FG P116)
360 lAiN = p1)7)*1.8+32. AIR INLET TEMP. DEC F.
370 IC 1.8*TC + 32.
380 TRF 1.8*TR432.
390 RTUCIR = STNCIR*CTRF 1G)*.2R4 ‘HEAT SUPPLIED BY CIRCULATING LIME
400 BTUOIL = F*1780. +
410 BTUCRI( = F#600. • HEAT TO CRACK OIL
420 RHA . 1 Sl 9 *(lSOS?._(3R25.+6.9R*(TA TNl0 )
430 BTIJAZR = A*(BHA+.0439*(TG1 6 OO ’
440 C = STNRT*CADPCT/(15*F*S
450 BTUCAL STNRT*CAOPCT* 14 .1
460 BTUCO2 STNRT*CAOPCI*(3.203* 0239*h I 1 6 O O
470 BTUSTN = STNRT*(1..0O186*CA0PCTC362 .94*284*1I I 6 O
4R0 BIULOS •242*D*(TG 0.)
490 RTUFXD = RTUCIR_BIULOS_RTUCALTU0 TL_8TTTTNBTD 2
500 BTUFG = (I.4056*TGi.172*TAI 204l 9 )*G
510 RTU8RN = BTUFC-BTUFXD
520 HTREL BTUBRN/1.E +06
530 BTULB = BTUBRN/F
540 O2FED = .Q 1332 *A+.00t58*6*O2 INFG ‘MOLES 02 10 GASIFIER/HR
550 HOXD BTUBRN/O2FED
560 150 RATIO = 2.9IE .4*EXP(9.76E_3*T
570 FRCO2 = 1./(RATIO + 1.) ‘FRACTION C02 IN PRODUCT GAS
580 5102 = CPCT/12 . +HPCT/4 . + SPCT/32. ‘STOIC 02. MOL/100 LB
590 R = 1.E4*O2FED/(F*ST02)
193
-------
Table E—]. Fortran Listing of
CAFE Heat Balance Programme
JHThOC CONTINUED
600 AA = 5.4
61$ 200 COXOP EXP(-AA)*(Rt.9 2)s(TCtI.336)
62$ COXDM 8.333E-4*CPCT.COXDP MOLE C OXIDIZED/ioo LB FUEL
630 HCCO2 169790. + .434*(TG - 1600.)
640 HCCO 4R525 • .7375*(TG-1600.1 • REACT HEAT. C TO CO
65$ HH2H2O = 106941. • 1.327*CTG- 16 00.) • REACT HEAT. H2 TO H20
660 BTUCCO2 = HCCO2aCOXDM*FRCO2 ‘BTU/lOPI LB FUEL FOR C TO COP
67$ STUCCO = HCCO*COXDM*C1.-FRCO2)
690 MOXDM=(BTUCCO2+BTtJCCO-C .s*cox *c I . FRCO?))s11OXD ,( .5*HOXD-)4H H2Q)
690 HOXDP 2 E2*HOXDM/H CT
700 CO2N = COXDM*FRCO2
71$ CON = COXDM-C02M
720 02H20 = OPFED*100./F -CO2M-COM/2.
