PB-211 438
STUDY OF CHEMICALLY ACTIVE FLUID BED
GASIFIER FOR REDUCTION OF SULPHUR OXIDE
EMISSIONS
J. W. T. Craig, et al
June 1972
DISTRIBUTED BY:
National Technical Information Service
U. S. DEPARTMENT OF COMMERCE
5285 Port Royal Road, Springfield Va. 22151
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ESSO RESEARCH
ABJNGDON
FINAL
OAF CONTRACT C PA 70*46.
WORK PERFORMED FOR
ENVIRONMENTAL PROTECTION AGENCY
CONTROL SYSTEMS DIVISION
STUDY OF CHEMICALLY ACTIVE
!j * - .'? * "
Bj&D GASIFIER FOE EED0CTION
SULPHUR OXIDE
By: J.W.T.Croig
G.L Johnes
G Moss
•* J.H.lQylor
D.E.Tisdol!
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* N O T ' I C C E
THIS DOCUMENT HAS BEEN REPRODUCED FROM THE
BEST COPY FURNISHED US BY THE SPONSORING
AGENCY. ALTHOUGH IT IS RECOGNIZED THAT CER-
TTAIN PORTIONS ARE ILLEGIBLE, IT' IS BEING RE-
LEASED IN THE INTEREST OF MAKING AVAILABLE
V.AS MUCH INFORMATION AS POSSIBLE.
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EPA-R2-72-020
Study of Chemically Active Fluid Bed Gasifier for Reduction
of Sulphur Oxide Emissions
$. Report Date
June 1972
6.
7. Autlior(s)
J.W.T. Craig, G.L. Johnes, G. Moss, et al
8. Performing Organization Kept.
No.
9. Performing Organization Name and Address
ESSO RESEARCH CENTRE
ABINGDON BERKSHIRE
ENGLAND
10. Project/Task/Work Unit No.
11. Contract/Grant No.
CPA 70-46
12. Sponsoring Organization Name and AddrebS
ENVIRONMENTAL PROTECTION AGENCY, OAP
Control Systems Division
Research Triangle Park, N. C. 27711
13. Type of Report & Period
Covered
Final
14.
5. Supplementary Notes
16. Abstracts
The first phase is presented of an overall project which has the objective of demonstrat-
ing and evaluating the Chemically Active Fluid Bed (CAFB) process in a converted power
generation boiler of approximately 150 megawatt capacity. The current work objectives of
this phase I study were as follows: a) purvey operating characteristics of two limestone
and one residual fuel oil to determine their suitability.-for the GAFB proce&e; b-) select
the best of the two stones for further testing; e) construct two new batch reactor CAFB
laboratory experimental units; d) measure the effects of important operating variables
on the CAFB process using batch reactor experiments with the selected oil-limestone com-
bination; e) operate for 200 hours of gasification a continuous CAFB gasification pilot
plant designed, constructed, and paid for by Esso. These experiments in batch reactors
and operation of a continuous, two-reactor, pilot plant have greatly increased the
understanding of the (CAFB) process for heavy fuel oil gasification and desulphurisation
in a fluidized bed of lime particles. The experiments have demonstrated a range of condi-
tions over which sulphur removal of 95% or more can be achieved while producing a hot
fuel gas which burns readily in a burner of simple design.
17. Key tt'ords and Document Analysis. 17o. Descriptors
Air pollution Limestone
Sulfur oxides Vanadium
Fuel oils
Combustion
Fluidized bed processors
Chemical reactors
Pilot plants
Gasification
Desulfurization
Regeneration (engineering)
17b. Idcniificrs/Open-Endcd Terms
Air pollution control
Chemically Active Fluid Bed (CAFB)
Air pollution control
17c. COSATl Field/Group 13B, 7D
Regional Center for Environmental Information
US EPA Region III
1650 Arch SI.
Philadelphia, PA 19103
18. Availability Statement
Unlimited
19. Security Class (This.
Report)
, UNCLASSIFIED
Security Class (This
Page
CUNCI.
ASSIF1FD
21. No. of Pages
349
22., Price
L
ORM NTIS-
USCOMM-DC M952-P72
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ESSO RESEARCH CENTRE
ABINGDON BERKSHIRE
ENGLAND
FINAL REPORT
OAP CONTRACT CPA 70-^6
JUNE 22 1970 to MARCH 1972
STUDY OF CHEMICALLY ACTIVE FLUID BED OASIFIER
FOR REDUCTION OF SULPHUR OXIDE EMISSIONS
J.W.T. Craig
G.L. Johnes
G. Moss
J.H. Taylor
D.E» Tisdall
JUKE 1972
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CONTENTS
Page No.
I SUMMARY 1
A. Fuel and Limestone Screening 1
B. Batch Reactor Variable Study 1
C. Continuous Pilot Plant Operation 2
II BODY OF REPORT
A. Introduction and Objectives 5
1. Background Information 5
2. Overall Project Objectives 7
3. Current Work Objectives 8
B. Apparatus, Materials and Procedures 8
1. Batch Reactors 9
a. The 1-A Batch Reactor 9
b. *4-A and k-B Units 13
c. Tall Batch Units 15
d. Flow Plan 16
e. Gas Analysis 16
2. Batch Unit Operating Procedures 18
a. Fresh Bed Tests 18
(1) Start-up 18
(2) Calcination 18
(3) Gasification 18
(U) Regeneration 19
(5) Shut-down 19
b. Batch Unit Cycle Tests 19
}. The Continuous CAFB Gasification 20
a. Flow Plan and Layout 20
b. Gasifier-Regenerator Unit 2k
c. Gasified Fuel Burner 26
d. The Fuel System 27
e. The Stone Handling System 27
f. Safety Provisions 27
g. Process Control 28
h. Analysers 29
i. Operating Procedures 29
Preceding page blank
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Page No.
4. Cold Model of Gasifier 30
5. Materials 30
a. Limestone 30
b. Oil 31
6. Experimental Programme 34
C. Results and Discussion 3^
1. Oil and Limestone Comparisons 3^
a. Oils 34
b. Limestone Comparison T^sts 35
2. Sulphur Removal During Gasification 40
a. Bed Processes During Gasification 40
b. Fresh Bed Sulphur Absorption Curves **•!
c. Temperature and Air/Fuel Ratio 44
d. Upper Temperature Limitation 30
e. Gas Velocity and Bed Depth in
Fresh Bed Tests 51
f. Lime Particle Size 5*t
g. Lime Replacement and Sulphur
Differential 58
h. Correlation of Fresh Bed and Cycle
Test Results 59
i. Continuous Pilot Plant Experienpe 63
3. Regeneration Behaviour 63
a. Fresh Bed Test 63
b. Cycle Test Regeneration 65
c. Continuous Regeneration 65
4. Vanadium Retention 69
5. Gasifier Product Gas Composition 71
6. Lime Attrition and Losses 72
a. Cyclic Batch Ted Bed Losses 72
b. Particle Size Changes 7^
c. Continuous Unit Experience 7^
7. Operation of Continuous Pilot Plant 77
a. Gasification and Sulphur Rsnoyftl 77
b. Solids Circulation 77
c. Gasifier Temperature Control 78
d. Regenerator Temperature Control 78
e. Limestone Feed 79
f. Refractory Condition J9
g. Agglomerations and Deposits 80
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D. Conclusions and Recommendations 86
1. Conclusions 86
2. Recommendations 90
III References 92
Appendix Title
A Chemistry of CAFB Gasification Process
B Results of work just prior to contract
C Batch Unit Fuel and Lime Comparisons
D Batch Reactor Variable Study with Fresh Lime Beds
E Batch Reactor Cycle Test Variable Study
F Correlation Functions for Sulphur Removal Efficiency
G Design Basis for the Pilot Plant Gasifier System
H CAFB Pilot Plant Alarm Systems
I CAFB Pilot Plant Operating Procedures
J Operation of Continuous Pilot Plant
K Burning Rate Model for Carbon on Lime in CAFB Gasifier
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I Summary
Experiments in batch reactors and operation of a continuous,
two-reactor, pilot plant have greatly increased our understanding of
the Chemically Active Fluid Bed (CAFB) process for heavy fuel oil
gasification and desulphurisation in a fluidised bed of lime particles.
These experiments have demonstrated a range of conditions over which
sulphur removal of 95$ or more can be achieved while producing a hot
fuel gas which burns readily in a burner of simple design.
A. Fuel and Limestone Screening
Preliminary screening tests in batch reactors built prior to
this contract established that a Venezuelan fuel oil available in the
United States could be simultaneously gasified and desulphurised
satisfactorily by the CAFB technique. A comparison of two limestones
suggested by NAPCA (now GAP) revealed that one, BCR 1691,was satisfactory
for CAFB operation, though it would require higher replacement rates
for equal sulphur removal because it contained less calcium oxide than
the previous standard, a U.K. Denbighshire stone. Preliminary indications
were that the stones were equivalent at equal CaO replacement rates.
A second stone, BCR 1690, was inferior as a sulphur removal agent even
at equal CaO replacement rates. It also was subject to greater attrition
rates and showed a tendency to agglomerate. No further work was performed
on BCR 1690.
B. Batch Reactor Variable Study
A batch reactor variable study in two new experimental units
screened the effects of major process variables with the Venezuelan
fuel oil (2.3$ S, 350 ppm Vanadium) and the U.S. limestone BCR 1691.
The effects on sulphur removal of limestone particle size, bed depth,
superficial gas velocity, air/fuel ratio, and temperature were explored
in tests with fresh lime beds. Bed depth, lime replacement rate,
the amount of sulphur exposure per gasification cycle, and particle size
were further studied in a series of cyclic tests using alternate cycles
of gasification and regeneration.
Gasification temperature and air fuel ratio were found to
interact strongly. Conditions that produced an adequate rate of carbon
burning near the air distributor were required to avoid rapid loss of
sulphur removal activity by excessive carbon laydown. With gasification
temperatures of 870°C or higher, air/fuel ratios as low as 20$ of
stoichiometric could be used with good (>95$) sulphur removal efficiency.
Temperatures down to 800°C could be used provided air/fuel ratio were
high enough i.e. above 27$ of stoichiometric. On the other hand, sulphur
removal efficiency declined at temperatures above 900"C.
Sulphur removal activity of fresh beds correlated with the number
average particle size of the lime.
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It appears that, at least at low lime utilisations, sulnhur
removal efficiency varies inversely with the square of particle si
However, in cyclic testing or continuous operation, consideration must
be taken of the effect of fines losses on performance. A limestone
with a large fraction of less than 600 micron particles undergoes
greater losses during calcination in the gasifier and gives poorer
performance than a stone with larger initial particle size but with
fewer losses. Extensive testing was done with a 300 - 3175 micron size
range stone but a size range of 600 - 3175 microns is preferred.
The effects of bed depth and superficial gas velocity on
sulphur removal efficiency can be correlated by their ratio, the
superficial gas residence time above the fuel injector. As a first
approximation, a first order reaction model expressed the variation
of sulphur removal with gas residence time in a fresh bed. However,
in cyclic tests, the deactivation of the lime, with sulphur exposure
must also be considered and gas residence time plays a less significant
part. The empirical equation which fits the batch unit cyclic test
data is:
Sulphur Removal Efficiency, %, = 100
1 -
where t = gas residence time, sec above fuel injector
m = Lime replacement rate, wt CaO/wt S
d = % Sulphur differential, wt 3 passed into 100 Ib lime.
This model implies a fluid bed depth of about 3.U ft for
sulphur recovery with stoichiometric replacement of BUR 1691 stone at
6 ft/sec velocity.
Vanadium is removed from the fuel oil by the lime during
gasification. In some cyclic test series vanadium removal from the fuel
was essentially complete. In other cases less than 100J6 removal was
experienced. Effectiveness of vanadium removal correlates with
regeneration temperature, which implies that a vanadium fixation step
is needed in addition to the preliminary vanadium absorption.
Lime losses from the bed in batch unit studies are sensitive to
bed depth. They vary from about 1 to 9 lb/hr - ft • at 6 ft/sec
superficial gas velocity. Similar loss rates have been observed in
the continuous pilot plant. In the pilot plant maintenance of cyclone
efficiency is a major factor in keeping losses low.
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Empirical equations were obtained to describe the variation in
composition of the gasifier product gas with temperature and air/fuel
ratio. The C0/C0p ratio increases with temperature as does the fraction
of hydrogen and hydrocarbons heavier than Ci).. Increasing air fuel ratio
increases the preferential oxidation of the fraction heavier than C^
C. Continuous Pilot Plant Operation
Three runs of the continuous gasifier pilot plant have been
completed. The first run demonstrated basic operability of the process
and showed that continuous gasification, regeneration, solids circulation,
and sulphur removal was feasible. An average of 90$ sulphur removal
from the fuel gas was maintained for a continuous 87 hour period while
producing a regenerated gas containing 8 to 10$ S0?. Boiler operation
showed that the gasifier produced a hot, combustible fuel which ignited
readily and burned with a clean luminous flame.
During this run it was observed that carbonaceous deposits
accumulated in the cyclones and gasifier outlet ducts during the run, and
this accumulation caused a gradual increase of pressure drop between
gasifier and boiler. The deposit was removed by air burnout, but
excessive temperature during this burnout damaged the initial cyclone dip
tubes made of stainless steel.
The second run was an attempt to operate the gasifier without
internal dip tubes in the gasifier outlet cyclones.
A continuous gasification period of 91 hours was maintained with
nearly 100$ removal of sulphur from the fuel gas. However, absence of
the cyclone tubes, coupled with high lime attrition rates caused by
modified high velocity fuel injectors, led to high lime losses and required
unreasonably high replacement rates.
The third run employed new cyclone dip tubes which could be steam
cooled during decoking operations. A run with over 200 hours of gasification
was carried out, using the steam cooling together with inert gas recycle
to control temperature when decoking.
This third run demonstrated the following features:-
• Continuous gasification periods up to 6k hours without decoking.
o Essentially 100$ sulphur removal at 100 to 200$ of
stoichiometric lime replacement rate and total bed depth
of 19 inches.
e Greater than 95$ sulphur removal at 70$ of stoichiometric
lime replacement and 22 inches total bed depth.
• Maintenance of regenerator gas concentrations of 6 to 9$ S0p.
• Automatic control of regenerator temperature by solids
circulation rate.
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e Regulation of gaslfier temperature by flue gas reoycle
rate.
• Trouble-free operation of the solids circulation system.
These pilot plant runs pointed out several areas which
require additional study to improve long term process operability.
• Decrease rate of solids deposition in cyclones and transfer
ducts.
• Reduce lime attrition and solids carry-over to boiler.
• Find improved material of construction for cyclone internals
to reduce time required for decoking by allowing higher
burn-out temperature.
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II Body of Report
A. Introduction and Objectives
Exploratory work at the Esso Research Centre, between 1966
and 1970, had shown a range of conditions voider which sulphur is
absorbed in a fluidised bed of lime in which fuel oil is undergoing
partial combustion and gasification. The possibility of regenerating
the lime also had been demonstrated. Thus the framework had been
established for a process to eliminate S0? pollution during combustion
of high sulphur fuel oils. The process is designated CAFB for
Chemically Active Fluid Bed.
The process employs a shallow (about 2 feet deep) fluidised
bed of lime particles in a gasifier to partially oxidise, to gasify,
and to desulphurise the heavy fuel oil. The hot gasified, sulphur-free
fuel passes from the fluid bed to equipment in which it is fully
combusted, such as the burners of a boiler. Calcium sulphide, which
forms in the gasifier fluid bed, is converted back to calcium oxide in
a separate fluid bed regenerator by contact with air. Hot solids
circulate continuously between gasifier and regenerator fluid beds.
A purge stream of lime removed from the regenerator is replaced by an
equivalent quantity of limestone to maintain activity of the bed for
sulphur absorption.
The sulphur leaves the system as a concentrated (about 10$) S0_
stream from the regenerator. This sulphur can be recovered in several
ways; for example, by conversion to sulphuric acid or by direct reduction
to elemental sulphur.
In addition to eliminating S02 pollution, the pro cess has the
additional benefit of removing vanadium before the final combustion
stage, and thus reducing or eliminating this source of high temperature
corrosion in boiler superheaters.
In June 1970, Esso Research Centre began work on a contract
with the United States National Air Pollution Control Administration
(now Office of Air Programs) to study the CAFB process for possible
applications in the United States. This is the final report on work
done under that contract in the period from June 22 1970 through
March 1972.
1. Background Information
The original CAFB work at Esso Research Centre was directed at
fully combusting oil fired fluid beds with the objective of achieving
sulphur removal along with the high heat transfer rates of fluid bed
boilers. It soon became evident that by using lime as the fluid bed
medium, benefits in pollution abatement could be obtained which
potentially were more important than the improved heat transfer.
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- 6 -
It was found that sulphur was trapped by lime both under fully
combusting conditions and under gasifying conditions, in the latter
case the oil is contacted with substoichiometric amounts of air and
the gasifier product is a hot, sulphur-free fuel gas which can be
burned in a conventional boiler.
Iftider fully combusting conditions the sulphur reacts with
lime to form calcium sulphate.
H2S + 202 + CaO - > CaSO^ + HgO
The sulphate can be reduced to SOp and lime with fuel oil.
With hydrogen the reduction reaction is :
+ E2 - > CaO + S02 + HpO
When fuel is combusted under substoichiometric conditions, the
sulphur is trapped by lime as calcium sulphide:
HgS + CaO - > CaS +
The lime can be regenerated by oxidising the CaS with air:
CaS + 3 00 - > CaO + S00
2
Each pair of reactions forms the basis of a regenerative lime
process for avoiding S02 pollution when using high sulphur fuel oil.
The gasifier process has the added advantage that it is potentially
applicable to existing power plant installations.
The early work was described In a technical proposal to NAPCA
in 1969. Later work obtained In the period between submission of the
proposal »nd beginning of work under contract CPA-70-^6 is summarised
in Appendix B of this report. A more complete discussion of CAFB
chemistry appears in Appendix A.
Although the CAFB process uses continuous circulation of solids
between two fluid bed reactors, the experimental work used a single
reactor which was cycled between gasification and regeneration conditions.
The earlier tests used reactors without cyclones and were restricted to
low gas velocities to prevent solids carryover. Later versions of the
apparatus employed internal cyclones and were able to use higher gas
velocities. Temperatures in the range of 800-900° C were found best for
sulphur removal during gasification. The regeneration of sulphided lime
occurred between 1,000 and 1,100°C.
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- 7 -
With the low gas velocity reactor it was necessary to underfire
that is, to burn some fuel in the plenum space beneath the air
distributor. If underfiring were not used, the velocity of the cold
entering air was insufficient to fluidise the solids in the vicinity
of the distributor, and poor desulphurising efficiency was obtained.
Even with underfiring, desulphurising efficiency declined severely as
air to fuel ratio decreased below 30$ of stoichiometric. When reactors
with cyclones were employed, higher total gas rates could be employed,
and underfiring was no longer required. Under these conditions, good
sulphur removal was obtained at air rates down to nearly 20$ of
stoichiometric.
When a single bed of lime is cycled between gasification and
regeneration conditions there is a gradual loss of sulphur removal
efficiency. If a portion of the lime is replaced with fresh limestone
at the beginning of each cycle, the efficiency will line out at some
level which depends on the quantity of lime replaced. A curve of lined
out efficiency vs. lime replacement rate was established which indicated
that replacement of about one mole of CaO per mole of sulphur fed
would maintain the sulphur removal efficiency at close to 100$. The
bed life tests on which this replacement study was based were made with
underfiring and about 30 to 35$ of stoichiometric air.
These life tests also demonstrated that the lime retained much
of the vanadium and some of the sodium from the fuel oil. About 20$
of the sodium was picked up. Vanadium retention was variable and
amounted to 100$ in some instances. A fuller presentation of results
appears in Appendix B.
2. Overall Project Objectives
Work under this contract constitutes Phase I of an overall
project which has the objective of demonstrating and evaluating the
CAFB process in a converted power generation boiler of approximately
150 megawatt capacity, situated in the United States. The overall
project is conceived as being a six phase operation, with the following
broad objectives for each phase:
Phase I • Laboratory Tests in batch CAFB reactors to select
suitable fuel and limestone combinations and to screen
critical process variables.
• Construction and commissioning of a continuous pilot
plant of about 7 million Btu/Hr capacity.
Phase II • Operation of continuous pilot plant to provide process
data and knowhow for design of conversion of commercial
unit.
• Development of conceptual design for conversion and
economic evaluation of process viability.
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- H -
Phase 113 • Full study design of proposed conversion.
• Additional laboratory and pilot plant studies to
answer specific design questions.
Phase IV • Detail design of conversion.
Phase V • Boiler conversion and commissioning
Phase VI • Boiler operation in normal duty while evaluating
long term performance of CAFB gasifier.
The Phase I experimental programme is now complete.
J. Current Work Objectives
The specific objectives of this Phase I study have been
as follows:
a. Survey operating characteristics of two limestones and one
residual fuel oil selected by NAPCA (now OAP^to determine
their suitability for the CAFB process.
b. Select the best of the two stones for further testing.
c. Construct two new batch reactor CAFB laboratory experimental
units.
d. Measure the effects of important operating variables on the
CAFB process using batch reactor experiments with the selected
oil-limestone combination.
e. Operate for 200 hours of gasification a continuous CAFB
gasification pilot plant designed, constructed, and paid
for by Esso.
B. Apparatus, Materials and Procedures
Apparatus and procedures used in this study are described in
this section. The apparatus includes
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- 9 -
1. Batch Reactors
Two batch reactors built by Esso were in existence when work on
this contract began. These units were used in the initial limestone
and fuel oil screening work while two new batch units were being
designed and constructed for the variable studies.
a. The 1-A Batch Reactor
The batch reactor used in the CAFB limestone and oil screening
tests is designated No. 1-A. The lower section of this unit was the
first CAFB reactor used extensively at the Esso Research Centre. In
those original tests, however, it was not fitted with cyclones, and
therefore was used only In low gas velocity experiments. Before the
current work began the unit was fitted with an expanded upper section
and an Internal cyclone. This cyclone, which is closed at the bottom,
traps fines but does not return them to the fluid bed. The quantity
of fines is withdrawn and measured at the end of each run period.
A drawing of the 1-A reactor is shown in Figure 1, and a photograph In
Figure 2.
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- 10 -
DRAWING OF I-A BATCH REACTno
RG.I
-------
ILLUSTRATIONS SIGNIFICANT TO TEXT MATERIAL
HAVE BEEN REPRODUCED USING A DIFFERENT
PRINTING TECHNIQUE AND MAY APPEAR AGAIN IN
THE BACK OF THIS PUBLICATION
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- II-
PHOTOGRAPH OF I-A BATCH REACTOR
FIG. 2.
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- lii -
The reactor is made of a carbon steel shell lined with a hieh
temperature castable refractory. The lower section, which contains the
fluid bed, is 7 inches internal diameter, by 33 inches high. The
upper section expands to 15 inches internal diameter and serves as a
disengaging space for all but the finest particle size range of solids.
\
The distributor is a "top hat" design with a raised section 5
inches in diameter. It contains sixteen horizontal holes distributed
around its circumference. Figure 3 is a sketch of the distributor.
BATCH UNIT DISTRIBUTOR
HOLE DIAMETER
125 in higti
. (08 in low
A P
A P
8 25
INCHES
RG.3.
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The 1-A unit has no provision for temperature control other than
by adjusting the quantity of insulation around the reactor. This
insulation was arranged to permit the unit to operate in the range of
800 to 900°C with air velocity of 4 ft/sec and air/fuel ratios in the
range of 20 to 30$ of stoichiometric. With constant air supply, the
unit tends towards 900°C at the lower fuel rate and 800°C at the higher
rate.
b. k-A and k-B Units
Two new batch reactors (designated 4-A and 4-B) were constructed
specifically for this programme so that the variable study could proceed
more rapidly. Their design was based to a large extent on that of the
1-A unit. Figure 4 is a drawing of the new reactor vessel with its
various inlets and connections. The reactor is made of a carbon steel
shell lined with a high temperature castable refactory.
CAFB BATCH UNIT NO.4. REACTOR
FIG. 4.
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- 111. -
The lower section, which contains the fluid bed, is identical
to that of the 1-A unit, being 7 inches in internal diameter by J5 inches
high. Fuel oil enters the reactor through a single £ inch nozzle
which protrudes one inch in from the reactor wall at a point fji inches
above the bottom.
The upper section was redesigned to permit installation of two
internal cyclones and to allow drainage of the cyclones to external
sample collectors during operation.
The distributor and plenum are the same as in the 1-A unit.
The distributor is a top hat design with a raised section, 5 inches
in diameter- It contains 16 horizontal holes distributed around its
circumference.
A low pressure drop distributor with 3/i6 inch diameter holes
has been used for most of the work, but a higher pressure drop
distributor with 1/8 inch holes has also been tested. Pressure drop
characteristics of the two distributors are shown in Figure 5.
BATCH REACTOR
DISTRIBUTOR CHARACTERISTICS
HIGH R DROP DISTRIBUTOR
NORMAL DISTRIBUTOR it HOLES
AIR FLOW(FT»/MIN)
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- 15 -
The k-A and 4-B reactors contain internal cooling coils which
allow reactor temperature to be controlled independently of fuel and
air rates. This was not possible in the 1-A unit.
Figure 6 is a photograph of the new units with their control
panel. These units have been designated No. 4-A (left hand unit
facing panel) and No. 4-B.
PHOTOGRAPH OF NEW BATCH UNITS
4- A AND 4-B
c. Tall Batch Unit
FIG. 6.
Several tests were performed in a third version of the batch
units, designated as No-3 unit. The reactor section of this unit is the
same diameter as that of the others, but the height is much greater;
17-ft compared with >ft in the straight sections of units 1-A, 1*-A and
k-B. The No.3 unit also has an expanded section at its top, but it
does contain an internal cyclone which returns fines to the bed. The
No.3 unit had been constructed by Esso prior to beginning work on this
contract.
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- 16 -
d. Flow Plan
A flow plan of the batch units is given in Figure 7. This simple
flow scheme applies both to the 1-A and to the 4-A and k-B units.
FLOW PLAN OF CAFB BATCH UNIT
Flare
Sample Gas
to
Analysers
Sample Gas
Pump
J Sample
Rome
CAFB
Botch Reactor
From Heated
txi Fuel Injector Air
Propone for start-up
FIG. 7
e. Gas Analysis
Product gas leaves the unit through the cyclone outlet and passes
to a flare outside the laboratory where it is burned. A portion of gas
is burned in a sample flame located just above the reactor, and the
combustion products are analysed for S02, 0^, CO and COp.
Most of the information obtained from the batch tests has been
based on gas analysis. During gasification, the desulphurising efficiency
of the bed has been calculated from carbon dioxide, oxygen and sulphur
dioxide contents of the fully combusted product gas. Because of the wide
range of sulphur compounds present in the gasifier product gas itself,
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- 17 -
this is the only practical method for measuring desulphurising efficiency.
During regeneration, the exit gas was monitored for carbon dioxide, carbon
monoxide, sulphur dioxide and oxygen.
Hie gas analysing equipment which was used in the tests is
listed in Table I.
TABLE I
GAS ANALYSING EQUIPMENT
Type
Manufacturer
so2
so2
so2
co2
CO
Infra-red Maihak
Infra-red Maihak
Conductimetric Wosthoff
Infra-red Maihak
Infra-red Maihak
Paramagnetic Servomex
Model
Uhor 6
Unor 6
Unor 6
Uhor 6
OA 137
Response
Continuous
Range
0-1,000 ppm
0-20 c/o uy vol.
0-1000 ppm
0-20 % by vol.
0-20 % by vol.
0-25 % by vol.
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- 18 -
2. Batch Unit Operating Procedures
Two types of study are carried out In the batch units; fresh
bed tests, and cyclic tests. Differing procedures are used in the
two studies .
a. Fresh Bed Tests
A new batch of calcined lime is used in each of the fresh
bed gasification tests. These tests were used to screen rapidly the
effects of a number of gasification variables such as bed depth, gas
velocity, particle size, and air fuel ratio.
(1) Start-up
The bed was heated to operating temperature by combustion of
propane (butane in the 1-A reactor) below the distributor and by
direct injection of kerosine through the fuel nozzle. First of all* ,
the bed space temperature was raised to 950" C by gas combustion below
the distributor. Then, with gas and air to the unit shut off,
approximately 9 ID of bed material was added. This was heated to
750° C using gas before switching to direct kerosine injection for
the final 200° C rise. These latter steps were repeated until all
the bed had been added and was at operating temperature. Although
gas combustion could be used on its own to raise the bed to operating
temperature, this route was avoided since it produced extremely high
gas velocities through the distributor which caused appreciable
attrition of the bed material.
(2) Calcination
The start-up procedure was also followed for calcination. Due
to the strongly endothermic nature of the reaction, 950* C could only
be attained when all the carbon dioxide was driven off. C0p emissions
were also checked by gas analysis.
Gasification
When the start-up procedure had been completed, the fluidising
air was set to the desired level, the kerosine supply was terminated,
and fuel oil was introduced at a predetermined feed rate. The sampling
flame burner and external flare were lit and the gas analysers connected.
Relevant data was recorded and bed samples taken from points 6 inches,
11 inches and 16 inches above the distributor at prescribed intervals.
Bed sampling was carried out under nitrogen to prevent any oxidation of
the sulphided lime. Samples of the gasifier product gas were also taken
for separate analysis by gas chromatography .
-------
- 19 -
Regeneration
For regeneration, the fuel supply was switched off and air
flow alone continued through the bed. The gas analysers were switched
to the sampling system for regeneration and the relevant data were
collected. A bed sample was collected after each regeneration.
(5) Shut-down
The bed temperature was allowed to drop to 700°C before the
bed was 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 had been removed,
the cyclone was emptied and all ancillary equipment shut off.
b. Batch Unit Cycle Tests
The 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 were performed at each set of operating conditions
to establish the lined-out sulphur removal efficiency for
those conditions.
Cyclic tests used the same calcination and start-up procedures
as thefraahbed tests. However, after the initial start, a series of
gasification and regeneration cycles followed each other. Sampling
and gas analysis procedures also were 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 was not possible to draw a regenerated
bed lime sample before temperature dropped 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 30
emission and fall of bed temperature, fluidising air was
stopped.
(2) Replacement limestone was added.
Fluidising air was resumed and when temperature reached
the desired point, oil feed was resumed.
Aso lids sample was withdrawn to provide a measure of
vanadium concentration between cycles and to return the
bed to ics proper level.
-------
- 20 -
3- The Continuous CAFB Gasification Pilot Plant
A continuous CAFB gasifier pilot plant has been constructed at
the Esso Research Centre to provide a demonstration of the CAFB process
under continuous operating conditions and to provide a means for studying
those operational variables which cannot be measured in batch reactors.
Features of the pilot plant are summarised here. The design basis and
construction details are discussed in greater detail in Appendix G.
a. Flow Plan and Layout
Figure 8 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 million 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.
CAFB PILOT PLANT FLOW PLAN
^
Cooling Tower
12,000 Imp gallons
Ston* F»»t
•r
Oil FMd ^cys!
Pumps /fl^
(3}
M«ttr A'
Fuel Injection
Air
(9)
Oi
Propor* to
Start-*
Fio.a
-------
- 21 -
Figure 9 shows the layout of the pilot plant equipment within
its building. 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 at the pit side. Electrical instrumentation and
manometers are mounted in a separate control cabinet. Gasifier blowers
are located in a separate blower house outside the main building, and
the cooling tower is mounted on the roof.
GENERAL PLANT LAYOUT
Regenerator Cyclone
Fuel Injector (3 off).
Air Supply to
Gaalfler.
V X X X X. X X XX
Air Supply to
Regenerator.
Circulating OH Supply
from 30x9 tank.
SECTION.
Electrical-
control
cabinet.
Off lea.
control
console.
iHeat Exchanged
Swinging jib
pillar crane.
szV
Boiler.
Blowers
for gaslfler.
To 54-0
high stack.
PLAN.
FIG.9.
-------
Figure 10 is an interior view of the completed pilot plant
showing rarifi«r unit, boiler and mechanical equipment console.
CAFB PILOT PLANT INTERIOR
FIG. 10.
-------
Figure 11 shows the pilot plant exterior with oil feed storage
tank, rear end of boiler, cooling tower, and flue stack.
CAFB PILOT PLANT EXTERIOR
FIG.IL
-------
b. Gasifier-Regenerator Unit
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. Figure 12 is a drawing of this gasifier-regenerator
assembly. 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. TJie regenerator
is 7 inches in diameter throughout.
Oui
Ins
Co
1
3'
**', -
supply
3"
Jfc-
•iwly-
lor metal cosing— \
utoting refractor y \ Cyclone for
stable r.froctory\\ «"»ifltr
\ \W /
\ \\ /
'/ 7 7\7 7 77 Y7 717
/
f
/
7
/
£
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/
r
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' GASIFIER
/ / / S j
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KVQvMtrQ
.Cyclone
\ f e4 Into
transfer
Connection between cyclones
Expansion bellows to
absorb vertical expansion
an gas outlets
Ges pulse
LAYOUT OF CONTINUOUS GASIFIER UNIT
Air
Air Supply
FIG. 12.
The gasifier air distributor consists of 32 nozzles each of which
contains six horizontal holes of 0.177 inch diameter. The nozzles are
stainless steel mounted on a refractory covered carbon steel plenum.
The regenerator distributor is of the same design as that used in
the batch reactors, a refractory top-hat shape with sixteen horizontal
holes. Hole diameter in the continuous unit distributor, however, is
0.25 Inch.
-------
Assembly of the gasifier unit steel framework and casting of
the high density refractory concrete was performed with the gasifier in
place at the Esso Research Centre. After pouring each lift of concrete,
it was vibrated to provide maximum density.
Figure 15 is a photograph showing the gasifier at an early
stage of construction. The plastic patterns which form the regenerator
and solids transfer ducts are in place together with the wooden pattern
of the lower gasifier bed. Three of the pre-insulated mild steel
encasing walls also are erected.
GASIFIER DURING CONSTRUCTION
FIG. 13.
-------
- 26 -
c. Gasified Fuel Burner
The standard oil burner of the 10 million Btu/hr boiler was replaced
with an experimental burns- designed to handle the hot gaslfier product.
A general arrangement of the gasified fuel burner is shown in
Figure 14. It will be seen that the air required for complete combustion
is introduced in two stages. Roughly about 10# of the air enters a crude
injector and is pre-mixed with the gas. The remainder of the air is fed
tangentially into a swirl chamber and emerges from an annular nozzle
which is concentric with the gas nozzle. The design of this burner was
purely empirical and was based on recommended pressure drops at the
nozzles. The main gas duct is sized to give a flow velocity of about
60 ft/second, the gas nozzle gives a pressure drop in the region of 3" w.g.,
and the air nozzle gives a pressure drop of about l»— 5" w.g., all under
full load conditions described in Appendix G.
MAIN GAS BURNER
FIG. 14.
A similar burner of about l/20th of the output was tested on a
batch unit prior to the construction of the large burner and was found
to give a satisfactory performance. The remarkably good behaviour of the
large burner, which required no modifications despite the 20/1 scale up,
suggests that the gas is of excellent combustion quality.
-------
A pilot light fired by propane gas is used for lighting purposes.
This pilot required a certain amount of development work since a spoiler
had to be fitted to prevent its extinction by the secondary air blast.
The whole burner system has been found to be reliable in operation.
d. The Fuel System
The burner which is used to heat the gasifier to its working
temperature is fuelled by propane gas. It is fitted to an end wall of
the gasifier and is angled downwards so that the hot gases impinge on
the fluidised bed. A standard commercial burner of about 700,000 Btu/hr
maximum firing rate is employed.
Three fuel-oil injectors are provided for use diring gasifying
operation and these penetrate the side wall of the gasifier opposite
the cyclones, as shown in Figure 12. These injectors are simple ^" bore
tubes made of a heat resistant iron-chrome alloy. In the first
continuous gasifier run the full diameter injector was used. In the
second run a restriction orifice drilled with a 3/16 inch hole was
placed in each injector tip to increase injection velocity. In the
third run, these inserts were drilled to £ inch diameter. The injectors
are positioned approximately k^" above the top surface of the distributor
and project into the bed about 3". Some air is injected through these
tubes with the fuel, to aid injection and prevent coking in the hot tube
ends. Each injector has its own fuel oil metering pump drawing its oil
supply from a heated Circulating main. It is also possible to switch
to an alternative distillate fuel supply which is used during the start-up
period in order to reduce the propane consumption.
e. The Stone Handling System
Standard commercial equipment, including an industrial vacuum
cleaner, is used for handling the stone outside the gasifier. The
stone feed hopper is mounted upon the lid of the gasifier, and a vibrator
is used to control tfee rate at which the stone enters the gas space.
Ihe hopper is pressurised by a nitrogen bleed during operation in order to
prevent the escape of gas from the unit and the consequent deposition of
tars In the feed line. This system has not proved to be entirely
•atisfactory. A more positive feed control would be preferred and
•edifications are planned.
f. Safety Provisions
The installation is provided with a comprehensive alarm system
which detects and indicates faults occurring in the boiler and its
cooling system, the gasifier and the experimental burner. In some
Instances, such as for example a flame-out in the boiler, or failure of
the pilot light, shutdown is automatic. Details of the alarm system -...i
of the various alarm actions are given in Appendix H.
-------
- 28 -
g. Process Controls
Figure 15 is a diagram of the pilot plant instrumentation
system. Automatic control boxes are used to regulate regenerator
temperature, regenerator bed level, and gasifier bed level. A
packaged pressurisation system maintains constant boiler cooling
water pressure and temperature. All other systems are manually
controlled by the process operator. Dashed lines in Figure 15 show
the indicators and control valves normally used by the operator.
Manometers indicate pressures and pressure differences in most
applications. In four instances pressure differences are also detected
by pneumatic delta pressure cells and transmitted to recorders. Pressure
switches also are employed in several locations to operate warning
lights for abnormal conditions.
CAFB PI LOT PLANT INSTRUMENT FLOW PLAN
Flu* on
— SMHc ftmm Lin*
•ItM Sl«ml
Fiq.is
-------
h. Analysers
On stream analysers monitor the compositions of the key gas
streams. Table II lists these analysers and their applications.
Table II
CAFB Pilot Plant Gas Analysers
•earn
Component
Analyser
Operating Principle
Range
Oas
er Plenum
CO,
Servomex OP 250
ffelhak Unor 6
Paramagnetic
Infra Red
0 - 25# by voJ
0 - 10JK by vol
Flue Gas
d at fire tube
co2
CO
so2
SO,,
Servomex OA 1J7
Malhak Unor 6
Malhak Unor -6
Malhak Uhor 6
Wosthoff
Paramagnetic
Infra Red
Infra Red
Infra Red
Electrical conductivity
of HpOg - S02 reactor
products in solution
0 -
by vol
0 - 20# by vol
0 - 2OJK by vol
0 - 1000 -nr
0 - 1000 ppm
srator
co2
S00
Servomex OA
Maihak Uhor 6
Maihak Unor 6
Paramagnetic
Infra Red
Infra Red
0 - 2.as oy /ol
0 - 10# b • voi
0 - 20# by vol
A sampling pump and filter system supplies each gas stream to
its appropriate group of analysers. A multipoint instrument records
outputs of all Instruments except the Wosthoff SOp analyser which has
its own continuous recorder.
i. Operating Procedures
Details of continuous unit operating procedures for start-up,
operation, decoking and shutdown are given in Appendix I.
-------
^. Cold Model of Gasifier
Before design of the continuous gasifier-regenerator unit
was finalised, a steel and plastic mock-up of approximately the same
dimensions was constructed. This unit was used for testing the bed
transfer system to determine its capacity and to determine the
effectiveness of alternate arrangements of the transfer ducts. A
photograph of the cold unit appears in Appendix Figure G-4.
5. Materials
Table III lists properties of the fuel oils and limestones
used in these CAFB studies. Basically there have been three stones
and two oils. However, different shipments from the same oil or stone
source have had minor variations. These differences are also indicated
in the table.
a. Limestone
The U.K. limestone used in early CAEB studies is a relatively
high purity material quarried in Denbighshire in North Wales. When
calcined it gives a lime containing 97-5# CaO. This stone was used as
a basis for comparison of the performance of the U.S. limestones. It
also was the material used in the continuous pilot plant operation
trials.
Two U.S. limestones were supplied by NAPCA (now OAP) for the
limestone comparison and selection study. These stones were shipped in
steel drums and crushed and sieved at the Esso Research Centre. Both
i.S. stones were of lower calcium oxide content than the U.K. stone and
represent classes of material more widely available in the United
States than the high purity stone. The stone designated 1690 in the
Bituminous Coal Research Inc (BCR) listing comes from Ogdensburg,
New York. It has the lowest calcium content of the three stones tested
and is high in both magnesia and silica. The stone BCR 1691, from Seneca Falls,
New York, also is high in silica but is low in magnesia. It is
intermediate in calcium content and is similar in composition to
Pennsylvania "cement rock".
After the preliminary comparison study, a larger supply of
BCR 1691 was obtained for more detailed variable studies. This main
supply had less silica and more CaO than the original sample, but in
other respects was quite similar in composition.
-------
b. Oil
The fuel oil used in early CAFB studies at the Esso Research
Centre, Abingdon, was a blend of Middle Eastern and Venezuelan stocks.
This had been chosen for the early tests because it was intermediate
in both sulphur and vanadium concentrations and because it was available
in large supply at the Research Centre. The initial sample supplied
by NAPCA (now GAP) was a Venezuelan heavy fuel oil from the Creole
Petroleum Co. Amuay refinery. When tests confirmed the suitability
of this fuel, a larger supply was obtained in drums for the variable
study. A quantity of oil from this same source in storage in the U.K.
was used for runs in the continuous pilot plant. Table III lists
inspection results of the oil samples and shows the minor variations
between the different Venezuelan shipments. The most significant
difference appears to be the lower vanadium content of the initial
one drum sample. The initial U.K. test fuel reflects the high sulphur
and lower vanadium content which is characteristic of its Middle Eastern
component.
TABLE III
Composition of CAFE Teat LlaBBtooes and Puel Olla
Limestone Composition
Fuel Oil Properties
. Stone:
jnent
3
3
wt. %
i0 released
as SO,
medium ppm
Denbiaiahlre BCR 1690 BCR 1691
(•) (b)
55.2
0.27
0.32
0.07
0.06
O.OT
43.4
-
26.
17.
13.
0.
5.
0.
35.
0.
7
5
5
83
10
19
5
59
45.
2.
14.
0.
1.
0.
35.
1.
0
O2
2
53
14
17
4
0
49
Z
8
0
i
0
37
0
< 10 <
.1
.08
.44
.47
.42
.14
.0
.85
10
ail
Property
Specific gravity
Klneaatlo viscosity
Cs at 140*P
180'P
210*F
Carbon % by wt.
Hydrogen % by wt.
Sulphur " "
mtrogen " "
Conradaon Carbon % by wt.
Asphaltenea % by wt.
Vanadium ppn
Nickel "
Sodlun "
Iron "
Venezuelan. Aauay
(c) (d) (e)
.95
-
54.7
-
85.9
11.5
2.22
0.41
7.5
6.3
270
-
13
4
.957
201
.955
221
U.K..
• 95
98.2
(75) Interpolated
. 41.4
85.9
11.3
2.35
0.35
11.6
7.1
366
4.3
36
3
39.3
85.6
11.4
2.48
0.26
10.9
4.8
345
40
35
4
26.2
84.8
11.6
3.03
.
9.32
3.72
140
20
37
1
(a) Initial sample of BCR 1691 used In screening tests.
(b) Larger supply of BCR 1691 used In variable studies.
(c) Initial sample of oil used In screening tests.
(d) Larger supply of oil used In variable studies.
(e) Large supply of oil used In pilot plant runs.
-------
6. Experimental Programme
Hie experimental programme of Phase I of this CAFB project
has included the following studies:
a Batch reactor oil and lime screening
b Batch reactor fresh bed variable studies
c Batch reactor cyclic test variable studies
d Continuous gasifier commissioning and operating trials.
Table IV summarises test conditions used in these separate
studies.
-------
Prog
Item
Table IV
Suamary of Phase I CAFE Experimental Condition*
Batch Unit Screening
Batch Unit Presh Bed VarUblea
Batch Unit Cyclic Variables
Pilot Plant Operation
Hkmber of Testa
Test Duration
Limestone Used
Oils Used
Gaslfler Temperatures
Regenerator Temperatures (peak).
Gaslfier Bed Depth (static)
Oaslfier Superficial Velocity
Solids Particle Size Range
Line Replacement Rate
Alr/Puel Ratio
Oil Peed Rate
60 to yx> mln
U.K.. Denbighshire;
U.S., 1690; U.S.. 1691
U.K.. yt 3; Venezuelan. 2.2% S
7'tO*C to 900* C
1000*0 to 1090'C
15.2 In to 18.1 in
3.5 to k.S Ct/sec
8OO/1200 microns;
6OO/l
-------
C Results and Discussion
This Riase I CAFB study has consisted of four activities*- eaoh
with its own programme of experiments and apparatus.
These activities are treated separately in the Appendix to this
report.. However, in this section, the discussion is organised around the
CAFB variables and their effects. An attempt is made to summarise and
interpret the findings In light of the total study rather than In sequence
of the particular experiments which produced the information.
1.
Oil and Limestone Comparisons
Tests to define the suitability for CAFB of oils and limestones
available in the United States were performed early In the program*, so
that larger supplies of the appropriate materials oould be obtained for
detailed variable studies.
a. Oils
One oil sample had been supplied as being typical of a heavy fuel
oil available on the East coast of the United States. This Venezuelan oil
contained a much higher vanadium to sulphur ratio (2?0 ppm/2.22£s = 121)
than the blend used In prior CAFB experiments (140 ppm /3.035*8 « 46).
Thus for an equal sulphur pickup, lime would be exposed to a ouch higher
vanadium level with the Venezuelan fuel. The test results In Figure 16
show nearly equal sulphur removal performances with the two oils in fresh
bed batch tests using UK lime in the 1-A test unit. Both gave better than
90# sulphur removal at calcium utilisations approaching 10%. On the basis
of this comparison, the Venezuelan oil was Judged to be satisfactory for
further study and a larger supply was obtained. All subsequent work,
including runs in the continuous gasifler pilot plant, has confirmed that
high levels of sulphur removal can be obtained from this oil In spite of
its high vanadium content. _
aomnm UPPVAL i
90
I '
;j
i
7C
a
6C
\\
i
_t'ji
• ; . -
I
.;
- •• r .
1
G-~*.-Q~
Q*-' ^{J*
^
... .. ,. ...
!-• —
;- - |
i . .
'
. ! ••
. ; ".'-
--I-..
s,:: -""•"
. T~_ .. V
-
~
• -
1
- '- '. .' .
- -.(..:..
..f.-.J
•0-^:-;
x I—"
D ^"O
u" D
1 — -
~" . | ^
..• :(.::.•:
i
- • • •
~--|— -
i
i
••- i • - •
: 1 .: :
' |: '•-
• r • •
1
1
....
• - |
~-|-:-r
I
*" 3»i*ol Fu.1 LlJH * SOU. Atr
22 O U.K. O.K. 27.0
O D U.S. U.K. 36.6
:
i '
* ""
0 -05 .10 .15 .so .25 .30
PWCTION OO REACTED
FIG. 16
-------
b. Limestone Comparison Tests
Samples of two limestones available in the North-eastern United
States were supplied. These stones were compared with each other and with
the UK stone in a series of cyclic tests in the 1-A batch unit. The UK
stone was subjected to 21 cycles of gasification-regeneration at conditions
listed below.
TABLE V
Conditions for Llaestone Comparison
Limestone Particle Size
Gasification Bed Temperature
% Stoichioraetric Air for Gasification
Superficial Air Velocity
Bed Depth (static)
Gasification Cycle Length
Maximum Regeneration Temp
Lime Replacement Rate (as limestone)
850 - 1200 microns
800 - 900 °C
25
4 ft/sec
15.5 In
60 min
1100 °C
0.75 wt CaO/wt S.
Figure 1? shows the results - Sulphur removal efficiency declined
for nearly ten cycles and then lined out at approximately ?2#. This
result agrees well with the curve of lime replacement versus sulphur
removal efficiency established earlier in other apparatus (Appendix B,
Figure B - 7). The data on which Figure 17 is based are listed in
Tables C - 1 and C - 3, runs 25 - 53.
BED LIFE STUDY
(U.K. FUEL AND STONE)
3?
o
ui
100
80
C 60
&
o
i 40
|
n 20
x x
X
FUEL SULPHUR
AIR RATE
MAKE-UP RATE
• 3-03%
• 23-25% OF STOICH.
" 0-75 wt» CoO/wt. S
10
CYCLES
no. 17
-------
Stone BCR 169! was tested under nearly Identical conditions.
However, to achieve an equal CaO replacement rate a higher actual stone
replacement rate was required since 1691 has only 8l# of the CaO content
of the UK stone. The efficiency vs cycle behaviour of 1691 was nearly
Identical to that of the UK stone. The lined out level of 1691 was also
sulphur removal.
These results are plotted in Figure 18 and listed in Tables C-l
and C-j5, runs 54-72. It was not possible to use the same replacement
rate in the initial cycles with limestone 1690. Preliminary tests showed
lime losses in the first few cycles to be much higher than 0.75 wt CaO/irt.S
Therefore the cyclic test was started at 1.5 wt CaO/wt S. This rate was
maintained for ten cycles (Figure 19) by which point the efficiency had
lined out at 80#. The high rate of solid losses encountered in the first
cycle declined in subsequent cycles permitting a reduction to 0.75 wt CaO/rt.
replacement rate in the final ten cycles. The sulphur absorption efficiency
then declined to about 64#. Data are listed in Tables 0-1 andC-3, runs
78 - 95.
CYCLE TESTS (U.S. LIMESTONE 1691)
100
a 80
o
I 60
i
0.
CO
UJ
ID
u.
40
20
0
X X
x x
X X
FUEL SULPHUR
MAKE-UP RATE
3 03%
0-75 wt. CaO/wt. S.
1
10
CYCLE
20
FIG. 18.
-------
CYCLE TESTS (U.S. LIMESTONE 1690)
IUU
a 80
Ul
s
S 6O
5
I
-I
3 40
_J
bJ
5
* 20
o
N 1 1 II
X
\
- *>J-JLirJL*7x
^x^ *
* *
-
_ X » MAKE-UP RATE, 1.52 WT. CoO/WT.S
• « MAKE -UP RATE, 0.75 WT. CoO/WT.S
^~ ^
II 1 1
0 5 10 15 20 29
CYCLE FK5.I9.
Limestone 1690 gave.good sulphur removal efficiency in its
first cycle, but in other respects it was inferior to 1691. Its low
onloium content (60# of limestone 1691) gave it a low sulphur capacity
on a total stone basis. It underwent severe attrition during calcina-
tion and in its initial gasification cycle. But even more serious it
tended to sinter and agglomerate at operating conditions. Figure 20
shows some of the agglomerates. This tendency to sinter and agglomerate
would present severe operating problems to a continuous CAFB unit. It
was not observed with any stone other than 1690.
-------
-38-
AGGLOMERATES FROM BATCH REACTOR AFTER
CYCLE TEST WITH BCRI690 STONE
FIG. 20.
-------
Lime losses during the stone comparison tests were estimated by
weighing the beds before and after the tests. Figure 21 shows lime
losses per oycle with the UK stone and the two US stones. The total
lime charge to the reactor was 8.2 kg.
HISTORY OF LIME LOSSES IN CYCLIC TESTS
1000
800
UJ
600
3 400
UJ
_i
fe
UJ
»-
200
U.S. LIMESTONE 1690
U.K.DENBIGHSHIRE STONE
U.S. LIMESTONE 1691
I
4 6 8 10 12 14 16
BATCH GASIFICATION CYCLE
18 20 22 24
F|G2|
-------
The UK stone settled down to a loss of about 140 g/cycle which
was 1.1% of the bed or 1.4# on fuel fed. The US stone l6°X) had a very
high loss In the first cycle, but lower losses thereafter. The loss
rite declined still further when replacement rate was reduced. During
the final cycles the losses averaged 244 g/cycle, which is ~yf> of the
bed or 2.5 wt % on fuel feed rate of 165 g/min.
Losses from stone 1691 started low, increased to a maximum at
about cycle 10 and then decreased again. The average rate of loss for
cycles 4 through 19 was 112 g/cycle which was 1.1# of fuel feed rate.
On the basis of these cyclic tests, limestone 1691 was selected
as the better of the two stones suggested by NAPCA. Except for requiring
a higher absolute make-up rate because of its lower CaO content, it
appears to be as good as the UK Denbighshire stone for CAFB operations.
On the other hand, stone 1690 does not appear very suitable for the CAFB
process. It would require an even higher replacement rate, and it could
cause serious operating problems.
2. Sulphur Removal During Gasification
Provision of sulphur-free fuel is the primary purpose of the
CAFB process. Consequently major attention has been focussed on the
factors which determine the sulphur removal efficiency, abbreviated SRE.
We have found that gasification temperature, air/fuel ratio, fluldlsation
gas velocity, bed depth, limestone particle size, lime replacement rate,
and the quantity of sulphur to which the IJme is exposed each cycle, all
affect the SRE. A number of variables interact with each other and
therefore must be considered together.
a. Bed Processes During Gasification
The following visualisation of the various processes which must
take place in the fluidised bed is an aid in interpreting the experimental
observations made in this study. The chemical processes which occur in the
CAFB gasifier are a complicated mixture of cracking, coking, oxidation,
gasification, and absorption. Oil enters the bed some distance above the
air distributor in a region where there may be little or no oxygen. The
oil contacts the hot lime particles, is partly vapourised, and undergoes
cracking and coking. Most of the organic sulphur compounds probably crack
to H2S which reacts with CaO in the upper portion of the bed to form CaS.
Lime particles with a coating of coke and heavy hydrocarbon eventually
circulate to the lower portion of the bed near the air distributor where
the coke is oxidised. This oxidation supplies heat to the system and
removes the coke that acts as a barrier to further sulphur abSorptibn*
Under some conditions, there may be enough oxygen in the vicinity of
the distributor, if temperature is high enough, to convert some of the
calcium sulphide to sulphate or even to oxidise the CaS to CaO and 30 .
In the latter case, the SO would be reabsorbed higher in the bed. ^
-------
b. Fresh Bed Sulphur Absorption Curves
The early part of a batch gasification cycle with a bed of fresh
lime usually gives an SRE close to 100$. The duration of this high
efficiency period and the subsequent variation of SRE with extent of
lime utilisation depends strongly on gasification conditions. In the
1-A batch reactor it was not possible to separate effects of air/fuel
ratio from those of gasification temperature as the two were strongly
cross-correlated. As there was no independent temperature control, the
gasification temperature increased as air/fuel ratio increased and
oxidised a larger fraction of the fuel.
Figure 22 shows 1-A unit sulphur absorption curves at several
air/fuel ratios and temperatures. The conditions and results are
summarised in Table VI and reported in detail in Appendix Tables C-1
and C-3. At the lowest air/fuel ratios the SRE fell immediately to a
low level and then began a further gradual decline. At intermediate
conditions there was a period of high SRE followed by a sharp drop and
a gradual decline. At the highest air/fuel ratio, the high SRE period
persisted longer and the subsequent decline was more gradual. The low
temperature-low air/fuel ratio conditions also were associated with high
levels of carbon on the lime at the end of the cycle.
TABLE VI
Conditions for 1-A Unit Fresh Bed Tests
Plotted in Figure 22
Run Symbol % Stoic Air Temp "C #C on Stone Lime Fuel
98 O 19.1 815 6.04 UK UK
96 D 19.3 800 - UK UK
99 A 20-5 810 5.59 UK UK
102 Q 22.8 860 1.71 UK UK
103 4- 24.8 865 0.15 us US
121 X 26.7 850 .87 UK UK
-------
SULPHUR REMOVAL QUOTES AT DIFFERENT
AIF7FUEL RATIOS
]oq
Run
98
96
99
102
103
121
Symbol
O
D
FRACTION CaO REACTED
% Stoich. Air Fuel
19-
19-
20.
22.8
24.8
26.7
.1
• 3
• 5
U.K.
U.S.
U.K.
Lime
U.K.
U.S.
U.K.
FIG.22.
Similar results, but at much higher levels of calcium utilisation,
were obtained in the 4-A and 4-B reactors where better insulation permitted
higher gasification temperatures. Figure 23 shows these curves. Again the
low temperature-low air/fuel ratio test suffered rapid decline in SHE. But
at 23»3?5 of stoichiometric air and 850°C, the lime maintained high SHE to
over 20$ calcium utilisation.
-------
Run Symb
24-4B
27-4B
6-4A
28-4B
•
X
o
A
Conditions for these tests are listed below.
TABLE VII
Conditions for Sulphur Removal Tests in No. 4 Batch Reactors
6 ft/sec Air Velocity
% Stoic Air Temp °C #C on Stone Lime Fuel
31.4 870 .04 US BCR 1691 Venezuelan
23.3 850 .38
19.7 845 4.68
14.8 780 11.88
n
ii
tt
100
S
I
60
40
RESULTS AT 6FT/SEC IN NUMBER 4 UNITS
SULPHUR REMOVAL CURVES AT DIFFERENT AIR/FUEL RATIOS
10 20
% CALCIUM UTILISATION
30
FIG .23.