730 HOXDM2 = O2H2O 2.
740 HOXDP2 = HOXDM2,(HPCT*5.E-3)
750 BTUH2O = HOXDM2$HH2H2O
760 NOEL = .01*F.(BTUCCO2+ STUCCO+BTtJH2O)/O2FED
765 PRINT .HOXDP.HOXDP2
780 610 FORMAT t4XsF5.1.4X .F6.2 .4X.F7.0 .AX,F8.0.4X.F1.0.4X,F5.1)
790 GO TO 50
800 600 FORMAT C/, 6Xs 1 HR. 7X. 5HHTREL. 4X . 6HBTU/L9, 6X . 6HBTU/02
810 800 STOP
820 END
194
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‘ Daole E-2 Gasifier Thernal Perforrance
Co iputed by Programme J}ITHOC
Run Time Percent Heat Release in Gasifier Percent, Oxjdised
Day-Flour Ic. Per lb Oil Btu/lb Mole 02 Mass Real
taxio 6 Mesure 1inc
1.2130 20.4 1.16 3071. 149479. 153831. ¶4.5
io,ooy 20.5 1.14 3241. 157451. 154603. ¶4.4
3,0030 ¶9.0 1.12 2974. 155696. 154216. 13.2 18.0
,1330 18.3 1.11 2925. 158537. 152165. 12.7 ¶4 .8
21.9 1.16 3080. 139622. 153441. 15.8 20.1
( ,o o 22.0 1.16 3074. 138590. 149615. 16.2 1.4
19.4 1.20 3134. 160568. 153010. 13.6 4.8
8,2300 20.8 1.17 3314. ¶58708. 153759. 14.8 23.3
8,1130 19.7 1.15 3102. 156458. 153065. ¶3.9 21.3
10.1330 21.5 1.04 3258. 150388. 154904. 15.4 ¶1.9
2,1230 20.8 1.19 3136. 150195. 154269. 14.7
5,0630 20.6 1.15 3170. 153128. 152576. 14.7 10.0
6,0 430 ¶7.7 ¶. 2845. 160022. 153325. 12.1 15.3
7,0330 20.3 1.22 3188. 155426. 152149. 34.5 ¶9.8
1,0930 19.9 1.14 3001. 150019. 154376. 14.0 18.5
1,10)0 ¶9.8 1.16 3046. 152706. 154112. 13.9 9.1
2,0030 20.5 1.19 3)35. ¶52254. 154476. 14.5 12.3
2,0830 20.6 1.20 3140. 151735. 154502. 14.6
4—, 2,1130 20.7 1.18 3097. 148491. 154774. 14.7 ¶1.3
3,0230 20.3 1.15 3014. ¶47617. 154447. 14.3 7.6
3 ,0630 19.5 1.14 3009. 153449. ¶54308. 13.6 6.9
7.07)0 19.5 1.15 3024. 153891. 154315. 13.7 ¶2.6
3,o8 o 18.6 1.12 2941. ¶57529. 154138. 12.8 13.1
3,1130 19.1 1.14 2993. 155765. 154238. 13.3 ¶6.7
1.06)0 20.6 1.24 - 3204. ¶54295. 54505. ¶4.6 ¶5.0
7.09 30 19.3 1.19 3000. ¶58475. 154277. 13.5 14.3
7.2230 17.6 1.18 3044. ¶72273. 153689. 12.0 18.6
7,2230 ¶7.6 1.17 3026. 171259. ¶54332. 11.9 39.6
7,2330 ¶7.7 1.19 3008. ¶73057. 154237. 12.1 36.5
8,1270 ¶9.7 1.15 3100. 156746. 154082. ¶3.8 40.9
8,2330 19.5 1.15 3211. 163932. ¶54310. 13.6 170
9.0730 ¶9.9 1.16 3251. 162581. 154638. 13.9 26.8
9,0830 20*7 1.17 3299. 158679. 154764. 14.6 24.8
9,1130 20.1 1.14 3204. ¶58843. 154660. 34.) ¶9.7
9,1830 20.3 1.12 3171. 155178. 154195. 14.3
9,2330 19.3 1.12 3167. 162757. 154400. 13.5 I5 5
10,0230 28.3 1.15 3241. ¶58165. ¶54339. ¶4,3 24.6
10,1 430 21.4 1.04 3261. 151476. 154883. 15.3 20.0
10,1530 21.4 1.05 3276. ¶52170. 154375. 15.3 11.1
p4,1330 20.4 1. 18 3134. ¶52815. ¶49320. 14.7 12.5
14,1530 20.0 1.18 3117. 155241. ¶48983. 14.4 18.8
8,0630 19.5 1.22 3161. 169463. 155029. ¶2.7 21.8
7,1530 ¶9.2 1.2) 3134. 162067. ¶52075. ¶3.5 34.0
2,0030 23.3 1.19 3232. 137605. ¶54295. ¶7.0 26.5
2,Ok)0 e 23.4 1.23 3339. 14) 778. ¶54685. 17.1
2.2
• Run 4 data
-------
ab1e E-2 Geatfier Thermal Perfox’4
Computed by programme 1 0C
Run.5
Day-Hour
2,2130
3 .0530
3,1530
3,2130
5,1030
6,0730
2,l530
12,1930
Percent
Stoic.
kir
T9.8
20.4
20.6
20.3
21.3
20.4
20.6
20 6
Total
Bt x 106
i. b
1.23
1.26
1.29
1.31
1.24
1.23
1.20
1.26
lb (
i .u
ffY3.