40
-------
These tests Indicate that levels of calcium utilisation
approaching }0# with maintenance of SRE over 90$ are possible if proper
gasification conditions are employed.
c. Temperature and Air/Fuel Ratio
Because the fuel handling capacity of a given size gaaifier
increases as air/fuel ratio decreases, efforts have been made to operate
at lowest possible air/fuel ratios and to define the limitations that
may exist. Use of cooling coils in the No. 4 units made possible the
separation of temperature and air/fuel ratio effects. The fresh bed
variable study with US fuel and limestone included tests over a range of
air/fuel ratios at 800°C and at 840 - 870°C. Results in Figure 24 show
that the higher temperature operations produced good SRE (measured at 3$ S
on lime) to below 20$ stoichiometric air. However, at 800°C the SRE at
5$ S declined at air/fuel ratios below 20$ of stoichiometric, but was equal
to the high temperature result at 25$ of stoichiometric air.
INTERACTION BETWEEN AIR/FUEL RATIO AND BED TEMPERATURE
100
o
LJ
O
i
oc
I 90
to
80
5% BY WT. SULPHUR IN BED
x 840-870°C
• 800 °C
1
1
15
20 25
% STOICHIOMETRIC AIR
30
FIG. 24.
35
Table VIII shows that in these tests also, the level of carbon
on stone increased sharply with lowered temperature and air/fuel ratio.
-------
- 45 -
TABLE VIII
Stone Carbon Content
in Temperature Comparison Tests
Run No. Gasification Temp
Average
°C
24-4B 870
27-4B 850
6-4A 845
8-4A 800
26-4B 800
Air/Fuel Ratio
% of
Stolchiometrlc
31.7
23.3
19-7
26.1
18.2
Carbon on Lime
at Cycle End
wt. %
0.04
0.38
4.68
1.59
7.12
The sulphur absorption curves obtained in batch tests suggest
that loss of sulphur removal efficiency can take two forms: (l) a very
rapid loss bearing little relationship to the amount of sulphur reacted
with the lime; and (2) a slow decline as the available lime is converted
to CaS. The data Indicate that the rapid loss in efficiency in some tests
is due to building of a layer of coke or carbon on the reactive lime which
acts as a barrier to sulphur absorption. Preferred CAFB operating con-
ditions are those which avoid build-up of this carbon barrier.
Figure 25 shows carbon accumulation plotted as the ratio of net
carbon deposition to oil fed versus air/fuel ratio, a form suggested by
R. Newby.17)
VARIATION OF CARBON DEPOSIT WITH AIR-FUEL RATIO
.05
O
3 .04
.01
02
g .01
BATCH UNIT
14
18 22 26 30 34
AIR FUEL RATIO, % OF STOICHIOMETRIC
JA
38
42
FIG.
-------
At high air/fuel ratios carbon accumulation is negligible, but
below some critical level the amount of carbon laid down increases rapidly
as air/fuel ratio decreases. Carbon levels observed in tests in the No. 4
units and the tall No. 3 unit were generally lower than those found in the
No. 1-A unit, but the cause of this difference is not yet clear* Since
temperature and air/fuel ratio were related in the batch units, similar curves
are obtained if carbon is plotted versus temperature. The general behaviour
of the data agrees with a mathematical model based on carbon oxidation
kinetics.
This model predicts carbon-on-lime concentrations of the same
magnitude as those actually observed and also predicts a steep rise in
carbon concentration as temperature and/or air/fuel ratio decreases.
The model,(described in Appendix K) assumes that coke is laid
down from the fuel oil at a rate proportional to oil feed rate and
Conradson Carbon concentration. It assumes that thin layer carbon
burning kinetics apply; i.e. that carbon is burned off at a rate
proportional to coke concentration in the lime, to oxygen concentration
in the burning zone below the oil injector, and to an oxidation rate
constant which increases exponentially with temperature. The effect of
temperature on C0/C02 ratio in the product also increases carbon burning
rate with temperature. The steady state carbon concentration is that
which produces a coke burning rate equal to the coke deposition rate.
Figure 26 compares measured concentrations of carbon-on-lime with curves
predicted by the model for 15 and 30# of stoichiometric air at 4 ft/Bee
superficial velocity.
-------
EFFECT OF TEMPERATURE ON CARBON CONTENT
OF BATCH GASIFICATION LIME
Lj
£
J
T
5
^
Q
C
t
14
12
10
8
6
4
2
0
• • • i I
PREDICTION BY BURNING RATE MOpEL
0 DATA, AIR/FUEL > 23% OF STOICHIOMETRIC
A DATA, AIR/FUEL < 23% OF STOICHIOMETRIC
OPEN POINTS 1
SOLID POINTS
V- i
\I5%STOICH.
- VAIR/FUEL
_\
i. \
\ -x
30% \
STOICH *
AIR/FUEL '
1
00 750
•A UNIT
N0.4 UNITS
*y
\ A o _
A A\ ~~
• ""^ . «. A OA. ""
800 850 900 950 IOCK
GASIFIER TEMPERATURE,°C FIG.26.
Although the model predicts the correct directional trends, the
data at low air/fuel ratio show much higher carbon levels than predicted
below about 850°C. Failure of several assumptions of the model may cause
this effect. For example, if a greater fraction of oil than Conradson Carbon
deposits as carbon on lime in the upper part of the bed, there may be an
oxygen limitation in the carbon burning region. In this case, the effect
of temperature on C0/C02 ratio may be very important in affacting efficiency
of oxygen utilisation. Also, once an appreciable fraction of the lime
surface becomes carbon coated, thin layer carbon burning kinetics no longer
apply, and burning rate ceases to increase in proportion with carbon
concentration. When the lime is completely covered, burning rate will be
independent of carbon concentrations. Under these unfavourable conditions,
carbon content of the lime does not reach an equilibrium level but increases
throughout the batch cycle.
-------
Evidence for this type of behaviour was obtained in 1-A unit
batch runs in which carbon content was measured on lime samples drawn
during a long gasification cycle. Figure 27 shows these data. In the
tests at temperatures of 850°C and lower, the carbon content Increased
throughout the cycle. At 860°C and 23.3$ of stoichiometrlc air, the
carbon content also increased. But at 860°C and 29.3# of stoichioroetric,
carbon level remained insignificant. . Increasing temperature to 865°C
permitted operation at 24.8# of stoichiometric air without increase of
carbon content.
CHANGE IN CARBON CONTENT OF LIME DURING
-BATCH GASIFICATION CYCLE I-A UNIT
40
80 120 160
GASIFICATION TIME.MINS
240 280
FIG. 27
-------
The scatter of data in Figures 25 and 26 and the large change
in behaviour over a narrow range of conditions shown in Figure 27
indicate that the temperature range below about 870° and the air/fuel
ratio range below 25# of stoichiometric is somewhat unstable. Results
in Nos. 3 and 4 units and particularly, experience in the continuous
pilot plant demonstrate that successful sulphur removal is possible at
air/fuel ratios below 20# of stoichiometric provided that temperature is
high enough. Figure 28 shows sulphur removal for a 15 hour period during
the first t^st run in August. IXiring this period air/fuel ratio averaged
17J6 of stoichiometric and SHE averaged about 95$• Gasifier temperature
varied between 8^0 and 88oeC.
LOG OF CAFB PILOT PLANT CONDITIONS
AT LOW AIR/FUEL RATIO
100
.- u O
4» 9 1£
3*£ 90
O
-t 18
-S u
o 5 .2
a: ^ f.
(A U
« o 2
16
1
li
u
0_
w
•2
.£ S
(/> 0>
p °-
O £
900
800
\r
I
7 12
Time, Hour
August 13
17
FIG. 28.
-------
Carbon concentrations were not measured on gaslfier solids
in August, but data from the third run in November-December show that
carbon concentrations on these solids averaged 0.1# with a measured
high of 0.4g^. These data appear in Appendix Table J-VII.
Because of its low surface area to volume ratio, the continuous
pilot plant loses a lower fraction of its heat and can sustain high
gasification temperature at much lower air/fuel ratios than the batch
units. It appears that operation quite close to the adiabatic limit of
about 15$ of stoichiometric air will be possible without excessive
carbon formation and loss of sulphur removal capacity.
The gasifier should be operated at 870°C or higher to provide
a safe margin away from excess carbon laydown. In the batch unit
variable study, conditions were selected to avoid undue carbon formation
except in runs where air/fuel ratio and temperature were under study.
d. Upper Temperature Limitation
Although the equilibrium concentration of 1198 and other reduced
sulphur compounds is very low at temperatures well over 1100°C,
(Appendix A), experience has shown that sulphur removal efficiency
decreases as gasifier temperature exceeds 900°C. Figure 29 shows this
basic temperature effect on SRE.
BASIC EFFECT OF BED TEMPERATURE
o
UJ
tr
CE
ID
100
90
80
70
60
5% BY WT. SULPHUR IN BED
60
800
1
850 900
TEMPERATURE, °C
-------
- 51 -
Batch test sulphur removal efficiency at 5 wt. % S on lime
falls rapidly as temperature rises above 900°C. The explanation for
this behaviour is thought to lie in the presence of an oxidising zone
close to the distributor where CaS and the incoming air can react to
form CaS04 and CaO + SOg. The amount of S02 released increases with
increasing temperature. It is likely that at temperatures above 900°C
the quantity of SC>2 released is too large to be completely reabsorbed
higher in the bed.
The optimum gasification temperature range is therefore 870 to
900°C to avoid excess carbon formation at the low temperature and local
release of SOg at the high temperature.
e. Gas Velocity and Bed Depth in Fresh Bed Tests
In this work, the practice has been to discuss gas velocity in
terms of superficial air rate - that is, the velocity which the air
supplied to the reactor would have at operating temperature and pressure
in an empty bed. The total gas velocity, of course, is different due to
the presence of cracked oil products and the bed material itself. Bed
depths also are normally expressed as static bed depths.
With other factors constant, and at an operating temperature of
about 870°C, the basic effect of increasing air velocity is to decrease
sulphur removal efficiency as Figure JO shows.
BASIC EFFECT OF SUPERFICIAL GAS VELOCITY
100
90
e
0 on
12 80
LJ
£ 70
Q.
UJ
55*
60
50
40
40
5% BY WT. SULPHUR IN BED
I
I
50 60 70 80
SUPERFICIAL GAS VELOCITY (FT/SEC)
90
FIG.3Q
-------
These data are from fresh bed tests with US fuel and limestone.
The SRE at 5% S in the lime remained close to 100$ over-the
velocity range of 4.3 to 6.1 ft/sec but decreased to 66# at 7.9 ft/sec.
Changing bed depth at 6 ft/sec produces a similar effect shown
in Figure 31 •
BASIC.EFFECT OF BED DEPTH
100
<£ 70
60
5% BY WT. SULPHUR IN BED
10-0
15-0
BED DEPTH (STATIC) INCHES
The SRE remained near 100$ for static bed depths of 20 inches
and 15.5 inches but declined to ?8# at 10 inches depth. At 8 ft/sec
the SRE was only 40# with a 10 inch bed but continued to improve with
depth and reached nearly 100# with a 20 inch bed.
-------
These gas velocity and bed depth effects can be correlated
by consideration of the superficial gas residence time in the bed.
Use of the residence time above the fuel injector gives a better
correlation than total residence time as would be expected if there
is little back-mixing of the gasified fuel downward from the injector.
Figure 32 is a curve of SHE versus superficial gas residence time above
RELATIONSHIP BETWEEN SUPERFICIAL RESIDENCE TIME AND
FUEL SULPHUR REMOVAL AT 5% BY WT SULPHUR
IN BED (FRESH BED TESTS)
100
80-
*
£
kl
I
bl
60
-------
The first order rate constant under these conditions (870°C, 5# S _-,
in lime, starting from a fresh lime bed of 300 to 3175 microns) was 33-9 sec ,
and the pseudo time value was (t - 0.051) seconds where t is the superficial
gas residence time above the fuel injector. The value of 0.051 subtracted
from t is empirical, but may be due to the time required for fuel vapourisation
and cracking before adsorption begins. Thus, the equation for sulphur
removal efficiency in the fresh bed test at about 870°C is:-
SRE = 100 Ql- EXP (33-9 ( t- .051))] (1)
A reversal of velocity effect was observed in tests of the high
gasification temperature of 950°C. At this temperature, sulphur removal
is poor as discussed earlier. However, increasing velocity from 6 ft/sec
to 8 ft/sec actually increased efficiency from below 20$ to around 45$.
This effect is not well understood, but possibly is due to increased bed
circulation rate which minimised local temperature rise in the zone near
the distributor and thus reduced re-release of SOg by local CaS oxidation.
Increasing bed depth at 950°C had the expected effect of improving
sulphur removal efficiency. The test at 20 inches static bed depth and
6 ft/sec superficial air velocity gave 82$ SHE at 950°C compared with
values below 20$ obtained with 15.5 and 10 inch beds.
f. Lime Particle Size
Early CAFB work had been done with narrow cut limestone size
fractions of 600 - 1200 microns and 800 - 1400 microns. Yield of these
narrow cuts is low in crushed stone, and for commercial economy a wider
cut stone is preferred.
A V8-inch to dust size fraction was considered from the stand-
point of cost, but at velocities in the gasifier the low size end of this
fraction would pass straight through the bed. Therefore a 300 micron cut
point was selected for the lower particle size. Three particle size ranges
were compared in batch unit fresh bed tests; 300-3175 microns, 600-1400
microns, and 1200-3175 microns. (The upper size of 1 /8-inch is approximate).
A fourth size range, 600- 3175microns, has been used in some of the con-
tinuous pilot plant operations because of difficulty in screening large
quantities of stone to 300 microns.
The results of fresh bed batch comparisons of the three selected
stone sizes appears in Figure 33. The 300-3175 micron material maintained
a high sulphur removal efficiency to nearly 30# calcium utilisation. The
600-1400 micron lime also gave reasonable performance but SHE with the
1200-3175 micron lime declined rapidly.
As Newby(f) observed, at low lime utilisation the relative
performance of the three stones varied inversely as the square of their
number average particle diameters. Simple reduction of particle external
surface area with increasing particle diameter cannot account for the
magnitude of efficiency loss observed. Some additional factor, such as
diffusional resistance, must be included to explain the observed behaviour.
-------
- 55 -
BASIC EFFECT OF PARTICLE SIZE RANGE
100
u
3
g eo
t
UJ
60
a:
40
20
X 300-3175 jt./f- 496
• 600-1400 ;i,£ • 682
o I200-3I75M.M • 1700
10
% Co UTILISED
20
VARIATION OF SULPHUR REMOVAL EFFICIENCY WITH
FUNCTION OF LIME UTILISATION
100
T
»V-w-
-x*
_L
AVERAGE DIAMETER
496
682
1700
J—L_J L_L
J_
SYMBOL
x
•
o
J I ' ' I
IO'4 6 8 10'* 2 3 4 5 6 8 K>~
A , 2. .[PARTICLE AREA x
V d I VOLUME REACTED SHELL THICKNESS
456 8 10"
.-a
, MICRONS'
FIG. 34.
-------
- 56 -
Figure }4 correlates test data from the three stones using the
assumption that reacting lime particles are spheres with a shrinking
core of unreacted material. The reaction rate at any extent of
reaction is assumed to vary with surface to volume ratio divided by
shell thickness. That is, SHE is a function of A x 2
V d
where A = total particle external area
V = total particle volume
D-d = diameter of unreacted core.
The fraction, U, of the sphere represented by the reacted shell
is related to particle diameter, D, and shell thickness, d, by geometry.
f d\ - 3 /d\2 + (d\
ID) VDJ \D)
d\ 3 (2)
U
and A = 6
V D
The points in Figure 34 were determined by:
1. finding U at a given SRE and D from Figure 33
2. finding d from equation 2 at U and D
3. finding A/V from equation 3
4. calculating A 2 .
V d
The relationship illustrated by Figures 33 and 34 shows that lime
particle size can be traded for extent of lime utilisation to obtain a
desired level of sulphur removal efficiency.
This relationship, of course, applies to fresh bed tests
at the base line experimental conditions of 6 ft/sec superficial gas
velocity and 15.5 inch static bed depth. As Table IX shows, increasing
the bed depth improves fresh bed performance with the large size stone
to an appreciable extent.
TABLE EC
Interaction between Bed Depth
and Particle Size Range
Static Bed Depth (inches)
% Fuel Sulphur Removed at 5% 10 15.5 20
by wt. S in bed (300-3175>> lime) 78.0 99-3 97.6
% Fuel Sulphur Removed at 5#
by wt. S in bed (1200-3175P lime) 34.9 37.1 74.5
-------
- 57 -
When adding fresh limestone under continuous operating
conditions, the effect of particle size is complicated by lime losses.
Cyclic test results on the effect of particle size differed from the
fresh bed tests in that the finest size lime make-up did not give best
results. Table X shows this comparison.
TABLE X
Effect of Limestone Particle Size
in Cyclic Test
Test No.
T-3(Cycles
1 to 14)
T-3 (Cycles
24 to
Replacement
Limestone
Particle Size
300 - 3175
600 - 1400
Limestone, BCR 169!;
Gasification Temp.
Air/Fuel
Superficial Air Velocity
Static Bed Depth
Gas Residence Time
above fuel injector
Replacement
Rate
wt. CaO/wt. S
2.38
2.37
61
68
Sulphur
Differential
/ wt.S fed i
100.x wt. lime
3.0
3.1
Fuel, Venezuelan 2.3# S
870°C
25% of Stoichiometric
6 ft/sec
15-5 inches
0.15 sec
This effect was due to the higher relative losses of the finest
particle size fraction during in-situ calcination of the stone. Separate
tests revealed an average loss of 52.5# on calcination of 300-3175 micron
stone compared with 42.7# of the 600-1400 micron stone. The weight loss
due to C02 only is 37#. The effective make-up rate of the wide range
stone, therefore, was less than that of the narrow range stone at equal
addition rates. This effect did not appear in fresh bed tests where the
reagent was added as precalclned lime.
In the third run of the continuous pilot plant, a limestone feed
of approximately 600-3200 micron size was employed because of practical
difficulties in screening the stone to the 300 micron cut point at high
rates. Unless the stone is predried to a very low moisture content, a
300 micron screen blinds very quickly and gives large quantities of dust
in the product.
-------
g. Lime Replacement and Sulphur Differential
Two of the important variables in continuous CAFB operation
cannot be studied in fresh bed batch tests but can be simulated in
cyclic batch tests. These are lime replacement rate and sulphur
differential. Replacement rate is the ratio of fresh limestone make-up
to the quantity of fuel sulphur fed to the bed. Sulphur differential
is the percent sulphur picked up on the lime between regenerations. In
cyclic tests it is convenient to speak of projected sulphur differential
(PSD) which is the increase in percent sulphur on lime which would result
if sulphur removal efficiency were 100$ throughout the cycle.
Projected sulphur differential is therefore:
(% S in fuel)(fuel feed rate)(gasification cycle length)
(wt. lime in gasifier)
Lime replacement rate has been known as an important variable for
some time. In the earlier CAFB batch studies it had been recognised that
sulphur removal efficiency of a lime bed declined with repeated gasifica-
tion-regeneration cycles but that the decline could be halted and a lined
out SRE obtained if a portion of the lime is replaced after each cycle.
The greater the replacement rate the greater is the lined out SRE. With
the UK test fuel and stone at superficial gas velocity of about 4 ft/sec,
a lime replacement of about 1.8 wt. S per wt. CaO maintained sulphur
removal at nearly 100# efficiency, and replacement at 0.75 wt/wt gave
12.% sulphur removal. An early test with U.S. stone BCR Ib91(page 36)
gave similar results, leading us to believe that the U.K. stone and
BCR 1691 would give equal sulphur removal when replaced at equal CaO/S.
The initial cyclic tests in the current study with Venezuelan fuel
and US BCR 1691 carried out at 6 ft/sec superficial air velocity and
300-3175 }i stone gave appreciably lower lined out SRE than those early UK
fuel and stone tests. A check run (Test T-12) in one of the new batch
units with the UK fuel and stone at conditions of the earlier tests produced
nearly the same SRE as predicted (98$ vs 92$). Thus the difference was
not due to a change in equipment.
The reduction in gas residence time caused by increased velocity
from 4 to 6 ft/sec, and the loss of small diameter stone during calcination
discussed on page 57 cannot explain all of the efficiency loss observed.
Additional factors must play a significant part.
But as yet, these factors have not been fully defined. Because of this
unexpected lower efficiency at mild conditions, the cycle test variable study
was extended to find conditions giving high levels of sulphur removal with
the BCR 1691 stone without excessive lime replacement rates and low gas
velocity.
Increasing bed depth at equal cycle length was found to improve
efficiency out of proportion to the change in residence time. This
increased benefit of a larger lime bed can be explained by the reduction
in PSD which occurred. Lime cycled through a lesser change in sulphur
concentration appears to retain greater absorption activity.
The effect of PSD was verified by reducing cycle length at constant
bed depth. Table XI shows this comparison.
-------
- 59 -
TABLE XI
Effect of Sulphur Differential (PSD)
Test No. Lime Replacement Rate PSD Lined out SHE
wt. CaO/wt. S ~W T"
T-l (Cycles 2.43 3-0 6l
1 - 14)
T-9 2.55 1.9 86
Limestone BCR 1&91; Fuel, Venezuelan 2.3# S.
Gasification Temp. 870°C
Air/Fuel 25$ of Stoichiometric
Superficial Air Velocity 6 ft/sec
Static Bed Depth 15.5 inches
Gas Residence Time 0.15 sec
above fuel injector
This large effect of sulphur differential implies that highest
possible lime circulation rates between gasifier and regenerator should
be used to improve sulphur removal. Heat balance considerations, of
course, will determine the minimum practical sulphur differential
h. Correlation of Fresh Bed and Cyq»Le Test Results
A single fresh bed test is much quicker and easier to perform
than a series of cyclic tests. It therefore would be very desirable to
predict cyclic test and continuous unit performance on the basis of a
fresh bed test. To explore this possibility, a function derived
from the first order rate relationship used for fresh bed tests
was modified to include effects of bed replacement rate and sulphur
differential. The derivation appears in Appendix F. The final
expression is: _
1
SRE = 100
1 - 0.10 k 9 m
(*)
10
where:
SRE = % sulphur removal efficiency
k = first order rate constant from fresh bed tests = 53.9 sec"1
6 = (gas residence time above fuel injector - .051) sec.
m = Lime replacement rate, wt.CaO/wt. S
d = Projected sulphur differential, %.
Since the rate constant was determined for a particular set of
conditions, this equation applies to the Venezuelan fuel - BCR 1691
limestone combination at temperatures in the range 840-900°C, air/fuel
-------
- 60 -
ratio of 20#,to 30# of stoichiometric, and limestone of 300- 3175 micron
particle size range. This function fits the cyclic test data with a
correlation coefficient of .85 but, as Figure 35 shows, usually predicts
slightly lower values for SHE than observed.
CORRELATION BETWEEN EXPERIMENTAL AND
CALCULATED LINED OUT EFFICIENCIES AND USING EQUATION.4.
100
60 80
CALCULATED EFFICIENCY(%)
100
FIG. 35.
-------
- 61 -
In order to obtain a more precise correlation of the actual
cycle test data, a regression was performed using a function which,
though similar to equation 4, permitted assignment of different weights
to the separate variables.
SRE % = 100
1 -
EXP
wtx
dz
(5)
This function contains four adjustable coefficients or exponents whereas
the constants in equation 4 were determined without reference to the cycle
test data. In comparison with equation 4, the term w contains the a x k
product while x, y, and z are weighting exponents for gas residence
time, lime replacement rate, and potential sulphur differential.
Regression of the cyclic test data with equation 5 yielded the
following result with a correlation coefficient of 0.94.
SRE
= 100
1 -
EXP
9.285
(6)
The experimental results indicate, therefore, that potential
sulphur differential has a more marked effect on lined out sulphur removal
efficiency than suspected and that gas residence time has less.
Efficiencies calculated for a variety of values of the independent
variables are listed in Tables KI through Kill. Figure 36 shows the effects
of residence time and make-up rate at the reasonable potential sulphur
differential of 2 wt. %.
The figure indicates a requirement for gas residence time
greater than 0.3 sec to achieve greater than 90$ sulphur removal with
stoichiometric replacement of BCR 1691 limestone at 300-3175^1 particle
size range.
Further work, especially with deeper beds, is necessary, however,
to prove the reliability of equation 6 for extrapolation.
-------
-62-
PREDICTED LINED OUT SULPHUR REMOVAL EFFICIENCYES(PS02OWT%)
(BCR 1691) (300-3175)1)
100
90 —
[ ]• MAKE-UP RATE
LBCoO/LB S
I-
EXP
0 01 02 03 04
RESIDENCE TIMt(SEC)
05
06
07
08
FIG. 36.
-------
- 63-
i. Continuous Pilot Plant Experience
Appendix J contains results of the three runs of the continuous
pilot plant. Figures J-l and J-16 show the variation of sulphur removal
efficiency with time during tests 1 and 3. In the continuous run period
of the first test, SHE ranged from about 80 to 100# with an average of
90$. Lime replacement rate varied from about 1.1 to over 3 moles/mole S.
Sulphur removal from the fuel gas was essentially 100# throughout
the second test in spite of beddepths as low as 11-12 inches w.g.
However, lime replacement was quite high, ranging from about 3 to 4.5
moles per mole of sulphur, and a portion of the sulphur reacted with lime
that passed through the boiler. In the final test run, sulphur removal
covered the range from 75# to 100# with changes in bed depth
appearing to be a major cause of the variations observed. There appears
to be little difficulty in achieving SHE of 95 to 100# at stoichiometric
lime replacement rate and gas velocity of 4 ft/sec with the Denbighshire
stone used in these pilot plant tests.
3- Regeneration Behaviour
The function of the regeneration step is to convert sulphur to
a concentrated S02 stream while restoring the lime to a form capable of
active sulphur absorption. During the regeneration step the lime must
be heated from gasification temperature of 850-900°C to regeneration
temperature of 1000-1100°C. Energy for this heating comes from oxidation
of carbon and sulphur on the stone. At the lower temperature range, this
sulphur oxidation proceeds to form calcium sulphate.
The two principal reactions are:
CaS + 2 02 > ca SCty
CaS + 3/2 Q2 > CaO + S02
At high temperatures a third reaction comes into play:
CaS + 3 Ca S(ty > 4 CaO +4 S02
a. Fresh Bed Tests
A typical gas composition profile for batch regeneration is shown
in Figure 37 (Fresh Bed Test 3).
(1) Figures in Appendix J are located on pages 291 through 334.
-------
EXIT GAS COMPOSITION DURING REGENERATION
(TEST 3)
30
-* 25
§
T
I
TEMPERATURE,°C
915 990 1060 1090 1130 970 870 800
10
TIME (MIN)
15
There is an initial high carbon oxide evolution as carbon is
burned off the lime, and the temperature rises. SOo then increases
and reaches a maximum at temperatures between 1060°C and 1100°C. Then,
as the reaction with CaS nears completion, oxygen begins to break through,
and C02 increases as carbon deposited higher in the reactor is burned.
When oxygen breaks through, the bed temperature starts to fall. At this
point all the residual sulphur in the bed is present as calcium sulphate.
-------
- 65 -
After each regeneration in the fresh bed test programme, a
stone sample was analysed for sulphur content. These analyses revealed
that the sulphur level varied between 0.55 and 3.14$ by weight. The
results of a correlation of sulphur contents with operating variables
are summarised in Table XII.
TABLE XII
Operating Variable Effect on Residual Sulphur Level
Gasification Temperature Inverse
Gas Velocity Direct
Air/Fuel Ratio Direct
Particle Size Range Direct
Bed Depth None.
The operating variables are listed in descending order of
importance.
The effects of gasification temperature, gas velocity, and
air/fuel ratio during gasification all can be explained by their
influence on the amount of exposure of sulphided lime to oxygen at
temperatures below 1000°C. If gasification temperature is low, there
must be more CaS oxidation to CaSO^ before temperature is high enough
to liberate SC^. At low air/fuel ratio there is greater lay down of
carbon, and oxidation of this carbon supplies heat to raise lime
temperature with less requirement for sulphate formation. Increasing
air velocity increases the average concentration of oxygen in the bed
during the low temperature period and this increases conversion to
sulphate. It is not understood, however, why particle size of the lime
should influence the amount of sulphate formed during regeneration.
b. Cycle Test Regeneration
In cyclic tests, loss of absorption activity with repeated
cycles did not reduce the ability of the stone to be regenerated.
However, increasing the depth of the fluid bed did reduce the efficiency
of regeneration. This result was caused by the greater heat loss from
the deeper bed. This heat loss affected regeneration indirectly by
forcing lower temperatures and higher air/fuel ratios during gasification,
and it also produced a direct effect by removing heat during regeneration.
c. Continuous Regeneration
In a continuous regenerator the total bed operates at approxi-
mately steady state conditions. However, the individual lime particles
must go through a cycle similar to that of a batch regeneration. The
-------
- 66 -
particle enters at gasifier temperature, is heated by carbon and sulphid
oxidation, and finally releases sulphur as SC>2 before returning to the
gasifier. Since air and gasifier lime both enter at the bottom and
leave the top of the regenerator, there is a gradient of increasing
temperature from bottom to top.
(1) Regeneration Selectivity
The regenerator temperature gradient reduces the selectivity of
CaS oxidation to CaO + SC>2 by increasing the proportion of CaSOjj, formed.
This behaviour is shown in Figure J>8 where oxidation selectivity in the
pilot plant regenerator is compared with calculated equilibrium curves.
Selectivity was calculated from regenerator gas analysis. It is evident
that the pilot plant regenerator required much higher operating tempera-
tures than equilibrium to achieve a given selectivity. It also appears
that somewhat better selectivity was obtained in Run 1 than in Run 3.
The cause of this difference is not yet known.
COMPARISON OF C0 S OXIDATION SELECTIVITY WITH EQUILIBRIUM CURVES
CAFB CONTINUOUS PILOT PLANT DATA
100
90
T
T
T
T
T
T
T
o
J
(/>
J
y
v
C/CoS- 278 0
MAXIMUM THEORETICAL
SELECTIVITY CURVES
80
TO
60
50
40
30
RUN I AUG II-IS o
RUN. I. AUG (6-18 A
RUN 3 •
960 980 IOOO 1020 1040 1060 1080 MOO
REGENERATOR UPPER BED TEMPERATURE,*C
1120
1140
1160
FIG. 38.
-------
- 6? -
This reduction in regenerator selectivity has several undesirable
effects. It reduces S(>2 concentration in the regenerator off gas thereby
increasing cost of sulphur recovery. It raises the required regenerator
temperature which probably accelerates lime deactivation. And finally it
increases the amount of sulphur recycled to the gasifier which requires
that the lime operates at a higher total sulphur loading.
(2) CaS Conversion
Not all of the calcium sulphide entering the regenerator is
oxidised. Samples removed from the regenerator show appreciable
quantities of unreacted CaS. It could be argued that since samples of
solids were taken from the reactor Itself and not from its outlet, they
do not show the true extent of conversion. On the other hand, solids
circulation rates, Figure J-30, estimated from sulphur concentrations of
gasifier and regenerator samples together with the S0? production rate
of the regenerator in some cases agree with, and in other cases are less
than circulation rates estimated from regenerator heat balance. This
comparison suggests that solids leaving the regenerator were not more
completely reacted than indicated by the samples. Table XIII lists
measured sulphur concentrations of gasifier and regenerator samples and
the conversions and selectivities calculated from them.
TABLE XIII
Regenerator Performance Based on Sample Analysis
Time
Day 1-1750
Day 2.1900
Day 3-0230
Day 3-0730
Day 5-0830
Day 7-1130
Day 8-2230
Continuous Pilot Plant Run 3
Regenerator Solids
Sulphide
2.45
2.67
2.53
2.40
1.87
1.19
.95
Sulphate
.95
.57
.64
.65
.84
1.18
2.55
Gasifier Solids
Sulphide
7.52
5.26
6.00
4.95
3.72
3-92
5.11
Sulphate
.09
.19
.15
.13
.19
.23
.24
Sulphur
Differ-
ential %
4.21
2.21
2.98
2.03
1.21
1.78
1.85
% Con-
version
of CaS
67.5
49.1
50
51.5
49.8
70
81.5
% Selec-
tivity
to CaO
83
85
86
80
65
65
44.5
Selectivities based on these solids analyses are in fair
agreement with the values calculated from gas analyses shown in
Figure J-35. Conversions range from 49 to 8l# and show a trend
of decreasing selectivity with increasing conversion.
Both the low conversions and the lower-than-equilibrium
selectivities show that considerable opportunity exists for improving
regenerator operation. One such improvement is a modified distributor
which will give higher pressure drop and better fluidisation. A mound
of agglomerated solids found around the distributor after each continuous
run indicated that fluidisation had been incomplete. This mound decreased
effective regenerator volume and probably caused poor contacting
Another possible cause of reduced regenerator efficiency was the'presence
of a vertical wall crack in the third run which could have allowed some
gas by-passing. This crack will be repaired before the next run
-------
- 68 -
Regenerator SQg Concentration
The time log of SC^ concentration in regenerator off gas from
the continuous ruiis appears in Appendix Figures J-5, J-lj and J-^l.
Concentrations of over 10$ were obtained for brief periods, but most of
the data from tests 1 and 5 were in the range of 6 to 8^ S02. In test 2,
the high limestone addition rate prevented achievement of high enough
sulphur concentrations on lime to give good regeneration. Prom experience
to date, a regenerator gas consistently containing 8# S02 appears to be an
easily achieved goal. Even higher levels are possible but will require
further improvements in regenerator design.
Figure ~%) compares experimental SC>2 concentrations with the
equilibrium curve. Points from alternate hours of pilot plant operations
in Runs 1 and 3 are shown. Run 3 concentrations are somewhat lower than
those from Run 1 and both sets of data fall well below equilibrium.
Formation of sulphate in the cooler portion of the regenerator is respon-
sible for the lower S02 concentrations. Apparently there is insufficient
residence time and contacting in the hot portion of the regenerator to
reach equilibrium by the solid-solid reaction, CaS + ^CaStty —> 4CaO + 4S02,
Changes which could improve selectivity of CaS oxidation and raise
S02 concentration are:
(l) Increase regenerator volume.
(2) Preheat regenerator air.
(3) Make regenerator gas-solids flow counter-current.
Each of these changes would add to the cost of a CAFB regenerator,
and whether or not they are justified remains to be established.
COMPARISON OF REGENERATOR SO, CONCENTRATION
WITH EQUILIBRIUM CURVE
CAFB CONTINUOUS PILOT PLANT DATA
16
14
3 10
MAXIMUM WITH AIM OXIDATION
NO CAMONONUMr
A« A
KUN I. AUO II-IB
NUN I *UO «•!•
NUN a.
O
A
•
trs icco loa IOBO IOTS
REOENCRATOR UPPER KO TEMPCRATURE. *C
(HO
IIZS
no n
-------
- 69 -
4.
Vanadium Retention
Ability of the fluidised lime bed to absorb vanadium during fuel
gasification is an important benefit offered by the CAEB process. If
complete vanadium removal could be assured, boilers using CAEB could
operate with higher superheater tube metal temperature and obtain improved
efficiency without high temperature corrosion. The experience with vanadium
retention has been variable. In the early studies, discussed in Appendix B,
vanadium retentions ranging from 30 to 100$ were observed. These early
tests employed underfiring, that is, some fuel was burned beneath the
distributor to increase fluidisation velocity at the bottom of the bed.
Complete vanadium retention was. obtained in these batch unit cyclic tests
conducted during the limestone comparison portion of the current study.
However, lower degrees of retention were measured in most cases during the
cycle test variable study with stone BCR 1691. Appendix Tables E
contain full data which are summarised here in Table XIV.
TABLE XIV
Summary of Vanadium Retention Results
Test
Tl(CltoCl4)
T3(CltoCl4)
T3(C15toC23)
T3(C24toC3l)
T3(C34toC4l)
T4(CltoC10)
T5(CltoC20)
T6(CltoC17)
T8(C7toC20)
T8(C21toC34)
T8(C45toC6l)
T8(c6?toC8o)
10
11
Average Peak'a) Resl- Bed Make-
Gaslfi- Regener- dence Loss u£
cation ation Time Rate Rate
Temp. Temp. (wt
(g/ Cap/
"C "C (sec) minT wtS)
Poten-
Vana-
dium
858
841
856
858
859
866
859
855
843
852
850
863
840
854
1024
1019
1034
1046
1028
1038
1017
1022
1008
1020
1015
1046
1026
1017
0.15
0.15
0.22
0.15
0.15
0.15
0.21
0.21
0.24
0.24
0.28
0.15
0.36
0.25
1.4
2.0
1.6
2.8
0.1
3.4
2.9
4.5
6.0
2.5
5.2
2.1
7.1
3.8
2.4
2.5
2.8
2.3
3.1
2.2
5.1
2.5
2.5
1.7
2.7
2.2
2.7
2.1
tial
Sulphur Loading
Differ- (gV/
ential gbed/
(wt %} cycle)
3.0
3.0
2.6
3.6
3.1
3.2
2.2
2.2
2.0
2.2
1.7
2.2
1.5
2.2
(a) Average of all cycles.
% Fuel
Vanadium
Removal
0.000465
0.000479
0.00048
0.000506
0.000502
0.000650
0.000406
0.000359
0.000315
0.000348
0.000263
0.000353
0.000243
0.000347
79
65
67
84
82
89
65
81
52
68
83
92
91
78
Regeneration temperature has given the most significant correlation
with vanadium retention, but the correlation coefficient is only 0.6, so
much of the variability remains unexplained. Figure 40 shows the variation
of vanadium retention with peak regeneration temperature.
A hypothesis to explain the effect of regeneration temperature is
that fixation of vanadium on the stone requires oxidising conditions and
that kinetic limitations affect these fixing reactions. This hypothesis
may also explain the apparent difference in temperature effect caused by
underfiring.
-------
- 70 -
VANADIUM PICK-UP V-MAX. REGENERATION TEMPERATURE
lOO
90
80
Q_
Z>
li.
38 50
40
30
1000
• BCR( 1691) RECENT TESTS
o BCR( 1691) COMPARISON TEST
X DENBIGHSHIRE(UNDERRRING)
A DENBIGHSHIRE(NOUNDERRRING)
J L
1050 1100 1150
MAX. REGENERATION TEMPERATURE,*C
1200
FIG. 40.
-------
- 71 -
As Figure 40 shows, tests using underfilling required higher
regeneration temperatures to obtain the same vanadium recovery as
tests without underfiring. If vanadium is fixed under oxidising
conditions there is greater opportunity for some of this fixation to
take place during gasification if there is a strong oxidising zone
near the distributor. With underfiring there is less oxygen in the
air at the distributor and less driving force for oxidation.
The recent test series did not bear out the early observation
that vanadium recovery correlates with iron pick-up. In these variable
studies the iron content of stone has been constant. The earlier
correlation, obtained under underfiring conditions, probably was due to
the effect of temperature on both vanadium fixation and iron pick-up
from metal parts of the unit.
Lime samples from the continuous pilot plant contain higher
vanadium concentrations than can be accounted for by the vanadium
content of the oil fed. Either there was selective loss of low vanadium
fines or there was an extraneous source of vanadium. This matter will be
investigated further in future runs.
5. Gasifier Product Gas Composition
Samples of gasifier product gas taken during the fresh bed test
programme were analysed by gas chromatography. Although the GC method
determines only the dry gas components, the missing quantities, H20 and
Xty hydrocarbons, were calculated from unit material balance considera-
tions. Composition results were examined statistically and rationalised
in the form of empirical equations which express the effects of variables
on the yield of different compounds. The experimental results and the
correlating equations are listed in Appendix P. Table XV, based on these
equations, shows variations in gasifier product composition which can be
expected over the normal range of temperatures and air/fuel ratios.
TABLE XV
Predicted Gasifier Gas Compositions
Air/Fuel Ratio; 15# of 20# of 25# of 30# of
Stoichio- Stolchio- Stoichlo- Stoichio-
metric metric metric metric
Bed Temperature; 850°C 900°C 850°C 900°C 850°C 900°C 850°C 9000C
Component
C02
CO
C < i
H2
H20
N2
C > 1
(*
C4
34 (*
* C
C0/C02 "
by vol.)
»
"
it
it
ti
wt.of
fed)
• WW J
8.2
9-5
20.2
6.0
0.7
54.6
20.1
1.16
6.4
12.6
19.1
6.9
0.5
54.0
21.7
1.97
8.7
10.2
15.1
6.3
1.4
58.3
19.6
1.17
6.8
13.3
13.8
7.1
1.0
58.0
21.9
1.95
8.9
10.4
11.7
6.4
2.3
60.3
15.6
1.17
7.0
13.7
10.4
7.2
1.6
59.9
17.5
1.96
8
10
9
6
3
62
10
i.
.8
.3
.0
.3
.5
.1
.2
18
7.0
13.5
8.3
7.2
2.4
61.6
10.6
1.93
Gas Velocity: 6 ft/sec. Bed Depth(Static): 24 inches.
Particle Size Range: 300-3175^.
-------
This table shows how the carbon monoxide-dioxide ratio varies
with temperature. At 850°C the ratio is close to unity whilst at 900°C
it has increased to nearly two. A high ratio is desirable at low air/
fuel ratios to increase the rate of carbon burn-off from the stone. As
expected, the hydrocarbon content of the gas increases as air/fuel ratio
is decreased, although it would appear that the greater-than-Cjj. fraction
reaches a maximum at around 20# of stoichiometric air. The water vapour
content of the gas increases significantly as air/fuel ratio is increased,
but the hydrogen remains fairly constant for a given temperature.
Increasing bed temperature from 850°C to 900°C, therefore, gives
the following directional changes:
C02
H20
c4
Increasing air/fuel ratio results in the preferential oxidation of the
> C^ fraction.
6. Lime Attrition and Losses
Loss of lime fines during continuous CAPB operation is undesirable
and should be minimised. Although such losses will not be great enough
under normal conditions to affect replacement lime requirements, they do
constitute possible problems in boiler tube fouling and particulate
emissions. The rate of lime loss in the continuous pilot plant depends
both on gasifier cyclone efficiency and on rate of fines generation.
However, the batch units do not return cyclone fines to the bed, so
cyclone efficiency is not a factor. On the other hand, the batch units
have expanded top sections which reduce fines losses by lowering gas
velocity.
a. Cyclic Batch Test Bed Losses
Data on bed losses from the cyclic test variable study with
BCR 1691 were analysed statistically and a correlation obtained with
bed depth and lime replacement rate. The lined out loss rate, after
seven cycles, was used in this correlation. Bed depth emerged as the
most important variable with lime replacement second. Gas velocity had
only minor effect over the range of 4 - 6 ft/sec studied. The regression
-------
- 73 -
equation obtained for the 6 ft/sec data was:
loss rate, gm/min = .00101 m1'11 h1'95
where m is lime make-up rate, gm/min
h is bed depth, inches.
Figure 41 compares the test data with this correlation.
(7)
COMPARISON BETWEEN CALCULATED AND ACTUAL
BED LOSS RATES(CYCLE TESTS)
2
4 6 8 10
CALCULATED LOSS RATE (g/min)
12
FIG.4I.
-------
The magnitude of the bed depth effect is somewhat surprising.
In part this increase in loss must stem from a higher attrition rate
as bed volume increases with height. The decrease in disengaging height
with increasing bed height also must play a part. But even with these
factors, a 1.95 power effect of bed height on losses is unusually large
for fluid beds. It is possible that operations were in a region where
increase of bubble coalescence with bed depth caused a rapid change in
the quantity of solids thrown violently into the gas space.
The nearly proportional effect of limestone make-up rate on bed
losses is reasonable because of the decrepitation loss that occurs during
in-situ calcination and because freshly calcined stone attrites more
rapidly than aged stone. Evidence for the decrepitation loss was obtained
during fresh bed calcinations where the total weight loss of the 300-3175
micron stone averaged 52.5$. The carbon dioxide loss due to calcination
is only 37$• The higher loss rate of freshly calcined stone was shown by
the fact that losses in the first cycle of a test series has been as much
as seven times the lined out loss rate. It is not yet clear whether the
change in loss rate with time is due to an increase in stone toughness or
to an initial loss of a softer fraction.
b. Particle Size Changes
The variation of lime particle size distribution through a typical
cycle test series with BCR 169! at 6 ft/sec and 15.5 inch static bed depth
is shown in Figure 42. In this test there was a trend toward Increasing
particle size with time. This trend is opposite in direction to that
reported in the Interim Report(^ ) for a test of BCR 1691 at 4 ft/sec.
Apparently the velocity difference was enough to Increase the rate of fines
loss faster than their rate of production. An opposite effect with
increasing bed depth appears in Figure 43. Here we see that average
particle size (measured after 20 cycles) decreased with increasing bed
depth. The greater bed depth apparently increased the rate of fines
production even more than the rate of fines loss.
c. Continuous Unit Experience
Lime losses in the batch unit at 6 ft/sec covered a range from
3 to 10 grams/min which is equal to 1.5 to 4.9 lb/hr-ft2. Losses in the
continuous pilot plant at 3 to 4 ft/sec, have tended to fall in the same
rang* and have depended to a large degree on the efficiency of cyclone
operations and in attrition produced by the oil feed inlets. The initial 40
hour period of the first pilot plant run had a loss rate of 1.44 lb/hr-ft2.
T.iis increased to 2.7 lb/hr-ft2 after the first decoking operation when one
of the cyclone outlet tubes apparently was damaged. In the final hours of
the test, rate increased to 9 lb/hr-ft2 when neither cyclone apparently
was operative.
-------
300
100
VARIATION IN PARTICLE SIZE DISTRIBUTION. TEST 3
T
• FRESHLY CALCINED LIME
o AFTER 7 CYCLES
A AFTER 14 CYCLES
1000
2000
PARTICLE SIZE. MICRONS
3000
4000
FIG.42
CHANGES IN MEAN PARTICLE SIZE WITH BED DEPTH(AFTER 20 CYCLES)
600-
400-
200
10
20
STATIC BED DEPTH (INCHES)
30
-------
- 76 -
The second pilot plant test is not considered to be representa-
tive of any practical operating condition. Two factors in that test
combined to produce entirely unacceptable lime loss rates. These were
elimination of cyclone gas outlet dip tubes altogether and use of high
gas velocities in the fuel distributors. This high gas velocity ( above
400 ft/sec) was intended to improve fuel atomisation and distribution
(which it did) but, in addition, it gave severe lime attrition. Losses
from the bed in this test ranged from 6 to 14 Ib/hr-ft2.
The final test run used cyclone gas outlet tubes which could be
steam-cooled during decoking, and, in addition, employed a controlled
temperature decoking procedure to avoid damaging the tubes. Nevertheless,
in this test too, cyclone operation failed during a portion of the test
when solids that had accumulated on material test specimens in the gas
transfer duct were dislodged and blocked the cyclone solids drain lines.
For the first 65 hours of operation, lime losses from the bed averaged
1.25 Ib/hr-ft2. During the test period between 1J30 and l8l hours while
cyclones were inoperative, the loss rate reached a high level of 3.7
Ib/hr-ft2-
This level of lime loss, even with good cyclone operation, is
higher than desired. It certainly would require particulate removal
equipment before the flue stack and could affect the rate of solids
accumulation on boiler surfaces. We suspect that unnecessary attrition
is taking place in the vicinity of the air distributor nozzles and have
designed a new distributor which will reduce the inlet jet velocity while
maintaining sufficient pressure drop to provide stable operation.
The method of fresh limestone addition in the pilot plant con-
tributed to the high lime losses also. The stone was merely dropped
through a connection in the lid. Though simple, this procedure caused
the fresh stone to fall through the rising gas stream which undoubtedly
carried a portion of it straight through the outlet duct. The pilot
plant is being modified to use a low level pneumatic lime feeder which
will avoid this source of dust losses.
-------
- 77 -
7. Operation of Continuous Pilot Plant
Three runs were made in the continuous pilot plant in 1971. It
was the intention in each of these runs to operate, if possible, for a
period of 200 hours of gasification, but the first two tests were
interrupted short of this goal. The third run lasted 2JO hours of
which 2Qk were gasification. The unit was shut down voluntarily at
the end of the third run.
Appendix J contains details of the continuous run log of
operations, results, and inspections following runs. Figures start on page 291.
a. Gasification and Sulphur Removal
The principal objective of the CAPB process:fuel gasification
with sulphur removal,was demonstrated throughout the three tests. The
product gas from the gasifier ignites readily and burns with a bright
luninous stable flame. Smoke free operation with about 1.5$ oxygen
in the boiler gas has been demonstrated for long periods. This is
better performance than found with normal fuel oil firing in this
boiler which requires about 3$ oxygen in the flue gas.
The variation of sulphur removal efficiency with time is
shown in Appendix Figure J-I6. Ability to achieve 100$
sulphur removal was demonstrated in a number of instances. In most
cases, this high removal efficiency was reached with lime replacement
rates of over 1.5 times stoichiometric, but removal of over 95$ was
reached with less than stoichiometric lime addition during the final
day of the third run.
b. Solids Circulation
Blockages in circulation of solids between beds were experienced
in the first run, but this problem was eliminated in runs 2 and 3.
Figure J-j50 shows Run 3 circulation rates as estimated from heat
balance. Higher circulation rates have been required than originally
estimated because the extent of calcium sulphate formation in the
reagent has been greater than expected. The sulphate forming reaction
releases extra heat and requires more solids circulation for
regeneration temperature control.
Circulation problems in the first run were due to obstructions
that formed in the transfer line where air was injected into beds
of defluidised sulphided solids. It had been intended to use flue gas
in the bed transfer pulse injector system. However, the system for
cleaning and compressing the flue gas proved unreliable due to plugging
by wet lime fines and air had been substituted. At points where air
entered the defluidised bed of sulphided lime there was partial
regeneration which appears to have resulted in momentary formation of
a liquid sticky phase and agglomeration of the solids.
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These blockages did not occur during periods of steady operation
in Run 1 but formed during interruptions to operations which occurred
for other reasons.
The problem was eliminated in Runs 2 and 3 by use of nitrogen
as the injection gas for the bed transfer system. No circulation
blockages were experienced in 91 hours of continuous operation in Run 2
or in 230 hours of operation in Run J. In Run 3 there were several
shutdowns and restarts for decoking, and in each case bed recirculation
was resumed without difficulty. At the end of Run } a small deposit of
agglomerated lime fines was found in the regenerator to gasifier
transfer duct. This deposit had not obstructed flow, but whether it
was stable in size or would have continued to grow is not certain.
Similarly, there were accumulations of agglomerated fines at the entries
to the solids transfer down comers both in the gasifier and the
regenerator. Neither of these accumulations had obstructed flow.
In cold circulation tests with the continuous pilot plant
mockup it had been observed that the bed transfer system was able to
move approximately two pounds of lime per cubic foot of air. The
transfer system in the hot pilot plant was considerably more efficient
with rates as high as 25 pounds per cubic foot of nitrogen being
achieved.
c. Gasifier Temperature Control
The gasifier temperature was adjusted by varying the proportion
of flue gas recycle to fresh air in the fluidisation gas. This
adjustment was made by the operator in response to the reading of the
bed temperature controller. While other conditions were stable, this
manual control was adequate to hold temperature within about 5 Deg. C
of target. However, when other factors, such as limestone or oil feed
rate, varied, there could be greater fluctuation of temperature. Even
so, it was possible to hold the temperature between 870 and 900°C which
appears to be adequate for the process. Automation of the reactor
temperature control would be a simple matter, but the cost of the
controller and valves does not appear to be justified at present.
d. Regenerator Temperature Control
An automatic controller for the regenerator temperature had
been installed prior to the first run but was not used until the third
run. In the first two runs, the temperature was varied by manual
adjustment of the solids transfer rate. In the third run this
adjustment was turned over to the automatic controller with excellent
results. Control could be maintained within plus or minus 5 Deg. C
of the set point, and response to set point changes was rapid. The
controller was able to compensate for minor changes in pressure
differential between gasifier and regenerator. Thus solids circulation
rate was more constant with the temperature controller in operation than
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it had been under manual control. The improvement due to the
controller can be seen by comparing uniformity of regenerator
temperatuic (Figures J-4, J-12, and J-29) before and after 2300
on the first day of Run 3 when the automatic system was activated.
e. Limestone Feed
A simple hopper with a vibrated transfer line leading
to an opening in the gasifier lid was used to feed limestone to the
gasifier. The vibrator on the transfer line was turned on and off
by a pressure switch operating on gasifier level. Solids feeding
was interrupted when it became necessary to refill the hopper.
A stream of air was admitted with the stone to prevent backflow of
gas to the feeder.
This system was not quite satisfactory and a modified system
is under study. Disadvantages of the old system were:
(1) Interruption of stone feed to fill the hopper affected
gasifier temperature;
(2) Dropping the stone through the gasifier lid increased
dust losses;
(3) Feed rate from the vibrator was variable and sensitive
to gasifier pressure;
(It-) Instantaneous feed rate could not be determined, only
averages over the periods required to empty a hopper.
In general the feed system required more operator attention
than is desirable. Between Runs 2 and 3 a pneumatic feed line had
been installed to inject limestone directly into the bed. However,
the vibrator used to feed solids from the hopper to the transfer
line did not prove to be satisfactory, and this system was not used.
An improved pneumatic feeder will be used in the next test.
f. Refractory Condition (1)
Details of the unit condition at the end of Run 3 appear in
Appendix J. In general the unit has stood up well. Some initial
roughening of the refractory in the bed region occurred during cold
circulation trials prior to the first run but little additional
errosion has been observed. Several vertical cracks have appeared
at the ends of the gasifier. There is also a vertical crack in the
lower portion of the regenerator. It appears that these cracks reseal
themselves in the upper, hot portion of the unit during gasification.
(1) End of run photographs start on page 326.
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It is probable that a crack remains open in the cold plenum region of
the gasifier and of the regenerator. The gasifier air distributor is
fixed to a closed metal plenum so that no bypassing is possible there.
Some bypassing through the crack may have occurred in the regenerator.
A new metal distributor fixed to the plenum has been designed for the
regenerator to eliminate possible bypassing in future tests.
No penetrations between gasifier and regenerator spaces have
appeared.
The interior cyclone surfaces have been roughened to a
considerable degree. At the end of each run the cyclone surfaces
have also shown a rough accumulation of lime and carbon scale. It
appears that maintaining a smooth interior cyclone surface will be
difficult.
The surfaces of solids transfer ducts have remained smooth
and intact.
g. Agglomerations and Deposits
Agglomerates and solid deposits have been encountered at
several points in the unit during the three runs. The cause of some
of these deposits has been identified and the deposits eliminated.
Others remain as potential operating problems.
(1) Regenerating Solids
The blockage of solids transfer lines in Run 1 has been
mentioned. This blockage was part of a more general problem area.
It appears that regeneration of sulphided lime produced a transient
liquid state in the CaO - CaS - CaSO^ system and that the liquid
formed can cement lime particles together if they are in close
proximity in a static bed. This transient liquid phenomena has been
observed by Consolidation Coal Co.(5) in their work on a COp acceptor
process. In Run 1 there were several instances where air was introduced
into a hot static bed of defluidised sulphided lime. In each of these
instances a lump of agglomerated lime particles resulted. In subsequent
runs it has been possible to avoid conditions leading to this
agglomeration in most cases. The only persistent problem area has
been the zone directly above the regenerator distributor.
In each of the three runs there has been accumulation of an
agglomerated lump directly above the distributor. This agglomerate has
been attributed to local poor fluidisation due to insufficient
regenerator distributor pressure drop and to the top-hat distributor
design which creates a defluidised centre region. A new regenerator
distributor has been designed to alleviate this problem.
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A slight deposit of cemented fines has been observed in the
regenerator gas space. At the end of Run 1 the deposit was located
just below the gas outlet. At the end of Run 3 there was no deposit
in the regenerator itself, but a small amount was found in the outlet
duct at the expansion joint. It is possible that there is some
critical temperature at which fines will adhere to the regenerator
wall and that the location of this temperature zone depended on
operating conditions. In Run 3 the regenerator temperature was higher
and more uniform than in Run 1. The quantity of this upper
regenerator deposit was small and is not considered a serious problem.
(2) Gasifier Air Distributor
In the first two pilot plant runs there had been negligible
increase of pressure drop through the gasifier air distributor.
The third run however encountered a steady pressure rise that lasted
throughout the operating period. The initial pressure drop of 9.5 inches
w.g. had increased to 33.8 in. by the end of the run. Inspection of the
distributor after the run revealed that all of the nozzles had some
degree of obstruction, 20$ of the holes were at least 2556 obstructed,
and about 2$ were completely blocked. The obstructions appeared to
consist of fine particles that had adhered to the periphery of the
nozzle bores. There fines probably came from the recycle flue gas.
The occurrence of distributor blockage in the third run but not
the first two is somewhat surprising as the quantity of fines in the
recycle flue gas was much less in the third run than in the others.
A cyclone had been installed in the flue gas returning line after the
second run. It had been effective in removing a great deal of solids
as the flue gas lines and the plenum itself were relatively clean after
the third run but had contained much solids after the first two. Pines
content of the recycle gas had been particularly high in Run 2. A
possible explanation for the greater adhesion of fines to the
distributor nozzles in Run 3 is the higher level of sulphur on gasifier
lime. In its passage through the boiler this sulphur was partly
converted to sulphate. Whether or not this sulphate would have rendered
the lime particles sticky at conditions in the distributor nozzles
remains to be established.
New distributor nozzles are being fabricated for the gasifier
to reduce lime attrition due to the distributor Jet velocity. Since
these nozzles will retain small orifices and will still be subject
to plugging by fines, a cyclone of increased efficiency will be
installed in the flue gas recycle line.
(3) Carbon in Gasifier Outlet
Deposition of carbon and lime on surfaces of the gas outlet
system appears to accompany gasification. In all of the tests to
date there has been a gradual build up of pressure drop between
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- 82 -
gasifier and boiler during gasification due to this deposition of
solids. Figures J-3, J-ll, and J-25 show the increase in total
gasifier pressure for the three i-uiis. Of course the total pressure
drop from gasifier to atmosphere is the sum of several pressure
drops. The breakdown of these various pressure drops appears in
Figure J-27 for Run }.
The initial pressure drops from gasifier to cyclone outlet
and from cyclone outlet to the boiler are approximately equal.
Both pressure drops increased during each gasification
period until the total pressure drop forced a decoking operation to
remove the cause of obstruction. In Run 1 the decokings were somewhat
uncontrolled and the stainless steel cyclone gas outlet tubes were
destroyed. In Run 3 a controlled decoking procedure was used and
the cyclone outlet tubes remained intact. The procedure used in
decoking was to first stop fuel addition and go to a very high flue
gas recycle rate for fluidisation until the sulphide of the lime
bed was converted to sulphate. Then the bed was slumped while a
flue gas-air mixture was added above the bed to burn out carbon
deposits in the cyclone inlets, cyclone bodies and tubes, and bifurcated
duct between cyclones and burner. When carbon burning was complete,
gasification was resumed. In some cases it was necessary to go to full
combusting conditions to raise or hold bed temperature before resuming
gasification. Steam cooling was used to control the temperature of
the cyclone outlet tubes during decoking in addition to the reduced
oxygen atmosphere provided by flue gas recycle.
The solids laid down in the gas outlet system are a mixture
of lime and carbon. Burning removes the carbon but leaves a loosely
agglomerated layer of lime. Therefore decoking alone is not
completely adequate and some form of soot blowing probably will be
required in a commercial installation. In the third pilot plant run
the solids were dislodged by rodding out through the duct at 1600 hours
of day 6. These solids appear to be soft and readily removed with
slight mechanical force.
There are insufficient data yet to establish a correlation,
but it appears reasonable that conditions of low air/fuel ratio and
low gasification temperature which increase the concentration of heavy
hydrocarbons in the gasifier product will increase the rate of
deposition in the outlet ducts. On the other hand, operation at
temperatures over 900° C, as on Day 5 of the third run, did not reduce
the rate of pressure increase in the cyclones. Neither did local
injection of air into the cyclone inlet ducts.
At the end of Run 1, the gasifier was shut down without a
burn out so that solids accumulations could be observed in place.
Large obstructions were found in the cyclone inlets which, at that
time, were square entry rectangular ducts. Between the first and
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second runs, the entry sections of these ducts were rounded to a
more stream-lined profile to reduce entry loss pressure drop.
This measure appears to have helped as pressure rise in the third
run was much slower than in the intact cyclone tube period of
Run 1.
The point of transition from the vertical cyclone outlet
tube to the horizontal gas duct represents another sharp change
of gas direction where solids accumulation is high. The outlet
system is being modified to provide a rounded transition at this
point.
The problem caused by solids accumulation in the gas
ducts was aggravated in the early runs by the relatively low head
available from the gasifier blower system. Although two blower
stages had been installed, the intention had been to use the first
stage only as a booster during startups. Therefore the flue gas
recycle line had been piped to the inlet of the second stage. With
this arrangement the flow rate of fluidising gas fell as back pressure
on the gasifier increased. This reduction had the effect of
decreasing air/fuel ratio and causing carbon laydown and back pressure
to increase even faster. The piping system was revised during a
decoking period on Day 6 of the third run to return flue gas to the
first stage blower and permit use of both blowers in series. The
result was much greater operating flexibility and the ability to
extend run length between decokings.
The effect of carbon deposition on gas passage pressure
drops is more severe in the case of a pilot plant than it will be in
a commercial unit with large size ducts. On the other hand, the rate
of accumulation of solids per unit of duct surface probably is
comparable. Thus the rate of pressure rise in a large unit will be
much lower than the pilot plant has experienced. The decoking
frequency of a large unit will probably be dictated by limitations
on the quantity of carbon to be burned in each decoking rather than by
pressure drop limitations.
(k) Boiler Tube Fouling
A large portion of the lime that passed through the gasifier
cyclones was deposited in the boiler at the end of the main firetube
and at the inlets to the first pass of if inch fire tubes. In Run 2
a large quantity of solids was removed from drains at the boiler back
end and the small tubes were only slightly obstructed at their inlets.
However, in Run 3 there was a progressive increase of pressure drop
across the boiler throughout the run, though less solids were withdrawn
through the drains. The rate of pressure drop increase accelerated
following the rodding of the duct on day 6 and it is assumed that the
cyclone dust return lines became blocked during that rod out. Blockage
of these lines rendered the cyclones inoperative and increased the
rate of lime loss from the gasifier to the boiler. The rate of solids
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removal from the boiler back end and flue gas final cyclone increased
sharply following this rodding as the solids recovery curves in
Figures J-19 and J-21, show.
The rate of pressure drop increase accelerated also as the
diameters of the tube inlets became progressively smaller.
In several cases the fire tube inlets were completely
obstructed by a crusty coating of fines. The crust could be broken
easily leaving a core shaped deposit in many of the tubes. These
cores penetrated for a short distance into the tubes but did not
extend the entire length. The longest core was 15 inches long with
the majority being less than 6 inches. The core material was
approximately the composition of calcium sulphate.
These fire tube deposits originate from highly sulphided lime
fines which convert partially to sulphate and become sticky during their
passage through the flame.
The deposits observed are not hard and could be removed
easily when the boiler had cooled. It may be possible to remove
these deposits under hot operating conditions by soot blowing
procedures, and tests of this possibility are planned for the future.
A deposit measuring probe will be inserted in the boiler for the next
test. Nevertheless, deposits are undesirable, and their rate of
formation should be minimised. Obvious steps to be taken are better
maintenance of cyclone operability to prevent excessive loss of lime
fines and reduced attrition to reduce rate of production of fines.
A third means of reducing deposits may be to reduce stickiness of the
lime by decreasing the sulphur level on gasifier solids. This &tep
may be accomplished by improving regenerator efficiency to increase
both selectivity and conversion level in regeneration. Such an
improvement also should reduce the lime replacement required' for good
sulphur removal efficiency.
(5) Test Specimens
A closed tube of silicon nitride was suspended in the
bifurcated duct above the left hand cyclone outlet in Run 2. A
second silicon nitride specimen was placed above the right hand cyclone
and a self bonded silicon carbide specimen above the left hand cyclone
in Run 3- The silicon nitride specimen was removed Intact at the end
of Run 2, but it developed cracks soon after and displayed extreme
brittleness. The silicon nitride specimen In Run 3 was completely
destroyed during the roddijig out operation on Day 6. It} too had
evidently become very brittle as no portions of it were recovered when
cyclone fines were examined after cyclone draining.
The self bonded silicon carbide specimen was removed Intact
after 104 hours of operation. Examination by the manufacturer revealed
that its strength and toughness had not deteriorated during the test.
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It had been subjected to three decoking operations. On the basis
of this test, arrangements have been made to install cyclone gas
outlet tubes of self bonded silicon carbide before the next pilot
plant trial. Durability of the material therefore can be evaluated
under realistic operating conditions.
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D. Conclusions and Recommendations
This Phase I study of the CAEB process has resulted in
the following conclusions and recommendations.
1. Conclusions
a. The Venezuelan fuel oil originally selected by NAPCA
(now GAP) performs in a satisfactory manner in the
CAFB process.
b. U.S. Limestone BCR 169! is suitable for CAEB applications.
It has only 8l# (initial sample) to 8<$ (large supply)
of the CaO content of the U.K. stone formerly used as a
standard, but at a high enough replacement rate it can give
nearly complete sulphur removal from the fuel.
c. The tested sample of U.S. Limestone BCR 1690 was inferior
to the U.K. stone standard and to BCR 1691. It contained
only 60# as much CaO as the 'original sample of BCR 1691,
it suffered greater attrition, it gave lower sulphur
removal efficiency even when replaced at equal CaO/S
rates, and it tended to sinter and agglomerate under CAFB
cyclic conditions.
d. Gasification temperature and air/fuel ratio interact in
their effect on sulphur recovery in a CAFB gasifier.
As gasification temperature decreases, high air/fuel ratios
must be employed to assure adequate removal of carbon
from the lime and avoid rapid deactivation toward sulphur
absorption. Gasification temperatures above 870*C permit
air/fuel ratios as low as 16$ of stoichiometric with good
sulphur removal. Gasification temperature as low as 800*C
can be used with air fuel ratios above about 27# of
stoichiometric.
e. Sulphur removal efficiency of a CAFB gasifier begins to
decline at temperatures above 900*C, possibly because of
partial regeneration in the vicinity of the air distributor.
f. The sulphur removal efficiency of fresh lime beds improves
as average particle size decreases. At low levels of lime
utilisation the lime utilisation at equal efficiency is
roughly inversely proportional to the square of number
average particle diameter. In cyclic tests on a continuous
unit, consideration also must be given to the effect
of particle size range on losses from the bed, especially
during in-situ calcination of the stone. A limestone with
lower particle size of 300 microns was less effective overall
than one with 600 microns minimum size because of the high
loss rate of ttie former.
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g. Hie effects of gasifier bed depth and fluidisation on
sulphur removal efficiency are correlated by their ratio,
the superficial gas residence time above the fuel injector.
As an approximation, sulphur removal follows a first order
rate expression.
h. Lined out sulphur removal efficiency with the large supply
of U.S. limestone BCR 169! was lower in batch cyclic tests
than predicted by the performance of the original sample
in the limestone screening tests. Greater limestone
replacement rate was required to achieve the same sulphur
removal efficiency. In part thds effect was due to greater
losses of the low particle size end of the 300- 3175 micron
stone selected for the variable study compared with 600-1^00
micron stone used in the comparison test.
i. Potential Sulphur Differential, the quantity of sulphur
(expressed as wt. % on lime) to which the lime charge was
exposed in each gasification cycle, emerges as an
important variable in the batch unit cyclic variable study.
Hie sulphur removing reaction rate appears to vary inversely
with the 1.9 power of this differential. The empirical
equation which correlates the results of the batch cyclic
tests is:
% Sulphur Removal Efficiency = 100
1 -
t-6
where t = gas residence time, seconds, in the bed above
the fuel injector.
m = lime replacement rate, wt. CaO/wt.S
d = % sulphur differential.
This relationship implies that beds in the region of J5 to k ft
deep will be required for >90$ sulphur removal with
stoichiometric lime replacement or less and gas velocities of
6 ft/sec.
j.1 Heat to raise lime temperature from the gasification level to
the regeneration level is supplied in part by oxidation of CaS
and in part by oxidation of carbon on the lime. At temperatures
below 1000°C most of the CaS oxidation produces CaSO^.
Therefore the lower the gasification temperature the
greater is the amount of CaSO^ formed and retained by the
lime after regeneration. Low gasification air/fuel ratio
increases the carbon content of the lime and reduces the
amount of sulphate formation. The quantity of sulphate formed
in batch unit regenerations also increases with gas velocity
and with average lime particle size.
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k. The pilot plant regenerator required temperatures
50 to 100*C higher than predicted by equilibrium
considerations to achieve a given level of oxidation
selectivity to calcium oxide. Contacting and kinetic
limitations appear to be responsible for the increased
temperature requirement. Temoeratures of 1070 to 1100°C
gave SOp concentrations of up to 10$ in the regenerator
off gas.
.1. Vanadium retention in the batch unit cyclic test series
correlated with temperature in the regenerator-higher
regeneration temperatures giving more complete vanadium
retention. Essentially 100$ retention of vanadium was
achieved with regeneration temperatures above about
1065*C.
m. Vanadium concentrations in solids removed from the
continuous pilot plant gasifier and regenerator beds were
several times higher than expected for uniform vanadium
deposition. These data imply selective loss of low vanadium
fines and accumulation of vanadium concentrates.
n. Composition of gasifier product varies with temperature
and air/fuel ratio in a manner that can be expressed by
empiriral equations. Increasing the temperature increases
CO/COp ratio, raises the hydrogen and greater-than-C^
fraction, and decreases water formation. Raising the
air-fuel ratio gives preferential oxidation of the greater-
than-C. fraction.
o. Lime losses in batch unit studies increased rapidly with
bed depth but were little affected by velocity in the k
to 6 ft/second region. Loss rates of 1 to 9 Ib/hr-ft^
were experienced. The continuous pilot plant encountered
lime fines losses ranging from 1 to 1^ Ib/hr-ft with
cyclone operation being the major controlling factor and
jet attrition by air admitted with the fuel being a
contributing factor.
p. The continuous pilot plant functions as intended by producing
a hot, combustible low sulphur fuel gas from heavy fuel oil
while delivering a concentrated stream of so from the
regenerator. A 2^0 hour pilot plant run was completed in
which 20^4- hours were at gasifying conditions. This test
employed 2.5# S Venezuelan fuel oil and Denbighshire limestone.
Sulphur removal from the fuel exceeded 9956 for long periods,
and 95$ removal was achieved with less than stoiehiometric
lime replacement.
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q. Introduction of oxygen containing gas into unfluidised
hot beds of sulphided lime causes agglomeration of the
particles, presumably by formation of a transient liquid
state in the CaS - CaSO. - CaO system undergoing
regeneration.
r The solids circulation system incorporated in the
continuous pilot plant performs in a satisfactory
manner. It delivered 2^0 hours of trouble-free
operation during the third pilot plant run prior to
planned run termination.
s Automatic regenerator temperature control by regulation
of solids circulation rate performs in a satisfactory
fashion.
t. Gasifier temperature control by flue gas recycle is
feasible and produces satisfactory temperature regulation.
u. Lime particles in flue gas recycle can cause gradual
blockage of the gasifier air distributor.
v. Gradual accumulation of a deposit of carbon and lime in
cyclones and ducts leading from the gasifier to the burner
forces periodic decoking operations. Periods up to 6k
hours between decokings have been demonstrated in the
pilot plant to date with cyclones functioning and longer
periods without cyclone outlet tubes. Decoking by
controlled burnout has been demonstrated by a procedure
which requires about two hours of interruption to
gasification.
w. A portion of lime fines from the gasifier adheres to the
entry sections of the pilot plant boiler tubes causing a
gradual increase in pressure drop through the boiler. The
scale is not hard and was easily removed following the
230 hour pilot plant run.
x. Operability and flexibility of the CAPB pilot plant
indicate that it is a useful tool for study of the CAPB
process under realistic operating conditions and that
it will be capable of providing answers to a majority of
CAFB process and engineering questions.
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2. Recommendations
a. Continued use of the Venezuelan fuel oil of the type
studied to date is recommended for the next stage of
CAFB pilot plant studies.
b. Use of limestone BCR 1691 in the next stage of pilot
plant studies is recommended.
c.
Gasification temperatures in the range of 8?0 to 900*0
are recommended as offering the most favourable range
for the majority of CAPE studies.
d. A limestone feed with a particle size range of 600 to
^200 microns is recommended for the majority of
experimental work in the next phase of pilot plant
studies.
e. It is recommended that greater bed depths and gas
velocities be explored in the next phase of pilot plant
work to minimise bed area and lime replacement requirements
for a given CAFB duty.
f. An improved air distributor is recommended for the
regenerator to avoid local areas of defluidisation and
to improve regenerator contacting performance.
g. Frequent sample analysis of all solids streams for
vanadium are recommended in future pilot plant studies
to establish a vanadium balance.
h. A redesign of the gasifier air distributor is recommended
to reduce local jet velocity and thereby to decrease lime
attrition rate.
i. An independent source of inert, oxygen free gas, is
recommended for use at any CAFB point where gas is injected
into hot sulphided solids which may be defluidlsed. The
pilot plant now uses nitrogen, but a commercial Installation
could employ an Inert gas generator. Dependence on flue
gas recycle is not recommended as it may be subject to
interruption when most needed.
j. A more efficient solids separator is recommended for the
flue gas recycle line to reduce possibility of plugging
the gasifier air distributor by recycled lime fines.
k. Close observation of the quantity of carbon laid down in
gas transfer ducts is recommended in future pilot plant
tests to establish conditions which moderate carbon laydown
and to permit prediction of carbon deposition under
commercial conditions.
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1. Provision of facilities for controlled decoking of
transfer ducts is recommended in commercial CAFB units.
m. Trial of self bonded silicon carbide as a material of
construction for pilot plant cyclone gas outlet tubes
is recommended.
n. Installation of a probe in the end of the pilot plant
boiler is recommended for study of the rate and nature
of solids deposits to be expected on boiler superheater
tubes.
o. Major process features of the CAFB process have been
demonstrated, and possible operating problem areas
having been defined, it is recommended that the continuous
CAFB pilot plant be used in an intensive study of CAFB
process variables and engineering features to solve those
problems and provide data needed for design of a 150
megawatt boiler conversion.
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REFERENCES
1. Moss, G: The Desulphurising Combustion of Fuel Oil in
Fluid!sed Beds of Lime Particles; Presented at First
International Conference on Fluid Bed Combustion, Hueston
Woods, Ohio, November 1968.
2. Moss, G: The Fluidised Bed Desulphurising Gasifier;
Presented at Second International Conference on Fluid Bed
Combustion, Hueston Woods, Ohio, October 1970.
3. Technical Proposal for Study of Chemically Active Fluid
Bed Technique for Reduction of Sulphur Oxides Emissions.
Report No. ER-rjM-69 by Esso Research Centre, Abingdon,
Berkshire, July 1969.
4. Study of Chemically Active Fluid Bed Gasifier in Reduction
of Sulphur* Oxide Emissions. Interim Report from Esso
Research Centre, Abingdon, Berkshire to Air Pollution
Control Office, U.S. Environemtal Protection Agency, under
Contract CPA 70-46.
5. Monthly Reports 1 through 16, July 1970 to November 1971.
Esso Research Centre to Office of Air Programs, U.S.
Environmental Protection Agency, under Contract CPA 70-46.
6. C-urran, 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 C., July 1st, 1969.
7. Westinghouse Research Laboratories TWenty-First Monthly Progress
Report to US Office of Air Programs under Contract No. CPA 70-9-
8. Weisz, P.B., and Robert P. Goodwin:
J. Catalysis, ^227-236.
9. Daizo Kunii and Octave Levenspiel:
Fluidisation Engineering p. 144, John Wiley, New York, (1969).
-------
- 93 -
APPENDICES
TO
FINAL REPORT
OAP CONTRACT CPA 70-46
STUDY OP CHEMICALLY ACTIVE FLUID BED
GASIFIER FOR REDUCTION OF SULPHUR
OXIDE EMISSIONS
APPENDIX SUBJECT
A CAFB Chemistry
B Results of work Just Prior to Contract
C Batch Unit Fuel and Lime Comparisons
D Batch Reactor Variable Study with Fresh Lime Beds
E Batch Reactor Cycle Test Variable Study
F Correlation Functions for Sulphur Removal Efficiency
0 Design Basis for the Pilot Plant Gasifier System
H CAFB Pilot Plant Alarm Systems
I CAFB Pilot Plant Operating Procedures
J Operation of Continuous Pilot Plant
K Burning Rate Model for Carbon on Lime in CAFB Gasifier
-------
- 94 -
APPENDIX A
CAFB CHEMISTRY
Chemistry of CAFB Gasification Process
In gasification, sulphur compounds react with lime under reducing
conditions. In contact with hot lime particles (800-900°C) the fuel oil
vaporises and cracks to liberate compounds such as HpS, COS and CSp.
Sulphur is then absorbed by the lime by reactions such as those listed
in Table A-I.
TABLE A-I
EQUILIBRIUM CONSTANTS AND HEATS OF REACTION FOR
DESULPHURISING REACTIONS DURING GASIFICATION
Reactlon 10gloKp (kcaVmol)
800°C 900°C 1000"C 1100°C ^Tvvnon
at 900 C
1 CaO + .BUS > CaS + J^O 2.97 2,91 2.82 2.70 -4.7
«- -^ ^
2 CaO + COS —=—±CaS + cO k.JO 4.42 4.13 3-83 -17.9
."* " f-
3 CaO + |CSP _* CaS + ^CO 4.91 4.63 4.28 3-90 -19-7
G* ,^»^^^^^^ b
The equilibria for these reactions are well to the right indicating
sulphur partial pressures of less than 200 ppm tip to 1100° C (2000°P) as shown
in Figure A-l
-------
- 95 -
EQUILIBRIUM LEVELS FOR SULPHUR COMPOUNDS
DURING CAFB GASIFICATION
-2
0.
0-6
-8
-10
-12
JJSUAL_ LEVEL £F JULPHUR_COMPOUNDS_WiTJi 4%_S
"FUEL AT 33% STOICHIOMETR~IC AIR
EQUILIBRIUM H2S LEVEL
COS LEVEL
800
9OO 1000
TEMPERATURE. °C
1100
1200
Figure A-l
With a fuel containing k% sulphur these equilibria imply a
desulphurising efficiency of greater than 90$ up to 1100°C. However other
factors limit the effective gasification temperature to the 800-900°C range.
In the shallow fluidised bed of the gasifier there is rapid
circulation of lime between top and bottom. There is evidence 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.
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 laboratory tests, but
in the continuous unit the average extent of reaction will only be to about
of the calcium in the bed.
-------
The lime reagent is regenerated by oxidation of its calcium sulphide
content. Again, several reactions may occur as shown in Table A-II.
TABLE A-II
EQUILIBRIUM CONSTANTS AND HEATS OF REACTION FOR
THE REGENERATION OF SULPHIDED LIME
Reactions
CaS +
5 CaS + 2 02
6 CaS + 3 CaSO.
10g10kp AH
(kcal/mol)
900 "C 1000 °C 110Q°C 1200°C 1300°Cat 1100'C
CaO + S0 16.11
13.01 11.8? 10.86 -109.5
CaSO,,
23.96 20.61 17.79 15.^ 13A6 -220.2
^ 4CaO + kSO -7-^ -^.09 -1.36 1.09 3.12 221.5
Both reactions CO and (5) are highly exothermic and their
equilibria under conditions of excess oxygen are far to the right. The
SOg partial pressures to satisfy their respective equilibria at different
temperatures are shown in. Figure A-2. (These are also the equilibrium SO,,
partial pressures for reaction (6)). c
Maximum SO partial pressure with air regeneration should,
therefore, be obtalnid at temperatures in excess of 1040*C (1900*F) when
the- reaction is conducted at atmospheric pressure.
EFFECT OF TEMPERATURE ON THC
SOt WW1TIAL PRESSURES FOR THi REACT I0»ft OF
OXYGEN WITH CALCIUM SUUPMIQi
MAXIMUM S0| LEVEL ATJAIIjUraLEJJSIN6
1000
TEMP€RATURe,"C
-------
- 97 -
APPENDIX B
RESULTS OF WORK JUST PRIOR TO CONTRACT
This appendix summarises results obtained by the Esso Research
Centre on the CAPB gasifier in the period between submission of the
Technical Proposal 13) to NAPCA and signing of the contract. These
studies were carried out with the UK fuel and Denbighshire limestone
whose properties are listed in Table III of the text. Batch fluid bed
reactors of 7-ln.i.d. were used in these tests.
A. Reagent Life
The effect of successive adsorption/regeneration cycles on
reagent activity during jartial combustion was studied at different
bed make-up rates and mean particle sizes. Absorption was carried out
at a stoichiometrlc air rate between 30 and 35$* a superficial gas-
velocity of 2.5 ft/sec, and a bed temperature between 800 and 900°C.
In each adsorption period the bed was exposed for a total of 100 minute.,
during which its sulphur content rose to between 3 and 4$. Regeneration
was carried out by shutting off the fuel supply and allowing air alone
to contact the bed. Bed samples were taken for sulphur analysis at
the beginning and end of each adsorption period, and a regular check
was made on bed losses; fresh material being added to the bed after
each cycle to make up for attrition and sampling losses. Bed make-up
rate was controlled by adjusting the number and size of the samples.
Varying mean particle size in a downward direction only was considered
since an increase would be unlikely to give a beneficial effect. This
was achieved in the low gas velocity batch fluldiser by fitting baffles
above the bed which gave better retention of the smaller particles and
also by changing the size distribution of the limestone feed to
600-1200 H from the original 850 - 1200 M-. These alterations gave a
reduction in mean particle size of lime in the bed from 810 to 635 p,.
Figure B-1 illustrates the effect that successive adsorption/
regeneration cycles can have on reagent activity.
These results show that, under the particular operating
conditions employed, the desulphurising efficiency dropped from 100$
to a fairly steady level of 60$ and that this steady level was reached
in less than 10 cycles. The rapid drop in efficiency is typical of
all tests where a steady level of less than 100$ was observed.
-------
- 98 -
BED AGEING EFFECT REPLACEMENT RATE 0-5 WTS C«0/WT. S.
IOO
> 80
UJ
o
fe 60
I"
S
SULPHUR
0 8
V
v*x»* • . •
1 .
-
- 1 1 1 1 1 . ...• .1 . >
--4 8 12 16 20 24
RUN NUMBER
FIGURE B-l
Figure B-2 summarises the results obtained in tests on the effect
of bed make-up rate and mean particle size.
EFFECTS OF«D MAKE UP RATES AMD
fP»RTICLE SIZE
100
-------
- 99 -
These show that at bed make-up rates in excess of 1.8 wts. CaO
per wt. of sulphur, a desulphurising efficiency of 100$ can be maintained
indefinitely and that at lower make-up rates, the efficiency falls off
rapidly. Reducing the mean particle size had no significant effect on
the steady level of desulphurising efficiency within the limited range
covered.
B.
Retention of Sodium and Vanadium by the Bed
During the life tests, the retention by the bed of sodium and
vanadium present in the fuel oil was studied.
Sodium pick-up was found to be consistently low, being less
than 20$ of the amount introduced.
The results obtained on vanadium retention were more encouraging
but also more diverse. Under conditions of partial combustion, the
efficiency of vanadium pick-up varied from test to test between the
levels of 35$ and 100$ and was not related to bed make-up rate (Figure
B-3).
VANADIUM PICK-UP UNDER GASIFYING CONDITIONS
T9-
a
s
£
25
FUEL VANADIUM
KO MAKE UP RATES
• I -3 WTS CoO KM WT. S
A 2-5 WTS CoO KR WT. S
x 1-7 WTS CoO PER WT.S
50 100 ISO 200
FUEL CONSUMED (Kg)
250
300
FIGURE
A correlation was found to exist however between rate of vanadium
pick-up and rate of iron pick-up (Figure B-4). The iron originated from
the metal components of the unit which protrude into the bed.
Subsequent work during the contract has indicated that the iron-
vanadium correlation is not one of cause and effect but that vanadium
retention and iron pickup were independently related to regeneration
temperature.
-------
- 100 -
NCUmONSHIP HMUN VIMMHUM AND IMM
PICK-UP MATES OMtMMtmCATION CYCLES
r
3o,
100
PICK-UP KATE (I/CYCLE)
FIGURE B-
C.
Gasification without Underfiring
All gasification tests previously reported had been carried out
in the low gas velocity batch fluidiser at superficial gas velocities
of less than 5 ft/sec. At these velocities, the combustion of kerosene
below the distributor was necessary to preheat the fluidisiag air and so
prevent the deposition of coke at and below the level of the fuel injectors.
The utilisation of 25# of the incoming air in this way had resulted in the
sulphur being evenly distributed, throughout the bed as calcium sulphide.
Tests carried out early in 1970 in. the tall. No.3 batch reactor showed
Miat underfiring was unnecessary when superficial gas velocities in excess
>,: 3 ft/secware employed. Operating without underf iring almost certainly
prolongs, distributpr life and, for this reason alone, it is the preferred
mode of operation. However, it also makes the bed more complex chemically
in that there is a transition from,calcium .sulphate to, calcium sulphide
with increasing distance from the distributor and, along with this, an
increase in total sulphur content. These effects result from the additional
oxygen entering through the distributor- This oxygen reacts with the
calcium sulphide below the level of the fuel injectors and at the
temperature of the absorbing bed (800-900*C) predominantly forms calcium
sulphate. The sulphur dioxide which is released during the less favourable
formation of calcium oxide is reabsor.bed as calcium sulphide above the
level of the fuel injectors.
The effect of this change in operatic**, on, -toe variables affecting
desulphurising efficiency has been studied to some extent. Results so
far show that an operating temperature between 800 and 900*0 is still to be
preferred. However, the influence of stoichioraetric air rate has changed
radically as is evident in Figure B-5 where the effects of stoichlometrio
air rate on the desulphurising efficiency of fresh beds are compared with
and without underfiring.
-------
- 101 -
STOICHIOMETRIC AIR RATE WITH AND
WITHOUT UNDERFIRtNG
100
80-
60
40
20
9% WT. S IN BCD
• 4-12% S FUEL WITH UNDERFIRIMG
A 2-62% S FUEL 3 FT/SEC
x 3% SFUEL WITHOUT UNDERFIRING
4 FT/SEC
_L
15
20 25 30
% STOICHIOMETRIC AIR
35
40
FIGURE B-5
Figure B-6 shows the manner in which desulphurising efficiency
varied with increasing CaO utilisation during the course of these high
velocity runs. Except for the run at 20.7# of stoichiometric air, the
lime maintained its activity to over 20# of utilisation.
100
*
o
u.
UJ
o
UJ
a:
a
to
90
80
70
60
50
SULPHUR REMOVAL CURVES, TALL BATCH REACTOR
Stoichiometric Air %
20 7
2I-.Q
23 7
26-3
29-8
JL
05 10 15
FRACTION CaO REACTED
20
25
30
FIGURE B-6
-------
- 102 -
It is evident that with no underf irlng high desulphurising
efficiency can be maintained at lower stoichiometric air rates. If
it is assumed that tihe drop In .efficiency previously observed with
reductions in stoichiometric air rate was due to the observed build-up
of carbon on the particles, then it is likely that the current
improvement is due to an increased amount of carbon being removed from
the bed In the oxidising zone resulting in a continual supply of carbon-
free stone to the gasifying region.
To be able to operate at the lower stoichiometric air -rate and
higher gas velocity is a great advantage since it means that the fuel
throughput is increased for a fluid bed of given cross- sectional 'area.
Multicycle tests reveal no significant effect of the absence
of underf irlng on the relationship between lime replacement rote and
the lined-out efficiency of sulphur removal. 2he line established In the
absence of underf Irlng, Figure B-7, is quite elmllar to that found with
underflrtag, Wgure B-2. The data in Figure -B-7 were obtained at 250
• of stoichiometric ttlr compared with 30-250 used with the underfiring
'tests in Figure B-2.
fiRFECT OFiH)*£-tJP RATE ON
SFFICIENCY
WTS CoO #€R
-------
- 103 -
APPENDIX C
BATCH UNIT FUEL AND LIME COMPARISONS
Appendix C contains experimental data from batch reactor
studies in the 1-A unit during preliminary limestone and fuel
comparison studies.
-------
O&BLE C-l
SHEET 1 of 6
Run
No.
22(c)
23(o)
25(c)
26
27
28A
289
28C
29
>?
32
Fuel
UK(a)
Limestone
UK
CAFB BftTCH:UNIT TEST CONDITIONS
1-A Unit Preliminary Fuel and Lintestorie Compari-sona
Limestone
Particle Size j
800 - 1200
Fuel Rate
(g/inin)
Air Rate
(L/inin)
"(d)
"(d)
"(d)
".(d)
"(d)
UK(d)
200
150
150
200
195
200
200
266(e)
200(f)
200
200
200
473
441
435
440
450
445'
464
459
420
480
471
454
Superficial air
vel. (ft/sec.)
4.3
4.0
>-9
3.9
4.1
4.0
4.2
4.2
3.8
4.4
4.3
4.0
(a) U.K. fuel J.QJif S
(b) U.S". fuel 2.22* S
(or) Fuel injected to' tfed ccfttrtf
(d) Use;d beef of lirtfc e^loisred in test
(e) l.^CFW air to fuel nozzle
(f) 0.5 CFM air to fuel nbzzle (other tests use 1.0 CFM air to fuel nozzle)
% Stoich.
Air
21.7
27.0
26.6
20.2
21.0
20. >
21.3
21.1
19.3
22.0
21.6
19-9
Temperature C
Initial
770
790
790
9
$50
840
860
790
915
730
900
Average
865
840
830
835
840
840
840
850
840
860
870
840
o
-------
TABLE C-l
SHEET 2 of 6
CAFB BATCH UNIT TEST CONDITIONS
1-A Unit Preliminary Fuel and Limestone Comparisons
Run
No.
33
34
35
36
37
38
39
40
41
42
43
44
45
46
47
48
49
50
51
52
53
Cycle
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
Air Rate
1/min.
426
432
439
448
444
442
442
423
424
412
423 '
415
438
414
414
413
424
422
418
426
419
Superficial
Gas Vel.
ft/sec
3-9
3-9
4.0
4.1
4.1
4.0
4.0
3-9
3.9
3.8
3-9
3-8
4.0
3-8
3-8
3-8
3-9
3-9
3.8
3-9
3.8
Fuel Rate
g/min
165
»
n
„
ir
n
it
n
n
it
it
n
it
n
it
n
•n
ir
tt
it
it
(Cycle Test, UK Fuel
% Stoich Feed Particle
Air Size u
9
23.7 850-1400
24.0 «
24.4 ,t
24.9 n
24.7 n
24.6
24.6 »
23.5
23.6 it
22.9 "
23-5
23. 1 ir
24.J it
23.0 it
23.0 tt
23.0 ..
23.6 it
23.5
23.2 »
23.7
230
and Limestone)
Bed Depth
(Static )
in.
15.5
_
15.4
_
15.6
-
15.2
-
15-7
-
15.2
-
-
-
15-2
-
-
-
15-2
Limestone
Make-up
Rate
g/cycle
400
it
n
n
n
n
if
n
n
n
n
it
it
n
it
ti
n
it
n
it
it
Average Bed
Te:np. (Absorption)
•c
860
880
870
870
880
890
890
880
900
880
880
870
880
880
890
890
880
880
880
880
880
Max. Bed
Temp. (Regeneration)
°C
1050
1060
1080
1070
1080
1070
1080
1070
_ . .
1070
1060
1070
1090
1080
1070
1060
1080
1070
1070
1070
I
5
60 minute absorption cycle
Normal, Low pressure drop, distributor
-------
CAFB BATCH UNIT TEST CONDITIONS
1-A Unit Preliminary Fuel and
Run Cycle
No.
5*
55
56
57
58
59
60
61
62
63
'64
65
66
67*
68
69
70
71
72
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
Air Rate
1/min.
441
447
461
411
422
414
408
419
433
415
415
419
416
394
413
411
407
403
394
(Cycle Test
Superficial Fuel Rate % Stoich
Gas Vel. g/min Air
ft/sec
3-9
4.0
4.2
3.8
3-9
3-8
3.8
3.8
4.1
3-8
3-9
3-9
3-8
3.6
3-8
3-8
3-7
3-7
3-7
165
165
165
165
165
165
165
165
165
165
165
165
165
165
165
165
165
165
165
24.5
24.8
25.6
22.9
23.4
23.0
22.7
23.3
24.1
23.1
23-1
23-3
23.1
21.9
23.0
22.8
22.6
22.4
21.9
Limestone Comparisons
UK Fuel and US Limestone 1691)
Feed Particle Bed Depth Limestone Average Bed
Size u (Static) make-up Temp. (Absorption)
in. Rate °C
600-1400
•
n
n
it
it
ft
it
ti
it
n
it
it
it
it
ii
V
If
It
ft/ cycle
16.5
500
500
16.0 500
500
500
500
16. 1 500
500
500
500
500
500
16.1 500
500
500
500
500
500
840
850
865
880
875
875
855
855
875
885
885
885
870
860
865
870
865
870
880
Max. Bed
Temp. (Regeneration)
•c
1050
1060
1060
1060
1070
1070
-
-
1090
1080
1080
1080
1080
1070
1080
1090
1080
1070
1080
60 minute absorption cycle
Normal, Low pressure drop,distributor
* 2nd attempt at cycle. In 1st attempt fuel pump failure after 21 min and spontaneous regeneration.
-------
TABLE C-l
CAPB BftTCH UNIT TEST CONDITIONS
SHEET 4 of 6
1-A Unit Preliminary Fuel and Limestone Comparisons
Run. Cycle Air Rate
No. 1/min.
78
79
80
81
82 '
83
84
85
86
87
88
89
90
91
92
93
9*
95
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
399
412
413
408
431
412
434
473
430
441
396
432
422
422
420
385
385
400
(Cycle Test UK Fuel and US Limestone 1690)
Superficial Fuel Rate % Stolch Feed Particle Bed Depth Bed Depth Limestone
Gas Vel. g/mln Air Size p (Static) (static) make-up
ft/sec In. Start in End Rate
Cycle Cvcle ir/evcle
3-6 165 22.2
3-7 " 22.9
3-7 " 23.0
3.7
3-7
3.7
4.0
4.3
3.9
4.1
3.6
4.0
4.0
22.7
24.0
22.9
24.1
26.3
23-9
24.5
22.0
24.0
23-5
4.0 23.5
3-9 " 23-3
3.5 " 21.6
3.5 " 21.6
3.7 " 22.2
600-1400
M
"
"
"
n
n
HI
n
u
n
H
n
M
"
it
n
ti
15-5 13.4
15-6 . 1800
— _ H
-
18.1
16.9
-
— .
- _
16.8 "
16.1 . 840
- _ •
- - "
- - "
15-5
15-9
- - "
14.2. "
Average Max. Bed
Temp. (Absorption) Temp. (Regeneration)
•c *c
845
850
830
845
840
855
860
870
865
870
860
880
890
890
' 890
645
830
865
1080
1060
1060
1070
1070
1090
1080
1080
1080
1050
1060
1070
1070
1070
1080
1070
1060
1060
§.
60 minute absorption cycle
Normal, Low pressure drop, distributor
-------
TABLE
CAFB BATCH UNIT TEST CONDITIONS
SHEET 5 of 6/
Zoapariaona
Run Puel Lime- Limestone Fuel Rate Total Air
No. Stone particle g/mlh Rate
size >i .1/mln
96 UK(3.05*8)UK
99
100
101
162
103
1W6
1C
*
n
Superficial % Stolen. Bed Depth (In) Average Bed Max. Bed Duration of
Gas Vel. Air (StaticJ Temp.(Absorption) Teqp.( Regeneration) Absorption
ft/sec. Start End *C *C (mln.)
600-1400 210
20$
195
245
184
175
165
173
165
142
206
206
165
206
441
433
435
540
444
435
446
440
43^3
454
Jg
4*5
3-9
3.9
3.9
4.8
4.1
4.1
4.2
4.1
4.0
4.2
j.B
a
4.0
19O 15-5 14.2
19.1
20.5
20.2
22.1
22.8
24.8
23.3
24.1
29-3
14.4
14.2
13.8
14.2
14.6
14.2
14.4
14.2
14.2
20.0
20.2
25.4 - 14.4
19.8 15.5 15.6
800
815
810
810
850
860
865
860
8^5
860
740
770
820
810
1090
1680
107,0
1096
1090
1090
1090
1080
1070
1050
1000
-
1010
-
300
126
120
90
120
120
240
120
120
205
60
30
30
120
i
High pressure drop distributor
(a' Run ended without regeneration, cooled under nitrogen blanket
-------
TABES C-l
Run
No.
110
111
112
113
114
115
116
117
118
119
120
121
Fuel
UK. 3-03*3
CAPB
BATCH (MIT TEST CONDITIONS
1-A Unit Preliminary Fuel and Limestone Comparisons
Lime-
Stone
UK
"
n
H
It
H
N
H
•
*
"
Limestone Fuel Rate
Particle (g/mln
Size Range
600- 1400 207
190
161
163
176
" 150
" 204
• 205
190
' 205
173
" 150
Total Air
Rate
401
430
M9
433
415
419
400
423
420
436
Superficial
Gas Vel.
3-5
3-8
4.1
3-8
3-6
3.8
3-6
3-8 '
3-5
3.6
3-7
3.9
% Stolen.
Air
17-7
20.7
25.2
24.2
21.5
25.6
19-3
19-9
19-3
18.9
22.3
26.7
Bed Depth
(Static
Start
15-5
15-5
15-5
15-5
15.5
15-5
15-5
15-5
15-5
15-5
15-5
15.5
(in)
End
14.2
14.0
14.4
14.0
14.1
14.3
14.0
14.2
14.2
14.4
14.3
Average Bed
Teap. (Absorption)
•c
810
820
855
825
as
B30
780
790
815
790
830
850
Max. Bed
Tenp. (Regeneration)
•c
1090
1085
1090
1060
1065
1090
1055
1060
1080
1O70
1080
1060
SHEET 6 of 6
Duration of
Absorption
(Bin. )
129
120
124
125
122
123
122
120
120
120
120
120
Distributor
Normal
"
N
"
High Pressure
Drop
N
•1
N
"
*t
-------
TABLE C-2
SULPHUR REMOVAL IN FRESH BED BATCH REACTOR TESTS
1 of 3
*OP
FEED SULPHUR REMOVED
1-A Unit
Run No.
Time, Mlns.
15
30
35
*5
60
75
90
105
120
135
150
180
210
240
270
300
330
EIOU) %
Carbon Rate
g/aln
11
77-3
79-3
-
-
_
.
83.2
-
-
.
82.7
*
83.8
-
80.7
.
-
82.5
-
22
100
-
95.0
95.2
95.1
_
94.7
-
90.4
.
84.7
86.1
-
.
_
.
-
92.5
.15
§2
100
-
-
-
-
93.7
92.7
-
.
-
92.0
.
84.1 .
-
79.1
76.3
78.3
90.5
.69
25.
86.8
62.3
-
64.1
64.8
61.6
64.5
61.5
60.9
-
-
-
-
-
.
.
-
61
7.83
26
-
-
-
-
-
95.7
- •
86.8
-
75
_
-
.
.
_
-
97.5
2.22
27. 28-A 28-B 28-C
75.5 ...
....
....
65.4
....
....
-
.
....
60.2
-
5M - -
43.4
.
....
62.5 -
2.84 .82 .89 1.27
30 32
- 83.7
- 74.4
.20
.31 6.68
(a)
ELO • Sulphur removed ef f loleooy at 10J< CaO reacted.
-------
TABLE C-2
SHEET'2 of 3
Run No:
Time (min)
10
20
30
40
50
60
70
80
90
100
110
120
130
140
150
160
170
180
190
200
210
220
230
240
250
260
270
280
290
300
E10(&) %
Carbon Rate
g/min.
77.5
73
7.23
(a)
SULPHUR REMOVAL IN FR2SH BED BATCH REftCTOR TESTS
% FUEL SULPHUR REMOVED
1-A Unit
9J3 99. 100 101 102
103 104
105
84
7.70
E10
83
81
91 95 86 85
9.76 H.58 - <.2 3.98 k. 69
Sulphur removal efficiency at 10% CaO reacted
106 109
88.5
87.6
84.5
82.6
82.4
82.7
77.8
73-9
75.4
76.9
78.4
75-7
74.6
72.6
73-4
73-5
69.4
_
76.6
69.8
78.7
76.1
72.6
67.1
70.7
70.6
69.8
69-3
68.9
68.7
-
72.7
74.6
74.6
74.3
72.7
72.7
74.2
73-9
73-5
72.7
71.6
72.7
_
-
-
-
_
-
-
-
-
-
-
_
_
-
_
_
_
99.1
96.8
91.1
86.2
86.0
85.9
85.1
84.5
83.6
83.2
83.6
83.6
-
_
-
-
-
_
_
-
_
-
-
-
_
_
-
_
-
• -
99.4
97-6
90.4
85-9
83.1
83.3
83.2
85-3
84.3
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
_
_
-
100
84.2
81.0
79-7
83.2
82.6
82.6
81.9
81-3
81.0
82.2
81.1
_
_
-
-
-
_
-
-
-
-
-
-
-
_
-
_
-
—
100
99-2
99-3
99-5
99.5
92.8
92.3
92.1
90.5
90.2
89.1
88.8
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
100
100
100
100
100
100
100
100
94.7
91.0
89.1
88.0
86.8
86.2
84.2
82.7
81.7
82.2
80.9
81.1
80.4
79.9
79-0
77-8
_
_
-
_
_
-
100
89.4 •
87-1
86.6
86.3
86.6
86.8
86.4
86.0
85-7
85-3
85-0
_
_
_
_
-
_
_
_
_
-
-
-
_
_
-
_
_
-
100
100
88.3
86.6
86.4
-
-
84.2
83.8
87.3
85-3
85.3
88.7
-
-
-
-
-
—
-
-
-
-
-
-
-
-
-
-
-
100
100
100
100
100
99-1
98.3
97.5
96.8
96.8
96.7
96.2
95.8
-
-
-
-
-
92.9
91.9
91-3
-
-
-
_
-
_
-
-
100
-
-
—
84.6
-
-
-
80.7
-
78.8
-
-
-
-
-
-
-
-
-
-
-
-
-
-
_
-
_
—
-
96.5 80.5
C-25
-------
TABLE C-2
SHEET 3 of 3
SULPHUR REMOVAL IN FRESH BED BATCH REACTOR TESTS
% FUEL SULPHUR
Run. No:
Time (min)
15
30
*5
60
75
90
105
120
EIO(A) *
Carbon Rate
g/mln.
110
85.6
79-7
78.6
75-7
75-2
74.5
71.6
-
73.0
7.57
111
82.6
83.8
78.8
78.8
77.4
75.9
76.4
77.0
2.54
112
100
88.2
85-5
85.6
84.7
85.6
87.8
84.1
85.0
l.ll
113.
100
95.5
91-5
91.1
87.5
86.1
88.3
87-9
87.5
1.67
l^A Unit
114
100
100
94.3
89.4
86.2
86.4
85.5
80.9
86.2
-
REMOVED
115.
100
100
100
100
100
99.0
98.4
94.1
98.8
0.55
118
84.5
83.5
81.9
80.0
82.6
79.0
80.8
77.6
79.5
6.89
U9_
86.4
84.7
84.3
83.3
84.4
83.1
82.7
83.4
83.5
6.61
120
95.3
91.2
90.7
90.3
89.1
88.4
87.0
86.8
88.0
4.17
121
100
100
100
100
100
100
100
99-8
100
0.59
1
M
M
to
1
(a)
E10 = Sulphur removal efficiency at 10J6 CaO reacted
-------
C-3
SHEET 1 of 3
Run
No.
Cycle
33
34
35
36
37
38
39
40
41
42
43
44
45
46
47
48
49
50
51
52
53
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
SULPHUR AND METALS PICK UP IN BATCH CYCLIC TESTS
(UK Fuel and Limestone)
Lime Metals Content^)
Sodium Vanadium Iron
ppm ppm ppm
1-A Unit
Bed Sulphur
After Regeneration
% by
-
-
-
-
_
_
_
-
-
_
_
—
75.7
83.5
73-2
-
77.9
70.8
69.7
69.1
91.0
92.5
93-7
91.7
87.8
_
67.1
-
-
_
_
—
_
_
_
_
_
_
-
-
-
_
_
_
_
.
_
-
_
.
.
—
79.3
82.6
74.1
74.1
78.1
75-6
73-0
72.5
94.7
94.1
95-1
_
-
_
_
-
-
-
82.5
—
_
.
_
-
-
-
-
-
93-8
100
93-8
91.4
90.2
83.0
-
' 70.6
-
-
-
73.8
81.3
76.7
80.2
70.0
76.9
72.9
70.9
72.0
72.2
160
180
170
200
200
180
160
230
150
230
240
270
_
260
260
230
220
340'
300
280
270
280
500
660
840
1000
1090
1340
1340
1600
2000
1900
2000
—
2400
2400
2600
2600
3200
3100
3100^
3100
640
680
670
780
790
910
890
890
820
1000
8*>
890
_
880
920
1000
1200
1300
1100
1000
1100
1.07
1.42
1.75
2.21
2.17
1.77
2.19
1.99
1.87
2.55
2.79
2.74
—
2.62
2.28
2.43
2.42
2.53
2.52
2.50
2.29
VM
I
(1) Measured after each regeneration
(2) All sulphur present as sulphate
-------
TABLE C-3
SHEET 2 of 3
SULPHUR AND METALS PICK-UP IN BATCH CYCLIC TESTS
- (UK Po*l,»nd US Limestone 1691)~~~
1-A Ifrit
Cycle
1
2
\
I
7
Fuel Sulphur Removed
min 3§ min 60 min
Lime Metala Content*
Sodium ppm Vanadium' PPB Iron ppm
Bed Sulphur Content (£ by wt)
Total 8 fft> min) ;, Total a fA.H.') Sulphate (A. H;)
IX
12
IS
us
-
89-3
81.6
- •
.
78.8
87.1
- • '
56.9
37.4
78,4
66. Q
Q&.4
70.2
73-1
77.2
-
73.7
76.9
98.8
87-1
83.3
84,7
87.5
83. 7
91.1
_
74.3
73-7
73-4
68.8
7*, z
79.3
71,3
73-7
72.3
. '
-
91.4
89.5
83.6
80.3
85.1
84.5
85-8
79.4
73-9
75.1
75.4
7te%
6|L%
7*«5
71.3
69.8
73.4
7*. 2
340
320
300
180
250
260
260
2JD
320
32Q
35ft
390
4600
4700
4500
4300
4800
5000
4900
4800
4700
4850
4840
4620
4610
5010
4780
4740
4850
4660
4650
3.02
4.63
5.44
4.10
5-76
4.60
4.66
4.00
4.47
4.62
4.50
4.73
4.65
3.68.
4,32
4.4?
4.49
4.26
4.16
1.98
2! 31
2.03
2.39
30
29
43
2.45
2.45
42
2.39
2.21
2.15
2.28
2.14
2.03
2.12
A. H. After Regeneration
* Measured after each regeneration
-------
TABLE C-3
SHffiT 3 of 3
SULPHUR AND METAE5 PICK-UP IN BATCH CYCLIC TESTS
(UK Fuel and US Limestone 1690)
Run
Mo..
78
79
80
81
82
83
84
86
87
88
89
90
91
92
93
94
A.R.
*
Cy«
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
Aft
Mea
% Fuel
min
98.8
92.1
83-5
82.9
90.8
88.1
85-9
87.7
86.3
86.0
78.8
77.0
68.0
65.6
75-9
72.7
Sulphur Removed
45 min 60 min
99-2 93-5
89.3 85.8
83.1 78.7
82.2 79-9
82.7 81.8
86.1 84.8
84.1 80.3
85.7 82.0
82.1 79.1
82.1 77.1
79-2 76.7
77-7 74.1
72.8 70.8
65.8
66.8
78.8
64.3
64.4
73.0
64.1
1-A Unit
Lime
Sodium ppm
500
460
420
480
310
290
250
290
260
230
260
230
260
240
250
220
260
260
Metals Content*
Vanadium ppm
210
360
400
430
570
880
760
730
900
1030
• 1280
1390
1340
1620
1570
2000
1870
2060
Iron ppm
1450
1240
2590
1430
1350
5060
5050
5120
5170
5040
5100
5070
4780
4950
4790
4900
4700
474O
Bed Sulphur Content (jt by wt)
Total 3 (60 min) Total S (A.R.) Sulph (A.R.)
3.34
4.11
4.37
96
83
71
00
3.72
3-57
3-70
3-87
3.86
3.89
3.95
4.04
4.13
4.30
1.58
2.41
81
74
89
93
56
38
43
37
85
15
94
95
,81.
2.29
2.29
2-13
1.58
2.26
1.81
1.71
83
93
56
38
43
35
84
97
84
82
72
2.01
2.14
2.04
After Regeneration
Measured after each regeneration
-------
TABLBC-*
GHfNOB IN PAJfflCUt SIZE OF LIME EWUMS CYQMc'
(UK Fuel
F»rtlol« sice
Microns
£300-1400
8*M»»
600-850
of Bad in Site Range
Unoalclned Stone
35.6
57.9
5.8
Q.Q
0.7"
Cycled
13 Cycles
*•>
Xt
10.9
1.8
Regenerated Line
gO Cyclea
3.0
3e!s
1^.9
8.7
-------
TABLE C-5
TABLE C-5
CHANGE IN PARTICLE SIZE DISTRIBUTION DURING CYCLIC TEST
(US Limestone 1691)
1-A Unit
% by wt. Bed in Size Range
Particle Size Range Cycled Regenerated Lime
Microns Uncalcined Stone Calcined Stone 3 Cycles 13 Cycles 19 Cycles ,_,
M
1200-1400 30.5 20.9 50.7 14.4 23.0 7"
850-1200 59.2 57-7 53-2 54.1 50.9
600-850 10.1 17.6 14.9 24.3 23.3
300-600 0 2.3 1.0 5.7 2,5
<300 0.2 1.4 0.2 1.5 0.3
-------
e-6
CHANGE IN PARTICLE .SIZE DISTRIBUTION DURING CYCLIC TESTS
(US Limestone 1690)
Unit
fay tft. led
B»oge
Size Range
Microns
1200-1400
,859-1200
.600-850
300-600
Cycled Regenerate* Line
iinoj t ^ stone
26.5
64.1
8.9
0.1
0.4
1 Cycle
29.0
55.4
13.8
1.2
0.6
5 Cycles
13.8
49.8
24.3
7.0
5.1
10 Cycles
14.1
49-9
24,3
8.3
3.4
18 Cycles
21.7
48.2
20.4
5-9
3.8
00
I
-------
TABLE C-7
COMPOSITION OP GASIFIER PRODUCT
DRY GAS BASIS
1-A Unit
Composition of Dry Gas, % by Vol.
Component Run No;
Oxygen
Nitrogen
Carbon MonoxJ.de
Carbon Dioxide
Hydrogen
Methane
C2
C3
.52.
0
64.0
10.4
10.1
4.2
4.4
5.4
0.3
0.2
65
0
57.0
10.3
9-7
10.0
6.5
5.8
0.5
0.2
86
0
62.0
8.3
10.9
7.4
5.1
5.4
0.8
0.1
-------
- 120 -
APPENDIX D
BATCH REACTOR VARIABLE STUDY WITH PRBSH LIME BEDS
The bulk of this appendix contains the data from the fresh
bed tests which formed the Initial part of the variable study using
limestone BCR 1691 and the Anuay Fuel Oil. It also contains the
results of commissioning testa on the *tA sad 4fi units.
As discussed in detail previously, the variable study was
divided Into two parts: fresh bed and cycle tests 00. Ideally,
cycle teats should have been used throughout since we wished to relate
the oJMnrgied effect to continuous operation. However, they are very
time oousaitng and a reasonable coverage of the varlaMsc in this manner would
have been prohibitive In both time and cost. Consequently, the two
part approach IMS adapted. The objective of the first part - the fresh
tad te*t programme - was to determine the basic effects of ttoe varlafcies
listed below and say Interactions which may have occurred befcween
tkWU
•Tj
Air/fuel ratio
'' Bad Temperature
Superficial Gas Velocity
Sen Depth
Uswticle Size Range
It was also Intended that these tests sxwtkd Allow some of the
above variables to te left oat of the second pant of the v*riable study,
the cycl*e test programme.
i
Previous studies had Indicated the geaamal mfjiion of oanditlons
for reasonable QSSB operation.. Therefore, the new variable study was
concentrated in a fairly narrow range. Ihe programme of fresh bed tests, whicb
were earrled out is gfcvsn In Table D-V. Ibis waa buaad on a sequential
eyamttMttnn of the veo^adtOtes. 9w design «a* .-suob that not only WAS
detailed Information on tte basic effeat of each "tmWtMkf obtained* but
also Indications of Inttrrarrtnnn irtilnh nnrlnt sfitnaan-^frr^lfilap- It Is
worth noting that a full factorial design had toaan aaa*3atoB*l but was
rejected because our existing Information on .the tksiswiQur -of the
variables was not sufficient to guarantee that the-remits £rom suoh a
design would be meaningful. When, as In our case, it is possible to
obtain results from individual runs before the entiie experimental
programme Is completed, sequential nuitfirlmmt dettl0is,are much more
efficient than factorial type designs.
The progwwse*s listed differs .slightly frjost the originally prepared
prograraae (If), in that operation of the ontts showed, firstly, that higher
air/fuel ratios then anticipated, were requteed to «lve bsd tenperatures in
excess df 900*C and, secondly, that BTtTc could;not be,achieved with 2Q<
stiitulHiasitrle air at the lowest g«s v«lontty (k .ttjfaui). Qonseouently, all
-------
- 121 -
tests previously planned at 25# stoichlometric air and 1000°C were
reprogrammed to yyf> stolchiometric air and 950° C and the test at
20J6 stoichlometric-air and k ft/sec gas velocity was abandoned.
On-going analysis of results Indicated that these alterations would
not detract from the value of the tests.
The wide range lime particle size was investigated to
determine the feasibility of using lower cost stone. Discussions
with limestone manufacturers had indicated that the lowest cost size
range of stone is the 1/8-down fraction which is nearly waste product
in many operations. All previous tests had been carried out within
the 600-1^00 \i, range.
In our data system each cycle in each unit has been given
a run number which is suffixed by the unit number. Tables D-II to D-IV
list fully the results from the commissioning tests and Tables D-VI to
D-XIV the results from the freih bed part of the variable . study.
-------
TABLE D 1
Teat Conditions (Comnisalooing Testa 4A it kE)
MUD Fuel I4M>- Limestone Fuel Rate Total Air Superficial % Stolen. Bed Depth (Static) Average Bad Ntt Bid Duration of
No. Stone Particle (g/mln) Rate Gas Vel Air (Inches) . Trip. Ataorptlon Teip.Regeneration Absorption
Size Range (1/mln) (ft/sec) Start End *C *C («ln)
1AA U.K.
3 4A U.K.
zvwi U.K.
^r"«}i--' .. : 2t ••
gy/kM U.K.
U.K. 600-1400
U.K. 606-lkOO
tf.K. 600-1VOO
O.K. 6oo-ivbo
15$
205
1&
ii&
• u ""
tljiQ
l»17
413
422
4.0
3.6
3.6
3^8
26.0
18.7
20.8
27.2
15.5
15.5
15.5
15-5
14.2
13.9
Ik.k
13.9
855
810
805
850
1070
10«>
1100
1050
120
120
150
175
i
M
*?