3102.
3198.
3195.
3277.
3237.
3105.
3026.
3216.
Btu/lb Mo).e 02
aiuL ed Mode 1
159492. 152692.
151492. 152409.
154496. 152704.
156095. 151379.
152803. 156889.
157959. 156234.
149565. 154505.
145809. 15437$.
154398. 154770.
iass
Heat
Thic e
14.0
14.5
14.7
14.6
15.1
14.3
14.6
i4.6
14.6
P iIitce
22.6
13.4
16.8
20.4
10.0
16.4
9.0
5.
14.1
13,0330
12,0830
17,1230
17,1730
21,0930
21,1830
22,0770
20.7
20.6
21.3
21.9
19.7
19.0
19.6
1.26
1.25
1.28
1.39
1.40
1.27
1.26
3165.
3129.
3187.
2968.
2977.
3066.
3079.
152395.
145675.
144341.
149845.
155948.
155877.
154165.
153613.
155509.
155854.
150096.
151771.
153294.
153987.
14.7
15.2
15.7
14.1
13.3
13.7
13.9
13.2
4.3
2.8
13.7
18.1
16.2
14.1
22,1830
19.8
1.29
3146.
162969.
552 3.
13.3
23.5
2 ,0730
19.2
1.27
3131.
162098.
153230.
13.4
25.0
25,0530
19.2
3085•
160530.
153597.
13.3
22.0
25,1530
26,05 )0
19.1
18.9
1.25
1.25
3048.
3046.
160349.
161566.
152916.
153916.
13.2
13.0
22.4
22.6
26,1830
13,0630
26,1130
20.8
18.1
1.26
1.24
3226.
3024.
153912.
166220.
153772.
154561.
14.8
12.4
15.0
28.0
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APPENDIX F
Gasifier Product Composition
197
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Appendix F
Gasifier Prdduct Coinpositi on
Samples of gasifier product were collected in metal containers and
analysed for composition several tines during run 5. Since the coke, tar,
and heavy hydrocarbon portion of the product is not measured by this
procedure an attempt has been made to estimate the quantity and composition
by material balance on the gas ifier itself.
The dry gas compositions measured were listed in Table 19 of the text,
and a sumeary of the material balance results and estimated heavy product
compositions was listed in Table 20.
The calculations made involve several assumptions. Examples of these
calculations and assumptions are presented in this appendix.
Data from the operating period 26.0330 are used in the examples.
• Air Plc i to Gasifier, 3C 4
231 To plenum and fuel injectors
2.8 with Stone feed
2 to cyclone inlet
2)5.8 8C .1
02 Feed = 2 5. 8 x .03)16 lb Mole/O2JHr
SCFM Air
7.819 lb Mole/Hr 2
• Flue Gas to Gasifier
Oxygen and C02 are measured on dried samples of flue gas in continuous
analysers. The rate of flue gas flow to the gasifier is measured by orifice
mater on wet gas. A correction is therefore needed for the composition change
due to water.
Dry Gas Rate (Wet Gas Rate) ( 0 / )
Where Y Vol.% C02 in Dry Gas
X = Vol.% H20 in Wet Gas
Z = Vol.% CC2 in Wet Gas
The ratio X/Z is assumed to be .818 based on the composition of the
fuel oil. This assumption implies that composition of the flue gas does
not change during passage through the flue gas scrubber.
198
-------
For the example period, Y = 14.1
Wet Gas Rate = 86 ScPM
Dry Gas Rate 86 ( 9)/(# 9i + .818) y .i scw
Dry Gas Rate = 77.1 x 12.17 lb Mole/hr
Inputs with Flue Gas
Mole/Hr CO 2 = Mole/Hr Flue Gas x Mole fraction 002
C02 12.17 x .141 1.717 Mole/hr
02 12.17 x .022 .268
12.17 x .837 10.19
Mole/Hr R 0 = Wet Gas Rate - Dry Gas Rate
H 2 O=(86-77.1)z 1.404
Inputs from Solids
Oxygen from regenerator sulphate
= (Lime Circulation, lb/Hr) (Fraction S as S04) - 2
32
02 from Sulphate = 833 x . 0137 = .713 Mole/Hr
Oxygen from fresh stone Carbonate
= (stone Rate, lb/Hr) (Wt. Fraction C02 in Stone)/44
= 32 x .434/44 = 0.316 Mole 0 2 /Hr.
Oxygen from CaO ÷ H2S CaS . .+ H20
=(Mole/Hr S to gasifier) (S removal efficiency)/2
=(409.1 x .0248/32) (.906)/2 = .144 Mole O2JHr
Total 0 Inputs to Gasifier
Air 7.819
Flue Gas 02 .268
CO 2 1.717
H20 .702
Sulphate .713
Stone C02 .316
CaO .144
11.679 lb Mole 02/Hr
199
-------
Nitrogen Inputs
Bed Air + Cyclone inlet flow + Stone feed Air
02 Rate x .797.21
7.819 x .79/.2]. 29.141* Mol/Hr N2
Bed Feed N2 Bleed .351 Mole/Er
Injector N2 .16 Mole/Er
Flue Gas — 10.19 Mole/Hr
Pressure Tapping Bleeds .03
Total N 2 Inputs 240.1145 Mole/Hr
Product Gas Rate N2 Rate/fraction N 2
2 40.1115 1.634 63.321 Mole/Er
Oxygen Outputs in Product Gas C oxides
(Fraction C02 + FrCO/2) (Product Gas Rate)
(.1088 + .0936/2) (63.321) 9.85 Mole 02 /Hr
2 Out as Sulphate
(lime Circulation) (Fr S as SOiL /32) x 2
- 825 x .0029/16 .1*9 Mole 02J11r
02 OUt as 1120
The oxygen leaving the gasifier as water is assumed to be the
difference between the 02 inputs and other known 02 outputs.
02 in Water 11.679 - 9.85 - .1*9 = 1.68 Mole/Hr
112 Inputs to Qasifier
Hydrogen in with fuel Fuel Rate x wt. fraction 112/2
1409.1 x 0.1114/2 23.319 Mole 112 /Hr
Hydrogen in with flue Gas 1.11011
Total 112 In — 214.723 Mole/Hr
200
-------
Hydrogen Output in Product H2 + Hydrocarbons
L (Mole/Hr Product)
As = 63.321
As CHi = 63.321
As C 2 Hk = 63.321
thjdrogen Out as H20
Moles H2 in Water =
H2 out in Water 2
Hydrogen Missing
= Known Inputs - Known Outputs
2 .723 - 16.602 - 3.356 k.765 Mole H2/Hr
Missing hydrogen is assumed to be in tars, coke and liquid portions of
product not measured by gas chrornatograph.
Carbon Inputs
With fuel = Fuel Rate x fraction C/12
= k09.1 x .856/12 = 29. 182 Mole C/Hr
With flue gas C02 input 1.717 Mole C/Hr
With limestone = Stone Rate x Mole fraction CO 2
= 32 x .l 31 /k4 = .316 Mole C/1 r
Total Carbon in = 31.215 Mole/Hr
(Mole fraction x
x.0642 =
x. 0609x.2
x .038lx2 =
Total
H2 Multiple)
1 . 065
7.712
9.825
16. 602
2 x Moles 02 in water
x 1.68 3.356
201
-------
Carbon Outputs
Carbon Out in Product Gas
Z Product gas x Mole. fraction x C multiple
As C02 63.321 x .1088 .3.6.889
As CO 63.321 x .0936 5.927
As CH 63.321 x .0609 3.856
As C 4flp= 63.321 x .0381 x 2 4.825
Dry gas total C 21.497 !4o) e/Hr
Carbon Out Regenerator
(Mole/Hr Regenerator Gas) x (Mole tract ion C02)
to Regenerator 2.902 Mole/Hr
N 2 in Regenerator Gas .95l Mole fraction
02 in Regenerator Gas .017 Mole fraction
002 out = 2.902 x _•• . Mol ./Ilx’
Carbon Missing
= Known Inputs - Known Outputs
= 31.215 - 21.497 - .052 9.666 Mole C/Hr
Like hydrogen, the missing carbon is assumed to be in the form of
coke, tar, and other heavy components of the product gas not measured by the
gas chromatograph.