-------
TABLE 0 XI
Sulphur Removal (Commissioning Teats)
Time(Min) : 15 30 Us 60 75 90 105 12O 155 150 165
% Fuel Sulphur
Removed In :-
Run 1/kA
Run 3/l|A
Run 21AB
Run 22/4B
100
70.9
99.7
100
100
7^.4
8UO
100
100
66.6
78.6
100
100
66.2
75-0
100
99.6
59-1*
-
100
99.2
67.1
-
100
98.2
56.8
-
99-1
96.1
57.9
-
98.4
-
-
-
96.6
-
-
65.8
95-0
N
W
I
91.8
-------
TABI£ D III
Regeneration Data (Unit *Al (Coir/.,issioning Tests)
Time
(ain;
1
Ij
2
3
4
5
6
7
8
9
10
»i
u
u»
13
1*
£5
^
;
17
16
19
j<|4
20
21
22
2"*
i^4'
25
25^
^
Run 1/4A
Exit Gas Composition (% by vol)
COg CO 0_2 302
19.2
-
8.6
1.2
1.0
0-7
0.2
0.1
0.1
0
0
-
0.2
a.5
*•*
3.6
1..6
-
-
.
_
_
—
.
.
.
.
_•
-
—
—
~
0
-
0
0
0
0
0
0
0
0
0
-
0
0
*
0
0
A
-
.
_
_
.
.
_
.
.
_
-
-
—
~
0
_
0
0
0
o
0
0
o
0
0
-
0
9.2
10.8
12.8
14.5 -
. *" _
-
-
.
— .
_ _
_
.
_ .
_ .
_ _
-
— . —
- -
— ~
i
Temperature
•c
890
-
910
9^0
980
1020
1025
1035
1040
XQ56
1060
.
1070
1000
950
910
870
-
-
-
-
-
_
.
-
-
-
.
-
-
-
~
Run 3/4A
Exit Gas Composition (%
:'''"
_
14.0
12.4
11.4
10.5
10.0
9.6
9-2
9.6
9.6 .
.
10.0
11.4
14.6
10.5
4f.5
-
-
3-5
2.9
2.5
-
2.2
2.0
5«i
-
7.0
6.5
6.2
-
5.2
.'
-
5.7
9-2
12.4
13-*
15-0
15.6
16.3
16.3
16.0
-
15.6
14>7
12.4
5-7
0.8
0.2
-
0.2
0.2
0.2
-
0
0
0
-
0
0
0
-
o
o
0
.
0
c
0
0
0
0
0
0
0
-
0
0
0
0
0
0
-
0
0
0
-
0
8.0
11.8
-
13.1
13.8
15.0
- i
16. o
by vol)
/
0
-
0
-
.
0
-
°.
0
.
o
-
0
-
_
10.1
-
13.6
-
14.1
- "
14.7
_ '.:
11.9
"
0.41
_
0.14
-
O.O4
-
Temperature
*C
860
-
870
900
920
940
960
970
960
1090
WOO
1
«.
1010
1020
idjo
1040
1060
1O80
-
1070
1070
£000
-
1080
1080*
D030
1000
940
910
—
870
OKTt
-------
Regeneration Data (Unit IB) (Cosnlsslonlng Tests)
TIM -
(•In) Exit GM Coapoaltlon (% by vol) TMperature Belt CM Composition (* tjr vol)
_ COz 00 QZ §02 *C OOg 00 £g a*
1
1*
2
3
4
4*
5
6
7
7*
8
9
10
10*
11
12
13
13?
14
15
16
16*
17
18
19
19*
20
21
22
22*
23
24
25
25*
26
27
28
28*
29
30
31
•xi i
•?•*•?
3?
33
34
13.4
11.8
10.9
10.0
-
9-6
9-2
8.8
-
8.4
8.4
8.4
_
8.4
8.4
8.4
_
8.8
9-2
_
.
11.4
12.9
16.0
.
6.6
4.2
3-5
3-2
3.0
2.9
2.5
2.2
7-5
5-1
6.0
5-4
4.8
4.2
3-5
8.4
10.8
12.2
13.6
-
14-5
15.0
15-0
_
15-5
15-5
16.2
_
16.2
16.2
16.2
_
15.5
15.0
_
.
12.7
10.0
3.5
_
0.6
0.2
0.2
-
0.2
0.2
0
-
0
0
0
-
c
0
0
0
0
0
0
0
0
0
_
0
0
0
-
o
0
0
.
0
0
o
_
0
0
0
-
0
0
0
-
0
0
0
-
0
0
0
-
0
0
0
-
9.0
12.1
13.2
14.0
14.9
16.2
_
0
.
0
-
0
-
0
-
0
-
0
-
0
-
o
-
0
-
0
-
0
-
0
-
9-7
-
12.2
.
13-2
-
13-4
-
13-7
-
13.0
_
1.2
-
0.38
•
.
-
-
B20 f°
«o : : : " 2S
5 : : :
- - - 1*5
1055
106°
1060
1000
1010 .... 950
1015 - - - - 9°°
1020 -
1025 -
1030 -
1040 -
1050 -
1060 -
1080 -
1080 -
1085 -
1090 -
1095 - - -
1100 -
1100 ....
1090 - ...
10» - - -
980
960 - - - ~ ~
920
890
-------
- 126 -
TflRT.p. n v
Fresh Bed Test Progranme
Test
1
2
3
k
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
22
23
2U
25
26
27
28
29
30
# Stoich
Air
30
25
20
15
25
30
20
25
25
20
25
25
20
20
25
25
20
20
30
30
30
30
30
30
25
25
25
25
25
25
Temp Gas Vel.
0 C ft/sec .
870
870
870
870
800
950
800
870
870
870
870
870
870
870
870
870
870
870
950
950
950
950
950
950
870
870
870
870
870
870
6
6
6
6
6
6
6
k
8
8
6
6
6
6
6
-6
6
6
if
8
6
6
6
6
8
8
8
8
6
6
Reagent Particle
Size (microns)
300 -
ti
tt
•«
«
n
•H
tt
tt
tt
600 -
1200 -
600 -
1200 -
300 -
it
n
n
it
n
600-
J.2GO -
3«0-
,300 -
600 -
1200 -
300-
300 -
0200 -
4260-
3000
1400
3000
1400
3000
3000
4*00
3000
3000
3600
4AOO
.3000
3000
3000
3®80
3WD
Static Bed
Depth (In)
15.5
15.5
15.5
15-5
15.5
15.5
15.5
15.5
15.5
15.5
15.5
15-5
15.5
15 ;5
10
20
10
20
15.5
15.5
15.5
15.5
40
20
, 15.5
15.5
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•20
10
20
-------
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15.5 11.9 8k5 1100 100
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19.9 U.( 610 1109 UO
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* 1090 IOX) 1060 970 10BO 800 9»0 970 99» 880 UOO 10*0 . 1080 1O90 1010 I HO 980 108D W OiO 1000 900 - UflO 1O» UUO ' 7TO 1O8D 900 UW)
T UX IOTO 1080 - 1060 750 9U 1000 lOli 810 97O 1100 . 1OW> IOX 9X 1100 890 IOM 890 81} 990 880 99O-. 9X «*> TV Who Mo lOfc-
* 9<0 IOX 1090 910 10k) - 860 »0» 10» 7» 880 KM . . 960 960 8(4 IOX 8W loko 8M T90 «n 8» - 9U Ko »U TJO 990 JU 1U>
* 9U 96» 1110 9X I0« 660 -CO IOM) IOJO TOO 8)4 970 - 87O 9?0 - 990- 97O 800 - 880 8DO - 890(UtJ8 - 871 HO IUD
U AMI 9X UX OW «0 - 790 loto 1060 - 790900. «M8tO- •>>- 9W - - . 7TO • 8hO T» T»0 . - »4 - UW
u «ao 860 loko »yj 870 - - 1060 790 • . 8ko - 8m 8» . 690- 890- - - no- - -- - No. MM
U 800 8X 970 10» . 109) 880 no - T8S 9X
15 . - 8X UX - - - IOX T*» - - - •§
it An IDS) ... 960 ....
IT . . 7W> IOM) ... 880 '.... ....
IS ... UU ... •» - .................". ....
19 / . . fto - . . 800
10 . . . Hi- . . . 780.-* ................ ....
II ...880... TiO (
a ...8»... •« ....
-------
- 13* -
Table D XXII
Stone Carbon and Sulphur Contents (Brush Bed Teat«)
Test Stone Carbon Content * Stan* Sulphur Content
After Absorption After Regeneration
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
22
23
24
25
26
27
28
29
30
0.04
0.38
4.68
11.88
1.59
0.06
7.12
0.65.
0
1.84
0.19
0.50
0.77
0.79
-
0
0.12
-
0.06
0.04
0.09
0
-
O.O8
0.24
0.24
0.12
-
0.23
1.00
2.43
1.19
0.98
2.02
0.35
2.01
3.1*
2.48
3-02
2.85
2.2V
1-97
1.78
3.13
0.98
1.13
2.36
2.57
2.03
2.56
2.28
2.16
* Sample taken from top OT
-------
- 135-
Component Conotntr«tlon» (% tgr »ol) *
°2
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
"2
57.8
6t>.8
57.5
59.1
6>i.5
62.7
53-5
56.1
62.9
62.0
55.1
58.it
60.U
63.7
55.2
57.5
63.0
57.9
62.0
59."t
58.lt
6lt.6
61. U
58.7
58.2
62.1
55.9
56.lt
*b
6.7
8.2
10.6
11.6
ll.lt
6.8
10.8
5.9
9.7
9-9
6.U
9.6
8.0
7.8
7.8
8.3
7.7
6.U
"t.9
5.3
5.0
6.5
7.0
9.0
7.3
12.2
7.3
6.1
00
15.3
6.0
7.8
3.3
6.3
13.7
5.*
17.8
10.8
9.lt
15.5
11.7
12.6
11.6
11.8
12.0
15.9
13.3
16.3
18.1
17.2
1U.O
12.6
9.2
12.2
-
12.0
16. U
(^
"t.7
5.5
6.5
6.9
3.5
3.6
6.6
"t.3
3.8
6.0
"t.7
6.3
5.9
3.8
5.1
5.1
2. It
2.3
3.6
3.3
3.6
3-3
U.I
5.lt
6.1
7.5
6.5
U.I
C2HU
3.5
5.1
6.0
7.U
U.I
2.0
6.7
3.5
3.8
6.1
3.9
5.2
5.3
3.2
5.1
5.1
1-9
1.9
2.3
1.3
1.3
1.6
U.O
5.8
It. 7
5.*
5.3
3.2
C2"6
0.2
O.U..
O.U
1.0
O.U
0.2
0.6
0.3
0.3
0.2
0.3
O.U
O.U
0.2
0.3
0.5
0.1
0.1
0.1
0.1
0.1
0.1
0.2
O.U
O.It
O.U
o.u
0.2
C3
0.2
0.7
1.0
3.3
l.U
0.1
1.6
0.3
0.6
0.8
0.3
O.U
O.U
0.1
0.3
0.7
0.1
0.1
0.2
0.1
0.1
0.1
0.8
0.8
0.2
O.U
0.2
O.U
CU
0.2
0.3
O.U
1.7
0.6
0.1
0.7
0.1
0.1
O.U
0.2
0.3
0.1
0.1
0.1
0.3
0.1
0.1
0.1
0.1
0.1
0.1
0.1
0.3
0.1
0.3
0.1
0.1
"2
10.5
9.8
8.7
5.7
5.U
10.U
6.9
10.7 '
7-3
7.2
10.6
7-7
9.2
7.8
8.3
9.2
8.7
9-0
11.5
12.8
12.7
9.9
8.U
8.0
10.6
U.2
9.8
10.1
% Fu.1 C
OxldlMd
63.6
26.U
3U.1
20.2
38.6
58.3
29-3
6U.6
51.6
>.5
51.5
U1.3
39.2
UJ.U
U8.8
39.3
75.6
67.6
62.1
68.0
72.1
65.7
U1.8
39-0
U3.5
.
UU.8
57.9
* Fuel
c > ck
0
33.5
21.7
25.2
20.1
17.2
18.0
0
12. U
2U.8
13.8
16.3
2U.8
31.5
8.6
22.5
10.0
8.U
2.8
12.3
U.8
10.7
25.U
15.2
17.7
_
11.2
12.1
% Pu»l H
Oxldlnd
9.U
28.0
U.I
8.7
15.0
2U.5
U.2
2.1
11-3
8.9
U.3
O.U
9.6
19.7
5.5
6.1
10.5
21.5
2S.O
11.5
17.0
30.9
17.6
11.0
13.0
_
10.1
5-8
* Mmsur*d it imblmt conditions
-------
- 136 -
The data.In Table D XI7 were subjected to regression analysis and
equations derived to describe the effects of operating variables on the
conversion of the fuel oil to the various components. These equations
wMeh have been used to extrapolate the data are listed below.
a) % toy wt C as CH.
,0.9 x V0.1
b) % by wt C as
- XR
°'28
c) % lay wt C as
d)
f )
% by wt C as C,
e) % by wt C as C.
by wt C >
1.25 x *
- x D
E67'6 x R°-5 x V°-3 x P0^ x
jlO.5
T2'2
x D
0'1
e6'0 x R2'1 x V°-9 x a6'1
°'9
g) % by wt C oxidised
h) CO/COp molar ratio
i) % by wt H as
1.1 X
R°'9 x T1'3 x 30-1
e8'2 x V°'5
0.000291 x e
0.00976 T
R°'9 x T2'* x S0'1
'* x IT-
J)
by wt H oxidised
e?5'5 x R?'2 x V2
T 6'7 x D0.7
R = Air/Fuel Ratio
T » Bed Temperature
V = Gas Velocity
D = Bed Depth
S = Number Average Particle Size
-------
APPENDIX E
BATCH REACTOR CYCIJS TEST VARIABLE STUD?
This appendix contains the data from the cycle tests which
formed the second part of the variable, study using limestone BCR 169!
and the Amuay Fuel Oil. In this part, the following variables were
examined.
Bed Replacement Rate
Gas Residence Time
Projected Sulphur Differential (P.S.D)
Particle Size Range
Our previous experience had shown that bed replacement rate
had an important effect on bed deactivation during cycle tests (^)
and, consequently, this variable was selected for investigation. The
fresh bed tests showed that gas velocity and bed depth could be
combined as residence time which had a marked effect on sulphur removal.
It was considered that this variable warranted further examination by
cycle tests. The third variable to be examined was projected sulphur
differential. It was not originally planned to look at the effects of
this variable since satisfactory sulphur removal was maintained in the
fresh bed tests to bed sulphur levels far in excess of the 3 to k%
proposed for the cycle tests. However, as the cycle tests progressed it
became apparent that this particular variable should also be investigated.
The fourth variable: particle size range, fell into the same category as
P.S.D.
On the basis of the results from the fresh bed tests, the target
levels for the other variables were set as follows:-
Air/fuel ratio 2F&, of stoichiometric
Bed Temperature 870°C
Although the fresh bed tests had shown satisfactory sulphur removal
down to 20$ stoichiometric air, 25# was selected on the grounds of more
convenient operation during cycle tests. Insufficient heat is liberated at
the lower air/fuel ratio for 870°C to be readily maintained during cycle
tests in the batch units.
The programme of tests was of sequential design with the objective
being to outline operating conditions for minimum make-up rate and
maximum gas velocity. Minimum make-up rate is desired to minimise
operating cost and maximum gas velocity to minimise capital cost of a
commercial gaslfler.
The test results are summarised in Table E-I and listed fully in
Tables E-II to E-XXVIII.
-------
- 138 -
The limestone comparison te%t is listed first. This test was
made during the programme to select between available U.S. limestones (k),
In this test, BCR l691 gave ?2# lined out efficiency at 0.75 wt CaO/wt.
sulphur. The lined but efficiency was identical to that obtained with
the Denbighshire stone at the same make-up rate. At this stage,
therefore, it looked as if the U.S. stone would behave similarly to
the U.K. stone under cyclic conditions. This surmise was supported by
results from the fresh bed tests. HdweVer, as the rest of the data In
Table_E-1 show, appreciable differences were revealed as the cycle
te"sts on BCR 169! progressed.
The first test in the current series (TlCl to C1H) was carried
out at a lime replacement rate of 2A wt CaO/wt. sulphur. This rate
would have given a lined out efficiency of nearly 10C# under the
original test condition's with UK stone, and on the basis of the
comparison test and the fresh bed studies, it was expected to give a
similar high efficiency with BCR 1691. In fact, the efficiency lined
out at only 6l£, which is lower than that observed in the comparison test
at lower make-up rate. A repeat test, T3C1 to Clk, gave the same
result*
These BCR 169! stone cyclic tests differed from the original
comparison test In a number of respects as shown below.
Comparison Test Recent Tests
Superficial Air Velocity, ft/se'c.
Lime Particle Size, microns
Pue\ Oil
Vanadium in Oil p'prti
Sulphur in Oil wt.#
Unit
k
600 -
U.K.
130
3.03
1-A
6
300 - 3200
U.S. (Amuay)
366
2.30
The next series of cyclic tests was conducted to isolate each of
these differences in turn in. order to establish which, if ariy df these, was
responsible for tthe lower efficiency -
Reducing superficial gas velocity to k ft/sec, fh tests Tl C15 to C19
and T3 C15 to C2£, Increased lined out suipvhur removal efficiency to
In tests T3 C27 to C31, make up with 600-11*00 micron stone improved
lined out efficiency to 69$. At first, this Heerris contrary to the fresh bed
test results. However, the difference cah be explained by the fact that the
wider particle size range stone suffers g-reate'r Ida's dtirlhg calcinatidn than
the 600-1^00 micron material. Separate tests show an average loss of 52. 5#
on calcination of the 300-3200 micron stone Cdftipared wi^ k2.1% of the
600-11*00 micron size range. The effective maketip rate of the wide range
-------
- 139 -
stone was, therefore, lower than that of the narrow range stone at
equal total addition rates. This effect did not operate in fresh
bed tests, because in those tests, equal weights of precalcined lime
were added to the reactor.
TestT^, cycles Cl to ClQ , was made with the U.K. fuel oil
containing J.OJfi sulphur and 130 ppm vanadium. Sulphur removal
efficiency was only 62$, the same as obtained with the U.S. fuel. The
cause of the difference, therefore, was not the fuel oil.
These tests indicated that the cause of lowered sulphur removal
efficiency in the cyclic tests was a combination of Increased gas
velocity and greater lime losses during in-situ calcination of the wide
particle size lime. No test of differences between experimental
reactors was made, as earlier comparisons had shown no differences
between the No.l and No. k units. It is unlikely that a difference would
arise with a different limestone.
Although reverting to a narrow particle size stone and reducing
gas velocity would improve sulphur removal efficiency, both of these
changes would increase costs. Therefore, efforts have been directed at
finding conditions which would allow use of the high velocity and wide
size range while improving sulphur pick-up and reducing replacement
requirements .
The tests discussed above were all conducted at a static lime
bed depth of 15-5 inches. When lime replacement rate at this bed
depth was increased to 3.09 wt CaO/wt. sulphur, the sulphur removal
efficiency improved slightly, to 65# (T5 CJ2 to
On further increasing the replacement rate to 5.08 wt. CaO/wt.S,
in a test starting with a fresh bed, an unusual reversal in trend was
noted. As usual in the first few cycles after starting with a fresh bed, the
efficiency declined. It fell to below 80$, but then began rising again.
The improved performance paralleled an increase in bed level as measured
by bed pressure drop. The bed level rise was due to insufficient lime
being removed between cycles to compensate the high replacement rate.
Level was allowed to rise to the equivalent of a 20 inch static
bed while the test series continued. The efficiency lined out at 98$ at
this level and replacement rate.
Reducing replacement rate to 2.5 wt CaO/wt.S, while maintaining
a 20 inch bed, gave a sulphur removal efficiency of 90# (T6 Cl to C17).
The improvement in performance with bed depth was not entirely
due to greater residence time. The test at 15-5 inches with k ft/sec.
aiperficial gas velocity actually had greater residence time than the
20 inch bed test at 6 ft/sec, but gave only lk% lined out efficiency.
A contributing factor was thought to be the reduction in sulphur
differential, the total sulphur pick-up per cycle, in the 20 inch bed
tests. This reduction came about because cycle length was held constant
while lime charge in the bed increased in going from 15.5 to 20 inches.
In the 15.5 inch bed tests the PSD was 2.6 wt.#. In the 20 inch tests'
this projected differential was only 2.2#.
-------
- 140 -
Subsequent tests were mainly directed at pursuing these trends
further with the objective of maintaining high sulphur removal under
lined out conditions at the lowest replacement rate.
Test 8, cycles 7 to 20, shows that an Increase in static
bed depth to 22.5 inches and a reduction in PSD to 2.0 wt.$. gave 97$
sulphur removal at the significantly lower replacement rate of 2.5 wt.
CaO per wt. sulphur. Attempts to pursue these trends even further
(Test 8 cycles ^5 to 61, Tests 10 and 11) were unsuccessful because
satisfactory regeneration temperatures were not reached. The deeper
beds Increased the heat losses through the walls of the units. This
heat loss is critical with: the small batch units but is not a problem
with the much larger continuous unit.
Test 8, cycles 67 to 80, and to a greater extent, Test 9
separate the PSD and residence time effects. The separation was
achieved by operating with a shallow bed to give a low residence time
whilst reducing the length of the gasification state of each cycle to
give a lower PSD. .Test 9, which Is similar in all respects to Test 1
with the exception of PSD being reduced from 3.0 wt$ to 1.95^* gave a
lined out sulphur removal efficiency of 864 as opposed to 61$. This
indicates that changes in PSD .are'more effective than changes In residence
time. Reductions In PSD are, however, limited by the amount of sulphur
required on the stone to maintain adequate regeneration temperatures.
A base case, Test 12, with U.K. fuel and stone, was carried out
at the end of the BGR 1691 cycle tests to ensure cycle test results from
the OAP batch units could be safely correlated with those from previous
units. The lined out efficiency of 89$ when compared with the 92$
expected from the relationship between make-up rate and lined out
efficiency determined previously'(Interim Report) indicates that the
OAPr batoh units have behaved similarly to previous ttftlts.
-------
TABLE E 1
Test
Comparison
Tl (Cl to Cl4)
Tl (CIS to C19)
T3 (Cl to C14)
T3 (CIS to C23)
T3 (C24 to C3D
T4 (Cl to CIO)
T3 (C32 to C4l)
T5 (Cl to C20)
T6 (Cl to C17)
T8 (C7 to C20)
T8 (Cdj to C34)
T8 (C45 to C6l)
T8 (C67 to C80)
T9
T10
Til
T12
Cycle Test Summary Table using
Fuel Limestone Absorption * Stoich.
Particle Temperature Air
Size *C
Range (>i)
UK (3-03*3) 600-1400
US (2-3*3) 300-3175
n i
n n
n n
n n
i
t
UK (3-03*3) "
us (2.3 *s) " '
n n
it it
n n
•I
it n i
n • n
It H
\ •>
It H
it n
it it
us (3.03*8 )uk(6oo-i4oo)
Super-
ficial
Gas Vel
BCR 1691
Bed Depth Residence Make-up Sulphur Lined Out
(Static) Time («) Rate Differ- Efficiency
(inches) (sec) (WtCaO/ ential *
(ft/sec)
.870
" 'n1
n
' " n
n
. •, „
< i
t • •
n
i
it
i •
N
• i i
•'•11
II
»' II
'1
It
II
It
\ '
II
It
N
25
n
"
"
n
• it '
« i
n t
H
" ,
< N <
I « •'<• v» .
* •&&
1 • '
1
30
, * '
M
35
27-5
30
4
6
4
6
4
6
6
6
6
•
- 6
;"6
.6
'6
i 'i
'..- ,6
4 -
5
4
15.5 '
n
n
•I t •
n
H
nit, t
n
( »
*
n
t
20"
H "
!',,.
22.5
1 » 1 it "
,k .|
"• ' 25.0
\\ , !5-5
15.5
22.5
20
15-5
0.22
0.15
0.22
0.15
0.22
0.15
0.15
0.15
0.21
0.21
0.24
0.24
0.28
0.15
0.15
0.36
0.25
0.22
Wt.S) (
6.75
2.43
2.76
2.38
2.75
'2-34
(6QO-l400ji)
2.24
3-09
5.08
St 50
•2; 50
•>N-70
• 2.69
2,'. 16
2.55
2.68
2.12
1.32
Nt.*)
3.7
3.0
2.6
3.0
2.6
3.1
3.2
5.1
2.2
2.2
2.0
i 1'
2.2
1.7
2.2
1.9
1.5
0.2
3.4
! 72 '
t ' '
61
' 75
61
74
68
62
65
98
*90
97
79
77
71
86
89
77
89
M
i
(a) Residence time in bed region above fuel injector.
-------
nau « ii
QUO (CicU •
Cy«l« fut\
ttm-
Stem
~ . I.
•••tor
rartldo .
ToUl Air
Ml*
(I/—)
* stoic*. M BMM (sutu)
hip.
Duration of Olotrl- Nik*-)*
1
5V*» 2
3V»» 3
5V* 1
3
•(2.3A) MM1691) 300-3175
*»
7
8
9
10
3V* u
36M 12
3V*» 1?
3V* tt
13
17
18
19
2J9
232-
238
221
221
2*0
2*6
.232
216
3
%
670
6*6
675
683
672
£?
671
666
681
660
6y*i
3.9
3-7
6.1
3.9
6.2
6.2
3-7
6.1
6.1
6.1
6.0
3.7
*.o
26.7
28
26,
28.
26.
25-
2*.
2t.%.
«*•>
26.*
3p;.»
28.7
13-3
1*.0
19.3
20.0
20.7
20.0
13-5
12.8
13.1
*.9
18.0
17.7
17.3
17.9
850
l*.0
15.0
13.-3
i*,o
1*.7
1*.0
Urge pueti of refractory viryliw U «li*
».0 29.} 1».7
•I* plocod onta 4Utrlhutor.
865
8*5
•30
•to
•35
•73
866
670
8to
865
870
96)
1020
103D
10*0
1030
1030
1030
10,0
1020
1000
10*0
: . . i
39
68
-------
o«u
B(2.3
«5A»
*6/U
«7/*»
«8/«A
»9/»«
50/*»
52/U
5V*«
.5
6
7
«
9
10
5V«» 12
5V»« 13
57/»» 1*
70/«B 15
71/1B 16
. 72/tt 1Y
7V*B 18
7V«B 19
TS/» 20
76/» 21
77/*« 22
78/»B 2)
79/*B 2*
ao/w 75
8l/«B 26
82/ln MteMn J
Total air
fata
(1/Blal
670
6*3
653
662
671
6**
65*
652
621
653
6*0
6J7
651
629
«51
«57
»52
»50
»))
*67
47*
*S8
*60
658
653
690
679
650
66e
660
670
6*5
661
676
628
6SK
709
6M
«76
6*0
652
>* ••* f:««
SiaBrtleUl * Stolen, a** DapUi
On. V«l Ur (laokc
(rt/m
6.1
3-7
5-6
6.0
6.1
3-7
5-8
5-9
5.6
5-9
3.7
3-7
3.8
$-7
*.o
«.l
*.l
«.l
3.9
«o
*o
*.<
*.2
6.0
5.9
CO
6.2
3-9
6.0
6.0
6.1
5.9
6.0
6.2
5-6
6.0
6.5
6.3
6.2
3-9
3-9
>
31.2
e-s
26.0
23-6
25.3
21.2
22.1
25.8
t*.2
23. 3
2J.7
2*.8
22.8
25-8
23.2
26.0
26.1
25-6
23.6
29.8
30.0
28.0
28.9
23-1
, "••
25-2
M.7
23.8
2*.a
H.I
23.*
H.8
2*.6
2*.S
21-7
23.1
250
2».»
23O
2T-»
25.2
Start
13-3
-
-
-
-
-
.
ISO
.
.
.
.
.
.
130
130
It.o
M.T
M.7
ISO
13-6
ISO
ISO
ISO
ISO
13.7
13.7
16.7
17.1
170
12.7
M.7
M-7
M.7
-
-
-
16.7
170
18.3
(StatU)
•) n
•*<
-
-
-
.
-
U.2
-
-
.
.
-
-
16.1
12.*
13O
13O
• M.7
M.O
M.O
13-3
M.7
M-7
l).»
M.7
ISO
15.7
15.6
16.5
16.5
11.7
12.*
12.0
13-1
-
.
.
16.0
16.7
17.«
£?(«£«£"«•> *aap.(».j.a.r»tlGo) »b.or,tl<.
r*>
860
830
833
830
850
825
835
830
8*0
«SO
8*0
8*5
835
8*5
8*0
830
850
8*5
8*5
870
865
813
860
855
855
865
tfo
860
860
860
860
860
850
863
8to
863
865
860
860
870
853
\ »»
980
1010
•1020
1030
10*0
1030
10)0
955
loeo
1030
1030
1020
10*0
W39
1000
1010
10*0
1030
1030
10*0
10*0
loto
10*0
970
10*0
10*0
1033
1030
10*0
10*0
1030
.
.
1030
10*5
1030
.
moot
10TO*
loao
uy
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•
•
•
*
81
•
•
•
81
•
8)
81
•
60
•
«
81
•
•
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81
•
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8)
•
•
m
•
•
•
•
•
•
•
81
•
•
81
•
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Ifioo
I la u ituapt to <*u:r -»• •rflclmt r»(intr*tlan, »ltdl» •i
U^tntvn KMcntd 1COC*:. Fu.1 n^piy tbm cut off u u-wi.
>t >urt at
i»tlea «lu> intll
-------
vau i iv
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C»ol»
69/U 8
Nu. CM '"'>ari' ttMHArbf BUtrl- Mto-i*
72/M
n/v B
77/U to
OMRV
rartiQM
•ujjp,
\ -\
U3(BOU69D 300-3173
•
•
•
•
•
•
•
•
•
N
•
•
•
m
m
m
m
m
ltV««l/
226
227
2J7
•9
2*8
2*ft
?*°
«a
246
299
; mmvf vmm» v*«*
(V-ln) Ut./Mi.)
635 5.8
6M 5-9
671
677
Z«
«»
678
652
661
6«3
.2
•3
•9
.0
•3
.0
.1
.0
Jtir • tMom*
Start
„ ,
29-
26.
26.
25.
27.
M.
25-
19.9
12.9
1«.4
1».0
13-6
tt.T
M.T
».7 ISO
t4-7 16.0
«3-5 16.5
T
•nd
12.7
12.9
12.5
12.0
»30
U-l
DO
u.o
1».5
W.9
(*C)
855
•fio
•TO
875
•75
aso
no
•60
•65
•so
(•C ) . ._.
1010
1030
1060
10*0
HBjO
M>
MtB
' MM
MOB
10*0
onrpiioi
n-
*5
•
•
•
55
•
*
*
•
•
(•/«*!•)
MMkl 1200
• ' •
- • •
• •
• .
• •
* •
• •
« *
• •
I
I
-------
SutConditle
fcr«le
51
Hut
No.
78/*A
79/»»
80/*A
8V»>
82/*A
83/*A
8*/*A
•s/**
86/»A
87/U
88/*A
89/«A
90/*A
91/*A
92/*A
9V»»
9*/»»
95/U
96/*A
97/*»
Orel* l*u«l LlBt- Lluestane
Stem Pirtlole
Slz. tan*
i wad.-jfa) os(Kmi69i) 300-3175
2 "
3 • .
* " " •
5 .
6
7
8
Q • " fl
10 '"
11 "'
IS
13 •"
1*
15 "
16 ""
17 - -
18 ' "
19
20 ""
fuel Rate
227
230
222
22*
232
2*7
227
222
21*
218
213
227
217
222
223
219
222
217
227
23»
ToUl Atr SttKrfleUl if Stolch.
Itata OM. Vel. Air
(I/Bin) (rt./Me. )
650
690
6*9
638
697
652
659
665
713
687
6*8
682
6*9
682
683
682
681
696
682
673
5-9
6.*
5-9
5-9
6.*
5-8
5-9
6.0
6.6
6.2
6.0
6.3
5-9
6.2
6.2
6.2
6.2
6.*
6.3
6.3
26.3
27.6
26.8
26.1
27.6
2*.2
26.6
27-5
30.6
28.9
27-9
27.6
27.*
28.2
28.1
28.3
28.1
29.*
27.6
26.*
Bed Deptfe
(ImbM
Start
1«.0
1*.7
15-3
15-7
170
18.7
18.7
20.0
17.8
20.0
16.0
16.0
18.0
18.7
20.9
20.0
20.*
20.0
20.0
18.7
(etatle) Aww* Bed A**r«f* Bed Du*mUflO of
i) T.^.p.lAb.orftloo) T»^>.(»t|po«p«tlon)A,Morptloo
•nd Cc) (C) ("in)
13-1
13-3
15.2
16.0
17.3
18.0
16.5
16.8
18.0
!*.«
!*.«
16.7
19-3
20.0
20.0
20.0
20.0
18.3
16.7
855 .
870
850
BIO
860
835
8*5
850
870
855
870 ,
865
855
850
850
860
855
865
875
880
1000
1030
1050
loao
JOBS
1030
1020
1030
1020
1010
980
1000
1010
1010
1010
1020
1C10
1C20
1C20
1C30
*5
1
*
"
"
11
"
"
*
9
9
9
9
9
9
9
9
9
9
•
Dletrl- Hrt«-up
butor wlcht
(llms torn)
HoTMl 2*00
I
*
I
-------
Fuel
ioyu
7
8
9
10
89/te ii
12
13
9V*B 15
16
17
Urn- Llm.too. Pu»l Itato Total Air ftvorfUUl * Steloh. M D*ptk (*UtU) AMI** M AMIW B*d
•Urn Mrtloto (•/•**.) '-MM IM*]-**!? * *lr ' (16*..) : Mi». UMiooipttai t^8ggnlittlUa»
9lM " " " ~~" " --....
03(8011691) 300-31T5
Duration of DUtrl- Mte-up
Kti*U<»
butor Ml
" 218
231
236
231
221
219
2i6
«*
231
2*2
275
275
238
238
226
218
217
U*dn.) ^t.
665 «
663 <
685 4
69* <
656 •
673 «
689 t
••**•;
679 <
655 4
671 4
661 <
«55 i
666 i
661 t
fix :
6*2 !
635 ;
£*•> •*
(.1 28.0 15.4
5.1 16.3 16.C
5.2 «6.6 17-3
5.3 27.< 19-2
5.9 *7-2 20.J
5.> 28.2 3O.1
6.* 29.3 20. (
5.2 26.* 20 <
s.o 'as.o 20.:
(.1 25.* 20.:
5.9 22.1 17-<
J.9 23.» 170
5.0 25.7 18-1
(.0 J5.5 18.<
5.7 25.8 20.4
5.8 27.0 20.4
}.8 26.9 20.<
t ' M
i 1».7
l 19.*
p ii.3
) 18.7
i 20.0
' 20.0
1 20.*
1 ?*•?
> 20.0
> 19.8
> 16.0
> 16.7
r ia.3
> i9-5
> 19.7
> 20.0
1 20.*
~~ ro
865
860
855
855
850
865
880
800
•55
860
•35
>°
s
8*0
850
855
=^1sr*r 5 "^ ^a-o
f8D ' «5 •€«
ioeo • •
use • •
1O2C • •
uec • •
103C
1A*C " '
V i»»e
uec • •
uec * "
9*
1015 '
102C .«
1030 • •
10S5 • "
1025 " »
1035
,^^ffl>
•i taoo
N
•
•
•
•
•
•
I •
• {
m
m
m
m
m
•
•
I
£
•
1000 • uMd ted aftted to ralM h*l*t to
• - -
-------
T»»t Conalttcna (Orel* l»at 8)
»» Cycle rial
»o. No.
W8/4* 1 03(2.3*8)
109/4* 2 •
110/4* J
111/** 4 •
113/4* S •
113/4* e •
U4/4* 7 "
115/4* 8
116/4* 9 •
117/4* 10
118/4* 11 *
119/4* 12 •
120/4* 1} •
121/4* U •
122/4* IS
123/4* 16
124/4* 17 "
125/4* 18 •
126/4* 19 •
127/4* 20 "
116/4B 21 "
U7/4B 22 •
118/4B 2}
119/4B 24
120/48 25
m/4B 26
122/48 27 "
12V4B 28
124/48 29
125/4B 30
12V4B 31
127/4B 32
128/48 33
129/48 3*
Lla»-
•tOM
BCR1691
fl
fl
•
•
H
*
N
•
ft
•
H
«
«t
*
H
"
H
•
II
ft
ft
"
"
ft
«
•
«t
•
It
It
-
••
Llaeatone
Particle
Size Range
300-3175
Fuel Rate
(s/«ln)
2O2
228
230
230
235
341
2*2
240
242
240
208
2O4
209
197
227
219
232
222
271
266
262
257
256
256
256
255
256
255
256
266
241
242
247
3»6
Total *lr
Rat*
(V«ln)
659
678
675
677
686
668
636
664
683
658
650
622
676
649
624
641
611
668
655
647
657
632
673
691
668
693
686
690
705
684
679
698
660
672
Superficial
Ou Vel.
(ft/aec)
5-9
6.2
6.2
6.1
6.2
6.0
5.6
5.9
6.2
5-9
5.8
5.6
6.1
5.8
5.6
5.8
5.4
6.1
6.0
5-8
5-9
5.6
6.0
6.3
6.0
6.3
- . J
6.3
6.5
6.3
6.1
6.3
6.0
6.1
% Stoleh.
Air
29.9
27-3
26.9
27.0
26.8
25.4
24.2
25-4
25-9
25-2
28.7
27.8
29.7
30.2
25-2
26.9
24.2
27.6
28.3
22.3
23.0
22.6
24.1
24.B
23.9
24.9
24.6
24.8
25-3
23.6
25.8
26.5
24.5
25.1
Bad Depth (Static)
(Inchaa)
Start End
16.0
15-7
16.0
17.3
16.7
16.5
20.0
20.7
21.3
21.7
21.3
22.7
22.7
22.7
24.0
22.7
24.0
22.7
24.0
24.7
20.0
25-3
24.7
22.7
22.9
23-3
23-3
23.3
22.7
22.7
23-3
22.7
22.7
22.7
15-3
15.1
15-3
16.7
16.0
16.0
19-5
20.3
21.3
21.6
21.4
22.0
22.0
22.9
22.7
23.0
24.7
23-3
23-3
24.0
18.4
23.3
22.4
22.7
22.7
22.7
22.7
22.7
22.7
22.4
23.3
22.7
22.7
22.4
Average Bed
Teap- (Abaorptloo)
(•C'
835
865
870
840
855
845
825
830
845
835
840
835
850
845
836
840
825
860
860
835
850
825
840
850
840
«55
860
865
865
865
850
855
850
855
Hue. Bed
Veep. 'Regeneration)
CO
955
990
1015
1010
1030
1035
1010
1000
1000
1010
990
10OO
1010
1010
1000
1010
1010
101O
1030
1025
1025
1010
1020
^025
1030
1010
1005
1040
1040
1040
990
1005
1030
1015
Duration of
Abeorptlon
(•in)
45
"
•
ft
(1
It
H
ft
n
H
•
•
m
m
m
ft
•
«
«
ft
it
•
•
"
n
N
ft
H
•
fl
H
fl
"
V
IMa>-ia>
MUM
(tiaeeto
(«/CTOle)
1200
900
-------
13V*
Stan
«(2.*b) 0X100691) 330-7175
13V*
13V*
137/*«
13V*
13V*
14V*
142/*
14V*
1«V*
128/U
129/4*
132/4*
143/U
150/4*
1S1/U
1S2/U
58
39
to
4*
*5
K
47
48
50
52
5>
S»
99
56
S7
61
6*
(6
66
67
68
69
70
71
72
73
7»
75
76
77
78
79
FiBl fete
(•/•in.)
245
295
2*
257
255
255
227
2*0
2*7
*V*
2K
206
209
JO)
208
201
aa
Z«
207
198
219
216
223
222
289
*»
^
'F
33»
219
232
236
m
2TB
afis
2»
235
265
267
2«
253
27*
269
267
-f :. .. ^
MBl »lr :
«•«•
(V-U)
665
T»
If*
7»
•W
666
OB
709
«T8
**8
«*»
-6
26,6
2T.6
2T.4
*•«
25.1
a.?
21.3
23.9
28-5
24-3
22.4
28.7
22.6
23-6
21.6
22.0
22.4
^X1
• B*d D>p<
2
23.2
22.7
22,7
2C-7
28,7
2>3
24.0
24.8
85,1
25.1
25-3
2>3
24.4
25-3
25-3
25-3
23-3
25.3
25-3
26.7
26.0
26.7
26.0
25.3
-
.
17-3
16.4
16.4
16.0
16.3
14.7
14.7
15-3
16.0
16.1
16.7
17-1
16.7
16.4
Ik (Mttte)
feu) ft
**
»-T
•.•
81.*
28.6
28.9
28.8
22.7
28.7
82-7
8C.7
22.7
23.7
24.8
24.0
24.0
24,5
23-3
24.4
29-3
29.1
29.3
25-3
25-3
••3
2J.3
23,T
26,0
26.1
25.3
-
.
17-1
16.4
16.0
16.0
16.0
14.0
l»-7
14.7
1S-7
16.0
16.0
16.3
16.0
16.0
*IM«9 Bid
*f. (AkMfptli
ro
»
•#
MB
860
8»
855
8*5
•35
8*0
tfS
890
8*5
830
850
855
860
8*0
«9B
8*5
860
8*5
8*5
850
ejo
<*5
8SP
«P>
•**»
8»
•T5
865
855
8*5
we
950
•to
880
865
865
870
875
860
870
870
*«. 8M.
•) MV iuinii
CO
1030
M20
-
1030
1035
1040
.
995
1025
101O
103B
WJO
10^0
UOD
MX»
I0*e
*»
]BD
1005
1010
imo
HU>
•BB
••88
up
UB9
970
1OOO
1B0P
IPBB
"w
ug»
1°3B>
10*0
1030
1039
1020
1030
UtfO
W60
106o
1060
1O70
1060
1065
ttmtlaa of
vtlcaJfttomrptKo
(«J»)
*5
•
m
m
9
m
m
*
*
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
»
90
•
•
•
•
3>
•
•
•
•
m
m
m
m
m
m
n
•
'
»Utrl-
butor
•«.!
*
•
•
•
•
•
•
"
•
M
•
•
"
•
•
.•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
Itote-t*
r1** ,
(itB-MtOM)
1200
•
•
•
•
•
A
•
*
•
»
•
•
•
•
•
•
•
1
• f
. •
•
•
•
•
•
U9>
•
•
•
•
•
ago
•
•
•
•
•
*
m
"I
* "
•
•
•
-------
»BU B n
T%at Coodltlona
T»it 9)
Nun Cyela
Mo. No.
159/4B 1 U
160/4B 2
161/4B 3
162/4B 4
16V4B 5
164/4B 6
16J/4B 7
166/48 8
167/4B 9
168/4B 10
169/4B 11
170/4B 12
171/4B 13
172/4B 14
17V4B 15
fuel Ll*m- LlMatona
atom Particle
Slza Rang*
(u)
3(2.3*9) BCR(l69D 300-3175
1 * 1
• t m
• M •
• N «
• • *
• • *
• w t
M It 1
* • t
» N '1
1 • t
"» « *
* m t
• ft
Pual Data
(|/.ln
227
223
231
201
214
218
221
219
223
223
223
223
231
229
233
Total Air
Hat*
641
649
643
650
658
661
658
660
649
646
663
648
653
628
657
Stawrtt
Oaa. \
(ft/M
5-8
5-9
5.8
6.0
6.0
6.1
6.0
6.0
5-9
5-9
6.1
6.0
6.0
5.7
6.0
lelal * Stolon.
'•1. Air
1C)
25-9
26.7
25-5
29-7
28.2
27.8
27-3
27.6
26.7
26.6
27.3
26.7
25.9
25.2
25-9
Bad Dapth
(Static)
(Inchaa
Start End
14.7
13-3
14.0
14.7
15.1
15-3
16.0
16.5
«.3
16.3
16.3
16.0
16.0
15.6
15-3
13-3
12.7
14.0
14.7
14.7
15-3
16.0
16.4
16.5
16.1
16.0
15-9
15.6
14.7
15.2
Avaraca Bad
Taap(Abaorptlon)
CO
845
850
84C
870
865
fffO
860
860
860
860
865
870
860
860
865
Hut. Bad
Taap(Dagana ration)
CC)
1000
1030
1040
1010
1040
1040
1040
1040
1050
1040
1050
1050
1035
1040
1020
Duration of Dlatrt-
AbaorpUon butor
(•In*
30 Monaal
•
* •
• •
M •
• •
• •
• •
• •
• II
• •
• •
• •
• m
N •
Nato-up
wl«ht
(llawatona)
(«/o»cla)
800
•
•
•
•
•
•
•
•
•
•
•
•
•
•
-------
HIM Cycl.
lie. No.
2
3
»
5
*
»
15
PtMl
Llm- UoutoM Awl H»U Total Air
MrUel« (t/mtn) %te
300-5175
118
101
3.9
3-8
*.0
3.8
*»»
Ha
136
119
4,0
9Urt
35.5
38.8
35.0
3B.9
3^-3
20.7
20.7
22.7
22.3
2B.7
*•>.
*.7
3».T
2>T
2B.7
18.»
20.1
M.O
28.7
21.0
2M
X.I
•e.o
28.3
22-9
CO
«
OoMtlen of DUtrl-
«b«or»tton tutor
«50»
«B9*
***
tfSf
a*>»
aeo*
loop
1020
MBO
"**
68
1000
1050
•HIM*
top (pi botlM of tad >«0'C.
-------
T«it Condition* (Cyel«
11)
Run
Ho.
Cyel*
Fuel
156/4A
159/4A
148/4A
159/4A
160/4A
161/4A 6
162/4A 7
16V4A 8
164/4A 9
165/4A 10
167/4A I?.
168/4A 13
169/4A 14
170/4A 15
171/4A 1C
17»/4A 17
LlM-
•tom
!8(l691)
•
•
"
•
*
•
*
"
"
•
•
•
•
1
•
•
LUeetone
Fartlcle
Size Rang*
300-3175
•
n
"
n
*
m
H
«t
M
•
H
"
"
M
II
ft
Fuel IUU
(•/•in)
150
170
175
185
180
164
187
195
193
166
180
194
185
188
189
199
196
Total Air
tat*
(l/'ln)
559
549
532
555
560
526
553
523
544
548
565
557
574
536
543
553
549
Superficial
OM V.X.
(ft/we)
5-2
4.9
4.8
5-0
5-1 •
4.7
5.0
4.7
4.9
5-0
5-2
5.0
5-3
4.9
4.9
5.0
5.0
% atolch.
Air
3*.l
29.6
27.9
27.5
28.5
26.2
27.1
24.6
25-9
30.3
£8.8
26.3
28.5
26.2
26.4
25-5
25-7
Bed Depth
Start
18.0
19-7
20.O
20.0
20.4
20.0
20.0
20.0
20.0
19-3
2C.O
20. 0
20.0
20.0
io.o
50.3
20.0
(Statlo)
bid 1*
16.8
19-3
IB. 7
19-7
20.O
20.0
20.0
20.0
19-5
19-3
19-3
19.6
20.0
20.0
20.0
20.3
ao.o
Average Bed
•p. (Abeorptlon) T
870
840
650
850
850
846
850
840
850
855
865
850
865
855
855
855
855
1
Max. Bed
••p. (tegeneretl
970
1000
1020
1O2O
1030
1020
1030
1030
1030
970
1010
1030
1015
102O
1030
1030
1C30
Duration of Dl«trl- Mike-up
on) Absorption butor iwl|ht
(llecetone)
(f/oyele)
55 Horaal
• ' 2000
• -, " looo
• It M
• • N
» m m
• • «
• • •
" * •
•
" • •
"
" •
• .
•
» • M
...
1
01
T
-------
Cy«l.
*>?•
•78/fc
1&/4A 10
183/4* 11
I84/4A IJ
185/4* 13
186/4* 14
fu.1
Llm-
stom
Pu*l taU
(•/•in)
U8
133
143
153
*iy
Condition! (Cycl« Tt»t 13)
TaUl Air Superficial
IUU Ou V.I.
(ft/we)
437
4H2
»3?
"39
•»5
42?
449
442
435
418
436
*?3
4«
»38
3-9
4.0
3.9
3.9
4.0
M
4.1
4.0
3-9
M
3-9
3-9
3.8
4.0
Astoleh.
Air
32.1
35-3
29.2
33-2
29.0
20.5
Bed Depth (3UU«)
(Inolwt)
Start End
CO
CO
Ounitln o( Dutrl-
ttiorftlon bulor «M|
(•!•) (ll«it«m)
16.0
16.2
MH
16.0
^f*w
»5.«
>5-5
15.1
»5.»
l»-7
16.0
16.0
16.0
16.0
16.0
16.2
16.0
1^.0
16.0
•""S
»5-3
15-2
U.I
15.1
>»-7
»5.3
16.0
16.0
»5-7
>«-9
835
*?
*o
BM
^rr
»«
8»
aw
«*?
845
860
850
B55
*9
^9
>?p
1010
*£
Inkn
T
-------
Run No.
TEST 1
Cycle % Fuel Sulphur Removed *
15 min 30 min 45 mln
52/4B
53/4B
5*/*B
55/*B
56/*B
57/4B
58/4B
59/4B
33/*A
3*/*A
35/*A
36/4A
37/4A
38/4A
60/4B
61/4B
62/4B
63/*B
64/4B
-
-
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
-
-
100
86.6
70.7
78.1
74.8
69.3
53-6
63.4
-
65.7
59.2
-
62.6
63.6
80.0
69.5
76.6
-
68.1
-
-
97.5
76.9
68.9
74.7
75-5
70.1
59-*
71.4
-
67.1
64.4
60.0
62.4
68/
77-0
69-7
76.3
76.1
68.1
-
-
97.8
76.0
72.9
75.7
72.6
65.9
59.6
77.5
63.1
58.8
56.5
62.5
60.2
-
70.7
77-6
74.4
76.1
68.6
-
-
60 min
_
-
-
-
-
-
-
-
-
-
-
-
-
-
69.2
56.1
71.3
75-6
68.6
-
-
Run No.
^^••^•^•^•B
44/*A
45/4A
46/4A
47/4A
48/*A
49/4A
50/4A
51/*A
52/4A
53/*A
54/4A
66/4A
56/4A
57/4A
70/4B
71/4B
72/4B
73/*B
74/4B
75/4B
76/4B
Cycle % Fuel Sulnhur Removed *
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
15 min 30 mln 45 mln
100
100
85.6
_
.
75-3
67.6
.
-
75.6
70.4
69.4
59-9
70.0
80.6
67.2
76.0
80.0
80.9
79-0
82.2
100
100
85.6
66.9
80.2
73-2
68.2
-
-
61.6
65.2
65.0
60.7
67.1
73.6
68.8
74.5
79.2
76.9
82.4
82.0
100
100
84.1
67.8
66.4
70.4
69.0
-
-
59.6
63-9
62.1
56.7
67.1
63.0
66.6
67.1
76.4
74.3
78.9
72.8
60 min
_
_
-
_
-
-
-
-
-
-
-
-
-
-
59-2
67.6
65.6
72.8
72.8
78.2
70.3
Run No.
______
77/4B
78/4B
79/*B
80/4B
81/4B
82/4B
83/4B
84/4B
85/4B
86/4B
58/4A
59/*A
60/4A
61/4A
62/4A
63/4A
64/4A
65/*A
66/6A
67/4A
-
TEST } (continued)
Cycle
_____
22
23
24
25
26
27
28
29
30
31
32
33
3*
35
36
37
38
39
40
41
.
% Fuel Sulphur Removed
15 mln
79-*
82.8
58.8
65.2
65.2
57.3
68.2
72.5
65.6
65.6
76.5
51-9
70.7
67-9
63.8
75.0
73-8
73-5
83.4
77.0
-
30 min
79.*
79-*
60.2
65.0
61.0
61.5
65.4
-
68.6
68.6
69-5
62.1
56.1
55-*
55-9
76.6
72.2
69.0
7*.*
70.7
.
45 Bin 00 mln
78.8 75.*
75-3 7*.8
60.5 -
59-0 -
58,2 -
62.0
58.8 -
71-5 -
65.6 -
68.2 -
53.8 -
50.0 -
-
54.4 -
53.1 -
61.2 -
65.5 -
59-8 -
61.2 -
72.8 -
_ _
* Instantaneous Sulphur Removal Eff
-------
aatotanr Reaoval tfficlenolW fcrcle tots 4.5. »nd 6)
TEST 4 ^"
Ron Ho*
68/4A
69/4A
70/4A
71/4A
72/4A
7*4*
74/4A
75/4A
76/4A
77/4A
-
*
-
-
-
-
-
-
-
-
Cycle %
1
2
3
5
6
7
8
9
10
-
-
-
-
-
-
-
-
-
-
Fuel Sulnbur Rene
_ 15 ain 30 Bin 45 Bin
98.8 98.6 90.2
81.7
72.0
77.1
72.3
80.0
60.0
67.5
-
6t).3
-
-
-
-
-
-
-
-
-
-
75.8 71.0
58.6 35.4
54.7 51-5
61.9 -
76.7
63.8 -
57.8
r
$9.3
-
-
-
-
-
-
-
-
-
-
ved * Run No.
60 «lft
78/4A
79/4A
80/4A
81/4A
82/4*
83/4*
84/4A
85/4A
86/4A
87/4A
88/4A
09^4*
9t>/4*
91/4*
9V4A
93/4*
94/4A
95/4*
96/4*
97/4A
Cycle
1
2
3
4
5
6
7
8
9
10
11
12
13
14
IS
16
17
18
19
20
Iftili
15 nip 30 Ml
98.9 98.3
85.1 .
78.6
80.8
84.2
89-5
86.2
-
90.6
-
71.9
85.7
84.8
90.8
-
99.5
95.7
97.8
8S.o
73.0
75-2
73.0
77^8
80.7
84.0
86.2
89-3
86.8
82.9
58.7
85.6
85.3
93-0
98.4
99.0
98.2
92-6
76.1
)huf Honored * -
Run No.
Cycle
J45_jgn 60 nto ~
91.2 - 98/4A 1
60.8
69.5
67.0
74.4
77.2
80.6
83.8
79.0
81.9
66.0
61.0
80.2
76.9
86.7
95.9
».0
-
87.6
74.9
99/4A
100/4*
101/4*
102/4*
103/4A
104/4*
105/4*
106/4*
107/4A
89/te
90/4B
9V4B
«/4B
83/4B
94/4B
95/48
-
-
_
2
3
4
5
6
7
8
9
10
11
a
£3
14
15
16
17
.
>
_
TBS; 6
j
15 n1
Smm •
99.3
98.2
98.4
98.8
98.1
98.1
94.9
96.8
-
96.1
99.0
100
99.0
92.0
-
63.4
-
.
.
_
Fuel Sulpbur Renoved »
JR 30 «ta *^ ffiff
98.2 98.0
98.2
95.9
96.4
96.4
98.7
89.7
95.2
94.9
92.5
76.0
96.7
90.9
89.4
91-5
89.6
-
_
-
—
98.1
93-2
94.1
96.3
96.6
88.9
93.8
91.0
93.1
80.6
95.8
85-3
90.0
95-0
' 85.6
-
_
-
—
6O Bin
-
-
-
-
-
-
-
V
-
-
-
-
-
-
-
-
.
_
«• Instantaneous Sulphur Itomoval Efficiency
-------
TABLE -E-XV
Sulphur Removal Efficiencies (Cycle Test
Cycle
1
2
3
4
5
6
7
8
9
10
11
12
13
14
16
17
18
19
20
% Fuel
15 min
95.4
98.2
98.4
98.4
95.8
89.0
-
99-5
-
100
100
99.5
98.3
99.6
99-6
99.4
96.4
99.2
99-3
Sulphur
30 min
95-8
98.4
98.8
95-2
-
82.8
97-7
98.8
96.0
98.9
99.6
99-5
99.1
99-2
93-9
99-5
89-5
99.2
97.5
Removed
45 min
99.5
99.0
98.2
91.6
87.0
82.8
97-7
99.0
96.0
98.9
98.4
99-4
94.8
96.7
94.2
98.7
90.2
99-2
98.0
*
Cycle
iBB.HBB«BBB._
21
22
23
24
25
26
27
28
29
30
31
32
33
34
36
37
38
39
40
% Fuel
15 min
99.1
100
99.5
-
-
-
61.3
78.7
-
85.9
-
84.2
80.7
73.7
83.7
79-0
91.0
94.7
94.0
Sulphur
30 min
99-1
97-4
97-0
_
_
-
78.4
80.4
76.9
74.7
86.9
84.7
74.2
74.5
78.6
82.2
88.1
92.6
-
Removed
45 min
97.1
97-2
_
-
-
76.4
78.1
71.9
72.3
78.9
82.3
70.0
74.5
74.8
83.7
85.5
91-7
90.1
*
Cycle
41
42
43
44
45
46
47
48
49
50
51
52
53
54.
56
57
58
59
60
% Fuel
15 min
86.0
73-3
73-6
98.5
88.4
80.3
78.5
72.2
81.1
T8.-7
85.1
87.6
80.5
75-3
71.0
75-7
72.7
74.6
84.5
Sulphur
30 min
«^— B^-^^—
78.0
75.6
74.7
93-8
79-0
78.0
77.1
74.6
79-3
77-7
79.7
80.9
79.9
74.0
72.8
74.5
76.6
76.2
83.2
8)
Removed
45 min
78.0
73-1
67.1
87.6
79.0
78.0
74.9
72.0
75-7
75.1
78.2
78.1
83.2
7.6.9
74.3
74.6
69.8
76.3
80.0
*
Cycle
61
62
63
64
65
66
67
68
69
70
71
72
73
74
76
77
78
79
80
% Fuel
*
15 min
85-5
99.4
70.6
61.9
79.1
57.9
99.0
67.7
-
-
-
100
78.8
63.5
80.9
69.6
62.1
75.4
77.0
Salphur
30 min
94.8
96.0
70.8
75-5
76.3
47.9
99.0
64.7
-
-
-
78.5
78.4
69.9
77.2
59.6
59-0
73-3
68.0
Removed
45 min
98.8
73-4
68.7
56.0
79-1
62.4
-
-
-
-
-
-
-
-
-
-
-
-
-
* Instantaneous Sulphur Removal Ef'fic-v
-------
TAEU5
Sulphur Removal Efficiencies (Cycle Tests 9, 10, 11 and 12)
TEST 11 i TEST 12
Qycie.
" 15 rain
1
2
3
4
5
6
7
8
9
10
11
13
13
14
15
-
-
100"
100
-
?7.2
87.2
84.3
85.3
82.1
91.6
89.1
90.0
89.5
92.4
-
-
-
-
Sulphur Removed
30 min 45 min
100
100
-
87.9
85.4
79-9
-
81.6
85.9
84.3 -
86.4
86.8 . -
88.8
-
-
-
-
*
Cyc]
1
2
3
4
5
6
7
8
9
10
11
12
. 13
-
-
-
-
"'"% Fuel
— 15 min
98.9
-
98.3
98.6
99-5
99.1
98.2
99-0
98.4
88.6
-
94.3
92-9
-
-
-
-
Sulphur
30 min
97.8
99-4
98.7
99-7
99-7
94.5
99-0
99.. 8
93-5
92.9
-
89.7
90.9
-
-
-
-
Removed *
45 min
100
98.5
98.6
Sft-7
94.7
93-4
98.2
99-5
84.9
86.5
-
89.a
88.9
-
-
-
-
Cycle
1
2
3
4
5
6
7
8
9
10
11
12
13
Vk
15
16
17
% Fuel
15 min
99.2
100
93-5
88.8
-
-
-
-
-
89.7
85.8
81.8
88.7
87.8
79-3
82.7
81.4
Sulphur
30 min
99-2
99-8
87.8
90.5
88.4
-
-
-
-
91.4
82.4
79-3
-
81.6
79-3
77.0
77-3
Removed *
45 min
99-3
94.0
82.8
91.1
88.9
-
-
-
-
83.0
75-7
71-7
81.7
76.7
76.0
-
72.6
Cycle
1
2
3
4
5
6
7
8
9
10
11
.12
13
14
-
-
-
% Fuel
15 min
98.1
98.4
100
95-6
97.4
98.6
92.7
87-7
86.6
-
87.5
-
91.3
86.0
- •
-
-
Sulphur
30 min
99-6
98.4
96.8
93-1
99-3
-
94.7
91.7
89-7
89.1
-
87.8
-
87.4
-
-
_
Removed
45 min
97-2
98.0
-
93-1
97.2
95-9
88.7
90.6
87-7
89.5
86.3
89.7
89-3
86.5
-
-
_
-------
• 157 -
Tilt 1
IMtk
Tilt 5
01.
1
2
3
k
5
6
7
8
9
.0
LI
12
13
Lk
15
.6
.7
8
9
Nu. SO.
(« tv vU)
k.3
3.2
C.O
6.7
7.5
8.9
7.5
-
-
k.O
8.k
8.k
8.0
8.k
k.9
7.k
6.9
8.7
8.3
Nu.Vwi
i rci
1000
965
1020
1030
lOkO
1030
1050
1030
950
1020
1030
lOkO
lOkO
1080
1000
1035
1020
1060
lOkO
p. Cr*l« Nu. 80. Nu. tap. Cyol» Nu. 80. Nu. *Mp.
Mtvvof) CO M*w« CO
1
2
3
k
5
6
7
8
9
10
* V
^
. 12
"0
Ik
15
16
17
18
19
20
81
22
23
2k
25
5.0
6.2
9.5
9.0
9-5
10.3
8.6
k.7
7-2
8.2
8.6
7.k
7.k
7.2
3.1
7.2
8.k
8.6
B.k
6.8
7.k
7.k
8.0
-
6.9
980
1010
1020
10|)
lOkO
1030
1030
955
1020
1030
1000
1020
10)0
1030
1000
1010
loko
1030
1030
lOkO
lOkO
1080
lOkO
-
960
26
27
28
29
30
31
32
33
>
35
36
37
38
39 '
ko
kl
9.0
9.k
7.8
6.k
9.6
8.6
6.0
6.8
3.5
7.3
9.0
8.6
8.6
8.0
7.8
8.2
990
lOko
1000
1010
1100
-•1070
1020
1030
970
lOkO
lOkO
1035
1030
lOkO
lOkO
1030
Oral* Nu SO. Nu. *Mp. Cyol*
« t* vfi) re)
1 k.) 1010 1
2 9.k 1030 2
3 9-0 1060 J
k 8.6 lOkO k
5 8.2 1050 5
6 8.6 lOkO 6
7 8.0 lOkO 7
8 7.8 1020 8
9 7.9 1035 9
10 7.2 lOkO 10
11
12
13
Ik
15
1C
If
18
19
20
Hu. SO. Max. Tcmi
« ta vol) CO
6.9
5.k
8.2
6.7
7.k
8.2
7.0
7.k
7.k
7.0
3.8
7-2
7.k
6.2
6.6
7.0
7.2
7.U
7.k
10OO
10 'C
10JO
loro
KXX>
10JO
10£0
1030
101 ">
1010
100C
1010
101C
1010
1CSO
101O
10ZO
lo;,,
10J"
-------
IM-
T»«t 6
*nt 8
Cycle
1
2
>
*
5
6
7
8
•>
10
11
u
I)
Ik
10
K.
17
Nu. SO-
M .!»- voi)
3.1
6.0
6.6
6.0
6.6
6.6
6.8
C.8
6.6
i.O
.5.5
6.8
7.0
fi.7
t,.3
D.6
8.6
""pcT*'
960
10W
1090
1OW
1020
1030
IdkO
lOkO
10*6
10W
988
1015
10W>
1030
1025
M*5
1035
Cral*.
1
2
3
k
5
6
7
8
9
10
11
12
13
Ik
15
16
17
18
19
80
21
88
83
2k
85
86
87
m*. so,
(f IV TOU
k.2
7.5
8.9
8.k
7.9
8.7
7.5
6.8
6.8
7.0
3.8
8.3
6.0
6.«
7.0
7.0
7.1
6.«
8.3
7.9
6.8
».3
5.7
5.7
6.6
6.8
6.5
**.M*
99)
98»
1015s
MM
1030
1095
1010
1000
1000
lOtt
10U
10W
low
1010
10*
10M
10t»
1MB
1090
1029
1020
101*
1020
1085
1096
1000
1005
CNU
28
89
30
34
9>
39
J*
35
36
97
98
99
kO
kl
k2
*9
kk
k5
k6
k7
k8
*9
90
51
5»
59
5k
m*. to,
Bt^\3
T.5
7.t
7.8
8.9
9.9
-
6.k
7.9
8.5
8.9
7.3
7.7
7.5
-
6.1
6.8
9.9
9-5
6.k
5.9
6.3
-
6.5
6.8
8.%
9-9
6.0
rar
10kO
10k»
MkO
«•
MS
1090
1019
1090
1080
988
1090
109S
lOkO
-
1000
1009
1010'
1090
1090
1090
1030
low
lOkO
W90
96*
«8»
M0»
^.
59
56
97
9*
59
60
61
68
63
6k
65
66
67
68
69
70
71
78
79
7»
75
76
77
78
79
8D
(?ff^l>.
5.»
-
6.9
6.3
6.0
"t.7
5.7
9.$
5.0
7,«
-
7.8
5>8
6.6
7,»
6.k
7*
a,«
5vO
7.5
6.5
7.5
7.5
8.9
7.9
8.*
i "^r"
1010
10W
1010
10W
1010
1020
10W
970
1000
10M
1035
1035
1090
1090
lOkO
1039
1030
1005
1030
1060
1060
106»
lOtO
1070
U60
M65
C)r«te. Hu. M.
1 1.8
8 7.8
3 8.2
» 7.8
5 7.2
6 7.1
7 6.k
8 8.2
9 0.2
10 5.6
11 6.C
12 6.*
1} 6.0
Ik 6.li
15 (..5
"^
in
WJ
IN
UN
m
1M
M
UM
UK
w
IW
1M
10*
M
m
-------
-169-
TAHt£ E XDC
Cycle Teat Regenerations
Teat 10
Teat 11
Teat 12
Cyole Max. S00 Max. Teap. Cyole
« tar vo?) PC)
Max. 30_ Nax.TeBp.
br TO!) PC)
Cycle Max. S00 Max. Temp
(* tg vol) PC)
1
2
3
u
5
6
7
8
9
10
11
12
13
3.5
U.2
5.0
5.5
7.5
6.6
6.8
U.6
5.8
6.2
8.7
5.8
6.6
1000
1020
1020
1025
1030
1030
10HO
1000
1030
1030
1050
1030
10UO
1
2
3
k
5
6
7
8
9
10
11
12
13
Ik
15
16
17
1.7
k.5
6.6
7.0
7.0
7.0
6.8
7.8
6.6
3.1
k.B
6.6
6.6
6.2
5.8
6.6
6.6
970
1000
1020
1020
1030
1020
1030
1030
1030
970
1010
1010
1015
1020
1030
1030
1030
1
2
3
k
5
6
7
8
9
10
11
12
13
Ik
3-7
-
e.k
9.2
7.6
7.8
6.2
8.6
9.2
8.6
7.7
7.8
8.1*
7.8
990
1010
1025
1<*0
1030
1030
1000
1030
1030
1010
1030
1035
10UO
10UO
-------
- 160 -
TABLE E XX
Metals Picfr-UD (BCR 169J1.(Cycle Test)
:ycle
1
2
3
4
5
6
7
8
9
10
11
12
13
14
Vanadium*
ppm
860
1250
1460
2040
1920
1620
1600
196OV
2600
2400'
2800
3100
2300
3400
Sodiunrt*
ppm
235
26O
300
335
295
2?0
295
300
271
270
265
255
300
29&
Iron*
ppm
7000
6900
6800
6900:
690O
6900-
6700-
6800-
7200D
6900-
6900
6900.
6900
6900
* After regeneration samples.
-------
- 161 -
Table E XXI
Metals Pick-up (BCR 1691) (Cycle Test 3)
Cycle Vanadium* Sodium* Iron* Cycle Vanadium* Sodium* Iron*
ppm ppm ppm ppm ppm ppm
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
460
920
960
1030
1100
1270
1920
1920
1980
2050
2100
2350
2660
2940
2900
1990
2910
2120
2720
2830
2830
245
235
270
220
220
205
235
215
210
255
3*5
215
235
255
300
340
330
295
375
330
300
^•^•^•^MI
7200
7300
7100
7200
7000
7200
6800
6700
6700
6600
6700
6600
6900
7300
6700
6700
6800
6700
6800
6700
6700
22
23
24
25
24
27
28
29
30
31
32
33
34
35
36
37
38
39
40
41
2580
2980
3750
3000
3600
3700
3800
4200
3600
3600
3900
2400
3300
3500
3900
4300
4900
4400
4300
4100
270
270
240
260
290
290
310
280
285
280
275
240
275
290
290
315
315
435
370
430
6700
6500
6300
6400;
6600
6500
6700
6700
6700
6500
6500
6300
6800
6500
6300
6500
6600
6600
6600
6500
* After regeneration samples.
-------
- 162 -
Table E XXII
»la Pick-up
Vanadium*'
PDTO
550
480
550
690
770
800
850
950
1000
840
(BCR1691) (C:
Sodium*
PP»
265
295
295
500
505
295
300
290
285
270
Cycle Vanadium* Sodium* Iron*
ppm
1 550 265 6600
2 480 295 7100
5 550 295 6900
4 690 500 6800
5 770 505 7000
6 800 295 6900
7 850 500 6800
8 950 290 6900
9 1000 285 6700
10 840 270 6400
* After regeneration samples.
-------
- 16?-
Table E XXIII
Cycle Vanadium* Sodium* Iron*
ppm ppm ppm
1 4?0 250 7000
2 450 280 7000
3 740 285 7000
4 1100 325 6900
5 1040 295 7000
6 1090 290 6800
7 1350 265 6900
8 1400 300 7000
9 1070 325 7000
10 1340 300 6700
11 1280 290 6500
12 1250 290 6700
13 1340 310 6700
14 1740 255 6700
15 1640 290 6800
16 1500 250 6700
17 1380 255 6500
18 1500 250 6700
19 1840 235 6300
20 1650 190 6700
* After regeneration samples
-------
- 164 -
TABLE E XXIV
Metals Pick-up (BCR 1691) (Cycle Test 6J
'ycle
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
Vanadium*
ppm
530
790
840
1060
1160
1170
1220
1666
1720
1940
2356
3300
5000
2555
2555
3200
3200
Sodium**
ppm
223
246
241
266
238
250
271
259
220
267
263
276
297
262
284
289
283
Iron*
ppm
6600
6500
6500
6300.
6500
6200
6200
6200
6$00
6400
6900
6600
6900
6900
7200
6600
6400
* After regeneration samples.
-------
Metals Pick-up (BCR 1691) (Cycle Test 8)
Cycle Vanadium* Sodium* Iron* Cycle Vanadium* Sodium* Iron* Cycle Vanadium* Sodium* Iron* Cycle Vanadium* Sodium* Iron
1
2
3
4
5
6
1
8
9
10
11
12
13
14
15
16
17
18
19
20
ppm
230
560
6.30
1080
1000
1390
1555
1510
1780
2180
1630
-
-
-
-
-
2440
-
550
318
320
334
311
302
308
305
300
295
325
-
-
-
-
-
313
-
ppm
6500
6800
6800
7100
6500
7200
6600
6500
6200
6200
6300
-
-
-
-
-
6000
-
21
22
23
24
25
26
27
28
29
30
31
32
33
34
35
36
37
38
39
40
ppm
2020
2040
2750
3120
3150
2690
3600
2250
2800
4190
2400
2900
3500
4200
4400
2900
3400
3800
3900
5700
ppm
297
312
327
310
313
-
316
302
303
296
273
260
28?
270
317
266
279
333
301
335
ppm
5900
5500
5700
5900
6300
-
6300
6100
6300
6200
6100
5900
6300
6200
6500
6100
6100
6200
6100
6100
41
42
43
44
45
46
47
48
49
50
51
52
53
5*
55
56
57
58
59
60
ppm
4900
4700
3900
3900
3900
4100
3900
4800
4200
4100
4000
4280
4680
3570
3500
3440
3500
4690
3900
3980
ppm
350
298
268
296
317
314
330
313
330
316
300
270
259
277
284
285
269
283
274
284
ppm
6100
6000
5400
6200
5700
6100
6000
5900
6200
6000
5900
6000
6200
6100
6300
6500
6000
6000
6100
6400
61
62
63
64
65
66
67
68
69
70
71
72
73
74
75
76
77
78
79
80
ppm
3880
3900
3700
3400
3900
4JOO
3000
4400
3400
4000
4300
4680
3900
4930
3700
4690
5100
4580
4160
3780
PP"
254
326
286
302
305
297
320
309
298
288
333
279
284
271
274
279
314
313
302
308
ppm
6100
6100
6300
6200
6300
6400
6000
6100
6100
6000
i
6700 M
ON
6500^
6200
6300
7000
6800
6500
6700
6400
6500
* After regeneration samples
-------
- 166 -
Table E XXVI
(BCR 1691) (Cycle Te«MQ)
1
2
3
4 •
5
6
7
8
9*
10*
11'
12
13
Vanadium*
pom
300
370
920
820
1070
1100
740
1600
1500
1400
2300
1900
2500
Sodium*
PPB
30V
324-
343-
299<
270
268
268
244
234
234
259
244
238
Iron*
ppB,
6400
6500-
TlUUk.
f ^W- -
6700-
6600
6300
600O
6200
6000
6200
6200
6100
6300
* After regeneration aaopl»»-
-------
- 16? -
Cycle
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
Table E XXVII
tals Pick-up
Vanadium*
ppm
310
600
620
920
1160
1420
1180
1470
2480
2500
2060
2260
2340
2140
2860
3120
3060
(BCR 1691)
Sodium*
ppm
301
363
303
334
279
304
273
280
292
284
256
287
253
247
255
256
264
(Cycle Te
Iron*
ppm
6500
6600
6400
6800
6600
6400
6000
6000
6200
6000
6000
5900
5800
5500
5700
5600
5600
* After regeneration samples
-------
- 168 -
Table E XXVIII
Bed Losses Through Cyclones (Cycle Tests)
Teat Cycles Losses Make-up/Lime
Total Rate wt % on. Total Rate
(g) (g/min) fuel (g) (g/mln)
1 1 110} 21.2 9.5 588 11.3
2-8 1081 }.0 1.5 4536 12.5
9-14 565 1.9 0.8 3780 12. 4
15-19 18»; l*.9 5.0 3780 10.1
3 1-7 2091 5.7 2.7 4536 12.4
8-ll* 1361 5.7 1.9 4536 i2\>
15-2} 2338 3.7 2.8 6048 9.6
24-3} 1145 2.8 1.} 529* 13.0
34-41 3478 6.7 3.3 9072. 17.5
4 1-10 3119 6.6 3.3 6804 14.4
5 1-10 3075 6.0 3-1 13600 26.5
11-20 5299 10.3 5-4 10578 20.5
6 1-10 42Bi 8.3 4.1 8085 15-7
11-17 2531 7.0 3-3 453> 12.6
8 1-10 5940 11.4 5.£ 6800 13-1
11-20 3348 6.4 3.3 6800 13>0
21-30 2176 4,2 1.9 5100 9.8
31-40 2900 5-5 2.7 62}^ 11.8
41-43 1539 9-8 5.0 151JL 9.6
44-51 3637 8.7 4.9 5289 12.6
52-61 3810 7.3 4.0 6800
72-83 2324; 5-3 2>5 6897
9 1-15 3727 6.7 3.8 7052. 12.8
10 1-7 4381 7.8- 4.9, 5666 10.1
8-13 1903 3.9 2.3 4408 9.0
11 1-9 872 1.8 1.8 2647 5.5
10-17 1536 2.6 1.9 3706 6.}
-------
- 169 -
APPENDIX P
CORRELATION FUNCTIONS FOR SULPHUR REMOVAL
EFFICIENCY
In Section II-C-2-h of the text, two equations, k and 6, were
presented to express sulphur removal efficiency in cyclic batch
gasification tests. The basis for these equations is presented here.
As a starting point, a first order reaction in sulphur
concentration was used for correlating sulphur removal efficiency in
fresh lime bed tests.
-ff = kS (p-l)
where S is sulphur concentration remaining in the gas, t is
gas residence time/seconds, in bed above fuel injector, and k is a first
order rate constant, sec ~1>
Solution of the first order rate equation gives:
d ln S = - k dt (F-2)
S/S6 = 1 = 1
(kt> (kt) (F-3)
10 2'^
where So is the sulphur concentration that would exist in the
gas if no reaction occurred. The equation can be rearranged to
express the efficiency of sulphur removal (SHE) where,
SHE = 100 (So - S ) = 100 ( 1 - | ) ; (F-?)
So
thus, SHE = 100
1-
t kt
(
10
(F-5)
In correlating fresh bed test data, better results were obtained
by modifying the equation to use a pseudo residence time 0, rather than the
superficial residence time t. The value 6 was defined empirically
0 = (t - .051), seconds (P-6)
-------
- 170 -
A possible reason for the improvement given by using 9 rather
than t is that a finite time is required for cracking the fuel oil
and liberating the sulphur in a form in which it can react with the
lime.
With this empirical modification, equation F-5 becomes,
SRE =
100
1 -
(P-7)
In order to apply equation P-7 to the cyclic test situation where
the lijne is partly deactivated by prior cycle*, the following assumptions
were made:
o Rate constant, proportional to lime replacement rate, m.
o Rate constant Inversely proportional to quantity of sulphur
'to which line is exposed in each cycle, d, the projected
sulphur differential.
Thus K_ = k_
2.305 2.303
m
k a m
d
(F-8)
where K is the effective rate constant for the cycled lime, k is
the rate constant measured in a batch test with fresh lime, nu is the
equivalent lime replacement rate for a fresh bed test, ^ is €he effective
potential sulphur differential in a fresh bed teat, and
a =
2.303
The first order fresh bed constants were measured at the time
at which the bed had reached a sulphur content xjf 5 wt.#. Thus the
effective replacement rate-was 3100/5 = 20 and the sulphur differential
was 5. The constant, a, thus becomes 5_ = 0,109,
2.5Q5,x 20
and equation P-7 becomes,
SRE =
100
1 -
(,109 k-Q ;m
m
d
(P-9)
Linear regression df the fresh bed*test;results-with U.S. limestone
BCR 1691 yielded 33.9 sec "-1 as the best value i for *the rate constant, k.
Figure 35 of the text compared experimental cydle -best data *»ith values
-------
given by equation (F-9)» text equation (7), and showed that in
general the experimental results were slightly more favourable than the
equation predicted.
To obtain a better representation of the actual cycle test data,
they were regressed with an empirical equation that allowed assignment
of different weights to the variables involved. The general form of
equation adopted was,
SRE = 100
1 -
1
Exp (
wtx my
(P-10)
where w, x, y and z are empirical constants.
The linearised form of this equation on which a computerised
regression was performed is,
log
'-log (1- SREJ
100-
log w + x log t + ylog m - z log d
(F-ll)
where logarithims are to base «.
The cyclic test data matched this equation with a correlation
coefficient of 0.9^, and yielded the following values for constants.
w = 9.285
x = 0.6
y = 0.95
z = 1.9
The data therefore show that sulphur differential with an exponent
of 1.9 has a much larger effect on sulphur removal than gas residence
time with an exponent of 0.6
The final empirical equation is;
SRE =
100
1-
Exp
9.258 t
.6
m
.95
1.9
(P-12)
Figure F-l shows the match between experimental and computed
efficiencies using equation F-12 which corresponds with text equation 6.
Values predicted using equation F-12 are listed in Tables F-I to III.
-------
- 172 -
Table F I
Prediction of Lined Out. Efficiencies from Cycle Teat Results
Residence Make-Up Potential Sulphur Lined Out
Time Rate Differential Efficiency
(seej (Wt.CaO/wt.S) (nfr.*)
0.1 1.8 2.0 66.707
0.2 1.8 " 81.026
0.5 1.8 " 87.951
0.4 i.fl- " 91.888
0.5 1.8 " 94.324
0.6 1.8- ** " 95-916
0.1 0.& " 39.950
0.2 0.8 " 53.708
0.3 0.8 " 62.492
0..4 0.8 " 68.775
0.5 0.8 " 73-537
0.6 0.8 " 77.280
0.1 1.3 " " 55.414
0.2 1.3 " 70.497
0.3 1.3 " 78.864
0.4 1.3 " 84.193
0.5 1.3 " 87.839
0.6 1.3 " 90.045
o.l 2.3 " 75-089
0.2 2.3 " 87.719
0-3 2.3 " 95.076
0.4 2.3 " 95.797
0.5 2.3 " 97.521
0.6 2.3 " 98,232
-------
- 173 -
Table F II
Prediction of Lined Out Efficiencies from Cycle Test Results
Residence Make-Up Potential Sulphur Lined Out
Time Rate Differential Efficiency
(sec) (Wt.CaO/wt.S) (wt.#) %
0.1 1.8 1.5 85.065
0.2 1.8 1.5 94.350
0.3 1.8 1.5 97.423
0.4 1.8 1.5 98.700
0.5 1.8 1.5 99.299
0.6 1.8 1.5 99.603
0.1 0.8 1.5 58.568
0.2 0.8 1.5 73.594
0.3 0.8 1.5 81.647
0.4 0.8 1.5 86.632
0.5 0.80 1.5 89.958
0.6 0.80 1.5 92.286
0.1 1.3 1.5 75.254
0.2 1.3 1.5 87.881
0.3 1.3 1.5 93.191
0.4 1.3 1.5 95.880
0.5 1-3 1-5 97.382
0.6 1.3 1.5 98.276
0.1 2.3 1.5 90.921
0.2 2.3 1.5 97-337
0.3 2.3 1-5 99-011
0.4 2.3 1.5 99.583
0.5 2.3 0.5 99.809
0.6 2.3 1.5 99.907
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- 174 -
Table F III
Prediction of Lined Out Efficiencies from Cycle Test Results
Residence Make-Up Potential Sulphur Lined Out
Time Rate Differential Efficiency
(sec) (Wt.CaO/wt.S) (wt.#) %
0.1 1.8 3.0 39.855
0.2 1.8 3-0 53.621
0.3 1.8 3-0 62.403
0.4 1.8 3.0 68.686
0.5 1.8 3-0 73.451
0.6 1.8 3.0 77.198
0.1 0.8 3.0 20.990
0.2 0.8 3.0 29.955
0.3 0.8 3.0 36.448
0.4 0.8 3.0 41.611
0.5 0.8 3.0 45.911
0.6 0.8 3.0 49.593
0.1 1.3 3.0 31.160
0.2 1.3 3.0 43.123
0.3 1-3 3-0 51.249
0.4 1.3 3-0 57-376
0.5 1.3 3-0 62.242
0.6 1.3 3,0 66.234
0.1 2.3 3.0 47.350
0.2 2.3 3-0 62.070
0.3 2.3 3-0 70.896
0.4 2.3 3-0 76.893
0-5 2.3 3-0 81.238
0.6 2.3 2.0 84.515
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- 175 -
CORRELATION BETWEEN EXPERIMENTAL AND
CALCULATED LINED OUT EFFICIENCIES(USING EQUATION FI2)
60 70 80
CALCULATED SRE ( %)
90
100
FIG.FI.
-------
- 176 -
APPENDIX G
Design Basis for the Pilot Plant Gasifier System
A. The Gasifier
The size of the experimental gasifier was determined by the decision
to adopt dimensions for the regenerator virtually identical with those of
the batch unit. This decision was taken on the grounds that this was a
proven design which had operated under regeneration conditions in a
satisfactory manner at the gas velocities which were envisaged.
Since experience had also shown that the S02 concentration in the
regenerator tail gas would be in the region of 10# by volume, the sulphur
throughput of the unit, under conditions which had actually been used,
was limited to 8.6 Ibs/hr. In the case of a 2.2% by wt. sulphur fuel
the fuel throughput was therefore restricted to 591 Ibs/hr. giving an
output of 7.1 MM Btu. gross or 6.68 MM Btu. net, per hour. At 20# of
stoiohiometric air and a gas velocity of 4 ft/sec, at 900°C this indicated
a cross-sectional area for'the gasifier of 4.8 sq. ft.
At an early stage in the design of the installation it was decided
to use the gasifier to fire a conventional packaged hot water boiler.
It was considered advantageous to show that the gasifier could be used
in conjunction with commercially available equipment ^nd the installation
lent itself readily to the"employment of flue gas recycle as a means for
controlling the temperature of the gasifier bed. The boiler also provided
a suitable test chamber for the development of the gas burner. This was
a necessary requirement since no burner handling this type of gas is In
current commercial use.
A schematic -layout of the completed plant is shown in
text Figure 9. Th® pressurised hot water boiler is
provided with a heat exchanger which transfers the heat generated to
an atmospheric pressure evaporative cooler. The-heat transfer system is
completely automatic and' maintains a constant boiler water temperature
regardless of variations in the heat input. As can be seen it was judged
advantageous to set the gasifier in a 6 ft'deep pit in order to simplify
the ducting between it and the burner.
The actual geometric configuration of the gasifier was very largely
influenced by the decision to build it as a monolithic structure in
refractory concrete. A major advantage of this method of construction
is that it allows heat to be conserved in the transfer system between
the two fluidised beds. It also allows the beds to be positioned very
-------
- 177 -
close to each other, so reducing the overall size of the unit. Dis-
advantages arise from the fact that the unit is very difficult to modify
once the concrete has been poured. In addition considerable reliance
is placed on the durability of the concrete, since the development of
cracks in the walls between the gasifying and regenerating beds could
adversely affect performance.
This precise form of construction is not recommended for general
use but was considered appropriate to the case of a small experimental
unit of limited life. Every effort was made to obtain the best advice
available in choosing the refractory material and in developing a
detailed design for the refractory block.
The geometrical configuration of the gasifier was determined largely
by the need to accommodate cyclones for the desulphurised gas stream.
It was found that the choice of a single cyclone resulted in an awkward
layout and increased the overall height, and therefore cost, of the
unit. For this reason two cyclones were used in parallel to give the
layout shown in text Fig.12 .Since the cyclones consisted merely of holes
cast in the refractory block it cost very little more to provide two
small cyclones in place of one large one, and the extra expenditure on
cores was more than offset by the saving due to the reduction in the
overall size of the gasifier.
The cyclones were positioned one on each side of the regenerator.
This allowed the use of a rectangular gasifier bed and resulted in a
compact design. Another advantage of this layout was that it resulted
in a simple ducting system for the transfer of bed material. Once the
cyclones had been positioned either side of the regenerator it seemed
reasonable to make use of the bed transfer system in order to drain
them of fines. In the arrangement which was adopted half of the fines
from the gasifier are returned to the gasifier bed, the remainder being
fed to the regenerator. The regenerator cyclone, which is an external
fitting, is used to drain fines from the unit.
Another question which had to be considered was the manner in
which the gasifier should be brought to its working temperature. The
small batch units are initially fired with LPG beneath their distributors.
This procedure results in an excessive pressure drop across the distri-
butor under starting conditions, and also imposes thermal stresses on
its structure. For these reasons it was decided to position the starting
burner above, the distributor of the gasifier in such a way that it acted
as a submerged combustor in the fluidised bed. Since with this arrange-
ment the underside of the distributor was always cold the distributor
could be constructed of steel.
In order to facilitate any necessary repairs or alterations to
the distributor it was designed for easy removal , and adequate ground
clearance was allowed beneath the gasifier for -chis purpose.
-------
Since refractories are not very good thermal insulators it was
necessary to incorporate a shell of insulating material between the mono-
lithic gasifier and its external mild steel casing. The large temperature
difference under working conditions between the refractory block and its
casing,resulting from, this arrangement,introduced a differential thermal
expansion problem. In order to allow for the "growth" of the refractory
block in the vertical direction it was found necessary to "float" the
lid of the gasifier on the refractory block instead of bolting it to the
flange of the casing. Differential expansion in the horizontal sense
was dealt with by leaving a gap between the refractory block and its
insulating shield.
The need to accommodate differential thermal expansion dictated the
designs of the offtakes from both the gasifier and the regenerator.
The simplest possibility which presented itself was to run these off-takes
through the lid of the gasifier so that they could "float" with the
refractory block and not be constrained in any way by the steel casing.
The design of the ducting from the gasifier cyclones to the burner is
shown in Fig. G-l. The main component is a Y shaped bifurcated duct
which runs horizontally and incorporates a 90" elbow for each cyclone
outlet connection. This duct is supported at three points. It is suspended
close .to each cyclone connection by a hanger attached to a frame fixed
to the gasifier casing, whilst the common duct rests on a roller which
allows horizontal movement along the axis of the boiler fire tube •
The connectiorobetween the bifurcated duct and the cyclone outlets are
spigoted and incorporate sufficient clearance to allow for the vertical
expansion of the gasifier block. Gas-tightness is ensured by the use of
thin walled steel bellows which enclose the spigoted Joints.
-------
- 179 -
SLAB INSULATION
^ REFRACTORY LINING
EXPANSION ALLOWANCE
BELLOWS SEAL
REFRACTORY LINING
SUPPORT FRAME
SUPPORT YOKES
SECTION ON A-A
CYCLONE OUTLET CONNECTIONS
GAS OUTLET PIPES
Secondary
Air
CYCLONES TO BURNER MANIFOLD
-------
- lot* -
B.
Des ign Flow Sheet
Figure G-2 is a schematic flow diagram of the gasifier-ri-egwoerator
system showing the mass flow rates of the major streams. The basic
assumptions on which this material balance is based are listed below.
Sulphur throughput
Fuel oil sulphur content
Air/fuel ratio
US limestone
% sulphur in lime* from gasifier
% sulphur in lime* from regenerator
Lime make-up rate
S02 in regenerator off gas
Carbon on gasifier lime*
* Lime taken as Ca + 0 + S in stone
8.6 Ib/hr
2.2 wt. %
20$ of stoichiometric
W>
&
1 mole CaO/mole S fed
10 mol. %
0.3 "Jt. %
MATERIAL BALANCE AND FLOWS
CONTINUOUS CAFB PILOT PLANT
(ALL FLOWS IN POUNDS/HI})
Fuel Go*
J57I
5-2 ft.2
18-in Deep b«d
4ft/MC
Supfrficial got vflpcity
4 5 ft/we
Fig G-2
-------
- 181 -
C. Principles of Gasifier Control
The results of the batch tests indicated that the gasifier would
function best at a temperature in the region of 850°C. It was also found
that the temperature of the regenerator bed should lie in the region of
1050°C. Since the regeneration reaction is exothermic and since the air
and stone entering the regenerator are heated to the operating temperature
by the heat released, it is evident that temperature control may be
obtained by varying the rate at which stone circulates between the gasifier
and the regenerator. A consideration of the chemical reactions involved
indicates that the regenerator temperature is, to some extent, self
regulating so that the system is inherently stable.
Calcium sulphide is oxidised by two reactions which may be regarded
as successive stages of oxidation. The combination of 1^ moles of ©2
with a mol of CaS results in the formation of CaSO^ which is either
oxidised by an additional 5 mol of 02 to CaS04 or else decomposes to
yield a mol of CaO and a mole of S02« Complete oxidation to CaS04
yields 220 Kcal/g mol of CaS whilst decomposition yields only 109V5
Kcal/g mol.
There is a thermodynamic limit to the concentration of S02 which
can exist in equilibrium with lime at any given temperature. Any excess
S02 will recombine to form CaSOj which in the presence of 02 will be
oxidised to CaS04« Since this temperature/ equilibrium concentration
relationship is known}a'it is possible to compute the heat released per
mole of CaS at any given temperature under conditions in which there
is no oxygen in the gas leaving the process. This has been done in
Table G-I.
Table
Heat released per mole of CaS
When oxidised in the absence of
excess oxygen
Regenerator Equilibrium Mol. Ratio Mols of Heat Released
Temp.°C Press. S02 304/30-5 02/Mol CaS Kcal/Mol CaS
Ats
970 0.053 1-62 1.81 177
980 0.064 1.20 1.77 169
99^ 0.082 0.74 1.71 156
1000 0.091 0.58 1.68 150
1004 o.097 0.48 1.66 14^
In addition to CaS the stone leaving a CAFB gasifier also carries
about 0.3# by weight of carbon. This carbon will largely be oxidised
to C02 under continuous running conditions but its effect on the heat
released will depend on the carbon/sulphur ratio. In a typical
situation the stone might contain 3.0#S and 0.j$C by weight. This
is equivalent to 0.267 moles C/mol CaS. On this basis the last
two columns of Table I may be adjusted to allow for the additional
oxygen consumption and heat release, 93.5 Kcal/mol.8.
-------
- 182 -
Op consumption and heat released per
mole of CaS, allowing for carbon
at 0. 267 moles/mol.
Regenerator
Temp.°C
970
980
994
1000
1004
Molar Ratio
Oa/CaS
2,04
2.002
1.932
1.897
1.872
Heat Released
Kcal/Mol.CaS
195
186
170
163
157
In a large unit only a very small fraction of the heat released will
be lost through the wall of the regenerator and, over the temperature range
considered, this loss will be almost constant. As a first approximation
it may be assumed that all of the heat produced goes to raise the temperature
of the stone and the air. If the heat capacities of the stone and the
air are both equal to 0.24 KOal/lCg.'C, and if the gasifier temperature is 850°C,
the regenerator temperature 1000'C and the air temperature 15°C then the
sulphur content of the stone for temperature equilibrium may be computed
as follows.
heat released by
oxidation/mole CaS
163 k cal
heat absorbed by air + heat absorbed by stone
•] '[
kg air
1.897 x 29 x
1000 x .21
kg stone v x
mol S '
P
.24
.24
x 985
150
163 kcal
Kg Stone/Mol S =
wt % S in Stone = %
62 + 36 x (K« stone/Mol S)
2.8
32 x 100
2.76 x 100
-------
Of this oxidised sulphur, the amount released as S02 is 0.845$ by
weight on stone. If the gasifier temperature is held constant and the
regenerator temperature is allowed to vary the results in Table G III
are obtained on the assumption that all the oxygen in the regenerator
air supply is utilised.
Table
Oxidised Sulphur Content of Stone
and Sulphur loss per pass
Regenerator Oxidised Sulphur Sulphur lost as
Temp. °C % wt. Stone SOp % wt. stone
970 0.71 0.32
980 0.82 0.4}
994 1.03 0.70
1000 1.14 0.84
1004 1.'26 0..99
If excess air is used the tendency will be to form more sulphate
and less S02> more heat will be generated whilst sulphur will be conserved.
The effect of this will be to lower the concentration of S02 observed at
any given temperature and to transfer oxygen from the regenerator to the
gasifier.
The inherent stability of the system is indicated by Table G-III which shows
that as the regenerator temperature rises the amount of sulphur lost from
the system as S02* for a given stone throughput, increases very rapidly.
If the sulphur input is constant this will result in a rapid approach to
the temperature at which the S02 output matches the sulphur input.
Variation of the stone transfer rate will allow this equilibrium temperature
to be varied. It is therefore desirable to have an accurate and reliable
means for controlling bed transfer rate in order to ensure that the
equilibrium temperature is low enough to avoid dead burning of the stone
and the imposition of undue thermal stresses on the structure of the
gasifier.
D. The Design of the Bed Transfer System
The system which is used to effect the transfer of bed material
between the two reactors employs gravitational dense phase transfer.
Each bed container is provided with a catchment pocket slightly above
the surface of the bed, and each of these pockets communicates via an
almost vertical duct with the side wall of its neighbour at a point
slightly above the level of the distributor. There is a
-------
- 184 -
short horizontal section at the bottom of each duct so that in the absence
of any external stimulus the two transfer ducts simply fill with static
bed material. An arrangement of this type is shown in Figure G-J4 whloh
gives a schematic layout of the first bed transfer test rig.
TEST RIG FOR BED TRANSFER TRIALS
P» PULSED AIR FEEDER
V- VIBRATORY FEEDER
AIR SUPPLY
AIR
Plg.G-3
Each horizontal connection is provided with a suitable tube running
parallel with its axis. When gas is introduced via one of these tubes
the particles in the horizontal duct are fluidised and expelled, being
replaced by gravitational flow down the vertical duct. As soon as the
gas ceases to flow so do the solids. By pulsing that gas flow and
varying the frequency of the pulses it is possible to control closely
the rate at which bed material is transferred. Very little gas is
required to operate the system. In the original test rig it was possible
to shift about two pounds of bed material per cubic foot of gas. In view
of the possible variation of sulphur content throughout the depth of the
bed, the transfer of material from the top of each bed to the bottom of
the other is advantageous from an operational point of view. The dense
packing in each duct, together with the continuous gas bleed which is
required to prevent blockage of the activating tubes, ensures that the
leakage of gas between compartments is minimal.
-------
- 185 -
The layout of the bed transfer ducts was largely dictated by the
geometry of the gasifying and regenerating compartments. It was not
possible, however, to produce a detailed design on theoretical grounds.
There was no way of predicting the dimensions and positions of the
catchment pockets which would provide satisfactory results for example,
and it would have required an excessive degree of confidence in the
outcome to have simply guessed the dimensions of the transfer ducts.
It was, therefore, decided at an early stage that the best procedure
would be to build a full scale cold model so that these finer points
could be settled experimentally. A view of the cold model is shown
in Figure GJ 4.
Fig.G-4
-------
- 186 -
The greater part of the structure was constructed of thin steel
sheet, but the bed transfer system was made of transparent plastic
so that the behaviour of the flowing solids could be observed. Plastic
construction also.lent itself to rapid modification.
The bed material which was used to test the cold model was a
crushed brick having the same size distribution and bulk density as
the lime which was normally used in the batch units.
Trouble was in fact encountered with the original design In which
the fines from the main cyclones were led directly into the bed trans-
fer ducts through br«ich pipes. It was found that the prevailing flow
of gas up the transfer ducts from the gas injectors carried fines into
the interstices between the coarser material up stream of the cyclone
branches and that this resulted In bridging. This difficulty was over-
come by adopting the configuration shown In Figure G»5.
SUCCESSFUL CYCLONE FINES RETURN ZONE
Solid* from
Fig. 0-5
The transfer duct was enlarged at the point at which It changed
direction horizontally to form a chamber having a vertical face which
was swept by gas bubbles rising from the Injector. The shape of this
chamber was such that a free surface of bed material existed within
it, so that the gas bubbling up the verticle face could only continue
on up the duct by re-entering this surface which therefore acted as a
trap for fines.
The fines from the cyclone entered the entralnment chamber through
its bubble swept face, and the mouth of the duct from which they drained
was positioned so that it was partially open to the voidage above the
free bed surface. This system was found to function well, and It was
-------
- 18? -
possible to observe entrained pockets of fines being carried towards
the delivery ports by the coarser bed material.
Tests made on the cold rig showed that the system could circulate
bed material at a maximum rate of 700 Ibs/hr. A cold check on the
completed gasifier confirmed this result. Indications are that the
transfer system consumes much less nitrogen under hot conditions than
it did cold when the requirement was about 1 cubic ft for every two
pounds of bed material transferred.
In the control system which has been adopted the pulse rate of the
gasifier to regenerator transfer line is controlled by a simple variable
propdxrtional band two term controller, according to the deviation of
the regenerator temperature from a chosen set point. The operation of
the return line from the regenerator to the gasifier is controlled by a
pressure switch actuated by the pressure drop across the regenerator
bed.
It has been found that this system is somewhat sensitive to
differences in pressure between the gas spaces of the regenerator and
gasifier and that it is necessary to ensure that these pressures are
within a few inches water gauge (w.g.) of each other. A butterfly
valve was fitted to the regenerator offtake for this purpose and was
adjusted manually. An automatic control loop will be installed for this
purpose in future runs.
E. Gasifier Construction
Construction of the gasifier unit itself consisted of bolting
together the base and prelined steel sides, building up the cores and
molds which form the interior voids, pouring the refractory concrete
interior and assembling the lids and distributors.
The base is mounted on concrete piers in a pit which houses the
gasifier unit. Each of the eight side plates was lined with a 5-inch
thick layer of 50 Ib/cu.ft castable insulation which was poured in place
on the ground before erection. All penetrations through the walls had
been precast in this insulating concrete. The bottom four side sections
were assembled on the base and lined with a -5--inch thick blanket of
expanded polystyrene sheet to provide the expansion gap. The wooden
lower sections of gasifier and regenerator cores were set in place along
the plexiglas cores for the transfer lines and catch pockets. The plastic
sections were filled with paraffin wax to provide rigidity.
Figure G-6 is a drawing of the plastic patterns.
Two views of the plastic form-work installed are shown in Figure G-7
and in text Figure I}.
-------
Regenerator.
Collection
Pocket
Gasifier
Side Wall
Cyclone
Drain.
Fines Return
From Cyclone.
PERSPEX PATTERN FOR TRANSITION DUCTS
VIEW ON ARROW. A.
Fig.G-6
-------
- 189 -
PHOTOGRAPH OF GASIFIER UNDER CONSTRUCTION
Fig. G-7
-------
- 190 -
The metal inserts for fuel injection, thermocouples, pressure taps,
sampling etc. were installed and grease coated to provida for easy
removal.
The first three lifts of refractory concrete were then poured.
The material was Durax C.1600 Aluminous Hydraulic Concrete obtained
from G.R. Stein Refractories Ltd., poured and vibrated to maximum
density.
Table G-IV lists properties of this concrete.
Table G-IV
Properties of Durax C.l600 Concrete
Maximum Service Temp. l600°C
Maximum Grain size 3/8 in.
A12 0 content 49.9 %
Fep 0^ content 0.6 %
Density, Cast and Dried 142
Density after firing to l600°C rj6
Compression strength after firing 6?27
Modulus of Rupture after firing 808 Ib/in
Thermal Conductivity at 600°C 7.615 Btu/hr.ft .-°P/ln.
When the first lifts had set, the upper four side plates
were bolted into place, polystyrene blankets, upper cores and inserts
placed, and the final three lifts of concrete poured. Precast concrete
forms were used in the start-up burner entrance, for the regenerator outlet
section, and for the cyclone outlets. Stainless steel inserts formed
the cyclone outlet tubes. Final assembly steps consisted of pouring
a layer of insulating cement on the top, seating the precast lid,
installation of refractory fibre packing in the Joints, and bolting on
the cover plate. Figure G-8 shows a photograph of the unit after the
first three lifts had been poured. A photograph of the completed
installation is shown in Figure G-9.
-------
- 191 -
PHOTQglAPH OF GASIFIER WITH LOWER CONCRETE
LIFTS POURED
-------
- 192 -
PHOTOGRAPH OF COMPLETED GASIFIER
Fig G-9
-------
APPENDIX H
CAFB Pilot Plant Alarm Systems.
I.. Alarm Actions
The installation is protected by a number of safety detection
systems and there are four main actions that may be triggered off by
an alarm depending upon which alarm is energised.
a. Alarm Action A
*
• Fire Valves Close on oil line flow and return
at entry point into the building.
• Oil circulation pump stops.
• Gasifier control panel is alarmed - the
consequences of this alarm may be controlled
and are described below.
• Interior and exterior bells ring inside and
outside laboratory.
• Red light comes up on auxiliary panel
located on laboratory wall close to main
door to JJA.
b. Alarm Action B
• Gasifier control panel is alarmed -
consequences of this alarm may be controlled
and are described below.
c. Alarm Action C
• Alarm light comes up on auxiliary panel.
• Interior and exterior bells ring inside and
outside laboratory.
d. Alarm Action D
• Alarm light only comes up on main control
panel.
e. Choice of main control panel action
There are various actions that may be preselected to follow an alarm
signal to the main control panel.
• An internal bell can ring or may be made
inoperative by a switch on the panel.
• Automatic shut down of all blowers, pumps,
compressors except burner air supply which
is manually controlled at all times.
Alternatively this procedure can be made
inoperative by a switch on the panel.
-------
- 194 -
• All alarm sensorr, may show as lights on
choice the panel when alarmed.
^ • The first alarm sensor only will show
as a light and any other alarms which
are energised as a result of the automatic
shut down will not show.
• The low hopper alarm has been arranged to
show a warning light only and cannot be
arranged to sound a bell or cause shut
down.
Generally it is proposed to run the unit with the alarm bell in
circuit and automatic plant shut down selected -this is necessary
because failure of the gas pilot flame could result in unburnt gas
forming in the boiler and ultimately an explosion.
f. Method of alarm display on main panel
The particular alarm light will be dependant upon the source of
the alarm . Table H-l lists the alarms and how they will be shown.
2. Alarm Systems
The installation is best considered as four main systems:
• The boiler and its cooling system
• The Gasifier
• The experimental burner on the boiler
• General alarms
a. . The, boiler and cooling system
The water in the boiler circuit is pressurised to 50 psi and pumped
through a heat exchanger. The secondary side of the heat exchanger is
pumped to an evaporative cooler located, on. the roof of the building.
The following alarms are installed in this system.
Failure of the cooler circulating pump will operate a
differential pressure switch across the pump feed and delivery lines -
Result - Action (A) above.
(2) High water temperature in the cooler water feed line -
set to operate at 1908F.
Result Action A above
•
(3) High water temperature in the boiler - set to operate
at 245eF.
Result Action B above.
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- 195 -
Low Pressure in the pressurisation unit - set to
operate at 4} psi.
Result - Action B above
b. Gasifier
The gasifier has many alarm circuits whose action has been
described earlier. The following parameters are monitored:-
(1) High temperature in the gasifier bed - set to
950PC. Shown and alarmed from Graardian
indicator on main panel.
(2) High temperature in regenerator bed - set to
HOCfC. Shown and alarmed from Leeds and
Northrup recorder.
(j5) Gasifier distributor blocked - shown on switch
. and set to 15".
(4) Regenerator distributor blocked - shown on
switch and set to 15".
(5) Regenerator bed low level - shown on switch
arid set to 10".
(6) Regenerator bed high level - shown on switch
and set to 24".
(7) Down stream blockage indicated by pressure
in gasifier gas space - shown on switch and
set to 10".
All these alarms Result - Action B
(8) Hopper low level will cause a warning light only
on panel.
Result - Action D
c. Experimental burner
The experimental burner will cause an alarm signal if there is
a failure of:-
pilot flame
main flame |f
low gas pressure set to ^ w»g«
low air plenum pressure - set to 2" w.g.
low pilot air pressure - set to .3 psi
Any one of these alarms will shown as alarm J on main panel.
Result - Action B
d. General Alarms
(1) Fire Detector
A fire detector is situated over the boiler and will alarm
at 150eP.
Result - Action A
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- 196 -
(2) Sump Level
in the event of thn build up of liquid in the pit to about
1" deep over the floor of the pit an alarm will sound.
Result - Action C
Emergency Stop Buttons
There are three red emergency stop buttons located
• Main door at 3A end of laboratory-
• Main door at main stores end of laboratory.
• Ad.jacent to ladder on side of the pit.
Result - Action A
Emergency Stop Button on Main Control Panel
Button located at centre of main panel.
Result - complete unit shut down
3- Restart Procedure
In the event of an alarm showing it is important to determine
the cause of the alarm and if it is safe to restart.
a. • Restart after Action A
Gasifier
The gasifier restarting procedure is described in the operating
section F Appendix I.
Auxiliary Panel
If the audible alarm has not been muted it
may be done by pushing the mute button on
the auxiliary panel.
The panel may be reset as soon as the alarm
signal has ceased and the red light will go
out.
The two fire valves must be reset by lifting
the lever and releasing the rewind mechanism
by raising the rubber covered lever on the
side of the panel and tightly winding up the
string in a clockwise direction until the
levers are set up as high as possible. The
rubber covered levers can now be switched
down.
The oil circulating pump is restarted by
depressing the green button on the side of
the starter located near the pump.
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- 197 -
b. Restart after Action B
Gaslfier
The restarting procedure is described in Appendix I.
c. Restart after Action C
Check cause of high liquid level in pit and determine if it is
safe to carry on.
d. Restart after Action D
Fill up hopper with limestone.
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- 198 -
TABLE H-I
Summary of Pilot Plant Alarm System
Indication
Source of Alarm
Auxiliary Panel Main Control Panel
Failure of water
pump.
circulating Red Light
Red light titled
Panel"
Auxiliary
5-
6.
7-
8.
9-
10.
11.
12.
14.
35-
High water temperature on
cooler feed line
High water temperature in
the boiler
Low Pressure in pressurisa-
tion unit
Gasifier high temperature
Regenerator high temperature
Gasifier blocked distributor
Red Light
and bells
None
None
n.b.
None
None
None
Regenerator blocked distributor None
Regenerator low bed level None
Regenerator high bed level None
17-
Downstream blockage on
gasifier outlet
Low Hopper level
Experimental Burner
Failure of:
Pilot flame
Main flame
Low gas pressure
Low air plenum pressure
Low pilot air pressure
Fire Detector
Sump Level
Emergency Stop Buttons
in building
Emergency Stop Button
on panel
None
None
None
Red light titled "Auxiliary
Panel"
Red light titled "Boiler
water or Burner"
Red light titled "Auxiliary
panel"
Red light shown on pressurisation
panel, and its own bell rings.
Gasifier high temperature
warning light.
Regenerator high temperature
warning light.
Gasifier blocked distributor
warning light.
Regenerator blocked distributor
warning light.
Regenerator low bed level
warning light.
Regenerator high bed level
warning light.
Downstream blockage warning
light.
Low hopper level warning
light.
Red light titled
Boiler water or
Burner
Red light
and bells
Red light
and bells
Red light
and bells
None
Red light titled Auxiliary
Panel
Nothing
Red light titled Auxiliary
Panel
None
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- 199 -
APPENDIX I
CAEB PILOT PLANT OPERATING PROCEDURES
A. Preparation for Running
1 Operation of Services
a. Oil Ring Main
Check tank contents.
Check outflow temperature is a bout 1^0/150°F.
Start compressor/oil pump on boiler by
switching on at panel. Do not run for more than
5 seconds a t first few runs in order to set thick
oil moving. Continue until temperature gauge on
outflow heater reaches 200/210°F. and then pump
may be left on.
The boiler heater circuit is now on automatic
operation and is ready for use.
b- Kerosene
Stored in 500 gal. tank - check that there
is enough for anticipated usage. First check
that isolating valve is turned off on pump
supply feed pipe. Then open valve at barrel.
c. Nitrogen
Stored in liquid Ng tank- check that there is
sufficient for anticipated usage.
Check all valves are turned off at the .
bleed locations on gasifler before opening valve
on manifold.
d. Propane
Stored in 2 banks of cylinders outside building.
Qommence runs on the bank with least available so
that replacement bottles may be ordered in
reasonable time. Warn Purchasing at least 7 days
prior to anticipated use and thereafter replacements
can be obtained on U8 hours notice.
Only turn on outside valve immediately prior to use
for safety reasons.
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- 200 -
e. Boiler and Cooling System
(1) Cooling System
• Open drain valve on cooling pump delivery which
bleeds water to waste.
• Go up on roof and check valve is open on water
feed to cooler - listen to check that water ia
flowing. If not - check all valves back to tower.
• Turn off valve on drain to conserve water until
heating takes place.
• Warn Services of an anticipated soft water usage -
maximum of 600 gph. Please give as much notice as
possible.
• Turn on secondary side i.e. cooler pump - hold button
in for at least 10 seconds to prevent a shutdown alarm.
• Turn on boiler circulating pump. No button to hold on
this circuit.
• Turn on supply to cooler fan located on end wall
(main stores end) n.b. fan is thermostatically controlled
and will not cut in until pond water reaches 100°P.
• Check temperature setting of automatic mixing valve
is set to l80°P and turn on electrical supply to valve.
(Controller set on wall by pressurisation unit).
(2) Pressurisation System
• Check water level in storage tank is at rubber band
marker and make up is turned on.
• Check nitrogen bottle pressure is above 500 psi.
• Turn on pressurisation system and check that pressure
reaches 50 psi approximately. On start up bell will
ring, cancel bell, then turn off N2 at cylinder.
(3) Boiler
• Check that main flue butterfly valve is at marked
position. (Do not take any notice of "open" and
"shut" marked positions as this was incorrectly
fabricated).
• Check that water pressure in boiler is approximately
50 psi.
• Check that any sampling lines are correctly installed
in flue.
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- 201 -
Preparation for Running Gasifier
The gasifier is assumed to be cold and empty of bed.
a. General Check
• Carry out general Inspection to ensure all main supply
lines are complete and that no sampling or other
instrumentation holes have been left open.
• Check condition of all drive belts.
• Check oil level in metering pumps, drain all air filters
for moisture.
• Check all sampling pumps work.
• Check all analytical equipment and calibrate if necessary.
b. Control Unit
Set alarm switch to "all alarms will show" position
Set alarm bell to "mute"
Set automatic shut down switch to "inoperative" position
Turn on main power supply.
"Burner or Boiler water" and "regenerator low bed"
warning lights should come on. The former alarm is caused
by pilot flame being out.
Turn on regenerator blower and check that it will shut
down if "automatic shut down" is selected. Reset alarm
panel. Replace switch to "shutdown inoperative"
Check that original two alarm lights show.
Check and fill manometers if required.
Start bleeds on injectors using direct N2 bypass to avoid
starting Compton compressors.
Set pilot burner air supply to 6 cfm, turn on propane at
external manifold, at plug cock on the wall, at valve over
pit. Start main burner blower to pressurise plenum.
Turn on pilot operating switch. The pilot should light and
lock on. The pilot should stay alight and the warning light
on the control panel go out.
Set up recorders to 20 psi air supply. Switch in charts to
check movements and recording pens are working. Turn off
charts until required. Switch is in back of panel - plug and
socket on each recorder.
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- 202 -
B Warm up
The gasifier is heated over a 3 day period with a rate of
temperature rise of about 12°C per hour.
It is important that heat be passed through the regenerator to make heat
up of the refractory as uniform as possible. To secure a gas flow through the
regenerator, the regenerator outlet valve must be open wide and the main
boiler outlet tube damper closed. Be sure there is a nitrogen bleed through
all injectors and pressure taps before starting.
The heat up sequence is as follows:
1. Start second stage air blower and line up 200 CFM air flow to gasifier.
Remote valves set as follows -
1st Stage outlet Closed
2nd Stage inlet Open
Flue Gas Recycle Closed
2. Turn on regenerator blower and set flow to about J CFM through
regenerator.
3. Open air flow to fuel injector lines. Set to 5 CFM on each
injector.
4. Start air flow to boiler main burners - Set to 900 CFM.
lj. Turn on propane supply up to gasifier start up burner. Set ganged
propane valve to "start" position.
C. Line up air flow to start up burner.
y. Turn on cooling water spray togas buuner casing.
8. Turn on electrical power to start up burner control and start burner.
9. Adjust propane air/fuel ratio to obtain about 5 times stoichiometric
air to fuel.
10. Monitor temperatures and adjust main and regenerator dampers to obtain
uniform temperature rise in gasifier and regenerator.