2Q&0 2 4ade in Gas ifler
The CQ made in the gasifier is taken as C02 out less C02 from lime
stone and flue gas recycle
CO 2 made=6.889- .316-1.717=4.856
CO made = 5.927
00/CO 2 = 5.927/4.856 1.221
H/C Ratio of Heavy Products
The H/C ratio of missing products is taken as 2 x H2 missing/C missing
H/C 2 x 4.765/9.666 .99
Results of these component material balance calculations for four
gas samples are suimnarised here in Table F-i and in Table 20 of the text.
202
-------
Table F-i
Summary of Gasifier Component Material
Balances
Time 22.1030 22.1730 26.0330 26.1770
02 In, Mole/Hr 11.797 11.715 11.679 11.240
02 Out, 10.08k 9.75) 9.999 8.931
02 to H 2 0 (Diff) 1.713 1.962 1.678 2.309
H 2 to 3.426 3.924 3.356 4.618
112 inputs 24.901 24.852 24.723 24.677
112 in products 20.472 19.792 16.602 12.301
112 missing 1.00) 1.136 4.765 7.758
C inputs 31.441 31.356 31.215 31.255
C in Products 24. 1 1 .39 22.544 21.549 20.
C MIssing 7.002 8.812 9. 66 10.
H/C in missing .286 .258 .986 1.43
CO/CO 2 Made 1.102 .97 1.22 1.36
H missing, % of feed 4.) 4.8 20.4 33..)
H ozidised % of feed 14.6 16.8 14.4 19.9
C missing, of feed 23.9 30.0 33.1 37.4
C oxidized, % of feed 35.8 k.o 37.0 31.9
203
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BIBLIOGRAPHIC DATA 1. Repori No. 2.
SHEET
3. Recipient’s Accessica No.
4. Title and S c&zk
Chemically Active fluid—Bed Process for Sulphur
Removal Thiring Gasification of Heavy Fuel Oil
5. Report Date
November 1973
6.
7. Aut oc(s) J.LT. Craig - C.L. Johns ., C. Moss, J.R. Taylo
*n,i_IL._! ._TIm I.11
B. Performing Organization Kept.
No.
9. Performing Orga.izatian Namr sad Address
Easo Research centre
Abingdon, Berkshire
10. Project/Task/Work Unit No.
ROAP ADD-BE
11. Contract/Grant Na.
Contract No. 68-02 .-0300
12. Sponsoring Organ4zaeion Name and Address
EPA, Office of Research and Development
NERC-RTP, Control Systems Laboratory
Research Triangle Park, North Carolina 27711
13. Type of Report & Period
Coycted
Second Phasa
14
iS. Supplementary N iea
lB. Absttsets The report describes the second phase of studies on the CA1 ’B process for de—
sulfurixing gasification of heavy fuel oil in a bed of hot lime. The first continuous
pilot plant teat with U.S. limestone 3CR 1691 experienced local stone sintering and
severe production of sticky dust during startup. Batch tests confirmed that 3CR 1691
produced mere dust than the purer Denbighebire or U.S. 3CR 1359 stones. With 3CR 1691,
10 time, more dust was produced during kerosene combustion at 870C than during gasif i-
cation/regeneration. The continuous pilot plant was modified to improve operability
ui4er dusty conditions: 332 gasification hours were spent in a second run with Denbigh-
shire and 3CR 1691 atones in six operating periods, the longest being 109 hours. SU1
fur removal efficiency was comparable for the two stones 1 ranging from 60 to 952. Re-
generator performance was less satisfactory than in earlier tests • A poor sulfur mate-
rial balance indicates need for improved analytical procedures. Total CAYB development
through a large d nstration test will probably take about 6-7 years and require
$3,320,000 in engineering effort.
Il. Key Wntds and Document Anslynis. ha. L)escriptora
Air Pollution
Reg ratioe (Engineering)
Sulfur
Desulfurisat ion
Gasification
Fuel Oil
Calcium Oxide.
Limestone
17b. Identiliers/Open-Ended Terms
Air Pollution Control
Stationary Sources
CAIB Proce.a
Chemically Active PluiAised Bed
Pluidixed Lime Bed
Heavy Fuel Oil
17€. COSATI Field/Group 21.3 131
1$. k aiIabi1iry &atrment 19. Security Class (ibis 21. No. of Pages
Report) 204
Un Int l ted . 22. Price
Pane
1JNCLASSIFIED
— 204 U5COIU.IDC I455*. 7z
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