"11. Adjust gas and air rate to burner to follow desired temperature schedule
on TC 7 of the L & N Recorder (gas space below Lid).
12. Do not let air rate to distributor of gasifier fall below 100 CFM.
When necessary, increase total air supply to gasifier. Remember
that air to distributor is difference between total air and air
to burner.
13. When gasifier temperature reaches ?00°C begin kerosene firing. Do
not start kero firing without technical man on unit with operator.
14. Kerosene should be cut in very slowly at first, one pump at a time,
to avoid upset to temperatures. Once kero feed is started, gradually
raise kero rate to follow temperature schedule to 850°C. Air rate
to plenum may have to be increased to provide sufficient air for kero
combustion. There should be at least 50 CFM air through the plenum for
each gallon/hr of kerosene feed.
-------
- 20} -
15. When gasifier temperature reaches 850°C begin reducing gas and
air to burner while increasing kero and air to gasifier to maintain
temperature. Gas rate should be reduced to about 100 CFM, burner
air to about 60 CFM, and total air about 550 CFM.
16. Fill limestone hopper while heat up is in progress.
17. Start air flow at 5 CFM to limestone feed system.
18. Start limestone feed vibrator at low setting (about 1) and begin
adding stone to unit. Do not start limestone addition without
technical man present.
19. Observe temperatures and fluidisation as soon as limestone addition
starts to be sure that stone is being heated and fluidised throughout.
If necessary adjust air and firing rates to obtain uniform heat up.
Avoid letting stone temperature fall below 800°C.
20. Gradually add limestone to reach specified bed level on manometer.
Do not exceed specified vibrator setting while adding bed, or fill
line will block.
21. Start Compton compressors, raise regenerator air flow to 15 CFM, and
begin bed circulation between gasifier and regenerator as soon as bed
depth reaches 10 in. w.g.
22. When bed level reaches 12 in.w.g. the start up gas burner can be
turned off. Start a k CFM air flow through burner cooling inlet to
keep burner head cool and block main air line to burner.
C. Initiating Gasification
When the gasifier is at operating temperature with a desired bed
level established and circulating, gasification can be started. Hie
procedure used is to switch from kerosene combustion to fuel oil combustion,
increase oil rate to the stoichiometric ratio to eliminate oxygen from the
system, and then increase oil rate to that required for gasification.
1. Starting point is with gasifier at about 850°C with at least 12 in.w.g.
bed depth, and good solids circulation in both directions between
gasifier and regenerator. Temperature is maintained by kerosene
combustion in the bed with excess air.
2. Increase air flow to main burner to 1100 CFM total of which 150 CFM
is passed through the premix nozzle.
3. Check that pilot flame is strong and stable. Do not proceed until it
is.
**• Check cooling system to be sure it is operating satisfactorily.
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- 204 -
5. Line up flue gas recycle through second stage blower and adjust
total flow to gasifier plenum to 260 CFM; J50 CFW flue gas recycle
and 230 CPM fresh air.
(Procedure was changed during 3rd pilot plant run to pass flue
gas recycle through both first and second stage blowers).
6. Stop oil pumps and switch pump suction lines to fuel oil
supply from kerosene supply.
7- Set centre pump to deliver 78 Ib/hr fuel oil (stoichiometric)
and start pump.
8. Set pumps 1 and 3 to deliver 1^0 Ib/hr fuel oil each, and
when gasifier bed temperature reaches 900°C, switch on both pumps.
Normally pumps 1 and 3 will be required about 15 seconds after
starting pump 2 at stoichiometric oil rate.
9. Main boiler flame should ignite within about 15 seconds of
starting pumps 1 and 3-
10. Raise pump 2 setting to deliver 1J50 Ib/hr fuel oil.
11. Adjust flue gas recycle rate to control temperature at desired
level.
12. Check that cooling tower fan starts when secondary water
temperature reaches the set point.
13. Begin limestone feed addition.
Ik. Bring all conditions to running specifications.
D. Solids Handling
1. BED REPLACEMENT
It is necessary to add fresh bed to the gasifier to maintain
efficiency of operation. Bed material will be drawn out of the
regenerator at 18/20 Ibs/hr using a vibratory feeder to an external
storage hopper on the floor of the pit. The gasifier bed level will then
drop and start a vibratory feeder which will add material from a storage
hopper mounted on the gasifier lid. The storage hopper has a low level
warning and may be replenished from a ground level supply of raw
material by a suction lift pipe.
It is most important that gas cannot enter into the hopper and
deposit tar which would stop the feed of material. The feed system
is therefore a sealed system valved in three places and with an inert
gas purge in the feed pipe to discourage back flow. During hopper
refilling the vibratory feeder must be turned off and the system must
be isolated by the air operated valve above the gasifier lid before
opening the valve on the material feed line and the valve on the vacuum
-------
- 205 -
blower. Refilling will be carried out when the low level alarm shows
a nearly empty hopper and the requisite amount of material as shown by
the full line on the storage hopper may be added without overfilling
the hopper. Immediately after filling close the valves on the feed
line and vacuum blower and then open air operated valve and turn on the
switch on the vibrator feed.
Summary of operating procedure for gasifier hopper filling:
F.H. is feed hopper located above gasifier.
S.H. is ground level storage hopper.
a . Low level alarm shows on console.
b . Turn off vibratory feeder switch and wait for line through
valve to clear.
0 . Turn off isolating valve by turning 3 way air cock to closed
position.
d. Open upper valve on vacuum blower - check lower valve is
closed.
e . Open valve on feed pipe near S.H.
f. Check S.H. level is at correct height.
g. Switch on vacuum blower until all material has gone into
"P.H. Low alarm should now go out.
h. Shut valves on vacuum blower and feed line.
i. Open air operated isolating valve.
J . Turn on vibrator switch.
k. Refill S.H. to correct level for next filling operation.
2. BED REMOVAL
Hot bed material will be removed from the regenerator by manually
operating the vibratory feeder which will feed the material into a
storage hopper - trials will have to be made at temperature to determine
if a low steady draw off can be achieved or if it will be necessary to
draw off at intervals. A nitrogen purge is available if required.
3. DRAINING REGENERATOR CYCLONE
This cyclone may be drained periodically as shown to be necessary
by experience. Normally the upper valve is open and hopper drain valve
closed. To empty hopper just shut upper valve and open lower valve,
tap the upper sides to free any fines adhering to the walls. Close
lower valve and open upper valve when emptying is completed.
^. GASIFIER BED SAMPLING
Samples may be taken from either of the gasifier drains by closing
lower valve, opening upper valve and bleeding in nitrogen for a brief
period. Then open lower valve with a suitable container in place and
collect sufficient material.
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- 206 -
E. Carbon Burn-out Procedure
The carbon-burnout procedure is initiated when the gasifier
pressure above the bed reaches 20 in.w.g. or when pressure drop
through the cyclone and duct system begins a rapid rise.
1. Preparations
a. Call out additional assistance if needed.
b. Connect burn-out gas supply fitting to gasifier lid.
c. Prepare gas sampling lines from plenum gas feed and lid
gas feed lines to 0^ and COp meters.
d. Shut off any air bleeds through lid
2. Bed Sulphation
a. If necessary, lower bed level by draining gasifier solids.
b. Reduce gasifier temperature to below 850 °C by increasing
flue gas recycle rate.
c. Line up steam to steel cyclone gas outlet tubes.
d. Switch fuel injector bleed gas from air to nitrogen at
2 CBW each.
e. Shut off fuel to gasifier by stopping pumps and closing
valves in fuel lines.
f . When main flame goes out, close air valves to burner and
close flue outlet damper. Close fresh air inlet to gasifier
blower .
g. Set flue gas recycle rate to maximum.
h. Set high nitrogen flow rate to lid. This flow can be
stopped later-
i. Control bed temperature to 900-950° C range by regulating
flow of fresh air to main burner. This air circulates back
through flue gas recycle line to provide oxygen for bed
sulphation. Proper oxygen concentration is about 10# Op in
gas to plenum.
J. When sulphation is complete, temperatures down stream of bed
will begin rapid rise and bed temperature will begin to fall.
Sulphation normally requires about twenty minutes. At this
point close air flow valve to gasifier distributor and reduce
regenerator air flow to about 3 CPM.
-------
Carbon Burn Out
a. Start flow of flue gas and air through gasifier lid and
observe effect on gas space and cyclone temperatures.
b. Adjust fresh air flow to main burner to obtain desired
0 content of gas to gasifier lid.
c. Adjust steam rate to cyclone tubes to avoid over-heating.
Hold cyclone tube temperature at 900°C or lower.
d. Gradually increase gas flow through lid to about
50 CPM. Continue close control of temperatures.
e. When carbon burning at cyclone tubes is complete, gas
flow and 0 content can be increased to complete burning
down stream of cyclones.
f. Burn out is complete when all duct temperatures begin
falling. Care must be taken not to allow bed temperature
to fall too far during burn-out or difficulty will be
experienced during restart.
Restart
a. Stop air flow to gasifier lid.
b. Close steam valves to cyclone tubes.
c. Open main flue dampers.
d. Open air flow to main burners.
c. Start pilot light.
f. Partly close flue gas recycle valve and open air suction
valve to gasifier blower.
g. Set gas flow to gasifier plenum to 260 CM.
h. Return fuel injector gas to air from nitrogen.
i. Start fuel oil injection at low rate to return gasifier
temperature to 850° C.
j. Initiate gasification by usual procedure except that no
switch over from kerosene is required.
-------
Operation of Continuous Pilot Plant
This appendix contains a summary of operations during the three
1971 runs of the CAFB pilot plant, data plots and tables, photographs
and a description of the unit at the end of the third run.
A. RUN 1
The continuous CAFB pilot plant was operated as a gasifier for 160
hours in August. During this period, a number of major process features
including continuous sulphur removal, production of a highly combustible
gas giving a clean luminous flame, gasifier and regenerator temperature
control, production of a concentrated SOp stream from the regenerator, and
continuous solids circulation were demonstrated.
However, several operating difficulties were identified which
prevented full achievement of the 200 hour run objective and which caused
several interruptions in gasification. To a large extent these difficulties
were of a mechanical nature which slight modifications remedied. Two
process related problems were also identified which required modified
operating procedures for their solution.
1. Description of Operations
a. August 1-5
The CAFB gasifier unit was preheated at the end of July, and a bed
of lime was established. Hot lime circulation tests were conducted during
the period of August 1-5- It was established that best circulation rates
could be maintained if pressure differential between gasifier and regenerator
was kept to below 5 inches w.g. A valve was installed in the regenerator
outlet line to facilitate control of this pressure difference. Solids
transfer rates of over 800 Ib/hr were reproducibly demonstrated during this
period.
«.,'
A test of the fuel feed system revealed that the shrouded injector
lines would give an unacceptably high pressure drop at start-up (when the
final section of injector line was cold). The fuel system was modified by
removing the inner injector tubes and allowing the fuel flow to pass through
the larger diameter tubes formerly used as shrouds. A stream of air injected
with the fuel keeps the velocity high and prevents coking in the inlet tube.
b. August 6-9
After making additional tests of the pilot flame system to verify
:onditions which assured its stability under hot operating conditions,
rasification trials were initiated on August 6th. Two tests were made
with the back boiler door open and an extension to the fire tube in place.
Only short runs could be made with this arrangement, since no flue gas
recycle was available, and ability to control temperature was limited.
The procedure used in changing from combusting to gasification conditions
was to:-
(1) raise fuel rate to the stoichiometric level for
approximately ten seconds to sweep air from the
gasifier with combustion products;
(2) raise fuel rate to 500 percent of stoichiometric
to initiate gasification.
-------
- 209 -
Smooth ignition of the main flame took place at once when gasification
started.
The first open tube gasification trial lasted seven minutes, the
second approximately fifteen minutes. Both tests were carried out long
enough to observe that the main burner flame was stable and that the
start-up procedure appeared to be safe. Both tests were terminated by
turning off fuel oil flow and returning to kerosene combustion when
gasifier temperature exceeded 920°C. During these brief tests it was
noted that regeneration had started as regenerator temperature began rising.
For the second test it was noted that sufficient sulphur had been deposited
in the gasifier bed to cause a sharp temperature rise when fuel flow was
discontinued.
The extension tube was removed and the boiler door closed after the
second open tube test. A gasification trial was then started with the
complete boiler, flue gas stack, and flue gas recycle system in operation.
The first continuous system trial lasted apprximately 12 hours. It
was terminated when a gradual increase of back pressure on the gasifier
was accompanied by a loss of air flow rate to the gasifier plenum. Gasifier
pressure had increased from an initial level of 8 inches to 14 inches w.g.
when the shutdown was made. Reduced capacity of the air blower was later
traced to belt slippage. Pressure rise in the gasifier was attributed to
build up of carbon between gasifier and outlet duct. The obstruction causing
this pressure rise was removed by returning the unit to kerosene combustion
conditions with excess air. The excess air apparently burned out any carbon
accumulation that had existed.
During the twelve hour operating period, good temperature control of
gasifier was maintained by adjusting relative flue gas and air rates, and
regenerator temperature was controlled by adjusting solids circulation rate.
However, S0p concentrations from the regenerator never exceeded 4.3# by vol.
Good solids circulation rate was maintained right up to the shut-down.
For a period following the shut-down of gasification, all air and
fuel flow to both gasifier and regenerator beds was stopped except for small
air bleeds through the injectors and pressure taps, and a slight bleed to
keep the regenerator distributor holes free of fines. When air flow was
resumed, an obstruction was observed in the regenerator bed itself and in
the regenerator-to-gasifier transfer line. The period from August 7th to
August 9th was spent in removing these obstructions and in re-establishing
solids levels, circulation rate, and temperature under full combustion
conditions.
c. August 9-11
Gasification resumed on August 9th at 15.30 hours and continued
until 10.00 on August llth. Gasification was interrupted because of two
operating difficulties: (1) pressure in gasifier had risen to 14 inches
w.g. and flue gas recycle rate was limited by reduced blower capacity at
the higher back pressure; (2) regenerator instrumentation indicated an
obstruction in regenerator outlet. This was subsequently traced to a
plugged pressure tapping which caused a false pressure reading.
When gasification stopped, a flow of fluidising air was maintained in
the regenerator. The gasifier bed was slumped and then pulsed with short
bursts of fludising air over a period of three hours to convert the bed
-------
sulphide to sulphate without excessive temperature rise. The bed was then
put on full combustion with excess air to burn out any carbonacious obstruc-
tions. Before resuming gasification it was necessary to remove an obstruction
that had formed in the gasifier-to-regenerator transfer line. During the
August 9-H operating period conditions were varied to increase regenerator
SOp concentration and to improve sulphur absorption in the gasifier.
Desulphurisation of over 90$ was achieved briefly, and SOp concentrations
over 10$ were reached. During this period it was apparent that the gasifier
cyclones were functioning as large quantities of fines were removed from the
regenerator cyclone drain.
d. August 11 - 15
Fuel oil was introduced to resume gasification at 22.40 hours on
August llth. Continuous gasification was maintained until 13.30 on August
15th when gasifier internal pressure had increased to 24 in. w.g. This
period of approximately 87 hours provided the most consistent demonstration
of gasification, desulphurisation, and regeneration in the first
test. Conditions maintained in this period are shown as a series of log
plots in Figures J-l through J-7.
Desulphurisation efficiency varied from 75$ to 97$ and averaged 90$.
The SOp concentration in the regenerator off gas averaged 7.5$ for the
entire period and exceeded 10$ a number of times.
Other average conditions for the period are listed below in Table J-I.
It was evident that something had happened to change performance of the
gasifier cyclones between the end of the August 9 - 11 period and beginning
of the August 11 - 15 period, as practically no fines were recovered from
the regenerator cyclone in the latter period. Fines evidently were lost
directly from the gasifier into and through the boiler.
Table J-I
CAFB Pilot Plant Average Run Conditions
11 - 15 August, 1971
Gasifier
Temperature 869"C
Bed Depth 15.6 in. w.g.
Bed Density ,^ 49-5 Ib/ft?
Superficial Gas Velocity^ ' 3.2 ft/sec.
Oil Feed Rate 361.4 Ib/hr
Stone Feed Rate 43.8 Ib/hr
Calcium/Sulphur Ratio 1.66 mole/mole
Air/Fuel Ratio 16.8$ of Stoichiometric
Sulphur Removal Efficiency 90$
Regenerator
Temperature 1047°C
Air Rate to Distributor 16.7 cu.ft./min.
SOp in off Gas 7-5$
Op in off Gas 0.47$
CDp in off Gas 2.75$
Solids Discharge Rate 10.2 Ib/hr.
(a) velocity of air and flue gas fed to plenum taken
at gasifier bed temperature and top bed empty
cross section area of 5.2 sq.ft.
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- 211 -
e. August 15 - 18
After termination of gasification on August 15th, the gasifier was
pulsed with air to sulphate the bed without excess temperature rise.
Kerosene combustion with excess air was then resumed to decoke the gas outlet.
Following this operation it was found that partial obstructions had developed
in the solids transfer lines, and efforts to remove them were unsuccessful.
Gasification with limited circulation was conducted from 2200 hours on
17 August until circulation stopped at 1900 hours on 18 August. Gasification
without circulation was continued until 01.00 on 19 August when fuel was
stopped because of gasifier pressure increase to 20 in. w.g. At that point
a controlled bed sulphation was conducted, but air flow was stopped before
burning off carbon above the bed in order to preserve any obstructions which
had formed. The unit was then cooled and opened for inspection.
2. Condition of Unit After Run
The unit was inspected for refractory cracks and was found to have no
new cracks since the initial firing-in period. However, some of the old
cracks in the lower portion of the unit appeared to have increased in size.
These enlarged cracks were in the bottom lift of concrete which was the
coldest portion of the unit. None communicated between gasifier and
regenerator. Some led out to the space between concrete and insulation.
There appeared to be no cracks in the transfer lines or in the cyclone walls.
The cyclone walls did show appreciable roughening.
Accumulations of solids were observed at several locations in the
pilot plant when the unit was opened. Samples of these accumulations were
analysed, and their chemical composition is reported in Table J-II. Their
location and general appearance is described in the following section.
a. Gasifier
Walls of the gasifier above the bed were coated with a very thin
layer of carbon. There appeared to be little carbon on the walls below
the bed level. There was a carbon-free area in the gasifier lid around
the limestone feed point. Apparently the air bleed in the limestone inlet
had prevented carbon formation at that point. The gasifier bed wall on
the fuel injector side showed a roughened appearance both below and above
the bed level. This contrasted with the smooth wall on the opposite side.
Mounds of loosely agglomerated lime particles were found beneath
and surrounding each of the fuel injection nozzles. These mounds were
removed from the gasifier intact, but in several days they lost strength
and disintegrated into lime particles. They appeared gray and relatively
free of carbon. A similar, but much smaller mound was found just beyond
the outlet tube where regenerator solids return to the gasifier bed.
b. Gasifier Outlet Ducts
The thin layer of carbon on the gasifier walls thickened in the
region of the outlet ducts leading to the cyclones. Formations appeared
to have grown across the duct inlets obstructing a large portion of their
areas. These deposits extended only a short distance into the ducts and
appeared to have formed in the eddies of the entrance region. These
deposits are the probable cause of much of the increase in pressure drop
between gasifier and cyclones during gasification. The deposits were hard
and dense^ .They ajgpear to consist mostly of carbon but contain some lime fines.
-------
- 212 -
c. Cyclone Outlet Tubes
Both cyclone outlet tubes, made of type 310 stainless steel, had
been burned away, and only short stubs remained. Portions of the tubes
had fallen into the cyclone bodies blocking the drain lines. Much of
the cyclone bodies were filled with loose lime particles.
d. Regenerator Bed
Walls of the regenerator in the bed space were clean. However,
two large solid lumps occupied the lower portion of the bed Just above
the distributor. Each lump obstructed a portion of the bed cross
section, but a passage remained between them. The material from which
these lumps were formed was quite hard and dense.
The regenerator walls above the bed were clean for a considerable
distance above the bed. However, a rough surfaced growth had formed Just
beneath the outlet duct at the regenerator top. This formation did not
reduce the opening below the outlet duct diameter, but it did form a
conical transition shape from the bed diameter to the duct diameter.
e. Solids Transfer Lines
The catch pocket in the gasifier wall was crusted over with a skin-
like layer of agglomerated solids. Otherwise the gasifier-to-regenerator
transfer line was clear. On the regenerator-to-gasifier side, a lump of
hard solid was firmly attached to the wall of the mixing pocket Just above
the agitator air inlet. Another hard obstruction was located at the bottom
of the solids down-comer where it entered the mixing pocket.
f. Gas Transfer Duct
The major portion of the gas transfer duct was clear except for a
thin deposit of carbon. However, a growth of carbon had formed at the
junction where gas streams from the two cyclones combined. The growth
acted as a fairing in the dead zone between branches of the Y. Other
small growths protruded into the ducts above the cyclones where the gas
changed direction from vertical to horizontal. Again the growth filled
what appeared to have been the down stream eddy spaces.
g . Burner
The steel inner liner of the burner up-stream of the main air mixing
point was badly burned on its lower surface but not on its top. It appears
that carbon particles may have accumulated in the bottom during decoking
operations and then burned at a high temperature. The burner face and
the stainless steel baffles installed for pilot flame stabilisation were
intact and in good condition.
h. Boiler
The boiler end space and lower tubes contained a considerable accumu-
lation of lime fines. Approximately 1^00 pounds of solids were recovered
from the boiler and its tubes. The accumulations were all loose and
probably could have been removed during operation had there been appropriate
drain points.
-------
TABLE J-II
Composition of Solid Residues from Pilot Plant
after First Run
Sample Location
Composition
CaO wt.#
MgO "
Sl°2 "
Na20 ppm
Fe20 wt.#
A120 "
Carbon "
Total Sulphur wt,
Sulphide Sulphur
Vanadium wt.#
Raw Stone
55.8
0.5
0.5
54
0.2
0.2
-
.*
wt.g -
-
Regenerator
Bottom Deposit
59.5
0.6
0.5
740
0.59
0.2
< .1
25.17
0.07
0.08
Regenerator
Cyclone Fines
80
0.6
0.5
195
0.49
0.2
0.25
8.62
7.26
0.15
Gasifer Cyclone Deposit
Inlet
24.5
0.6
7
540
0.89
5.9
46.4
2.87
2.59
0.25
Outlet
70.2
0.6
0.7
145
0.94
0.4
16.9
8.14
7.90
0.57
Boiler Tube Deposit
1st Pass 5rd Pass
70.1
0.6
0.7
210
0.94
0.5
.25
5.7
1.55
0.4
74.8
0.6
0.7
220
0.61
0.5
1.18
5.08
5.52
0.5
-------
-214 -
Mechanical Trouble Areas
Three operational difficulties of a primarily mechanical nature
contributed to run interruptions and non uniformity of operating conditions.
These were: (1) Slippage of drive belt on gasifier air blower;
(2) Blockage of limestone addition line.
(3) Plugging of flue gas compressor suction line.
Slippage of the drive belt on the gasifier blower was a purely mechanical
problem easily corrected when recognised. However this slippage decreased
blower capacity during early stages of the gasification run and forced
premature shutdown after moderate gasifier pressure rise on August 7 and
August 11. Tightening the drive and application of a belt dressing eliminated
the problem after August 12.
Periodic plugging of the limestone addition system was a more
persistent problem that never was completely eliminated during the run. In
early stages of the run it was aggravated by use of flue gas recycle from
a Nash compressor as bleed gas. This flue gas was wet and caused the stone
to become sticky and prone to plug the line. Replacement of the flue gas
with air and installation of a clearing rod with block valve and stuffing
box greatly improved the operability of the system but did not eliminate
occasional stoppage, especially when addition at a low rate was desired.
A larger stone addition line and a greater air bleed rate was provided
for the next test to improve reliability of stone addition. The stone flow
rate interacts with gasifier temperature control, since stone addition
imposes a substantial heat load on the system. Improving stone addition
uniformity eases the task of gasifier temperature control.
Periodic plugging of the suction line to the flue gas compressor
eventually led to abandonment of the use of flue gas for instrument bleeds
ind injection gas to the solids transfer system. Replacement of this flue
gas with air eventually created additional problems. The plugging agent
'r\ the flue gas recycle line was lime fines which built up at the point in
the line at which temperature fell below the dew point of the flue gas.
The flue gas compressor suction had been equipped with a water scrubber
to knockout the lime fines. However, the scrubbing water cooled the inlet
line enough to cause condensation of water vapour at the inlet to the
scrubbing section. This inlet then quickly plugged. Although a more
sophisticated scrubbing system with filters and preheat could be provided,
it was decided to use nitrogen from the ERCA site supply as the gas for
bleeds and solids transfer injectors.
4. Process Operating Problems
Two process features have the potential for creating operational
iifficulties, but results of the first continuous pilot plant test indicated
that control of these features could be achieved. These process problem
areas were:(l) Formation of deposits of carbon and lime fines in critical
operating points between gasifier and boiler, (2) Agglomeration of
sulphided lime particles when regenerated in a defluidised state.
-------
-215 -
5. Process Variable Relationships
There was no attempt in the first run to carry out a programme of
variable studies. The objective was to hold most variables at
relatively constant levels, and this precluded much possibility of
variable correlations. However, experience in operating the unit
demonstrated interdependence of a number of factors.
a. Limestone Feed Rate
Limestone feed rate influenced most aspects of gasifier operation,
and it was unfortunate that this feed rate often was erratic. With a
constant limestone addition rate, the unit was self stabilising and
tended to line out at a constant set of conditions. However, sudden
changes in feed rate caused variations which propagated throughout the
system. Increasing limestone feed rate increases heat demand of the gaslfier
both to heat the stone and to achieve calcination. To provide additional
heat, the air/flue gas ratio must be increased or gaslfier temperature, will
fall. Increasing the quantity of fuel combusted, together with the CO-
released by calcination, increases the gas velocity from the gaslfier.
This in turn increases the pressure drop through the gas ducts and cyclones
and raises gasifier pressure. This change in gasifier pressure relative
to regenerator pressure affects solids circulation rate between the beds and
in turn affects regenerator temperature and S0?evolutlon rate. An interrup-
tion to bed addition allows gasifier bed depth to fall. Finally, a prolonged
change in limestone addition rate affects activity of the bed for sulphur
recovery and changes sulphur removal efficiency.
b. Sulphur Removal Efficiency
Sulphur removal efficiency varied from over 9856 near the beginning
of the test period to a low of 75# near the end, with fluctuations between
80 and 98# throughout the test. The major cause of fluctuations is believed
to be variations in limestone addition rate. However, there is not yet a
clear correlation due to variation of other factors such as sulphur content
of the circulating lime. It is evident that sulphur removal efficiency of
over 97# was obtained with about 1.5 to 1.7 mol Ca/mol S lime addition
rate. Following an interruption to bed addition early on 15 August, bed
depth decreased, and efficiency fell to about 75#.
c. Gasifier Temperature and Air/Fuel Ratio
Gasifier temperature was self-stabilising so long as fuel rate, air
rate, and limestone addition rate were constant. There was no automatic
control device for temperature, so any changes required were made by the
operator. The principal cause of the temperature variations that did
occur was interruptions to limestone addition which required adjustment of
the ratio of flue gas recycle to fresh air addition.
The average air/fuel ratio was \1% of stoichiometric which gave a
rich, highly combustible fuel gas. Even lower air/fuel ratios, down to
15$ of stoichiometric, were used during periods of no limestone addition.
Substantial cooling by flue gas recycle was being used even at these low
air fuel ratios which means that, from a heat balance standpoint, even lower
air/fuel ratios are possible.
-------
TABLE J-III
Cowoaltlon of Regenerator Samples
Renoved from Pilot Plant during Run 1
Sample
Composition
CaO Mt.£
MgO
Sl°2
ppm
wt.*
Carbon "
Total Sulphur wt.£
Sulphide Sulphur wt.S
Vanadium wt.Jt
'-8-0130
87.7
0.6
1.0
185.
0.9
0.4
3.37
2.20
0.75
12-8-0730 12-8-1430 12-8-1730 12-8-2030 13-8-0430 13-8-0730*
Total
89.7
0.6
0.8
195
0.69
0.4
4.36 4.91 3.13 3.33 3.15 2.38
3.45 3.73 0.73 2.04 0.84 0.96
0.68 0.6o 0.68 0.67 0.77 0.73
brown
79.4
1.1
2.5
525
1.05
1.8
2.61
2.05
0.91
White
88.0
1.1
0.2
42
.16
0.2
2.35
0.90
0.18
14-8-1030* 15-8-0030 15-8-0830
Total
94.1
0.6
0.9
295
0.3
0.4
5.10
4.44
0.8
Black
90.5
0.6
0.7
420
0.36
0.4
8.52
8.15
0.65
White
82.7
0.6
0.3
51 -
.11
<0.2 - - £
5.36 4.38 4.73
4.41 3.20 4.57
0.09 0.77 1.02
(a) Sample subdivided by colour of particles
-------
- 21? -
d. Regenerator SO,-, Concentration
This test provided the first data on CaS regeneration under
continuous operating conditions. The results show that continuous regenera-
tion is feasible and that high (of the order of 10$) SOp concentrations
can be obtained below 1100°C. The SCL concentrations achieved averaged
about 50# of the equilibrium values for the corresponding temperatures
as shown in Figure J-7.
Composition of regenerator solids samples removed during the run
are listed in Table J-III.
e. Sulphur Balance
The following sulphur balance was obtained during the period of
August 11 - 15 in the pilot plant:-
Sulphur Input 361.4 Ib/hr. fuel x .0248 fr S = 8.97 Ib/hr
Sulphur Out in fuel gas 8.97 x 0.1 =0.90 Ib/hr
Sulphur Out Regenerator as SOp = 6.82 Ib/hr
Sulphur on Solids from Regenerator = 0.4 Ib/hr
Sulphur on Fines Passing to Boiler = 0.64 Ib/hr
Unaccounted for Sulphur = 0.21 Ib/hr
Sulphur Balance 100 x 8.76/8.97 = 97.3$
6. Unit Modifications
The following modifications were made to the pilot plant between
Runs 1 and 2.
a. A nitrogen supply was installed for gas bleeds to avoid need
for flue gas compression and to provide an inert gas source
when the boiler was not operating.
b. The limestone feed line was increased in diameter with
provision for additional air bleed.
c. Entries to the gasifier outlet ducts were rounded to
provide smooth entrance regions.
d. View ports were installed in the lid, and provision was
made for directed air injection to burn carbon at the
outlet ducts.
e. Cyclone outlet tubes were removed.
f. The burner nozzle shape was modified to eliminate
accumulation of loose carbon.
g. Provision was made for draining lime fines from the
boiler end during operation.
h. Fuel injector nozzles were reduced in diameter.
Li..
-------
- 218 -
B. RUN 2
Heating of the CAFB pilot plant for its second test began on October 1,
1971. Gasification started on October 7 and terminated on October 11
because of excessive lime losses from the unit.
Most of the mechanical problems which affected operations in August
were absent in October. These items included blower belt slippage, plugging
of lime transfer lines, and blockage of the limestone feed line. However,
a high rate of lime loss from the gasifier interacted with many features
of the process to reduce unit performance and prevent obtaining information
in the region of principal interest.
Desulphurisation was essentially complete throughout the test period,
as measured by S0? content of the boiler outlet gas. However, much of the
sulphur was absorbed or reacted with lime that passed through the boiler
rather than through the regenerator. Also, limestone replacement rate was
higher than would be desirable for a commercial plant.
The range of operating conditions is shown in a series of log plots
in Figures J-8 through J-15.
1. Description of Operations
Temperature of the gasifier interior was raised to 900°C by firing with
propane and kerosene between October 1 and October 4.
Lime addition, using material removed from the bed after the August
run, started on October 4. By October 5, a bed level of 15 inches w.g. had
been established and bed circulation was started. Circulation was uneven
however, and attempts to find a set of conditions giving steady circulation
were unsuccessful. The difficulty was traced to a crack which had developed
in the regenerator distributor. This distributor was replaced on October 6.
However, under hot conditions it was not passible to obtain a perfect seal
between the regenerator base and distributor with the result that some air
bypassed the distributor holes. Good fluidisation and solid circulation were
obtained when gas flow to the unit was resumed. Circulation calibrations and
final unit checks were completed and gasification initiated at 21.14 hours
on October 7.
There were no interruptions to gasification for the 9J- hour period of
operation through 16.30 on 11 October. Sulphur removal from the flue gas was
essentially complete for the entire period. However, lime losses, and the
consequent limestone make-up rate were very high, and less than half of the
sulphur was being removed as SOg in the regenerator off-gas. Thr remainder
^•as absorbed on lime which either was removed from drain points on the
boiler or was lost as fines with the flue gas.
Attempts to reduce lime losses to a reasonable level by variation of
operating conditions met little success, and therefore the run was stopped
at the end of 91 hours.
-------
- 219 -
2 Lime Losses
Figure J-8 is a time plot of average limestone feed and lime removal
rates for various periods of the run. Since lime losses were higher than the
target lime replacement rate, no lime was removed from the regenerator other
than periodic analytical samples. As Figure J-3 shows, the rate of lime
loss varied from about 20 to 50 Ib/hr during the run. Limestone replacement
rates as high as 135 Ib/hr therefore were needed to maintain gasifier bed
level. In the early portion of the run, bed level had been allowed to fall
from an initial level of 18 inches w.g. In the mid portion of the run,
the level was stabilised at 13 to 14 in. w.g. while temperature was maintained
at 870-900°C. This condition gave the highest fines losses.
In the final portion of the test, bed level was allowed to fall to 11
inches w.g. and gasifier temperature was lowered to below 850°C. These
changes reduced the fines losses, but not enough to give satisfactory
performance. As Figure J-9 shows, the calcium oxide/sulphur ratio varied
from about 2.5 to 4.6 mole/mole.
The major cause of the high lime fines losses during this run was
elimination of the central cyclone outlet tubes from the unit.
However, a contributing factor was the use of high velocity air jets in
the fuel injectors. After Run 1, a 3/16" diameter nozzle had been inserted
in each fuel injector to improve dispersion of the oil. During start-up it
was observed that air velocities of about 10 CFM injected with the fuel gave
improved combustion in the bed and these high velocities were maintained through
the gasification run. These conditions produced sonic jet velocities at the
fuel inlet and evidently produced serious lime attrition.
The method of adding limestone directly through the gasifier lid also
contributed to the losses as some of the stone was blown through the outlet
ducts before it even reached the bed.
3. Gasifier Outlet Pressure Drop
The first run encountered a gradual rise in pressure drop between the
gasifier and the transfer duct. Examination of the unit at the end of
that test indicated that most of the resistance had been due to an accumulation
of lime and carbon at the entrance to the gasifier outlet ducts leading to
the cyclones. In the October run, the rate of pressure rise was much lower,
and for over 70 hours appeared to be negligible. However, for the final
20 hours, the pressure did rise at a significant rate (Figure J-10).
4. Regenerator Operation
The high lime loss rate in the October run caused much of the sulphur
absorbed in the gasifier to be removed with the solids that passed into the
boiler. This factor reduced the sulphur content of lime going to the
regenerator and reduced regenerator efficiency.
Consequently, the sulphur : dioxide content of the regenerator off gas
remained below 4# during much of the run and failed to exceed 8# at all.
"The selectivity of regenerating sulphide to oxide also was lower than in the
first run, because the quantity of sulphide oxidised to sulphate in heating
the solids to regenerator temperature represented rtost of the sulphide
available. Figure J-15 shows the log of regenerator S02 output
-------
- 220 -
(as equivalent sulphur Ib/hr) and of CaS to CaO oxidation selectivity. This
selectivity generally was below 50# compared with values from 50 to 80#
htained in the first run when the lime entering the regenerator contained
More sulphur.
Figure J-14 shows the oxygen and carbon dioxide concentrations in the
regenerator gas. Both were higher than experienced in the first run due
to the lower sulphur on lime. The very high oxygen level observed early
on October 11 was obtained during a test in which solids circulation was
reduced to a very low rate. This test indicates some passage of gasifier
product through the transfer line into the regenerator since there could
not have been enough carbon on the slowly circulating lime to account for
the significant level of COp observed in the regenerator gas during this
period.
It is evident from regenerator performance during this run that
regenerator efficiently is strongly affected by sulphur concentration in the
incoming lime and that high sulphur concentrations improve regenerator
efficiency and reduce the relative amount of sulphur recycle.
5. End of Run Inspection
At 1700 hours on October 11, gasification was terminated and the bed
was converted to sulphate with a continuous flow of air and recycle gas
(while adding stone to control temperature).
The unit was held at temperature by gas firing for approximately 24
hours and then was shut down and allowed to cool.
When opened, the gasifier was found to contain no lumps or solid
dejosits such as those observed around the fuel injectors after the first
test. Gas transfer ducts, cyclones and the lime circulation passages also
were clean. Of course, the period of hot operation after terminating
gasification would have removed any carbonaceous deposits but the absence
of lime agglomerates in the gasifier and transfer lines shows that the use of
nitrogen rather than air for pulse gas in the transfer system was effective
in avoiding agglomerates.
The lower portion of the regenerator did contain a single large deposit
which covered the centre portion and about half the perimeter of the
distributor. This deposit tapered upward and away from the port through
which solids entered from the gasifier. Inspections showed that the distributor
plate was not well seated at its base, and evidently had bypassed a large
portion of the fluidising gas around one side of its rim. It appears that
t'r.i bypassing had reduced the pressure drop through the distributor holes
enough to permit defluidisation on one side of the distributor. The defluidised
lime particles then had agglomerated, forming the deposit found at the end of
the run.
It is evident that incomplete seating was achieved when the distributor
was replaced under hot conditions Just before starting gasification. Care
in seating the distributor together with provision of an air flow to the
regenerator throughout heat up should avoid this problem in the future.
A considerable quantity of fine lime was found in the flue gas recycle
line, blower, and in the gasifier plenum. However, no obvious pressure
drop increase had occurred in these svstems duririP- the course of t.ViP run.
-------
- 221 -
UNIT MODIFICATIONS PRIOR TO RUN 3
A fairly extensive series of modifications was made to the pilot
plant prior to Run 3 to correct problems which had arisen.
Gasifier
a m Bed Feed System
A modification was made so that the limestone addition would be
introduced directly into the fluidised bed through the sidewall rather than
dropped through the lid with the possibility of some of the finer material
being blown straight out before reaching the bed. Due to limited time this
system could not be developed to give complete reliability and so the
original system of gravity feed through the lid was left available in case
of trouble.
b. Fuel Injectors
The hole in the end of the fuel injectors was increased from 3/16"
dia. to 1/4" dia. to reduce the fuel and air velocity.
c . Cyclone Tubes
Double walled type 310 stainless steel cyclone tubes were fitted with
provision for steam cooling. Each tube was instrumented in the interspace
with thermocouples.
2. Gasifier Services
a. Flue Gas for Decoking
An additional piping system was added so that the flue gas could be
directly recycled into the gasifier top space through the lid to provide
a means of controlling the carbon burn-out in the ducts.
b . Flue Gas Recycle Cyclone
A cyclone was fitted on the flue gas recycle line to reduce the dust
returned to the blowers.
c . Outlet Gas Clean-up
A knock out chamber followed by a cyclone was built into the main
flue between the boiler and stack to reduce the stack solids emission.
3. Test Specimens
Arrangements were made to obtain and test two alternative materials under
gasifying and controlled burn-out conditions with the aim of determining
their suitability for possible use as cyclone outlet tubes. Silicon nitride,
and self bonded silicon carbide test pieces were fixed in the bifurcated duct
above the outlet from each cyclone.
-------
- 222 -
D. CAFB - OPERATIONS LOG RUN 3
19 11.71 The propane gas burner was started up at 18.00 for the gradual
heating-up process. Some hot gas leakage around the lid but
these were sealed with fireclay.
20.11.71 The heating continued on schedule.
21.11.71 A test was made of the regenerator destributor flow/pressure drop
characteristics.
The gas burner failed and the flame detector was cleaned and
replaced.
A blocked pressure tapping in the regenerator was cleared.
22.11.71 General checks were made on pressure measurement and the bed
pulser control equipment. The regenerator to gasifier pulse
controller gave only a continuous signal instead of an intermittent
one and a wiring fault was detected and rectified.
The three metering pumps were started on kerosene. The No.2 pump
kept cutting out and affecting No.l pump which was on the same
electrical circuit. No.2 pump was isolated and the unit ran on
pumps 1 and 3-
Lime feed was started but the modified direct injection feed
system was inclined to block. The original gravity feed through
the lid system was replaced.
23.11.71 The limestone feed was continued, but there was some difficulty
in maintaining the bed temperature.
The Compton compressor broke its connecting rod bolt and was
stripped down to await new components.
The unit was shut down for a short period to permit a modification
to the electrical supply for the metering pumps to give them
individually fused supplies.
The first stage blower drive pulley became loose and threw a
belt. Pulley was relocated and the blower restarted.
24.11.71 There was still little progress on getting the bed to fluidise.
The kerosene was burning above the bed and modifications to the
fuel injectors did not help in the combustion within the bed.
It was considered that the bed material may have had insufficient
fine particles for fluidisation at available gas velocities. Some
lime was reground and gradually fed into the unit after some of
the bed had been drained off.
25.11.71 Gradually the bed began to fluidise, and reground limestone was
successfully added to build up bed height. The regenerator air
rate was increased but fluidisation and transfer was not good.
The nitrogen system for bed transfer and general inerting
requirements was checked out, and it was discovered that the
suppliers had omitted their daily visit to refill the bulk
nitrogen storage which by now was almost empty. The gasifier was
switched over to air for bleeds and pulsers to conserve nitrogen
for other inerting requirements on the site when nitrogen was
essential.
The Compton compressor was repaired and checked out satisfactorily.
-------
- 22J -
6.11.71 The left hand cyclone had been deliberately plugged at its
lower end where it drains to the regenerator to gasifier transfer
duct, to prevent back flow of material during the bed filling
operation. Now that the bed height was up to its operating
level the plug was withdrawn so that the cyclone fines would be
returned to the gasifier bed.
After various adjustments to pulser pressure and bleeds, the
transfer system operated well together with good fluidisation in
the regenerator.
At 05.30 the liquid nitrogen storage tank blew its safety valve
which failed to reseat and caused loss of the tank contents.
The liquid Np suppliers came and rectified the valves, repiped
the system to ensure that the bursting disc would operate
independently of the safety valve, and refilled the system. The
cause of the trouble was shown to be an incorrect filling procedure
by the tanker used the previous evening. The boiler extension
tube was removed and the boiler rear door closed. Some trials
were made on the main flame pilot. By starting up at low air flow
rates to the burner the pilot was established and could be main-
tained while the air rate was gradually increased.
All external cyclone dust collection drain systems were installed
and sealed to minimise dust emission during collection.
The cyclone cooling steam lines were completed and checked out.
At 19.30 a low cooling water pressure switch shut the plant down.
Examination showed that the soft water make-up flow meter had
seized up and cut off the flow of cooling water. The meter was
by-passed and the supply resumed.
The burner-to-boiler joint was sealed with fireclay and asbestos
rope after the unit and ductwork had reached their operating
conditions.
27.11.71 The analytical sampling system was completed in readiness for
(Day 1) gasification.
At 01.30 gasification started - the main flame ignited first
time - as usual.
The first few hours were spent in getting the unit to settle
down and the transfer rate of bed from the gasifier to regenerator
was regulated to obtain good regeneration.
A boiler fire tube temperature of 1086°C was measured by suction
pyrometer through the boiler rear door.
After 22 hours of gasification the regenerator S0~ level had
reached 9.6$ . The automatic regenerator temperature controller
was cut in which held the temperature steadily within 5°C of the
set point.
28.11.71 A further measurement of the boiler fire tube back end temperature
toy 2) showed an increase to 1139°C over the first 24 hours of operation.
The gasifier bed had gradually been increased to 19" and associated
with this then was some difficulty in controlling the regenerator
temperature - possibly due to a higher rate of bed transfer. The
nitrogen pressure in the pulser system was reduced and this helped
to line out conditions.
The rate of transfer was very sensitive indeed to the relative
pressure between gasifier and regenerator. Irregularities in gas
space pressures produced irregularities in bed transfer which in
turn upset the regenerator temperatures.
The regenerator drain would not run slowly at the required rate
of 15 Its/hour and intermittent manual removal was necessary. The
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- 224 -
drain was rodded out before each removal to clear away a crust in
the entry to the drain.
The silicon nitride test specimen showed some irregular temperature
variations giving values up to 1090°C.
The gas sampling flame was tried out but thee was not sufficient
gas pressure to give a good flame. A Drager tube test on the boiler
flue gas gave a value of 100 ppm for NO .
X
29.11.71 Some plugging was experienced with the limestone bed feed system
(Day 3) and very delicate balancing of the pressures between the gasifier
and limestone hopper was necessary to ensure a steady flow.
A manometer was added to the boiler flue to provide information
on the pressure drop across the boiler water tubes.
Gradually the gasifier gas space pressure had risen to 17" and
with the blower fully open any subsequent pressure increases cut
back on total air and flue gas to the gasifier.
The gasifier distributor has shown a gradual increase in pressure
drop since the start of the run.
At 18.50 the bed was sulphated by controlled oxidation using full
recycle of flue gas and limestone feed to control the temperature.
At the completion of sulphation the bed was slumped and a controlled
carbon burn-out was completed using flue gas recycle and nitrogen
injection into the gasifier top space. The cyclone tubes were
cooled with steam to prevent them overheating. The oxygen concentra-
tion was controlled by air addition to the boiler through the
burner secondary air supply.
The carbon burn-out was completed and gasification resumed at 22.05.
30.11.71 The next few hours were spent in letting the unit settle out but some
(Day 4) irregularity in the gasifier and regenerator pressures produced
fluctuations in the bed transfer which in turn caused variations in
the regenerator temperature.
It had become apparent that the limestone feed had a considerable
quantity of material larger than 1/8". At this stage it was
decided to pass the material through an additional sieve to remove
these larger particles to prevent their accumulation in the lower
part of the gasifier bed.
The gas space pressures had risen to 16.5" and preparations were
made to start another carbon burn-out which commenced with
sulphation of the bed at 21.18 followed by a burn-out. The silicon
carbide test specimen rose in excess of 1200°C whilst the silicon
nitride specimen temperature became erratic.
1.12.71 At 01.20 gasification was resumed and conditions soon steadied out.
(Day 5) Trials on varying the air rate to the cyclone bleeds produced down-
- stream temperature changes particularly on the silicon carbide
specimen. The thermocouple in the silicon nitride sample stopped
functioning reliably.
Conditions steadied out with an optimum pressure differential of
about 5" between gasifier and regenerator which gave a steady
controllable transfer rate.
The limestone feed rate was high at this period showing about 70 Ibs/hr
with only a small increase in bed levels. The corresponding fines
removal rate from the regenerator cyclone was also higher than normal
and possibly this batch of limestone was dustier than usual.
At 08.20 some difficulty was found in draining the regenerator -
the skin across the drain seemed very hard and had to be hammered
to break it away.
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The bed feed became sluggish. Replacing the feed vibrator
controller improved matters.
At 16.30 there was some difficulty with the regenerator temperature
controller. When temperature exceeded the set point by too great
a margin, pulsing stopped and a steady injector gas flow resulted.
This action caused a decrease in solids circulation rather than
the desired increase. The remedy was to increase temperature set
point momentarily to restore pulsing and to increase gasifier to
regenerator pressure slightly which increased maximum solids flow
rate.
Conditions were steady with excellent sulphur removal in the
gasifier and good regeneration but gradually the gasifier pressure
had risen up to 20.5" with corresponding decrease in the air and
flue gas flow rates.
Preparations were made for a further burn-out by the usual technique
of limestone addition and flue gas recycle for sulphation followed
by a controlled carbon burn with the gasifier bed slumped.
The burn-out started at 17.3^ and at 19.16 the unit was shutdown
to permit draining of the left hand cyclone prior to removal o_ the
silicon carbide test specimen located above this cyclone outlet.
It was thought that this hanging specimen could be acting as a
collector for carbon and lime and when removed some of this material
might fall into the cyclone and block the fines return line.
Gasification was resumed at 21.30.
2.12.71 Unit steady after the burn-out apart from some temperature
(Day 6) variations in the gasifier (a). These temperature swings were
accompanied by 30°F variations in the fuel oil preheater temperature.
The burner throat temperature Jumped up to 1200°C for a few hours.
The bed feed vibrator again became sluggish, and it was replaced by
one removed from the regenerator drain. This did hot help much in
improving the rate of bed feed.
At 06.30 the gasifier pressure had again built up to 19" w.g. with
the blower fully open, and another burn-out was necessary. Due
to unreliable vibrator, this burn-out was done without the cooling
aid of limestone feed into the gasifier. Previously the regenerator
air flow had been reduced after the sulphation of the gasifier bed
to about 3 cfm during the burn-out period. On this burn-out,
which started at 06.55> it was decided to leave the regenerator
fluidised but this caused a difficulty in that the air supply came
up the transfer duct into the gasifier and upset the control of
the gasifier during the carbon burning. The air rate was reduced
to 9 cfm and so re-established the differential pressure between
gasifier and regenerator to stop the back flow of air.
In spite of the absence of cooling limestone addition this burn
was controlled with flue gas recycle, nitrogen addition and steam
to the cyclone tubes.
During the burn many red hot particles and sparks were seen
passing into the boiler.
These last two burn-outs were short in duration with small pressure
recovery which indicated that the amount of carbon was probably
small and the pressure rises were most likely due to inorganic
material.
(a) Temperature variations later traced to uneven fuel rate
caused by low tank level.
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- 226 -
A sieve analysis of the feed limestone showed it to have a major
portion in the Il80 to 1400 micron range with a very small quantity
greater than 1400 microns. There were considerable variations in
the limestone from batch to batch.
Whilst still on combusting conditions before resuming gasification,
the first stage blower was run to provide an additional boost of
air with the intent of blowing some of the lime deposits out
of the system.
At 11.05 gasification was resumed, but the gasifier distributor
pressure drop was 23" leaving very little margin for additional
resistance from further duct obstruction. Plans had been put in
hand earlier to rerun some of the pipework so that both blowers
could be used with flue gas recycle, and at 15.20 the unit was shut
down to permit final connections to be made.
During this work the bifurcated duct was rodded out from the rear
access plates on each side and some material fell down into the
cyclones.
At 18.35 the, unit was put back on gasification with both blowers
operating. It was observed that there could be some relationship
between pressure rise on the ducts and high gasifier temperature
and, therefore, the gasifier temperature was lowered to 870 - 880°C.
The cyclone air bleed rates were cut back to a lower value because
of the possibility of the higher temperatures induced by these
bleeds causing the fines to pass through a sticky phase and so
accelerate the rate of pressure rise.
The bed depth was built up slowly but the vibrator was not
completely satisfactory, and this impeded the rate of bed addition.
3.12.71 There was a gradual rise in the gasifier pressure of 1" in 3 hours
(Day 7) but otherwise all ran quite steadily with the new blower system.
At 11.20 some samples were taken from the top of the stack for
particulate emission tests and NO samples taken from the boiler.
About mid-day the SOg level began to rise in the boiler. There had
been a alow bed feed into the gasifier in the preceding hours, and
so some material was withdrawn from the regenerator in order that
fresh bed could be added without increasing the bed depth. After
about 2 hours, the SOp level had dropped back in the boiler.
During the day two further particulate emission samples were taken
from the top of the stack.
The rate of pressure rise was the greatest across the cyclones and a
further short burst of air was added through cyclone probes. There
was some very slight improvement in the pressure drop in this area.
The limestone feed again proved erratic and the surface bed level
fell during the evening with corresponding increases in the boiler
SOp concentration.
The regenerator cyclone had not collected much solids since the
previous burn-out and it appeared that, the gasifier cyclones may
have become blocked when the bifurcated duct was rodded in the
morning.
4.12.71 The gasifier distributor pressure drop continued to rise and the air
(Day 8) bleeds with the fuel were increased to break up any larger bed
particles which could nave segregated to the bottom of the bed and
contributed towards the higher pressure drop. There was not any
significant improvement with this change in fuel air rate.
The increased quantity of material removed from the boiler and its
cyclone suggested that the gasifier cyclones were not effective at
that time.
-In 1pi?t1nn to t.hp nvrOnnp Inlpt. rtnnt.fs was tried for a brief
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- 227 -
period to determine whether this technique was more effective
than air injection but there was no significant change.
The gasifier to regenerator solenoid valve failed, and the
standby valve was switched into circuit.
At 14.55 the bed was sulphated prior to another burn-out. The
burning had not finished after 82 minutes, but because of the drop
in the bed temperature the unit was put onto combusting conditions.
Both the cyclone drains were emptied out, giving 63 Ibs. of fines
from the right hand cyclone and 40 Ibs. of fines with some lumpy
material from the left hand side.
12 71 '^ie k°iler S02 level rose during the period that the limestone
' > feed was shut-off for hopper refilling.
Day 7) iphe regenerator temperature control became erratic for a few
hours apparently caused by pressure fluctuations, which in turn
caused irregular bed transfer.
The boiler tube pressure drop continued to rise. Attempts were
made to drain the various outlets of material but without much
success.
The regenerator drain was found to be blocked at 06.4-5 and all
attempts to clear it were unsuccessful apart from allowing small
quantities of material through. Some material was withdrawn from
the gasifier in order to maintain the fresh bed feed.
During the course of the morning it was decided to try and get the
right hand cyclone fines return line functioning so that the dust
emission to the boiler could be reduced. The unit was sulphated
at 11. 15. When this was complete, the bed was slumped1 whilst the
cyclone was drained and the gasifier to regenerator transfer line
drained off to provide access for rodding out the fines return line.
The silicon nitride test specimen holder situated in the bifurated
duct over the right hand cyclone was removed, and it was possible
to see into the red hot cyclone and observe the rod come up the
fines transfer line.
At 15.28 gasification was resumed but still there was not much dust
collected in the regenerator cyclone.
The back end of the boiler which had previously been bypassing some
hot gas due to the pressure rise across the fire tubes seemed to
improve during the cooldown and behaved reasonably after this
period.
The gasifier bed level gradually rose to 17" and boiler S0p level
remained below 100 ppm although regenerator SO level was slow to
rise after the burn-out.
6-12. 71 The bed feed rate was set to just maintain a steady bed level and
(Day 10) with this slow feed the rate of rise of boiler tube pressure was
steadied.
The oil fuel tank was switched to the full compartment* and the
additional head produced an increase in oil flow rate which in turn
caused the burner in the boiler to go sub-stoichiometric for a few
minutes. Under these reducing conditions the sulphate in the
boiler liberated SOg.
The unit continued to run smoothly and with a bed feed rate around
1.5 wt. CaO/wt of sulphur the removal efficiency was greater than
As the boiler fire tube pressure drop continued to rise very slowly
the air and fuel rates were backed off to compensate.
At 16.30 the unit was shut down at the completion of over 200 hours
gasification.
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- 228 -
E. Examination of the Continuous Gasifier and Auxiliary Equipment
after Run 3.
At the end of the run, at 1630 on Monday, 6th December, the unit
was shut down by stopping the fuel pumps, followed by all blowers as soon
as the main flame had gone out. Not enough air was put into the unit at
this time to sulphate the bed or burn any carbon from the system. The uniti
then sealed and allowed to cool. The intent was to leave the unit with any
deposits in place as they had been at the end of operations. It is apparent,
however, from the trace of the temperature recorder which was left operating,
that about j£ hours after shutdown some air entered the gasifier and
burned off much of the carbon residue. The gasifier gas space thermocouple
which had fallen to 500°C, began rising and reached 570°C. The bed
temperature also stopped falling for a period of about 10 hours as heat
was evolved. It is probable that the air entered around the gasifier lid
and leaks developed during cooling and that a draught was set up through
the boiler and regenerator outlet flues.
1. Gasifier
a. Gasifier Distributor
The nozzles of the distributor had shown an increasing pressure
drop from 9.5?' water gauge to 33.8" during the duration of the test.
Examination showed that all the nozzles had some degree of obstruction
both within the bore and sometimes at the exit as well. 20# of the
holes had at least 23$ of their area obstructed and about 2# were completely
obstructed. The obstruction in the bore was formed by fine material adhering
around the periphery of the bore, and the external obstruction was caused
by a thin flaky deposit attached to the stainless steel nozzle. A
photograph of several nozzles appears in Figure J-36.
The fine material deposited in the holes of the nozzles was carried
in the flue gas recycle and thrown out of the gas stream by changes in
gas direction and turbulence at the nozzles. The flaky material round the
outside of the nozzles could have been formed from fines within the bed
being locally regenerated by the air from the nozzles and so passing through
a sticky phase at the higher temperature and adhering to the nozzle.
The distributor refractory covering was coated mainly around the
periphery with a whitish deposit which in some areas was thick enough to
obstruct the holes in the distributor. In other areas the air from the
nozzles had cut a channel in the deposit which may be seen in Figure J-37»
Apart from the deposits mentioned above,.the refractory concrete
was in good condition with one or two hairline cracks and a few chipped
edges.
b. Gasifier Refractory Concrete
(I)- Surface Condition
Generally the surface finish had not deteriorated very much since
the first room temperature fluidisation trials in which the surface became
quite rough particularly immediately above the bed. The rough surface is
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- 229 -
not uniform around the gasifier faces and could be caused by the method of
pouring and vibrating the concrete which caused a concentration of
large particles near the surface which were then easily removed by
abrasion.
During Run 5 a black flaky deposit was laid down on the lower
gasifier faces, particularly on the short chamfered faces at the corners.
Fig J-37 shows a light coloured zone on the bottom of the gasifier walls
which dips away around nozzles which were blocked and at a regenerator to
gasifier materials return line. This colour change could indicate the
extent of the oxidising zone present in the gasifier with the dips
corresponding to regions without oxygen.
The refractory concrete was cracked down both the gasifier short
sides after the critical firing process, and these two cracks have
increased in width a little during subsequent operations. The worst
crack is in the lowest concrete lift which is around the cold penum. Here the
concrete is put into tension by the expansion of the hot concrete immediately
above it. The cracks in the hot zone will almost close up at operating
temperature but those in the plenum will stay partially open. The gasifier
distributor and plenum is one integral subassembly and hence there
cannot be any leakage of air through this crack to bypass the distributor.
Because of air leakage during cool down, wal^s of the gasifier gas space were
relatively clean.
The lower area of the gasifier was blackened to some extent and the
two cracks in the walls had a carbon and lime deposit around them.
There were some other small cracks on the faces which often
terminated at the Junctions between the various lifts in the concrete.
This limitation of cracking shows the usefulness of this form of construction.
(2) Start-up Burner Quarl
Flame from the start-up burner enters the gasifier through a downward
angled, precast, refractory tube or quarl.
There was a large deposit of lime particles in this quarl, most of
which fell out once the lower portion was moved. Some finer material had
adhered to the walls especially adjacent to the burner itself.
The lower section of the quarl developed a crack during the last run
but the piece was still firmly held in place by the concrete around it.
Gasifier Lid
The concrete slabs on the gasifier lid were generally clean apart
from a build up of carbon and lime around the limestone feed point. A
few fragments of concrete had broken away from the hot face of the
concrete but as these were lying on top of the slumped bed it evidently
happened during cool down or whilst the lid was being lifted out. There
was a considerable quantity of carbon and lime deposited in the annulus
between the lid blocks and the gasifier wall.
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- 230 -
(4) Materials Transfer Ducts
• Gasifier to Regenerator
This duct remained in good condition with no signs of any blockages.
The refractory became a little brown during the course of all the
operating time but
• Regenerator to Gasifier
It was not possible to inspect all the internal bore of this
transfer line but there was a small deposit of agglomerated fines on the
flat rear face of the transfer line. This deposit was covered by a skin
which presumably formed when the material passed through a sticky phase.
The composition of a sample of the deposit is listed in Table J-IV.
The entry to this transfer line in the regenerator wall was
partially obstructed by a build up of fine white material on the
regenerator wall which spanned across some of the opening.
The refractory concrete was generally good with no evidence of
cracking or spa11ing.
c. Gasifier Bed
Figure J-^8 shows the gasifier bed after removal of the lid. The
debris on the surface fell from the lid or the lid joint either during cool-
down or actual removal of the lid. Most of the carbon particles were
trapped between the lid concrete blocks and the gasifier wall.
The bed was uniformly withdrawn by vacuum tube until three dark zones
cou}.d be seen corresponding to the fuel injector outlets. The bed surface
at this level appears in Figure J-39. These darker areas were probably
formed at shut dfawn by oil cracking on the hot slumped bed - the
inequality in the size of the zones could possibly be caused by the
position of each metering pmip in its stroke at shutdown.
There was not any visual evidence of particle size segregation in
the bed but sieve analysis on the lower portion of the bed showed 55.6#
greater than 1400 microns whilst only 26.6$ of the upper bed portion was
greater than 1400 microns.
2. Regenerator
a. Distributor
Figure J-40 is a photograph of the regenerator distributor after
the run.
The distributor was constructed from refractory concrete with
sixteen •£•" diameter holes. After the 200 hour run when the distributor was
removed a piece broke off at the corner. Possibly a crack formed during
the run, but when the distributor was in place it was held tightly together
and would not have bypassed any significant quantity of air. Six of the
-------
Table J-IV
Composition of Solid Samples Removed from Pilot Plant After Run 3
Sample Location
Composition
Deposit Prom Deposit from Annular Deposit Bed by
Regenerator Bottom of Deposit from end No.3
to Gosifier Regenerator from of No.2 Fuel
Transfer Regenerator Fuel
Line Injector
Obstruction
from L.H.
Cyclone
Injector Sollda
Drain
Inner Core Outer Flue das Regenerator Flue Oaa
from Plre Surface Cyclone Cyclone Cyclone
Tube
Deposit
of Plre
Tube
Deposit
Drained
Pines
Hr.
Pines
198jHr.
Pines
Boiler
Drain
Sollda
Loss on Heating
to 900*C, wt %
Total Sulphur wt %
Sulphate Sulphur wt %
Vanadium wt %
Carbon wt %
CaO wt %
(2.39)1 ;
22.84
13. W
0.35
1 1.78
17.95
16.20
0.6?
0.77
20.09
19.16
0.17
67.2
4.15
19-93
17.9*
-
-
1.96
20.38
20.38
0.32
2.3*
20.20
20.20
0.27
6.58
4.01
3.07
0.60
3.18
9.65
7.06
0.25
*.55
3.79
3.01
0.63
3.*1 i
5.86 g
1.93 '
0.83
5*. 06
(a) wt. gain
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holes were plugged with a fine deposit of white powder on the inlet
side and loosely plugged on the outlet end with coarser material.
Examination of the records shows that the pressure drop in the
distributor was about l£" before any bed was added to the gasifier but
rose to 6" when the bed was added suggesting that some obstructions
were formed during the bed addition. During this period the air flow to
the regenerator was kept below 2 cfm to reduce the cooling effect upon
the regenerator refractory. This air rate may not have been adequate to
prevent blockage of some of the holes.
The top face and edges of the distributor showed signs of
abrasion with the removal of the finer material leaving the larger
particles exposed.
b. Regenerator Bore
The bore was partially obstructed by a block of agglomerated fines with
local pockets of coarser material inbedded in it. The obstruction was strong
and showed some layering in its construction as if it had built up, formed
a surface skin, and subsequently had another layer deposited. Figure J-4l
shows a view of this deposit. The analysis, listed in Table J-IV, indicates
that it is largely calcium sulphate. It contains enough vanadium to
suggest that it contains solids from late in the run and was not formed Just
at the run beginning.
Higher in the regenerator there was a further whitish annular deposit
shown in Figure J-42. It was about £ inch thick.and extended about 14 inches
in height. It partially obstructed the entry to the solids transfer
port from regenerator to gasifier. Its composition (Table J-IV) also
corresponds with calcium sulphate. Its lower vanadium content suggests
that it was formed from fines which had not remained in the gasifier long
enough to pick up much vanadium. Above this annular deposit there were
local areas of white deposits in the regenerator and in the refractory lined
ducts leading from the regenerator. Figure J-4} shows these deposits at
the regenerator outlet.
In previous tests there had not been any evidence of deposits
in the outlet ducts. These deposits were most probably held together by
some of the fines in the bed material passing through the sticky phase and
acting as a cement to bind other fines and coarser material.
c. Regenerator Concrete
There were some small cracks in the concrete walls but it is unlikely
that any leakage could have occurred at running conditions. The cold zone
around the plenum had two cracks which had been filled with fire clay prior
to the test run, and these have remained well sealed during the test.
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;>. General Penetrations
The penetrations into the gasifier and regenerator fluid beds for
temperature and pressure measurement were formed of 27$ chromium type kk6
steel which withstood the conditions with some scaling» These components
were not replaced during the total of approximately ^60 hours of gasification
and many hours of combustion.
The pulsers for bed transfer control also fabricated from the same
material were in excellent condition but have been cooled by the nitrogen
passing through them.
The stainless steel fuel injectors and bed feed tubes were
undamaged. The former were cooled at all times by either fuel or
nitrogen, and the latter were set back into the refractory concrete to
minimise heat pick up.
The fuel injectors had black shiny deposits around their exposed
ends, and the centre injector had a 2" long extended shroud of black
material which had built up at its end. This deposit contained 67$ carbon.
The two air injection tubes which passed through the lid and extended in
front of the cyclone inlets were used to blow air into the inlet to raise
the temperature in an effort to reduce carbon deposition. This technique
certainly affected downstream temperatures, but there was no clear evidence
of it being effective in reducing carbon deposition. One trial was made with
steam injection in one tubs. Subsequent examination showed that the tube
end had broken away and dropped into the gasifier bed. These tubes were
both heavily scaled.
k. Gasifier Cyclones
The right hand cyclone was rodded out towards the end of the test
run, and apart from a small deposit near the bottom which was not removed, the
cyclone and its fines return line were quite clear. (Figure J-WO
The left hand cyclone had not been cleared out and was blocked to a
level of about 13" above its drain off point (Figure J-^5)
The material in the cyclone consisted mostly of fine black free flowing
particles in the upper layer with a few larger pieces of agglomerated fine lime
pieces and carbon lying on this layer. As the material was removed the colour
gradually lightened until at the bottom of the cyclone it was quite white.
This accumulated material was prevented from flowing into the open
transfer line by a very hard piece of agglomerated white fine material about
2-|" x 2" x l£" which was lodged across the entry of the fines return tube.
It is not clear how this obstruction formed but it is unlikely to have fallen
from the bifurcated duct because experience shows duct deposits to be less
dense and hard. The analysis listed in Table J-IV corresponds with a
composition of 76.5$ Ca SO^, l8.9# CaO and k.5% CaS.
-------
Th'j cyclone entries, shown in Figure J-U6, unfortunately had been
burned reasonably clean leaving some granular deposits at the bottom of
f/ach duct. The entires to these ducts were radialized after the first run and
made smooth by buttering with a fine refractory paste which stayed in place.
The cyclone gas outlet tubes were redesigned following the melting of
the tubes in the first run. Double walled stainless steel (type 310) tubes
were used. Each tube was divided into 2 segments through which steam could be
parsed. An internal guide was incorporated to direct the steam down one
segment and return up the other. The outlet steam vented to atmosphere.
The tubes were tested to 5 psi prior to installation and after
removal retested. The left hand tuoe was scaled quite heavily but still
was leaktight. The right hand tube was also scaled but leaked around the
lower welded joint between inner and outler skin. Close examination showed a
line crack in the weld zone. There is some differential expansion between the
Inner and outer skins due to the outer skin being half embedded in refractory
concrete and not being in contact with the hot gas. This expansion would
Kive rise to bending and tensile stresses in the joints between inner and
outer skins which together with some corrosion could be the cause of this
failure.
The double walled tubes were both instrumented internally with four
ehromel-aluii.el thermocouples located at various levels within the tubes.
]n the left hand tube 2 of the thermocouples stopped working during the test
but the right hand tube instruments were satisfactory all the time.
^. Bifurcated Duct
The hot gas ducts leading from the two cyclone outlets became
partially obstructed at the junction between the vertical outlet from the
cyclone and the horizontal duct to the burner. This Junction included a
right angled change in direction where material was thrown out by the gas
to accumulate on the corner. Examination were made by removing the end access
covers, and at one of the burn out operations the ducts were rodded clear. It
wa:; during this operation that material fell into the cyclones. Another
contributory cause of deposits in this area was the test specimens of alternative
cyclone outlet tube material which were hung in the bifurcated duct at this
junction. These specimens acted as collectors for material and therefore one
specimen was removed half way through the run whilst the other one was accidently
broken during the redding operation.
There was some accumulation of carbon/lime deposits at the junction
I. . veen the two horizontal ducts prior to their entry into the main burner
section.
The expansion joints between the gasifier and the bifurcated
duct were in good condition without any evidence of cracking or deposition of
material.
fc>. Burner
a. Premix Section
The premix section was constructed from type (310) stainless steel
in the hot zone with ceramic fibre insulation. There was some flaky 1/16"
thick carbon deposits on the bore of the inner hot duct but no evidence of any
deterioration.
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- 235 -
b. Main burner
This burner was constructed from stainless steel in the hot zone
surrounded by an external annulus of cooling air from the secondary air
supply. The thermocouple which was inserted into the hot gas stream gave
erratic readings after 160 hours of gasification, and the protruding
end of the stainless steel clad thermocouple had completely disappeared at
200 hours. There were times during the operation, particularly during
and after a burn out when this temperature apparently went above 1200°C for
some hours probably due to local carbon burning off. During gasification there
is some oxygen available in this zone from the premix air supply.
There was a thin deposit of carbon/lime in the burner hot gas duct
but no evidence of any build up of material in the burner throat.
7. Boiler
a. Fire Tubes
The front end refractory around the burner has not deteriorated during
the test, and the stainless steel deflector plates in the pilot flame area
were in good condition. There was a deposit of fine lime around the area
of the boiler together with deposits in the boiler corrugated fire tube.
At the rear end of the fire tube there were very local white deposits
(Figure J-^7) adhering to the gas tube in random positions. It is assumed
that these were formed from fine material in the flame which went through the
sticky phase and gradually built up at random locations. The deposit was hard
but could be removed fairly easily. The greatest deposits were at the entrance
to the first tube pass where the majority of the tubes were partially obstructed
by a hard thick crusty coating of fines. In some cases the tubes were
completely obstructed. The crust could be broken off fairly easily and left a
cone shaped deposit in many of the tubes which could be easily withdrawn.
The longest plug removed was 15" with the majority less than 6" long.
Figure J-^8 shows the largest of these cores. The core composition
(Table J-IV) corresponded with a high content of calcium sulphate. The vanadium
content is in line with that expected for the average lime to vanadium feed
ratio.
The gas exit ends of the tubes of the final pass did not show any
crusty formation but many of them, particularly in the lower portion had a
deposit of fine material which was easily removed.
Figure J-^9 is a general view of the open boiler backend. The firebrick
arch separating the passes in the boiler showed some flaky deposit of lime, and on
the horizontal section beneath the main outlet there was a deposit of finer
material. The asbestos seal between boiler passes looked reasonable
although there was evidence of gas bypassing near the end of the test run due
to the build up of pressure across the tube passes. The boiler paint work was
discoloured at the lower portion due to this hot gas leaking through the
internal seal and overheating the casing locally.
-------
- 256-
b. Stack and flue gas cyclones
The flue gas cyclone was added to the system before the third run
to reduce particulate emissions and the system worked well. The stack was
clear with a very small deposit of lime particles on the horizontal
section at the bottom of the stack.
The flue gas recycle cjtclone.;was also added to reduce the fines
returned to the blower system. During the run material was drawn off
but some fine material was still passing and causing distributor blockages
in the gasifier-
8. Test Specimens
After the experience in the first gasification run when the stainless
steel cyclone outlet tubes were melted, investigation was made into the use ol
alternative materials. Two refractory materials were chosen for sample tests
by hanging the specimens in the bifurcated duct Just above the cyclone outlet
with thermocouples placed under the specimens.
In this run the temperature during burn out was controlled by
using flue gas recycle together with nitrogen injection to supplement the flue
gas if required. The oxygen level was controlled by passing air into the
system at the burner. This air returned to the top of the gasifier with the
recycled flue gas.
The sample temperatures could not be controlled other than by
gas recycle whilst the stainless steel cyclone tubes were steam cooled as
well.
a.
Silicon Nitride
A closed ended sheath of silicon nitride was fixed above the right
cyclone outlet. The recorded temperatures in this sample were rmther erratic.
Initially the temperature lagged on the main gas stream but then exceeded
the gasifier top space temperature. Much of this was caused by the air
injection into the cyclone inlets and subsequently erratic behaviour arose
during and after carbon burn outs when the indicated temperature rose above
1200°C.
When the bifurcated duct was rodded clear at the second burn out on
day 6, the test specimen was broken and fell into the cyclone. No trace of
fragments was recovered. A similar specimen which was tested in the
second run became very brittle. The resulta from an analysis by the manufacture
show that the lime reacted to form a mixed metal silicate which leached out the
silicon leaving a brittle material.
b. Self Bonded Silicon Carbide
Self bonded silicon carbide is a recently developed dense material
available from the U.K. Atomic Energy Authority. Our test specimen was a f inch
o.d. x \ inch bore open ended tube. This tube was hung above the left hand
cyclone in the bifurcated duct rfith a thermocouple fixed Inside. The specimen
-------
- 237 -
temperature again showed considerable variations with the gasifier top space
temperature - particularly after a carbon burn out operation.
After 104 hours of gasification in Run 3 the specimen was removed
intact to prevent its accumulating lime and carbon and contributing to duct
pressure drop. The tube appeared by visual inspection to be undamaged. It
was returned to the suppliers for more detailed examination. Some strength
tests were conducted on test sections cut from each end of the specimen,
and these were compared with a control specimen. Sections were also cut and
examined under the microscope again comparing with a control specimen. The
manufacturers say that the material is unaffected by its exposure to
gasifier conditions.
9. General Service
a. Boiler
The boiler and its associated cooling system have been trouble
free while maintaining a steady boildr water temperature throughout the
test.
b. Oil Supply
The rate of oil supply to the gasifier became a little irregular
when the oil quantity in the external oil storage tank became lower than
500 gallons. The system acted as though the circulating ring main began
picking up air bubbles at this tank level.
c. Gasifier controls
The control, analytical and recording equipment functioned reliably
apart from some minor electrical troubles.
d. Bed Handling
The alternative method of direct injection of the bed material was
not sufficiently reliable due to the limited time available for its development.
The use of a vibrator for controlling rate of bed feed was unreliable for
consistent input because it was very sensitive to the end constraints
of the pipework. Investigations are in hand for alternative methods of bed
feed control and input.
The vibratory regenerator drain did not function as a steady slow method
of bed removal because the pipe became plugged unless a fast flow of material
was maintained all the time. Investigations are in hand to improve the present
system which in this run relied upon manual draining at intervals.
e. Blowers and Compressors
The gradual rise in the syfctem resistance entailed modifying the
blower pipework so that the flue gas recycle system could be used with both
blowers in series.
-------
f. Limestone
It is not practicable to prepare locally the quantities of
limestone required with the smallest particle size below 800 microns
due to sieve blinding caused by the dampness of the material as received
from the quarry.
-------
- 239 -
F. Computer Programmes for Pilot Plant Data
Several programmes were written in BASIC language for the 0265
time sharing computer system to aid in processing and analysing CAPB
pilot plant data. These programmes are summarised and listed here
for reference.
1. Data Filing Programme
Programme JHTPPF was written to put a number of pilot plant
data items into file for future reference. It was intended that the
process operator would update the files periodically with hourly sets
of data readings. Because of file size limitations in the G-265 Mark I
system it was necessary to string several files together and to build
into the filing programme instructions to keep track of which file
was in use and the location of the last data item filed.
The data are entered with the names D(I) where I goes from 1
to 25. The individual data items have the significance listed in
Table J-V below.
Table J-V
Data Listed by Pilot Plant File Programme
Data Designation Data Item
D (1) Day. Hour
D (2) Gasifier Temperature, °C (Top of Bed)
D (3) Regenerator Temperature, °C (Top of Bed)
D (4) Gasifier Air to Blower, SCFM
D (5) Flue Gas Recycle, SCFM
D (6) Air to Injectors and Start Up Burner, SCFM
D (7) Air to Stone Feed, SCFM
D (8) Pressure in Gasifier Above Bed, in.w.g.
D (9) Pressure at Regenerator Air Meter, in. w.g.
D (10) Pressure at Regenerator Nitrogen Meter, in.w.g.
D (11) Regenerator Air Flow, Meter CFH
D (12) Propane to Boiler Pilot, SCFM
D (13) Oil Rate, Gallons (Imp)/Hr.
D (14) Oil Meter Temperature, °F
D (15) Nitrogen to Regenerator, (Meter CFH)
D (16) Stone Feed Rate, Ib/hr.
D (17) Oxygen in Plenum Gas, Vol %.
D (18) Oxygen in Boiler Flue Gas,-Vol %
D (19) Oxygen in Regenerator Gas, Vol %
D (20) S02 in Boiler Flue Gas, ppm
D (21) S02 in Regenerator Gas, ^ amp Meter
D (22) GOo in Regenerator Gas, ^ amp Meter
D (23) COp in Boiler Gas, ^ amp, Meter
D (24) Spare
D (25) Spare
-------
- 240 -
Programme to File CAFB Pilot Plant Data
Basic Language - G-265 System
JHTPPF
100 REM JHTPPF FILES CAFB PILOT PLANT DATA
110 REM REVISED 17/12/71 TO USE MULTIPLE FILES
120 DATA 7
130 REM NUMBER OF DATA SETS TO BE ENTERED. CHANGE EACH RUN.
140 FILES JHTFIU JHTFI2IJHTFI3UHTFI4J JHTFI5UHTFI 6
150 SET:!,1023
160 READ:l*H
170 RESTORErH
180 READ:H*P ,
190 LET PI = P+25
200 IF Pl>1001 THEN 390
210 LET S = P * 1
220 LET HI = H
230 SETlH.P+1
240 READ N
250 REM NUMBER OF DATA SETS TO BE PROCESSED THIS TIME
260 FOR J = 1 TO N
270 FOR I = 1 TO 25
280 READ D
281 IF I = I THEN 285
282 GO TO 300
285 PRINT D*P*1
290 REM D IS EACH 'OF THE DATA ITEMS TO BE READ IN
300 WRITE:H,D
310 LET P = P + 1
320 NEXT I
330 LET PI = P+25
340 IF PI > 1001 THEN 440
350 NEXT J
360 RESTORE:'H
370 WRITElH, P
380 GO TO 510
390 PRINT "FILE NO "I Hi " IS FULL"
400 LET H = H+l
410 SET:l,1023
420 WRITE:I,H
430 GO TO 170
440 PRINT "FILE NO "*HJ " IS FULL"
450 LET H = H+l
460 SET:1*1023
470 WRITE:!, H
480 LET P = 1
490 SET:H, 2
500 GO TO 350
510 PRINT "END OF DATA SET"
520 PRINT"THE NEXT DATA LOCATION IS FILE MlHJ" LOCATION"! P*l
600 END
-------
- 241 -
2. CAFB Velocities, Ratios, and Compositions
Programme JHT PPP uses data in files to compute various
parameters of interest in CAPB operations. It can be run during the
pilot plant operation to obtain a print out of recent operating
conditions and it can be used to prepare summary tables at the end
of a run. The current version of the programme prepares two tables.
The first lists boiler and regenerator gas compositions computed
from the analyser meter readings as well as sulphur removal
efficiency in the gasifier and selectivity of CaS oxidation to
CaO in the regenerator. These values are computed from feed rates
and gas compositions. The second table lists gasifier superficial
air velocity, the air/fuel ratio (percent of stoichiometric), the
calcium/sulphur feed ratio, and regenerator SOg output as pounds of
sulphur per hour and as percent of sulphur fed;
-------
-242 -
Programme to Compute CAFB Pilot Plant Gas Compositions,
Efficiencies, Feed Ratios, and Regenerator Performance.
Basic Language - g26j) System
JHTPPP
100 REM JHTPPP PROGRAM CALCULATES CAFB RESULTS FROM DATA. IN FILES
110 REM REVISED 17/12/71 TO USE MULTIPLE FILES
120 DATA 86* 2.48
130 REM PCT CARBON AND SULFUR IN FUEL OIL
140 DATA 237
150 REM NUMBER OF DATA SETS TO BE ENTERED* CHANGE EACH RUN.
160 FILES JHTFIIJ JHTFI2* JHTFI3* JHTFI 4IJHTFI 5UHTFI 6
170 DIM D(25)
180 DIM A(10>
190 READ Cl*C2
200 READ N
210 LET HI = 1
220 PRINT USING 1560
230 PRINT
240 PRINT
250 PRINT USING 1570
260 PRINT USING 1580
270 PRINT USING 1590
280 PRINT USING 1600
290 PRINT
300 PRINT
310 LET K = 0
320 LET H = HI
330 SET:H, 2
340 FOR L = 1 TO N
350 FOR I = I TO 25
360 READrH, D(I>
370 NEXT I
380 REM COMPUTE GAS COMPOSITIONS
390 REM REGENERATOR C02
400 LET Z2 = (D(22)/I0> - 1
410 LET A2 = 1.1608245E-03 +1;1574973*Z2-.27330138*Z2*2
420 LET A2 = A2+.075628236*Z2t3-.00837005*Z2t4+.00036410274*Z2t5
430 REM REGENERATOR S02
440 LET Z3 = (D(21)/10)-l
450 LET A3=.012783171*3.1254254*Z3-•63157951*Z3f2
460LET A3 = A3*.14834157*Z3t3-.015486024*Z3t4+6.1794915E-04*Z3»5
470 REM BOILER C02
480LET Z4 = (D(23>/10)-1
490 LET A4=.014181778*2.321528*Z4-.51096758*Z4t2
S00LETA4=A4+.I3242664*Z4t3-.0 I3988354*Z4r4 + 6.0512861E-04*Z4t5
510 LET F = D(13)*9.15
520 LET Bl = (I200/32)*C2*F
530 LET B2 = Cl*F/I00
540 LET B3 = 6*0(12)
550 LET 84 = Bl*A4/(B2+B3>
560 REM B4 = PPM S02 IN BOILER GAS IF NONE REMOVED
570 LET A<1) = 100 *(1-D(20>/B4)
580 LET A(2) = A4
590 LET A(3) = 0(20)
-------
- 24} -
Programme to Compute CAFB Pilot Plant Gas Compositions,
Efficiencies, Feed Ratios, and Regenerator Performance.
Basic Language G565 System
JHTPPP CONTINUED
600 LET Yl = D<19>/!00
610 LET Y2 = A2/100
620 LET Y3 = A3/100
630 LET 0(11) = DC!1)/60
640 LET A = DC1 1 >*<406.9+D<9>>/406.9
650 LET DOS) = D<15)/60
660 LET Nl= D(|5)*(406.9+D<10)>/406.9
670 LET XI = .2I*A/(A+N1)
680 LET G = (1-X1)/(l-Yl-Y2-Y3)
690 LET C = ((X1/G)-YI-Y^!-C3/2)*Y3)/«2*Y3)
700 LET A(4) = 100*<1-
760 LET A(10) = Y3*G**A<9>*A(2)»A<3>»AC5>*A(6)»A<7>,A(I)*A(4>
850 LET K = K+l
860 IF K > 40 THEN 880
870 GO TO 1010
880 FOR P2 = 1 TO 20
890 PRINT
900 NEXT P2
910 LET K=0
920 PRINT USING 1560
930 PRINT
940 PRINT
950 PRINT USING 1570
960 PRINT USING 1580
970 PRINT USING 1590
980 PRINT USING 1600
990 PRINT
1000 PRINT
1010 IF LOC(H) > 1001 THEN 1030
1020 GO TO 1050
1030 LET H = H+l
1040 SETSH* 2
1050 NEXT L
1060 FOR P2 = I TO 20
1870 PRINT
1080 NEXT P2
-------
- 244 -
Programme to Compute CAFB Pilot Plant Gas Compositions,
Efficiencies, Peed Ratios, and Regenerator Performance.
Basic Language - GL\6 System
JHTPPP CONTINUED
1 100
1 1 10
1 120
1 130
1 140
1 150
1 160
I 170
I 180
1 190
1200
1210
1220
1230
1240
1250
1260
1270
1280
1290
1300
1310
1320
1330
1340
1350
1360
1370
1380
1390
1 400
1410
1420
1430
1440
1450
1460
1470
1480
1490
1500
1510
1520
1530
1540
1550
1560:
1570:
1580:
1590:
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
LET F
LET
LET
LET
LET
LET
USING
USING
1630
1640
D(I3) * 9.15
Fl = (60*100+29>/(360*13.92>
F2=4.75886E-03
F3=30.7489/C2
K=0
HI
H =
SET:H, 2
FOR L = 1 TO
FOR I = 1 TO
READ:H, 0(1)
NEXT I
LET
LET
LET
LET
N
25
AS = D(4) *CD(I7)/21)*D(5)+ D<6>
F = -D(13> * 9.15
A(l ) = (F1*A5/F>
A6 = D(4> + >D<5) * D(6)
LET
LET
406^9)
1370
1 TO 20
A<2) = F2*A6*CD<2>*273>/(D<8>+
A(3) = F3*D(16>/F
PRINT USING 1650,0(1>,A(2>>A(1>,A(3)*D(24>,D(25>
LET K = K+l
IF K>40 THEN
GO TO 1470
FOR P2 =
PRINT
NEXT P2
LET K=0
PRINT USING 1620
PRINT
PRINT
PRINT
PRINT
PRINT
THEN 1490
USING
USING
1630
1640
IF LOC(H) >
GO TO 1510
LET H = H+l
SET:H» 2
NEXT L
FOR P2 = 1
PRINT
NEXT P2
PRINT "END
TIME
1001
TO 5
OF DATA
CAFB
BOILER
02
VOL
SET"
GAS COMPOSITIONS AND EFFICIENCIES
GAS COMP REGENERATOR GAS COMP
C02 S02 02 C02 _S02 SULPHUR RC6EN
VQL PPM ,VOL VOL VOL REMOVAL CAO/tt
-------
- 245 -
Programme to Compute CAFB Pilot Plant Gas Compositions,
Efficiencies, Feed Ratios, and Regenerator Performance.
Basic Language - G265 System
JHTPPP CONTINUED
1600: DAY.HOUR PCT PCT PCT PCT PCT PERCENT PERCEIB
1610: 000.0000 000.0 000.0 00000 000.0 000.0 000.0 0000.0 0000.0
1620: CAFB FEED RATIOS AND REGENERATOR SULPHUR
1630:DAY.HOUR VELOCITY AIR/FUEL CA/S RATIO REGENERATOR SULPHUR
1640: FT/SEC PCT STOIC MOL/MOL LB/HR PCT OF FED
1650?4'4''J'*d'^^'* 0 0 • 0 * W w w + w w * • W * w w it + w w
1660 END
-------
- 246 -
5. Circulation Rates of Solids
Programme JHT PPW was prepared to estimate the circulation
rate of solids from gasifier to regenerator on the basis of a heat
balance using measured temperatures and gas compositions.
Prom gas compositions and flow rates the programme computes
inlet and outlet gas enthalpies. The amount of CaS oxidised to
CaO + SO is determined from S02 production and used to estimate heat
produced by this reaction. The measured C02 concentrations allows
computation of heat release by carbon burning. These compositions,
together with oxygen disappearance allow calculation of the amount of
CaS converted to CaSO^ and the heat released by that reaction. Heat
losses through the regenerator wall are estimated from wall temperature.
The programme then computes solids circulation rate by the equation,
Inert Solids Rate = ^Heat in - ^Heat out
Enthalpy of Solids out - Enthalpy of Solids in
\Entering Gas Mass x Enthalpy
Entering SulphupfCarbon Mass
Heat of CaS Oxidation to CaO-»
Heat of Carbon oxidation to CO,
of Heat in =1 Entering Sulphuw-Carbon Mass x Enthalpy
/ Heat of CaS Oxidation to CaO+CaSO,.
L?
of Heat out = V~Heat transferred through walls
Exit gas Mass x Enthalpy
Table J-VI below lists the nomenclature of principal variables
in the programme.
Since the programme assumes a steady state heat balance and
makes no allowance for heat accumulation in the reactor walls or bed
itself, results computed during periods of rapidly changing regenerator
temperature are not valid. Where solids samples were available,
circulation rates were calculated from sulphur balance measurements
and compared with the heat balance estimates in Figure J- 50.
Reasonable agreement was obtained.
-------
- 24? -
TABLE J-VI
Nomenclature in Programme to Compute Solids
Circulation Rates by Heat Balance on Regenerator
Variable
S> T
750)
zi
Z2
Z4
Z5
Z6
Z7
Z8
Z9
Tl
T2
T4
T
H2
H3
H4
H5
H6
H7
H8
H9
Ml
M2
M4
M5
M6
M7
M8
M9
Ql
Q2
Wl
W2
W^
Meaning
atement
Off Gas Mol/hr
Outlet C00 "
II QQ^ II
It Q 2 „
N2
CaSO,
Inlet CaS "
•" Air Ib/hr
" N "
Gasifier Temp Deg K
Regenerator Temp " "
Gasifier Temp Deg F
Regenerator Temp " "
Inlet Air Temp " "
Inlet Air Enthalpy Btu/lb
n N ii n
Outlet L " "
II Q^ « II
It /-.H II II
ii SO " "
Inlet CaO2 " "
Outlet CaO " "
Outlet CaSOj, Enthalpy Btu/lb
Inlet Carbon " "
Heat Released by CaS > CaSO^
TT J. 1 .. . — J 1 /"!_ O X- /"i-»/\ '
neao rexeaseu oy i/ao > v^au
Heat in with Air + Np
" " " CaS + Carbon
Heat out with Off Gas
Heat released by C > CO
Heat Loss from Walls
Enthalpy Gain of Circulating CaO
Heat inputs except from
Circulating CaO
Heat outputs except from
Circulating CaO
Heat Available to heat
Circulating CaO
CaO Circulation
Solids Entering Regenerator
Solids Exit Regenerator
Btu/hr
"
n
11
n
n
n
Btu/lb
n
n
n
Ib/hr
II
II
-------
JHTPPW
- 248 -
Programme to Compute CAFB Pilot Plant Solids
Circulation Rates from Heat Balance
Basic Language G-265 System
100
1 10
120
130
150
160
1 70
180
190
200
210
220
230
240
250
260
270
280
290
300
310
320
J30
340
350
360
370
380
390
400
410
420
430
440
450
460
470
480
490
500
510
520
530
DATA 86*2.48
DATA 250
DATA -18-21464,.24642255*-2.588I903E-07*7.69R1252E-09
DATA -1.5706847E-12*-15.982317*.25671377, -2.4932232E-05
DATA 3.0113785E-08*-6.942916£-I2*-12.452016*.20982979
DATA 2.8205499E-05*-3.6796678E-09*-3.4144662E-13*-12.87373
DATA . 19510153.7.2269052E-05
DATA -2.554598E-08*4.2340585E-12*-9.4090856*.14176497
DATA 3.416l995E-05*-7.7396955E-09*9.5608306E-13*-l11.5*.322
DATA 7.79E-05,3470,-90*.2455*l.45E-04,2075*- 193*•At 1.96E-04
DATA 17500*396000,197000*72*12*44*64*32*28*108000*2.75*1162*136
FILES JHTFIUJMTFI2JJHTFI3JJHTFI4JJHTFI5* JHTFI6
DIM 0(25)
DIM A(50)
READ C1*C2
READ N
FOK I = I TO 49
HEAD A(I)
NEXT I
PRINT USING 1350
PRINT
PRINT
PRINT USING 1360
PRINT USING 1370
PRINT USING 1380
PRINT USING 1390
PRINT
PRINT
LET HI = 1
LET K
LET
SET;
FOR
FOR
= 0
H = HI
H* 2
L = 1 TO N
I = 1 TO 25
READrH* D(I)
NEXT I
KEM COMPUTE GAS COMPOSITIONS \
REM REGENERATOR C02 ^
LET Z2 = (D(22)/10) - 1
LET A2 = 1.1608245E-03 +1.1574973*Z2-.27330138*Z2t2
LET A2 = A2+.075628236*Z2t3-.00837005*Z2*4+.00036410274*Z2»5
REM REGENERATOR
LET Z3 = (D(21 )/
S02
LET A3=.0 12783 171*3. 1 254254*Z3- . 631 57951 *Z3t 2
540LET A3 = A3* . 1 4834 1 57*Z3 t 3- . 01 5486024*Z3r 4*6. 1 79491 5E-04*Z3t 5
550 REM BOILER C02
560LET Z4 = (D(23)/10)-l
570 LET A4=. 014181 778+2. 321528*Z4-.5l096758*Z4t2
580LETA4=A4+. 1 3242664*Z4t 3- .0 1 3988354*Z4» 4+6.051 2861 E-04*Z4t 5
5PPI I FT F = 00 T**Q. » «^
-------
- 249 -
Programme to Compute CAFB Pilot Plant Solids
Circulation Rates from Heat Balance
Basic Language G-265 System
JHTPPW CONTINUED
600
610
620
630
640
650
660
670
680
690
700
710
720
730
740
750
760
770
780
790
800
810
820
830
840
850
860
870
880
890
900
910
920
930
940
950
960
970
980
990
I
1010
1020
1030
1040
1050
1060
1070
1
LET
LET
LET
LET
REM
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
LET
Bl = C 1200/32>*C2*F
B2 = Ct *F/100
B3 = 6*U(12)
B4 = Bl *A4/
BA = PPM S02 IN BOILER GAS IF NONE REMOVED
Yl = D(19>/100
Y2 = A2/100
Y3 = A3/100
DC I 1 > = D< 1 1 )/60
A = D(l 1 >*<406.9«-D<9»/406.9
DC15) = D< !5)/60
Nl= 0(15>*<406.9+D( 10) )/406.9
XI = .21 *A/(A+N1 )
G = C 1-X1 )/(l-Yl -Y2-Y3)
C = ( -Yl-Y2-<3/2)*Y3)/C2*Y3>
Zl
Z2
Z3
Z4
Z5
Z6
Z7
Z8
Z9
Tl
T2
T3
T4
T =
G*(A+N1 >*60/387
Y2*Z1
Y3*Z1
Y1*Z1
Zl -Z2-Z3-Z4
Z3*C
Z3 + Z6
60 *A*29/387
60*Nl*28/387
D(2>+273
D(3>+273
D(2)*l .8*32
D(3)*l .8*32
80
H2=A(1
H3=A<6>+A(7)*T+A(8>*Tr2+A(9>*Tt3+AC10)*Tt4
H4=A(6>+A(7>*T4+AC8>*T4T2+A(9:>*T4t3+A(10>*T4T4
H5=A(11>+A*T4t2+AC14)*T4t3+A(15)*T4tA
H6 sA(16)+ACI7>*T4+A(18>*T4t2+A<19>*T4t3+A(20>*T4t4
H7=AC21)+AC22>*T4+AC23>*T4t2+A<24)*T4t3+A<25)*T4fA
H8=A ( 2 6 ) + A ( 2 7 ) * T 1 + A ( 28 > *T 1
H9=A(26)*A(27)*T2*A(28)*T2t2*A(29)/T2
M | =A ( 30 ) + A ( 3 1 ) * T2 +A ( 32 > * T2 t 2 + A ( 33 ) / T2
LET M2=AC34>+A(35>*T2+AC36)*T2t2+A<37>/T2
LET M3=Z6*AC38>
LET M4=Z3*A(39)
LET M5 = Z8*H2*Z9*H3
LET M7=Z2*A(42)*H6+Z3*A(43)*H7+Z4*A(44)*H5*Z5*A(45)*H4
LET M8=Z2*A<46>
LET M9=A(47)*(T4-A(48>>
LET Ql = H9-H8
LET 02 = M3*M4
LET Q3 = M7+M9
LET QA = Q2-Q3
-------
- 250 -
Programme to Compute CAFE Pilot Plant Solids
Circulation Rates from Heat Balance
Basic Language 0-265 3y«t«m
JHTPPW CONTINUED
1100
1 1 10
1120
1 130
I 140
1 150
1160
1170
1180
1190
1200
1210
1220
1230
1240
1250
1260
1270
1280
1290
1300
1310
1320
1330
1340
1350:
13601
1370:
1380:
1390:
1400:
1410
LET Wl = Q4/QI
LET W2»W1*A(41>*Z2+A<40>*Z7
LET W3 «W1+A(49>*Z6
PRINT USING 1400*0(1), W1,W2»W3
LET K=K+1
IF K>40 THEN I 170
60 TO 1300
FOR P2 « 1 TO 20
PRINT
NEXT P2
LET K«0
PRINT USING 1350
PRINT
PRINT
PRINT USING 1360
PRINT USING 1370
PRINT USING 1380
PRINT USING 1390
PRINT
PRINT
IF LOC(H) >
GO TO 1340
LET H = H+l
SET:H» 2
NEXT L
CAFB SOLIDS CIRCULATION RATES
FLOW OF HOT SOLIDS IN POUNDS PER HOUR
1001 THEN 1320
TIME
DAY.HOUR
#**.**##
CAO CIRCULATION
END
*****
TOTAL SOLIDS FLOW
FROM TO
SASIFIER GASIFIER
-***** *****
-------
- 251 -
4. Support Programmes
Several other simple programmes were written to assist in
data editing, listing, and handling. In addition considerable
use has been made of a 0.265 system programme for data plotting
on the teletype. This latter programme was used in preparing the
hourly log plots of Run 3 data and results.
Support programmes include the following:
a. JHTTEM - Lists selected groups of original
data from files.
b. JUT 99 - Initialises files for starting new data
additions at first location.
c. JHT RD - Prints initial data items (Date.Hour)
of each data set.
d. JHT COR - Permits correcting individual data items
in files.
-------
- 252 -
G. Data Tables, CAPB Pilot Plant, Run 3
-------
- 253 -
TABLE J-VII
Sample
Day
1 -
1 -
2 -
2 -
2 -
3 -
3 -
3 -
3 -
3 -
S :
5 -
5 -
5 -
7 -
7 -
7 -
8 -
8 -
9 -
9 -
9 -
10 -
10 -
10 -
10 -
Time
Hour
1730
1730
1900
1900
1900
0230
0230
0230
0430
0715
0730
1430
0800
0800
0830
1130
1130
1130
0730
0730
0100
0100
2130
0530
0530
1230
1230
Sample
Location
Regenerator
Gasifier Upper
Regenerator
Gasifier Upper
Gasifier Lower
Regenerator
Gasifier Upper
Gasifier Lower
Gasifier Lower
Gasifier Upper
Regenerator
Regenerator
Gasifier Upper
Gasifier Lower
Regenerator
Regenerator
Gasifier Upper
Gasifier Lower
Gasifier Upper
Gasifier Lower
Gasifier Upper
Gasifier Lower
Regenerator
Gasifier Upper
Gasifier Lower
Gasifier Upper
Gasifier Lower
i of CAPB
Sulphur
3.40
7.61
3.24
5.45
5.31
3.17
6.15
5.93
5.06
5.08
3.05
3.16
3.91
3.86
2.71
2.37
4.15
4.10
3.65
3.86
4.81
5.47
3.50
5.35
4.96
5.50
5.17
Solids Samples
Run 3
Composition, wt %
Sulphate
Sulphur
0.95
0.09
0.57
0.19
0.25
0.64
0.15
0.20
0.14
0.13
0.65
0.88
0.19
0.26
0.84
1.18
0.23
0.27
0.19
0.20
0.32
0.35
2.55
0.24
0.26
0.38
0.30
Carbon
0.0
0.13
0.01
0.36
0.26
0.0
0.31
0.27
0.37
0.45
0.0
0.04
0.06
0.05
0.0
0.02
0.15
0.12
0.22
0.23
0.01
0.01
0.10
0.0
0.18
0.12
0.09
Vanadium
0.52
0.62
0.65
0.78
0.92
0.93
0.72
0.82
1.15
1.15
0.92
0.95
1.1
1.5
0.7
1.05
1.5
1.5
1.4
1.45
1.4
0.97
1.15
1.2
1.8
1.4
1.7
-------
- 254 -
TABLE J-VIII
fnrnnosition of CAFB G^qif^er Product 088
Samples, Run 3
Sample Time
Day Hour 4-0500
Composition, vol %
°2 °
N2 54.2
co2 10.3
CO 13.0
Hg 10.2
CH^ 5.5
P 2i *
CgHg 0.4
C^Hg 0.02
Ctl f\ f\r\
•y*o \)*\J*f
CL 0.08
7-H50
0
57.8
9.4
9.0
9.7
6.0
4.1
0.2
<0.01
0.2
0.3
8-2230
0
63.8
11.4
6.5
7.0
4.8
3.7
0.3
<0.01
0.1
0.09
-------
- 255 -
TABUS J-DC
Sheet 1 of 20
Raw Data from Run 3
Day:Hour
Gasifier
Temp.
°C
Regenerator
Temp
•c
Gasifier
Air to Flue Gas Air to Fuel Injectors
Blower Recycle + Start-up Burner
(SCFM) (SCFM) (SCFM)
.0230
.0330
.0430
.0530
.0630
.0730
.0830
.0930
.1030
.1130
.1230
.1330
.1430
.1530
.1630
• 1730
.1830
• 1930
.2030
.2130
.2230
.2330
2.0030
2.0130
2.0230
2.0330
2.0430
2.0530
2.0630
2.0730
2.0830
2.0930
2.1030
2*1130
2.1230
2*1330
2*1430
2*1530
2*1630
2*1730
2*1830
875
880
900
889
880
881
878
870
872
890
890
900
905
885
882
888
888
898
880
875
860
882
870
908
868
872
875
867
862
869
870
875
878
868
872
878
875
882
882
886
882
895
975
980
101 1
1085
1 108
1088
1088
980
1003
1050
1040
1032
1052
1060
1105
1087
1085
1082
1063
1100
1088
1087
1090
1080
1085
1088
1044
1052
1032
1080
1085
1088
1095
1100
1098
1070
1078
1088
1095
1096
170
170
220
210
200
200
200
200
200
200
210
190
205
220
220
215
215
205
210
210
210
205
210
220
215
215
215
215
215
215
215
215
215
215
215
215
215
215
215
210
205
70
65
55
65
60
60
60
63
63
68
68
71
71
51
60
54
54
60
60
55
55
50
50
50
50
50
50
50
50
50
50
50
50
50
50
50
50
50
50
50
50
18
18
18
19
19
18
.0
.0
• 0
.3
.3
.0
.0
.0
18.
18.
18.0
18.0
18.0
18.0
1R.0
15*8
15*9
15*9
15*9
15.6
15.9
15*6
15.6
18*0
18*0
18*0
18*0
18*0
18*0
15*0
15*0
15.0
15*0
15*0
15.0
15.0
15.0
15.0
15.0
15*0
15*0
15.0
15.0
-------
- 256.-
TABIE J-EC
Raw Data from Run 3
Day:Hour
Gasifier
Temp.
°C
Regenerator
Temp
°C
Air to
Blower
(SCFM)
Oaslfier
Flue Gas
Recycle
(SCPM)
Air to Fuel Inje<
•f Start-up Burnei
(SCFM)
2*1930
2.2030
2.2130
2.2230
2.2330
3.0030
3.0130
3.0230
3.0330
3.0430
3.0530
3.0630
3.0730
3.0830
3.0930
3.1030
3.1 130
3.1230
3.1330
3.1430
3.1630
3.1730
3.230S
4.0005
4.0130
4.0230
4.0330
4.0430
4.0530
4.0630
.0730
.0830
.0930
• 1030
• I 130
.1230
4.1330
4. 1430
4.1530
4.1630
4.1730
888
898
902
910
900
870
880
870
860
885
865
870
870
870
865
865
870
877
885
880
898
909
880
880
880
890
886
880
910
899
885
900
910
890
888
915
910
910
912
910
915
1098
1100
1098
1097
1098
1098
1098
1090
1095
1100
1095
1098
1098
1098
1085
1095
1095
1095
1100
1095
1098
1082
1090
1092
1090
1030
1020
1025
1045
1050
1095
1065
1084
1092
1039
1085
1092
1098
1095
1081
1095
205
205
200
210
200
195
200
210
200
200
200
200
200
190
185
190
195
190
190
190
190
180
190
210
200
210
210
210
210
200
195
200
210
210
205
210
205
205
200
200
200
50
50
50
50
50
50
50
50
50
50
50
50
50
50
50
50
50
35
35
40
45
45
50
55
65
60
60
55
55
60
60
60
55
55
55
55
55
55
55
50
52
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
13.8
13.8
13*8
20.1
31.5
2J.0
21.0
33.0
33.0
33.0
33.0
33.0
30.0
30.0
30.0
30.0
30.0
25.6
19.3
18.9
19.1
20.4
20.8
18*4
-------
- 257, -
TABIE J-1X
Raw Data from Run 3
Sheet 3 of 20
Day:Hour
Gasifier
Temp.
°C
Regenerator
Temp
°C
Gasifier
Air to Flue Gas Air to Fuel Injectors
Blower Recycle + Start-up Burner
(SCFM) (SCFM) (SCFM)
4.1830
4.1930
5.0130
5.0230
5.0330
5.0430
5.0530
5.0630
5.0730
5.0930
5.1030
5.1130
5.1230
5.1330
5.1430
5.1530
5.1630
5.1730
5.2200
5.2230
5.2330
6.0030
6.0130
6.0230
6.0330
6.0430
6.0530
6.0630
6.1130
6.1230
6.1330
6.1430
6.1500
6.1830
6.1930
6.2030
6.2130
6.2230
6.2330
7.0030
7.0130
898
902
900
895
900
890
885
885
910
890
920
928
920
920
920
905
880
860
925
927
855
920
885
895
900
875
883
882
961
-
_
_
913
_
_
883
860
882
900
883
883
1098
1094
980
1010
1070
1090
1085
1090
1095
1092
1095
1080
1098
1098
1094
1089
1 130
1060
1015
1084
1010
1058
1095
1095
1092
1095
1098
1098
780
-
_
_
1 104
_
_
1088
1088
1090
1098
1095
1100
195
200
210
205
200
200
200
200
200
190
210
200
195
190
185
180
150
150
190
175
195
195
180
160
165
165
165
160
200
210
210
165
185
240
220
190
190
205
205
210
210
55
50
60
60
60
60
60
60
60
60
50
50
55
55
55
60
75
70
70
75
54
40
40
50
55
55
55
60
65
55
50
50
50
60
70
85
92
80
80
77
77
18.6
17.9
21.0
21.0
18.0
18.0
18.0
18.0
18*0
18.0
18.0
18.0
18*0
15.0
15.0
15.0
15.0
15.9
18.0
16.5
15.0
18.0
18.0
18.0
18.0
18.0
18.0
14.5
14.5
11.2
11.5
11.4
19.0
18.0
18.0
18.0
18*0
18*0
18.0
18*0
-------
- 258 -
TABLE J-IX
Raw Data from Run 3
Sheet 4 n
Day:Hour
Gasifier
Temp.
°C
Regenerator
Temp
°C
Air to
Blower
(SCFM)
Gasifier
Flue Gas
Recycle
(SCFM)
Air to Fuel Injc;
+ Start-up Burne
(SCFM)
7.0230
7.0330
7.0430
7.0530
7.0630
7.0730
7.0830
7.0930
7. 030
7. 130
7. 230
7. 330
7- 430
7. 530
7. 630
7. 730
7. 830
7. 930
7.2030
7.21 30
7.2230
7.2330
R.0030
8.01 30
8.0230
8.0330
8.0430
8.0530
8.0630
8.0730
8., 3 8 30
8.0930
8. 1 1 30
B. 1230
8. 1 330
8. 1 430
8.2030
8.21 30
8.2230
8.2330
9.0030
890
888
883
901
880
885
870
870
878
880
897
888
888
887
890
884
-
885
875
872
867
868
863
858
874
862
864
855
863
863
874
882
880
872
892
863
928
882
882
870
875
1 100
1 100
1098
1098
1098
1 100
1 100
1098
1099
1093
1090
1092
1090
1089
1090
1091
-
1090
1092
1090
1097
1 100
1 108
1 100
1 100
1 100
1 103
1090
1070
1070
1081
1068
1065
1068
970
1062
870
1055
1058
1068
1070
210
210
210
210
210
205
205
200
200
200
200
200
195
195
195
190
190
190
190
180
180
180
180
180
180
190
190
190
190
190
190
190
190
190
190
190
180
200
200
180
180
77
79
79
77
77
77
77
75
77
78
77
80
125
85
85
87
90
90
95
95
95
95
95
95
90
90
90
90
90
85
87
87
80
82
82
82
120
100
1 10
100
1 10
18.0
18*0
18*0
18.0
18*0
18.0
18.0
17.0
1 7.2
18.0
18>0
1 8*0
18.0
18.0
1 7.8
1 7.8
1 7.8
1 7.8
16*2
1 6.3
1 6. 5
1 6.2
20. 8
20.8
18*9
1 8* 4
1 9« 2
1 7*8
17.9
18*0
19.0
1 s * v
1 9*0
I 9.0
19.0
1 f • %J
19.0
1 9. fl
1 f • V
18.0
1 R. 0
1 O • V
ta.a
1 O • V
1 a. a
1 O • f
18.0
-------
- 259 -
TABLE J-EC
Raw Data from Run 3
Sheet 5 of 20
DayrHour
Gaslfler
Temp.
°C
Regenerator
Temp
°C
Gaslfier
Air to Flue Gas Air to Fuel Injectors
Blower Recycle + Start-up Burner
(SCFM) (SCFM) (SCFM)
9.0130
9.0230
.0330
.0430
.0530
.0630
.0730
.0830
.0930
.1030
.1130
.1530
.1630
.1730
.1830
.1930
.2030
.2130
.2230
.2330
0.0030
0.0130
0.0230
0.0330
0.0430
0.0530
0.0630
0.0730
0.0830
0.0930
0.1030
0*1130
0.1230
0*1330
0*1430
0*1530
8.1630
878
882
878
882
880
885
868
868
878
878
868
880
-
880
872
888
860
858
881
869
869
866
871
877
881
876
889
901
885
875
875
882
864
868
868
872
850
1070
1070
1068
1050
1065
1065
1055
1058
1068
1065
1068
720
-
1062
1068
1060
1070
1087
1062
1061
1063
1068
1078
1085
1069
1069
1068
1066
1072
1068
1068
1070
1069
1068
1070
1070
1068
200
195
200
195
200
185
185
190
190
190
185
190
200
190
190
195
190
200
200
180
190
190
190
190
-
190
190
200
200
180
190
190
190
185
185
185
140
90
90
100
100
100
100
1 15
120
110
110
107
80
85
95
95
88
90
94
92
95
97
97
97
95
-
95
95
90
85
90
95
89
90
90
90
90
105
18.0
18.0
18*0
18.0
18.0
18.0
18*0
18.0
17.8
17.0
17.9
17.6
18.0
18.0
18.0
18.0
18*0
18.0
18.0
18.0
18.0
18*0
18*0
18*0
18.0
18.0
18*0
18*0
18*0
18.0
18.0
18*0
18.0
18.0
18*0
18.0
-------
- 260 -
TABUS J-DC
Raw Data from Run 3
Day:Hour Gaslfier Regenerator Np Meter Regenerator Propane Oil Rate
Pressure Gas Meter Pressure Air to Pilot
Pressure
(inch.W.G) (inches W.G) (Inches W,0) (3CFH) (SCTM) (gali/hr
f
f
£
f
£.
f
f
r
f
f
(
f
f
c
£
£
.0230
.3330
.3430
.0530
.3630
.3730
.0830
.0939
. 103'4
• 1 130
. 1230
. 1330
• 1430
• 1530
.1630
. 1730
• 1830
. 1930
.2030
.2130
.2230
.2330
J.0033
>.0130
>.0230
>.0330
>.0430
J.0530
J.0630
>.0730
>.3830
'.0930
>. 030
>. 13:1
>. 230
>. 330
>. -430
>. 533
>. 630
>. 730
>. 830
9.00
9.00
1 .03
1 .00
1 .30
1 .25
1 .25
1 .20
1 .00
10.80
1 1 .50
12.00
12.30
11.50
1 1 .50
12.00
12.03
1 1 .50
12.013
1 1 .75
12.50
12.50
12.80
13.00
13.00
13.00
12.80
12.80
12.80
13.00
13.00
12.50
12.70
12.70
12.70
12.70
12.73
12.73
12.73
13.00
13. 13
36.0
36.0
36.3
36.3
36.3
36.0
36.0
36. fl
36.0
36.0
36.3
36.0
36.3
36.0
36.0
36.3
36.0
36.0
36.0
36.0
36.0
36.3
38.0
38.0
38.0
38.3
38.0
38.0
38.3
38.0
38.0
38.0
38.0
38.0
38. 0
39. 9
39.0
39.0
39. ft
39.0
39.0
83.0
83.0
83.0
83.0
R3.0
83.3
83 O
83.0
83.0
83.0
83.0
«3.0
83.0
83.3
83.0
83.3
83.0
83.0
83.0
S3. 3
83.3
83.0
83.0
83.0
83.3
83.0
83.0
83.0
83.3
83.0
83.0
83.0
83t0
83.0
83.0
83.0
83.3
83.9
83.3
83.9
83.0 1
1020
1 10a
800
700
990
910
833
870
790
883
880
800
840
951
951
065
35fl
195
203
395
095
130
103
082
107
137
089
102
107
093
010
915
210
210
840
1 73
330
833
130
910
1070
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.3
2.0
2.0
2.3
2.0
2.0
2.3
2.3
2.0
2.0
2.3
2.3
2.0
2.0
2.0
2.0
2.3
2.3
2.0
2. 0
2.0
2.0
2.0
2.0
2.9
2.0
2.0
2.3
2.0
2.0
2.3
2.0
2.9
33.
37,
41
41
41
41
41
41
41
41
41
41
41
41
41
41
41
4|
41
41
41
41.5
41.5
41.5
41.5
41.5
41.5
41.5
41-5
41
41
41
41
41
41
,5
,5
,5
,5
,5
,5
4|.5
-------
TABLE J-DC
Sheet 7 of 20
Raw Data from Run
Day:Hour
Gasifier
Pressure
Regenerator Propane
Air to Pilot
Regenerator Np Meter
Gas Meter Pressure
Pressure
(inch.W.G) (Inches W.G) (inches W.G) (SCFH) (SCJW)
Oil Rate
(gall/hr)
2.1930
2.2333
2.2133
2.2230
2.2330
3.0030
3.0133
3.0230
3.0330
3.3433
3.0533
3.3633
3.0730
3.0830
3.0930
3. 1330
3.1 133
3.1230
3.1330
3.1430
3.1630
3.1733
3.2335
4.0005
4.0130
4.0233
4.3333
4.0430
4.0530
4.0630
4.3733
4.0830
4.3930
4.1030
4.1 133
4.1233
4.1330
4.1433
4.1533
4.1633
4.1733
13.10
13.33
13.83
13.53
13.53
13.53
13.53
14.03
14.30
14.33
14.30
14.30
14.03
14.53
14.53
14.53
14.73
14.73
15.30
15.30
15.53
16.53
-
14.30
14.33
14.30
14.33
13.83
13.50
13.50
14.50
14.33
14.00
14.50
14.03
14.50
14.53
14.53
14.53
15.03
15.33
39.3
39.0
39.0
39.3
39.3
39.3
40.3
40.0
40.0
43.0
40.0
43.3
40*3
43.0
40.0
43.0
41 .0
41 .0
41 .3
41 .0
41 .3
41 .0
41 .0
41 .3
41 .3
41 .0
41 .0
41 .3
41 .0
42.3
42.3
42.3
42.0
42.0
42.3
42.0
42.0
42.3
42.3
42.3
42.0
83.3
83.0
83.3
83.0
83.3
83.3
83.3
83.0
83.3
83.3
83.3
83.3
83.0
83.0
83.3
S3. 3
83.3
83.3
83.3
83.3
83.3
83.3
83.3
83.3
83.3
83.3
83.3
83.3
83.0
83.3
R3.3
83.3
83.3
83.0
83.3
83.0
83.0
83.0
33.3
83.3
83.3
850
843
895
743
910
755
870
963
1333
1313
1310
993
880
1353
1353
1383
963
993
1363 '
983
925
753
1323
1 133
932
1010
983
853
1210
1213
1393
980
1273
813
923
1323
943
933
983
923
1373
2.3
2.0
2.3
2.3
2.0
2.3
2.3
2.3
2.3
2.3
2.3
2.0
2.3
2 . 3
2.3
2.3
2.3
2.3
2.3
2.0
2.3
2.3
2.3
2.3
2.3
2.0
2.3
2.0
2.3
2.3
2.3
2.0
2.3
2.3
2.0
2.3
2.3
2.3
2.3
2.3
2.0
41 .5
41 .5
41 .5
41 .5
41 . 5
41 .5
41 .5
41 .5
41 .5
41.5
41 .5
41.5
41 .5
41 .5
41 .5
41-5
41 .5
41 .5
41.5
41 .5
41 .5
41 -5
41 .3
41 .3
41.3
41 .3
41.3
41,3
41.3
41.3
41 .3
41 .3
41 .3
41.3
41 .3
41.3
41 .3
41.3
41.3
41.3
41.3
-------
- 262-
TABIJE J-EC
Raw Data from Run 3
Meter Regenerator Propane
essure Air to Pilot
Day:Hour Gasifler Regenerator
Pressure Gas Meter
Pressure
(Inch;W.G) (inches W.G) (Inches W.G) (SCFH)
Oil Rate
(SCFM)
4. 1833
4. 1930
5.0130
5.0233
5.0333
5.0433
5.3530
5.3633
5.073;)
5.0933
5. 030
5. 130
5. 230
5. 330
5. 430
5. 533
5. 633
5. 730
5.2200
5.2233
5.2330
6.0030
6.0133
6.0233
6.0330
6.0433
6.3530
6.3633
6. 133
6. 230
6. 333
6. 430
6. 503
6. 830
6. 930
6.2030
6.2133
6.2230
6.2330
7.0030
7.3133
5.50
6.00
3.03
2. 5U
3.00
3.23
3.53
3.50
4.03
4.00
4.53
4.40
5.33
5.33
6 . 43
7-50
8.53
20.50
4.80
5.03
6.03
7.03
6.50
7 . .30
7.50
8.03
8.33
9. 10
4.00
..
_
v
18.73
-
_
8.33
6*00
6.80
7.03
7-53
7.50
43.0
43.0
43.0
43.0
43.0
43.3
43.0
43.0
43.3
43.0
43.0
43.0
44.0
44.0
44.0
44.0
44.0
44.0
44.3
44.3
44.0
44.0
44.0
44.0
45.0
45.0
45.0
45.0
45.0
45.0
45.0
45.0
45.0
46.0
46.0
46.0
46.0
46.0
46.3
46.0
46.0
83.0
83.3
83.0
83.0
83.0
83.0
83.3
83.3
83.3
83.3
83.3
83.0
83.3
83.3
83.3
83.0
83.3
83.3
33.3
83.0
83.0
83.0
83.3
83*3
83.0
83.0
83.3
83.3
83.0
83.0
H3L&h
sJSjPi
83.3
83.0
83.3
83.0
83.3
83.0
83.0
83.0
83.0
1040
870
_
-
920
1333
1020
1020
1013
873
981
905
1314
990
•1001
893
895
904
-
1010
1021
1304
965
912
1 101
882
1042
1032
813
995
1098
1 103
610
_«
660
97R
928
910
921
937
848
2.3
.2.3
2.3
2.3
2.3
2.3
2.3
2.3
2.0
2.0
2.3
2.3
2.3
2.3
2.0
2.0
2.0
2.0
2.0
2.3
2.3
2.0
2.0
2.0
2.0
2.0
2*0
2.3
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
41.3
41.3
39.4
41 .71
43.3
39. H
43.5)
43.3
39.3
39.4
39.4
39.4
39.4
39.4
39.4
39.4
39.4
39.4
39.2
39.2
39.2
39.2
39.2
39.2
39.2
39.2
39.2
39.2
42.3
42.2
42.2
42.2
42.3
42.2
42.2
42.0
42.3
42.2
42.2
42.2
42.2
-------
- 265-
TABIE J-IX
Sheet 9 of 2(
Raw
Day: Hour Gasifler Regenerator
Pressure Gas Meter
Pressure
(Inch.W.G) (inches
7.3233
7.0333
7.0430
7.0 530
7.063(3
7.3730
7.0830
7.0930
7. 330
7. 130
7. 230
7- 330
7. 433
7. 530
7. 630
7. 730
7. 830
7-5:) 46
7.53 46
7.63 46
8.13 46
8.13 46
8.33 46
8.30 47
8.00 47
8.53 47
8 .83 47
9.50 47
9.73 47
9.70 47
9.80 47
9.83 47
9.83 47
_r 47
7. 930 20.50 47
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
7.2030 21 .00 48.
7.2133 21.53 48.
7.2230 21.50 48.
7.2330 21.03 48
•
8.0030 22.00 48.
8.0130 21.53 48.
8.0230 22.03 48
•
8.0330 21.83 48.
8.0430 22.03 48.
8.0530 22.00 48
8.0630 22.00 48
•
•
8.0733 22.33 48.
8.0830 22.30 49.
8.0930 22.53 49.
8.1130 21.50 49.
8.1233 21.50 49
8.1330 21.80 49
•
•
8. 1430 21 .20 49.
8.2i/)30
8.213d
8.2230
8.2330
9.0330
6.33 49
5.23 49
5.50 49
6.30 49
5.50 49
•
•
•
•
•
Data
from Run J5
N2 Meter
Pressure
W.G) (inches W.
0
0
0
0
0
0
0
0
0
0
3
0
3
0
0
0
3
0
3
0
3
3
0
0
0
3
0
0
3
0
0
0
0
3
0
0
3
0
0
0
0
83
83
.43
83
83
83
83
83
83
83
83
83
83
83
83
83
83
83
83
H3
8"1
R3
83
«3
83
S3
83
83
83
83
83
83
83
83
83
83
S3
83
83
03
83
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
0
0
3
3
3
'.3
3
3
3
3
0
0
3
0
0
3
3
3
0
0
3
3
3
0
3
0
3
0
3
3
3
3
3
3
3
3
1
3
3
3
3
Regenerator Propane
Air to Pilot
G) (SCFH)
874
853
830
824
849
859
864
847
831
782
781
857
948
942
958
970
954
954
944
936
914
931
882
881
901
990
953
899
933
794
782
831
823
826
824
893
«•
863
820
816
81 4
(SCFM)
2
2
2
2
2
2
2
2
2
p
2
2
' 2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
•
•
•
*
•
•
•
•
•
•
•
•
•
t
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
•
3
3
3
3
3
0
3
n
3
0
0
0
«
3
3
3
0
0
0
3
3
3
0
0
0
3
0
0
0
0
0
0
0
0
0
0
0
0
3
0
0
Oil Rate
(gall/hr)
42.2
42.2
42. 2
42.2
•'.2.2
42. P
42. r»
' ; • !
42. 2
4^.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
33.0
33.0
33.0
33.0
33.0
33.0
38.9
38.9
38.9
38.9
38.9
-------
- 264 -
TABLE J-DC
3hi«t jfl
Raw Data from Rim
Day:Hour
Np Meter Regenerator Propane
Pressure Air to Pilot
9.0130
9.0230
9.0330
9.0430
9.3530
9.3630
9.3733
9.0830
9.0930
9.'1333
•1 130
1530
1 630
. 1730
. 1830
.1933
.2030
• 2130
.2233
.2330
• 3333
• 31 30
13.'3 2 33
13.3333
13.3433
10.3533
10.3633
10..17 33
13.0830
10.0933
9.
9.
9.
9.
9.
9.
9.
9.
9,
9.
10.
10.
13.
13,
10.
13-
030
130
230
333
430
533
630
Oasifier Regenerator
Pressure Gas Meter
Pressure
(inch.W.G) (inches W.G) (Inches W.O) (SCFH) (3CPM)
Oil Rate
r-
16.23
16.20
16.20
18.03
18.03
18.30
18.00
18.00
18.50
19.0
19.03
15.53
18.00
18.30
18.33
18.70
18.70
19.13
1 9.6il
1-9.23
.33
19
19
19
19
19
19
23
19
70
53
53
7'0
20
R3
23.33
23.23
23 . 23
23 . 33
2vl.53
23.50
21
21
00
00
49.0
50.0
50.0
50.0
50.0
50.0
53.0
50.0
50 . V,
50.0
50.0
50.0
50.0
51.0
51 .0
51 .0
51 .0
51 .fl
51.3
51 .0
5 1 . 0
51 .0
51 .0
51 .0
51 .0
52.0
52.0
52.3
53.3
52.0
52.0
52.3
52.0
52.0
52.0
52.3
52.0
83.3
83.0
83.0
83.3
R3.3
83.0
83.3
83.0
83.3
83.0
83.0
83.0
83.3
83.0
83.0
83.0
83.0
83.0
83.0
83.0
83.0
83.3
83.0
83.0
83.-0
83.0
8 3. '3
83.0
83.fl
83.3
83.«)
83.fl
H3.3
83.0
83.0
83.0
83v0
833
812
762
1038
1352
971
1030
995
992
980
972
..
952
940
984
102!?
•957
867
860
833
833
813
810
860
850
853
840
853
848
857
851
855
847
846
843
831
835
2.0
2.0
2.0
2.0
0
3
2
2
2.0
2.0
2.0
2.0
2.13
2.0
2-. 0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.3
2.3
2.3
2.0
2.3
2.0
2.0
2.0
2.0
2.3
2.0
2.0
2.5?
2.0
2.0
38.9
38.9
38.9
38.9
38.9
3R.9
3B.9
33. 9
38. 9
38.9
38.6
38.6
38.6
38.6
38.6
3R.6
38. 6
38.6
6
6
6
6
6
6
6
6
38.
33.
33.
38.
33.
38 .
3B.
3H.
35.0
35. *
35.0
35.0
35.0
35.0
35.0
35.0
-------
•- 265 -
TABLE J-DC
Sheet 11 of 2 !
Raw Data from Run 3
Day:Hour
Oil Temp
N2 to
Regenerator
(SCFH)
Q in Og In
Plenum Gas Boiler Gas
Stone Feed
Rate
Ib/hr (% by vol.) (% by vol).
l.iJ23(?)
1.0333
.0430
.3533
.0633
.0730
.3830
.0930
*1030
= 1 133
.1230
.1330
.1430
.1533
.1630
.1730
.1830
-1933
.2030
.2130
.2230
.2330
2.0330
2.0130
2.0230
2.0330
2.0430
2.0530
2.0630
2.0730
2.0830
2.0930
2.1030
2.1 130
2.1230
2.13313
2.1430
2.1530
2.1630
2.1733
2.1830
235
200
193
175
195
185
203
185
185
195
233
185
190
190
175
213
175
1 75
195
175
185
233
230
230
200
203
175
190
200
200
200
190
193
235
205
203
175
200
200
235
205
51
51
51
51
51
51
51
51
51
51
51
51
51
47
47
47
47
47
47
47
47
47
56
56
56
56
56
56
56
56
56
56
56
56
37
37
37
37
37
36
36
33.3
33.0
33.3
33.0
33.3
33.3
34. 3
34.3
34.3
34.0
35.3
34.0
34.0
47.3
47.0
47.0
47.0
47.0
59.0
59.3
59.0
59.3
79.3
79.3
79.3
134.3
104.3
66.3
66.3
66.3
76.3
76.3
76.3
59.3
59.0
59.3
59.3
45.3
45.3
45.0
45.3
_ I
6.81
3. 23
3.33
}. 7 '
• )
3. 3 •]
3. 3 ;
3.83
.j. " '1
3.» i
J. 10
j . '") ."*•
3.33
3. v|
o. M
3.^3
3. 13
3.?3
3.33
3.33
3.33
3. 10
3.53
3. 13
3.53
3.53
3.23
3.50
3.23
UJ
ca
o
o
o
Q£
O.
UJ
-------
- 266 -
TABUS J-IX
Sheet 12
Day:Hour
2
2
2
2
2
3
3
3
3
3
3
3
3
3
3
3
3
3
3
3
3
3
3
4
4
4
4
4
4
4
4
4
1930
2333
2130
2230
2330
0030
0133
023-3
0330
0430
0530
0630
0733
3830
0930
1030
1 130
1230
1330
1430
1 630
1733
2305
0005
0130
0230
0330
0433
3533
0630
0730
0830
4.0930
4.
130
230
330
430
530
630
Oil Temp
730
200
233
233
193
233
2'30
200
210
200
203
205
210
200
200
200
200
230
200
230
200
230
203
200
200
200
210
203
203
230
233
200
203
1 75
230
230
230
203
190
200
200
200
Raw
Ng to
Regenerator
(SCFH)
36
36
36
36
36
36
36
39
39
39
39
39
39
39
39
39
39
39
39
39
39
39
28
28
28
28
28
28
28
28
28
28
28
28
28
28
28
28
28
28
28
Data from Run ~$
Stone Feed 0- in
Rate PleHum Qas
Ih/hr (% by vol.)
45.3
45.0
33.0
33.0
33.0
33.3
33.0
33.0
33.0
33.0
47.0
47.0
47.0
47.0
47.0
53.3
53.0
53.0
53.0
31 .0
31 .3
31 .0
31 .0
31 .0
109.0
109 . 0
24.0
24. 0
24 . 0
24.0
24.0
24.0
24.0
24.0
46.0
46.0
46.0
46.0
46.0
41 .0
41.0
Og in
Boiler OM
(% bar vol),
3.23
3.43
3.20
3.50
3.50
3.:13
3.20
3.50
3.33
3.53
3.83
3.83
4.33
3.53
8.50
4.33
4.33
3.53
3.53
5.80
3.40
3.53
_
3.50
2.53
2.83
2.50
2.50
3.03
2.53
2.23
3. 10
3.60
3.63
3.10
3.50
3.50
3.00
3.30
3.03
3.00
-------
- 26? -
TABLE J-IX
Sheet 13 of 20
Raw Data from Run
Pay:Hour
Oil Temp
CP)
N2 to
Regenerator
(SCFfl)
°2 to °2 ln
PleRun Gas Boiler Gas
Stone Feed
Ib/hr (% by vol.) (% by vol).
4. 1830
4.1930
5.0130
5.3230
5.0330
5.0430
5*0530
5.0630
5.0730
5.0930
5. 1030
5. 1 133
5. 1230
5.1330
5.1430
5.1530
5. 1633
5. 1730
5.2200
5.2230
5.2330
6.0330
6.0130
6.0233
6.0330
6.0430
6.0530
6.0630
6.1 130
6.1230
6.1330
6.1433
6.1500
6.1830
6. 1930
6.2030
6.2130
6.2230
6.2330
7.0030
7.0130
203
233
200
203
203
233
233
230
230
180
200
203
203
185
200
180
180
183
205
175
200
193
172
233
173
175
210
175
200
180
230
200
180
203
180
200
203
203
195
203
200
28
28
28
28
28
28
28
28
28
28
28
28
28
28
28
28
28
28
33
33
33
33
3U
30
30
30
30
30
25
25
25
25
25
35
35
35
35
35
35
35
35
41 .3
41.3
51 .0
51 .3
51 .0
64.3
64.3
64.0
64.0
37.0
37.0
37.0
37.0
37.0
37.3
37.3
58.0
58.0
58.0
58.3
58.3
6.0
6.3
6.3
6.0
6.0
6.3
6.0
6.0
6.0
6.0
6.0
6.0
56.0
56.3
56.0
56.0
56.0
42.0
42.0
42.0
P. 33
3.33
3.33
3.33
3.33
O £i ,-jt
.• ' " I
9.30
3.33
3.33
3.53
2.83
•? . 90
3.33
3. in
4.23
2.83
4.73
5. 73
2.33
3.43
1 .80
3.30
2.00
1 .60
2.40
2.40
2.03
2.33
2.03
2.03
2.10
-------
- 268 -
TABLE J-IX
Sheet 14
Day:Hour
7.0230
7.0330
7.0430
7.0530
7.0630
7.0730
7.0830
7.0933
7.
7.
7.
7 .
7 .
7.
7.
7.
7.
7-
7.
7.
030
230
330
433
530
630
73:-)
830
930
203;-)
PI 30
7.2230
7.2330
"..0030
01 30
0230
0330
0430
0530
rt.0730
8.0830
8.0930
8.113j
1230
1 333
1 430
2030
2130
2230
2330
0033
Oil Temp
(**)
203
200
235
180
200
200
193
190
230
200
200
200
200
185
200
180
195
200
200
205
205
205
235
1 75
195
202
200
190
195
202
200
200
204
194
198
204
204
200
200
235
200
Raw
Ng to
Regenerate:
(SCFH)
35
35
35
35
35
35
35
35
35
35
35
125
125
125
125
58
58
58
58
58
58
58
58
58
58
58
58
58
58
58
58
58
58
58
58
53
60
60
63
60
60
r Data from Run 3
Stone Feed
Rate
Ib/hr ('
42.0
42.0
42.0
59.0
59.0
59.0
59.0
36.0
36.0
36.0
36.0
36.0
36.0
28.0
28.0
28.0
28.0
28. 3
28.0
28.3
28.0
47.0
47.0
47.0
47.0
47.0
64.0
64.0
64.0
28.0
28.0
28.0
28.0
28.0
28-0
40.0
40.0
40.0
40 . 0
40.0
52.0
Op in
Plenum Gt
% by vol
-
— , .
— .
- •
'-
•
-
4.5
4.3
4.5
4.0
3.8
4.0
3. 7
3.4
3.4
3.4
3.2
3.3
3.5
3.5
9.0
3.0
3.8
3.6
3.4
3.5
3.7
3.7
3. R
3.8
3.8
5.4
6.0
4.0
3.R
3.5
4.0
3.5
3.R
••
02 in
is Boiler Gas
.) (% by vol).
P. .00
2.23
2.00
1 .80
2.03
2. 33
P. 23
2.43
°. 30
P. 30
2.03
2.30
2.10
2.«?3
-
2.30
-
2.23
1 .80
2.00
2.03
2.30
2.23
2.03
2.70
2.5)0
2.03
2 . *fyty
P . '33
2.13
1 -90
2. 13
3.00
2.R0
6.00
3.63
5.20
3.00
3. 10
2.50
2.53
-------
Day:Hour
3233
0330
3433
0533
3633
0730
3830
0933
1033
1 1 33
1533
1630
1730
1833
1933
2033
9.2130
9.2233
9.2330
19.3130
10.0230
10.3330
13.3433
10.3533
13.3630
13.0733
11). 08 33
10.3933
13.
13,
13,
13,
10.
10,
033
1 33
23.)
J33
433
533
630
Oil Temp
200
23(3
230
233
233
200
1 95
233
230
200
1 93
1 80
233
233
233
230
230
185
213
230
233
235
185
233
233
233
233
233
230
194
20'.)
195
198
197
195
- 209 -
TABIE J-EC
Sheet 15
Raw Data from Run 3
N2 to
Regenerator
(sera)
50
50
53
50
53
53
50
51
53
85
85
85
85
«5
85
85
85
85
85
85
85
85
85
85
103
130
133
133
71
71
71
71
71
71
71
Stone Peed
Rate
. Ib/hr
02 in
Plenum Gas
(% by vol.)
52.3 3.8
52.3 3.8
52.0 3.5
52.3 3.8
30.0 3.8
30.0 3.2
33.0 3.2
33.3 3.2
33.3 3.8
30.0. 13.5
33.3 13.6
30.0 13.4
26.3 13.4
26.0
26.0
26.0
26.3
26.0
26.3
26.0
26.3
28.0
28.3
28.3
28.3
28.3
28.0
28.3
5.0
5.3
5.3
5.3
5.3
5.3
5.0
5.3
5.0
3.2
3.3
3.8
3.8
3.8
9.3
3.3
3.0
3.3
3.3
3.2
-
3.5
3.3
9.0
3.5
3. 1
3.3
3. 1
3.5
3.0
3.0
3.2
1.8
0 in
Boiler Gas
(% by vol).
2 . 73
2.23
2.33
2.23
2.30
2 . * 3
?.:•) 1
2. 53
P • P,^
-
2. 5T
-
- . .
2 . 6.71
2. I?"
2. in
1 .83
1 .23
1 . 8'=)
2.33
2. 33
1 . 8T
2.33
1 .93
1 . "33
1 . 63
1 .73
2.23
2.30
.63
. 63
-33
.33
. 33
• TIT
.00
.93
-------
270 -
TABLE J-IX
Sheet
Day:Hour
02 in
Regenerator
Off Gas
(% by vol)
Raw Data from Run 3
SO In S02 in COg in COL in
Boiler Gas Regenerator Regenerator Boiler Gas
Off Gas Off Gas
(ppm) (^ amp) (p, amp) fo amp)
0230
• 0330
• 0A30
• 0530
0630
0730
• 0830
'0930
1030
1 130
1230
1330
I 430
1530
1630
I 730
1830
1930
2030
2130
2230
2330
2.0030
2.0130
2.0230
.0330
.0430
.0530
.0630
.0730
.0830
.0930
. 1030
• 1 130
1230
1330
2.1430
2.1530
2.1630
2.1730
2.1830
2.
2.
2.
2.
2.
2.
2.
2.
2.
2.
2.
13
.0
.0
• 0
• 0
• 0
.0
.0
• 0
.0
.0
.0
.0
• 0
.0
.0
• 0
.0
.0
.0
.0
.0
.0
.0
• 0
.0
.0
.0
• 0
• 0
.0
• 0
• 0
.0
• 0
.0
'0
.0
.0
.0
.0
.0
10
10
20
20
40
60
250
325
250
120
80
340
150
220
280
280
280
380
380
150
100
120
0
0
0
0
0
0
0
30
0
0
0
0
0
8
0
4
0
0
0
10
12
10
10
26
14
30
27
10
35
43
34
38
38
46
57
50
49
47
25
46
52
52
49
50
42
38"
23
35
32
1 6
45
46
42
47
48
29
4!
46
45
44
30
38
15
35
27
20
36
12
100
39 •
44
60
56
57
50
37
51
40
40
67
50
38
42
35
39
56
53
56
66
94
33
68
67
83
60
64
100
85
72
70
69
74
74
84
84
85
82
83
83
82
81
82
82
82
84
84
84
84
83
84
85
84
84
84
83
83
83
84
82
83
84
84
83
84
82
83
82
83
82
82
82
SI
-------
- 271 -
TABLE J-IX
Sheet 17 20
Day:Hour
02 in
Regenerator
Off Gas
(% by vol)
Raw Data from Run 3
SO In S02 in (X>2 in COg in
Boiler Gas Regenerator Regenerator Boiler Gas
Off Gas Off Gas
(ppm) (p, amp) (jj, amp) (g, amp)
2.1930
2.2030
2.2130
2.2230
2.2330
3*0030
3.0130
3.0230
3.0330
3.0430
3.0530
3.0630
3.0730
3.0830
3.0930
3.1030
3.1 130
3.1230
3.1330
3.1430
3.1630
3.1730
3.2305
4.0005
4.0130
4.0230
4.0330
4.0430
4.0530
4.0630
4.0730
4.0830
4.0930
4.1030
4.1 130
4*1230
4.1330
4.1430
4.1530
4.1630
4.1730
.0
.0
.0
.0
• 0
• 0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
1 .0
.0
.0
.0
• 0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.3
.0
0
4
0
8
3
160
120
120
120
120
1 1
104
80
64
80
28
20
24
60
0
0
0
0
24
4
8
8
0
0
0
0
0
8
8
0
8
0
0
0
0
0
47
48
48
54
54
52
50
46
46
48
42
48
46
42
39
39
39
40
42
41
34
40
40
30
22
28
22
24
28
34
42
38
45
43
32
45
47
46
48
45
45
60
44
54
46
46
44
52
64
70
52
54
70
74
68
62
62
68
76
79
80
34
78
38
46
38
38
40
39
38
39
42
40
40
37
39
37
37
37
37
35
46
82
82
81
82
81
82
82
82
83
82
82
82
80
80
81
80
80
82
82
77
83
81
-
84
85
82
82
85
81
83
84
82
80
81
82
80
80
82
83
83
83
-------
- 2?2 -
TABI£ J-IX
Sheet 18 of
Raw Data from Run 3
Day : Hour
02 In
Regenerator
Off Gas
(% by vol)
SO In
Boiler Gas
(ppm)
S02 in
Regenerator
Off Gas
(n amp)
C02 in
Regenerator
Off Gas
U amp )
C02 in
Boiler Gas
(ji amp)
A. 1830
4.1930
5.0130
5.0233
5.0330
5.0430
5.0530
5.0630
5.0730
5.0930
5. 1030
5.1 130
5.1230
5. 1330
5. 1430
5. 1530
5. 1630
5. 1730
5.2200
5.2230
5.2330
6.0030
6.0130
6.0230
6.0330
6.0430
6.0530
6.0630
6. 11 30
6.1230
6. 1330
6. 1430
6. 1500
6.1830
6. 1930
6.2030
6.2130
6.2230
6.2330
7.0030
7.0130
• 0
.5
.0
.0
.0
.0
.0
.0
.0
.0
.0
• 0
.5
.7
• 3
.3
4.5
17.5
• 0
1.2
.2
.0
.2
.3
4.8
.4
.3
.5
-
_
_
—
-
_
_
1.0
1 .0
.5
.3
.5
.4
0
0
0
16
0
0
0
0
8
32
160
152
88
-
13
32
24
M
0
0
192
1 12
1 12
96
352
320
320
440
_
_
_
576
224
208
200
1 12
-
. -
45
44
14
22
38
28
28
40
44
41
44
39
40
39
40
39
35
40
37
24
28
-
38
48
36
40
43
43
_
_
24
31
39
45
42
48
48
45
60
54
48
38
20
42
41
46
30
44
30
33
38
35
15
13
23
20
48
_
41
42
32
46
40
36
20
24
22
22
26
27
83
82
80
82
84
84
82
82
85
86
84
82
81
85
84
83
86
85
84
86
86
81
86
80
86
84
84
"
85
86
85
U v
84
w ^
—
-------
• 273 -
TABLE J-EC
Sheet 1^ of 20
Day:Hour
02 in
Regenerator
Off Gas
(% by vol)
Raw Data from Run 3
SO in S02 in COg in C02 in
Boiler Gas Regenerator Regenerator Boiler Gas
Off Gas Off Gas
(ppm) (n amp) (^ amp) (p, amp)
7.0230
7.0330
7.0430
7.0530
7.0630
7.0730
7.0830
7.0930
7.1030
7.1 130
7.1230
7.1330
7.1430
7.1530
7. 1630
7.1730
7.1830
7.1930
7.2030
7.2130
7.2230
7.2330
8.0030
8.0130
8.0230
8.0330
8.0430
8*0530
8.0630
8.0730
8*0830
8.0930
8.1130
8.1230
8.1330
8.1430
8.2030
8.2130
8.2230
8.2330
9.0030
.4
.4
.3
.4
.5
.6
.1
.3
.3
.3
.2
.2
.2
.2
2.0
.2
.3
.3
.5
.4
.4
.2
.3
.2
.2
.6
1.5
.0
.0
• 0
.0
.0
.0
.0
.0
.0
2.0
.0
.0
.4
.2
80
152
208
288
160
24
40
12
4
0
200
128
64
32
_
376
a*
368
360
256
160
176
208
128
64
0
16
60
48
8
0
0
8
64
24
32
800
160
224
1 12
72
46
48
49
53
48
47
44
52
49
54
50
49
47
52
16
47
43
45
36
43
35
43
36
45
42
37
38
45
48
46
48
50
50
50
12
50
10
46
46
38
46
27
26
27
23
25
27
30
46
20
34
20
50
52
51
52
55
55
55
57
57
57
27
24
29
20
27
38
40
44
38
24
26
23
30
14
32
14
14
18
21
28
-
84
85
85
86
86
85
85
85
85
86
86
85
84
-
86
-
86
87
86
87
85
86
86
80
85
86
86
86
85
86
86
84
84
74
84
80
84
84
86
85
-------
TABLE J-IX
Sheet
Raw Data from Run 3
Day:Hour
02 in
Regenerator
Off Gas
(% by vol)
S02 in S02 in COg in COg in
Boiler Gas Regenerator Regenerator Boiler Gas
Off Gas Off Gas
(ppm) (\i, amp) (^ amp) (^ amp)
9.0130
9.0230
9.0330
9.0430
9.0530
9.0630
9.0730
9.0830
9.0930
9. 1030
9.1 130
9. 1530
9.1630
9.1730
9.1830
9. 1930
9.2030
9.2130
9.2230
9.2330
10.0030
10.0130
10.0230
10.0330
10.0430
10.0530
10.0630
10.0730
10.0830
10.0930
10. 1030
10. 1 130
10.1230
10.1330
10.1430
10.1530
10. 1630
.2
.2
.0
.3
.5
• 3
.6
.4
• 4
-
.4
15.0
1.0
• 2
.3
• 3
.3
.0
.0
.0
• 0
.0
.0
• 0
_
.0
.0
• 0
.0
.3
.3
.3
• 3
.2
.2
.2
.4
160
192
120
136
96
144
192
176
152
_
168
-
_
24
184
232
88
80
32
8
32
48
64
48
104
48
80
24
32
64
64
96
48
48
44
24
136
45
49
45
38
36
40
38
40
39
—
37
12
39
41
42
42
42
46
48
48
48
52
5ft
48
50
48
48
48
44
42
41
44
40
42
46
41
20
22
40
26
26
40
35
28
32
39
13
12
25
19
26
32
34
30
36
34
28
22
22
24
24
26
24
26
25
26
34
28
33
25
28
85
86
86
86
86
84
86
86
85
85
_
84
86
86
87
88
86
86
86
87
86
86
86
w w
87
86
84
86
86
86
87
aa
uo
88
88
88
w w
88
-------
- 275 -
TABLE J-X
CAFB Gas Composit1ons and Efficiencies
Sheet 1 of 5
Run 3
TIME
DAY. HOUR
.0230
• 0330
.0430
.0530
.0630
.0730
.0830
.0930
.1030
. 1 130
. 1230
. 1330
. 1430
.1530
.1630
. 1730
. 1830
. 1930
.2030
.2130
.2230
.2330
2.0030
2.0130
2.0230
2.0330
2.0430
2.0530
2.0630
2.0730
2.0830
2.0930
2.1030
2.1 130
2.1230
2.1330
2.1430
2. 1530
2.1630
2.1730
2.1830
BOILER GAS COMP
02
VOL
PCT
6.8
6.8
3.2
3.0
2.8
3.8
3.7
3.5
3.3
3.3
3.8
3.3
3.2
3.2
3.0
2.8
3.0
3.0
3.0
2.9
3.1
3.0
3. 1
3.2
3.2
3.0
3.0
3.0
3.0
3.0
3.0
3.0
3.1
3.5
3.1
3.5
3.5
3.2
3.5
3.2
3.6
C02
VOL
PCT
1 .7
1 .7
4.4
4.4
4.7
3.8
4.1
4. 1
3.8
3.5
3.8
3.8
3.8
4.4
4.4
4.4
4.4
4. 1
4.4
4.7
4.4
4.4
4.4
4. 1
4.1
4. 1
4.4
3.8
4. 1
4.4
4.4
4.1
4.4
3.8
4. 1
3.8
4.1
3.8
3.8
3.8
3.5
302
PPM
1?
10
20
20
40
60
250
325
250
120
80
340
150
220
280
280
280
380
380
150
100
120
0
0
0
0
0
0
0
30
0
0
0
0
0
8
0
4
0
0
0
REGENERATOR GAS COMP
02
VOL
PCT
5.0
• 0
.0
.0
.0
1 .0
.0
• 0
• 0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
1 .0
.0
.0
13.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
C02
VOL
PCT
1 .7
2.3
.5
2.1
1.5
1 .0
2.2
.2
10.0
2.4
2.8
4.3
3.9
4.0
3.3
2.2
3.4
2.5
2.5
5. 1
3.3
2.3
2.6
2.1
2.4
3.9
3.6
3.9
5.0
8.8
1 .9
5.2
5.1
7.1
4.3
4.7
10.0
7.4
5.7
5.4
5.3
S02
VOL
PCT
.0
.6
.0
.0
3.9
1 .2
4.7
4. 1
.0
5.7
7.2
5.5
6.2
6.2
7.8
10.0
8.6
8.4
8.0
3.7
7.8
9.0
9.0
8.4
8.6
7.0
6.2
3.3
5.7
5.1
1 .7
7.6
7.8
7.0
8.0
8.2
4.5
6.8
7.8
7.6
7.4
SULPHUR
REMOVAL
PERCENT
99.2
99.2
98.7
98.7
97.4
95.8
83.0
77.9
82.6
91.5
94.4
76.4
89.6
85.3
81.3
81 .3
81 .3
74. 1
74.6
90.2
93.3
92.0
100.0
100.0
100.0
100.0
100.0
100.0
100.0
98.0
100.0
100.0
100.0
100.0
100.0
99.4
100.0
99.7
100.0
100.0
100.0
REGEN
CAO/CAS
PERCENT
. «2
5.5
. 1
. 1
32.7
10.4
41 .P
32. 4
.2
49.0
62. 7
53.=.
5>-. A
58. 1
6*.*
80.9
74.9
69.2
66.2
37-7
68. 0
73. S
75.5
68.9
71.4
64.4
57.0
33.7
56.4
67.4
52.2
76.0
75.6
78.3
73.7
76.0
64.8
77.7
76.9
74.7
71*8
-------
- 276 -
TABLE J-X
CAFB Gas Compositions and Efficiencies
Run 3
Sheet 2 of '
TIME BOILER GAS COMP
02 C02 S02
VOL VOL PPM
DAY.HOUR PCT PCT
REGENERATOR GAS COMP
02 C02 S02 SULPHUR REGEN
VOL VOL VOL REMOVAL CAO/CA
PCT PCT PCT PERCENT PERCEN
2. 1 930
2.2030
2.2130
2.2230
2.2330
3.0030
3.0130
3.0230
3.0330
3.0430
3.0530
3.0630
3.0730
3.0830
3.0930
3. 1030
3.1 130
3.1230
3. 1330
3. 1430
3. 1630
3. 1 730
3.2305
4.0005
4.01 30
4.0230
4.0330
4.0430
4.0530
4.0630
4.0730
4.0830
4.0930
4. 1030
4. 1 130
4. 1230
4. 1330
4. 1430
4. 1530
4. 1630
4. 1730
3.2 1
3.4 I
3.2 !
3.5 1
3.5 1
3.0 1
3.2 1
3.5 1
3.0 ]
3.5 ]
3.8 1
3.8 i
4.0 1
3.5 ]
8.5
4.0
4.0 :
3.5 i
3.5
5.8
3.4 !
3.5
-
3.5
2.5
2.8
2.5
2.5
3.0 :
2.5
2.2 !
3.1 !
3.6 !
3.6 !
3.1 !
3.5 1
3.5 ]
3.0 ]
3.0 !
3.0 1
3.0 I
13.8
13.8
13.5
13.8
13.5
13.8
13.8
13.8
1 4. 1
13.8
13.8
13.8
13.2
13.2
13.5
13.2
13.2
13.8
13.8
12.4
14. i
13.5
-
1 4.4
1 4. 7
13.8
13.8
14.7
13.5
1 4. 1
14.4
13.8
13.2
13.5
1 3.8
13.2
13.2
13.8
14. 1
1 4. 1
14.1
0
A
0
8
3
160
120
120
120
120
1 1
104
80
64
80
28
20
24
60
0
0
0
-
24
4
8
8
0
0
0
0
0
8
8
0
8
0
0
0
0
0
.0
.0
.0
.0
• 0
• 0
.0
.0
.0
.0
• 0
.0
.0
.0
.0
.0
.0
.0
.0
.0
1 .0
.0
.0
• 0
• 0
• 0
.0
.0
.0
.0
.0
.0-
• 0
.0
.0
• 0
• 0
• 0
.0
.3
.0
4.3
2.8
3.7
3.0
3.0
2.8
3.5
4.7
5.4
3.5
3.7
5.4
5.9
5.2
4.5
4.5
5.2
6.2
6.5
6.7
2.0
6.4
2.3
3.0
2.3
2.3
2.5
2.4
2.3
2.4
2.6
2.5
2.5
2.2
2.4
2.2
2.2
2.2
2.2
2.1
3.0
8.0
8.2
8.2
9.4
9.4
9.0
8.6
7.8
7.8
8.2
7.0
8.2
7.8
7.0
6.4
6.4
6.4
6.6
7.0
6.8
5.5
6.6
6.6
4.7
3.1
4.3
3. 1
3.5
4.3
5.5
7.0
6.2
7.6
7.2
5.1
7.6
8.0
7.8
8.2
7.6
7.6
100.0
99.7
100.0
99.4
99.8
88.9
91.7
91 .7
91.8
91.7
99.2
92.8
94.2
95.4
94.3
98.0
98.6
98.3
95.8
100.0
100.0
100.0
_
98.4
99.7
99.4
99.4
100.0
100.0
100.0
100.0
100.0
99.4
99.4
100.0
99.4
100.0
100.0
10(9.0
100.0
100.0
73.5
69.2
72.5
79.8
78.9
75.8
75.0
73.5
76.4
71,5
62.7
819.0
79.4
68.2
60.6
60.5
63.4
69.1
74.3
73.4
48.1
71.6
54.4
40.7
26.4
36.3
26.5
30.1
36.1
45.4
58.1
51.9
61.8
59.1
42.8
61*5
64.6
63.2
66*0
62.2
63*8
-------
- 277 -
Sheet 3 of 5
TABI£ J-X
CAFB Gas Compositions and Efficiencies
TIME
DAY. HOUR
4.1830
A. 1930
5.0130
5.0230
5.0330
5.0430
5.0530
5.0630
5.0730
5.0930
5.1030
5.1130
5.1230
5.1330
5. 1430
5.1530
5.1630
5.1730
5.2200
5.2230
5.2330
6.0030
6.0130
6.0230
6.0330
6.0430
6.0530
6.0630
6.1 130
6.1230
6.1330
6.1430
6.1500
6. 1830
6.1930
6.2030
6.2130
6.2230
6.2330
7.0030
7.0130
BOILER GAS COMP
02
VOL
PCT
3.0
3.0
2.0
3.0
2.5
3.0
3.0
3.0
2.8
2.0
3.0
3.0
3.5
<••
2.8
2.9
3.0
3.1
4.2
2.8
4.7
5.7
2.0
3.4
1 .8
3.0
2.0
1 .6
• -
-
.-
-•
2.4
-
-
2.4
2.0
2.0
2.0
2.0
2.1
C02
VOL
PCT
14. 1
13.8
13.2
13.8
1 4.4
14.4
13.8
13.8
14.7
15.0
14.4
13.8
13.5
—
14.7
14.4
14.1
—
15.0
14.7
14.4
15.0
15.0
13.5
15.0
13.2
15.0
14.4
1 -
- -
T
- -
14.4
-
-
14.7
15.0
14.7
14.4
-
-
S02
PPM
0
0
0
16
0
0
0
0
8
32
160
152
88
_'
13
32
24
_
0
0
192
1 12
1 12
96
352
320
320
440
-
- •
- •
-
576
- •
-
224
208
200
1 12
96
16
Run 3
REGENERATOR GAS COMP
02
VOL
PCT
.0
.5
.0
.0
.0
.0
.0
.0
.0
.0
.0
.0
.5
.7
.3
.3
4.5
17.5
• 0
1 .2
.2
• 0
.2
.3
4.8
.4
.3
.5
'
.
-
_
-
-
—
1 .0
1.0
.5
.3
.5
• 4
C02
VOL
PCT
3. 1
2.9
4.3
3.7
3.1
2.3
1 .0
2.6
2.5
3.0
1 .7
2.8
1 .7
1 -9
2.3
2.1
.5
.3
1 .2
1 .0
3.1
-
2.5
2.6
1 .9
3.0
2.5
2.2
-
-
-
-
- '
-
-
.0
.3
.1
. 1
.4
-5
S02
VOL
PCT
7.6
7.4
1 .2
3.1
6.2
4.3
4.3
6.6
7.4
6.8
7.4
6.4
6.6
6.4
6.6
6.4
5.7
6.6
6.0
3.5
4.3
-
6.2
8.2
5.8
6.6
7.2
7.2
-
-
-
-
-
-
-
3.5
4.9
6.4
7.6
7.0
8.2
SULPHUR
REMOVAL
PERCENT
100.0
100.0
100.0
98.9
100.0
100.0
100.0
1 00.0
99.5
97.9
89.3
89.4
93.7
' ! -
99. 1
97.9
98.4
-
100.0
100.0
87.2
92.8
92.8
93.2
77.4
76.8
79.4
70.6
-
-
..- •'
61 .6
-
r
85.3
86.7
86.9
9
SB. y
58-5
54. ri
54. 1
53. F
55.2
53.3
55.0
-
-
29.6
38.5
•
52.8
68*8
62.6
58.0
60.3
59.9
-
-
-
-
«•
-
-
29.5
41 .4
51.7
59.8
56.8
65.9
-------
- 278 -
Sheet 4 of 5
TABLE J-X
CAFB Gas Compositions and Efficiencies
Run 3
TIME BOILER GAS COMP
02 C02 S02
VOL VOL PPM
DAY.HOUR PCT PCT
REGENERATOR GAS COMP
02 C02 S02 SULPHUR REGEN
VOL VOL VOL REMOVAL CAO/CAS
PCT PCT PCT PERCENT PERCENT
7.0230
7.0330
7.0430
7.0530
7.0630
7.0730
7.0830
7.0930
7. 1030
7. 1 130
7. 1230
7. 1330
7. 1430
7. 1530
7. 1630
7. 1 730
7. 1830
7- 1930
7.2030
7.2130
7.2230
7.2330
8.0030
8.0130
8-0230
8.0330
8.0430
8.0530
8.0630
8.0730
8.0833
8.0930
8. 1 130
8« 1230
8. 1330
8. 1430
8.2030
8.2130
8.2230
8.2330
9.0030
2.0
2.2
2.0
1 .8
2.0
2.0
2.2
2.4
2.3
2.0
2.0
2.3
2. 1
2.8
-
2,.0
_
2.2
1 .8
2.0
2.0
2.0
2.2
2.0
2.7
2.0
2.0
2.0
2.0
2.1
1 -9
2. 1
3.0
2.8
6.0
3.6
5.2
3.0
3.1 1
2.5 1
2.5 1
_
1 4.4
14.7
14. 7
15.0
15.0
14. 7
1 4. 7
14. 7
14.7
15.0
15.0
1 4.7
14.4
•-
1 S.ld
_ '
15.0
5.3
1 5.0
15.3
14.7
15.0
15.0
13.2
14. 7
15.0
5.0
5.0
14.7
5.0
5.0
4.4
4.4
1 .7
4.4
3.2
4.4
4.4
5.0
4.7
80
152
208
288
160
24
40
12
4
0
200
128
64
32
16
376
-
368
360
256
160
1 76
208
128
64
0
16
60
48
8
0
0
8
64
24
32
800
1 60
224
1 12
72
.4
.4
• 3
.4
.5
.6
.1
.3
.3
.3
.2
.2
.2
.2
2.0
.2
.3
.3
.5
.4
.4
.2
.3
.2
.2
• 6
1 .5
.0
.0
• 0
.0
.0
.0
.0
.0
.0
2.0
.0
.0
.4
.2
-5
.4
.5
.2
.3
.5
.7
3.0
1 .0
2.0
1 .0
3.3
3.5
3.4
3.5
3.8
3.8
3*8
4.0
4.0
4.0
.5
.3
.6
• 0
.5
2.3
2.5
2.8
2.3
1*3
1 .4
1 .2
1 .7
.4
1 .9
.4
.4
• 8
1 .0
1.6
7.8
8.2
8.4
9.2
8.2
8.0
7.4
9.0
8.4
9.4
8.6
8.4
8*0
9.0
1.7
8*0
7.2
7.6
5.8
7.2
5.7
7.2
5.8
7.6
7.0
6.0
6.2
7.6
8.2
7.8
8.2
8.6
8*6
8.6
.6
8.6
• 0
7.8
7.8
6.2
7.8
_
89.9
86.4
81 .2
89.7
98.5
97.4
99.2
99.7
100.0
87.2
91.8
95.8
97.9
-
75.9
76" 4
77.4
83.6
90.0
88.5
86.7
91.8
95.4
100.0
99.0
96.2
96.9
99.5
100.0
100.0
99.5
95.7
98.0
97.8
41 .9
89.3
85.0
92. 8
95.3
62.8
65.6
67.1
72.3
65.7
65.0
59.7
77.1
65.4
76.7
66.6
82.9
79.1
87.4
20.8
73.8
67.7
71.0
57.3
69.0
55.4
59.5
48.8
63.0
56.5
51.3
SB. 3
65.1
71*0
66.7
66.2
69.6
68.7
70.5
5.1
70.6
_
60.4
61*7
51.9
64.8
-------
- 279 -
TABLE J-X
CAFB Qas Compositions and Efficiencies
Run_2
Sheet 5 of 5
TIME BOILER GAS COMP
02 C02 S02
VOL VOL PPM
DAY.HOUR PCT PCT
REGENERATOR GAS COMP
02 C02 S02 SULPHUR REGEN
VOL VOL VOL REMOVAL CAO/CAS
PCT PCT PCT PERCENT PERCENT
9.0130
9*0230
9.0330
• 0430
• 0530
• 0630
• 0730
• 0830
• 0930
• 1030
• 1130
9*1530
9*1630
9*1730
.1830
.1930
.2030
• 2130
• 2230
9.2330
10.0030
10.0130
10.0230
10.0330
10.0430
10.0530
10.0630
10.0730
10*0830
10.0930
10.1030
10.1130
10*1230
10*1330
10*1430
10*1530
10.1630
2.7
2.2
2.3
2.2
2.0
2.8
2.0
2.5
2.8
„
2.5
, •
2~6
2*1
2.1
l.R
1.2
1 .8
2.0
2.0
1.8
2.0
1.9
1.8
1.6
1.7
2.2
2.0
.6
.6
• 3
• 0
• 0
.0
• 0
• 9
14.7
15*0
15*0
15.0
15.0
14.4
15.0
15.0
14.7
. ,
14*7
m •»
» •*
•m
14.5
15.0
15.0
15.3
15*6
15*0
15.0
15.0
15.3
15.0
15.0
15.0
15.3
15.0
14*4
15*0
15*0
15*0
15*3
15*6
15*6
15*6
15.6
15.6
160
192
120
136
96
144
192
176
152
••
168
m
*•*<
24
184
232
88
80
32
8
32
48
64
48
104
48
80
24
32
64
64
96
48
48
44
24
136
.2 1
.2 1
• 0 S
• 3 1
• 5
• 3 i
• 6 {
• 4
• 4
3.5 i
.4 (
15.0
1.0
.2
.3
.3
.3
.0 I
.0
.0 1
.0 1
*0
• 0
.0
-
.0
.0
• 0
• 0
*3
• 3
• 3
.3 S
.2
.2
.2
• 4
.0
.1
2.5
.4
• 4
2.5
2.1
.6
1*9
2.5
2*4
.3
.2
1.3
.9
1.4
1.9
2.0
1.7
2*2
2.0
.6
.1
.1
i
• 3
.3
.4
• 3
.4
.3
.4
2.0
.6
.9
.3
.6
7.6
8.4
7.6
6.2
5.8
6.6
6.2
6.6
6.4
•
6.0
.6
6.4
6.8
7.0
7.0
7.0
7.8
8*2
8.2
8.2
9.0
8.6
8.2
m
8*6
8.2
8*2
8.2
7.4
7.0
6.8
7.4
6.6
7.0
7.8
6.8
89*5
87.7
92.3
91.3
93.8
90.4
87.7
88.7
90.0
77.7
89.0
.
•
98.4
88.2
85.1
94.5
95.1
97*9
99.5
97.9
97.0
95.9
96.9
93.3
97.0
94.9
98.4
97.9
95.9
95.9
93.9
97.0
97.0
97.3
98.5
91.6
61 .3
62.3
65.:-:
51.2
4R.6
57.4
53.9
54.9
54.??
-
52. P
.
53.7
57.7
57.8
59.3
61.1
67.4
69.5
71.5
71. P
75.6
70.9
67.4
-
71. e
69.4
69*8
69.3
64.6
59.1
57. R
64.4
56*5
60.7
64.9
5R.7
-------
- 280 -
TABIE J-XI
Sheet 1 of
CAFB Peed Ratios and Regenerator Sulphur
Run 3
DAY.HOUR
• 0230
• 0330
.0430
• 0530
.0630
• 0730
• 0830
• 0930
.1030
• 1130
.1230
.1330
• 1430
• 1530
.1630
• 1730
• 1830
.1930
• 2030
• 2130
• 2230
• 2330
2.0030
2.0130
2.0230
2.0330
2.0430
2.0530
2.0630
2.0730
2.0830
2.0931
2.1030
2.1130
2.1230
2.1330
2.1430
2.1530
2.1630
2.1730
2.1830
VELOCITY
FT/SEC
3*39
3*34
3.91
3.89
3.67
3.65
3.64
3.66
3.66
3.79
3.92
3*72
3.93
3.78
3.89
3.76
3.76
3.74
3.74
3.66
3*61
3*58
3*60
3.85
3.66
3.67
3.68
3*62
3.60
3.62
3.63
3.65
3*66
3.62
3.64
3.66
3.65
3*67
3.67
3*61
3.53
AIR/PUEL
PCT STOIC
24.2
24.0
25.3
21.8
20.8
20.9
20.9
20.9
20.8
20.9
22.0
20.0
21.4
22.3
22.4
21.8
21.8
21.0
21.4
21.3
21.4
21.0
21.5
22.5
22.0
22.0
22.0
21.7
21.7
21.7
21.7
21,7
21.7
21.8
21.7
21.8
21 .8
21.7
21 .8
21.3
20.9
CA/S RATIO
MOL/HOL
REGENERATOR SULPHUR
LB/HR PCT OF FED
• 35
• 35
• 21
• 08
• 08
.08
• S3
.53
.53
.53
.53
• 93
.93
• 93
• 93
2.58
2*58
2*58
3.40
3.40
2.15
2.15
2.15
2.48
2.48
2.48
• 93
• 91
• 93
• 93
.47
.47
.47
.47
•01
• 57
• 01
• 01
3.35
• 91
3*36
3.09
.01
4.47
5*80
4.03
4.82
5.39
6.80
9.86
8.31
9*05
8*61
3*61
7.76
9.02
9.21
8.33
8.76
7.14
6*19
3.25
5.T6
5*31
1 .69
6.63
8.82
8.05
6.23
8.82
4.38
5.38
8.17
6.44
7.29
• 2
7.6
.1
• 1
35*6
9.6
35.7
32.8
.1
47.5
61*6
42*8
51.2
57.2
72.2
104.7
88*2
96*1
91.4
38*3
82.4
95*8
97.8
88.4
93*1
75.8
65.7
34.5
• 61.1
56.4
18*0
70.4
93.7
85.5
66*1
93.7
46*6
57.1
86*7
68*3
77.4
-------
- 281 -
TABLE J-XI
CAPS Feed Ratios and Regenerator Sulphur
Run 3
Sheet 2 of 5
DAY.HOUR
2.1930
2.2030
2*2130
2.2230
2.2330
3.0030
3*0130
3.0230
3*0330
3.0430
3.0530
3.0630
3.0730
3.0830
3.0930
3.1030
3.1130
3.1230
3.1330
3•1430
3.1630
3.1730
3.2305
.0005
.0130
.0230
• 0330
.0430
.0530
.0630
• 0730
• 0830
.0930
.1030
• 1130
1230
• 1330
• 1430
• 1530
1630
• 1730
VELOCITY
FT/SEC
3.55
3*58
3.52
3.68
3.52
3.36
3.46
3.55
3*45
3.47
3.41
3.42
3*42
3.29
3*21
3.28
3.35
3*10
3.12
3.17
3.37
3*41
3*52
3.73
3.88
3.98
3.97
3.89
3.99
3.85
3.73
3*85
3.95
3.87
3.75
3.81
3.73
3.73
3.69
3*61
3*62
AIR/FUEL
PCT STOIC
20.8
20.9
20*4
21 .3
20.4
19.9
20*4
21.3
20.3
20.4
20.5
20.5
20.5
19.5
20.1
19.6
20.1
19.2
19.2
19.6
19.9
20.0
19.4
22.1
22.1
23*1
23.0
22.9
23.0
21*8
21*3
21.9
22.9
22.9
21.9
21*9
21 .4
21.3
21.0
20.9
20.7
CA/S RATIO
MOL/MOL
.47
.47
.08
.08
.08
.08
• 08
.08
.08
.08
.53
• 53
.53
.53
• 53
• 73
• 73
• 73
• 73
• 01
• 01
• 01
• 02
.02
3*58
3.58
.79
.79
.79
.79
.79
.79
.79
.79
.51
.51
• 51
• 51
• 51
• 34
• 34
REGENERATOR SULPHUR
LB/HR PCT OF FED
6.30
6*29
6.75
6*52
7.92
6.31
6*92
6.95
7.28
7.60
6*43
7.62
6.49
6.78
6.14
6.30
5.69
6.11
6*96
6.28
4.52
4.72
5.94
4.60
2.44
3.74
2.57
2.56
4.45
5.74
6.77
5*38
8.54
5.22
4.08
6.89
6.74
6.49
7.20
6.25
7.28
66.9
66*8
71.6
69.?
84. 1
67.0
73.5
73.8
77.3
80.7
68*3
80.9
68*9
72*0
65.2
66.9
60*4
64.9
74.0
66.7
48*0
50.1
63*4
49.1
26*1
39.9
27.4
27.3
47.5
61.2
72.2
57.4
91.2
55.7
43.5
73.6
71.9
69.2
76.8
66.7
77-7
-------
- 282 -
TABLE J-XI
CAFB Feed Ratios and Regenerator Sulphur
Run 3
DAY.HOUR
4.1830
4.1930
5.0130
5.0230
5.0330
5.0430
5.0530
5.0630
5.0730
5.0930
5.1030
5.1130
5.1230
5*1330
5.1430
5.1530
5.1630
5.1730
5-
5.
5.
6.
6.
6'
2200
2230
2330
0030
0130
0230
6.0330
6.0430
6.0530
6.0630
6 . 1 1 30
6.1230
6.1330
6.1430
6.1500
6.1830
6.1930
6.2030
6*2130
6.2230
6.2330
7.0030
7.0130
VELOCITY
FT/SEC
3.54
3.54
3.87
3.79
3.70
3*66
3.64
3.64
3.72
3.52
3.75
3.64
3.61
3.50
3.42
3.37
3.10
2.78
3.73
3.63
3.37
3.35
3.10
2.99
3. 13
3.06
3.08
3.07
3.90
-
-
-
3.27
-
-
3.79
3.82
3.93
3.99
3.95
3.95
AIR/FUEL
PCT STOIC
20.3
20.7
22.8
21.7
21.4
22.0
21.5
21.5
22.0
20.6
22.6
21.7
21 .4
19.7
20.0
19.6
16.9
15.4
21 .3
19- 7
21 .6
21.4
19.5
18.0
18.2
18.5
18.2
17.7
19.4
20.2
19.9
15.9
18.3
23.3
21 .4
19.7
19.6
20.7
20.7
21.2
21 .2
CA/S RATIO
MOL/MOL
• 34
• 34
• 75
.69
.73
2.22
2.17
2.17
2.22
.27
.27
.27
.27
.27
.27
.27
.99
• 99
2.00
2.00
2.00
.55
.55
• 55
-55
• 55
• 55
• 55
.52
• 51
• 51
• 51
• 52
>80
• 80
• 81
• 81
• 80
• 35
• 35
• 35
REGENERATOR SULPHUR
LB/HR PCT OF FED
5.
3<
7.11
5.84
• 04
• 10
12
72
3.74
5.99
6.68
5.32
6.43
5.18
5.94
5.65
5*89
5.09
4.60
6*42
• 21
3.04
3*86
5.36
6.82
5.95
5.29
6.74
6*67
2.97
4.02
5.23
6.31
5.91
6.37
75.9
62.3
• 4
1.1
56*4
42.1
41.2
66.0
75.4
59.5
71.9
57.9
66*4
63.2
65.9
56*9
51.5
71.8
2.4
34.2
43*4
60.2
76.7
66.9
59.4
75.8
75*0
31.2
42*2
54.6
65*9
61*7
66.5
-------
TABIE J-XI
Sheet 4
CAFB Feed Ratios and Regenerator Sulphur
Run 3
DAY.HOUR
7.0230
7.0330
7.0430
7.0530
7.0630
7.0730
7.0830
7.0930
7.1030
7.1130
7.1230
7.1330
7-1430
7.1530
7.1630
7.1730
7.1830
7.1930
7.2030
7.2130
7.2230
7.2330
8*0030
8.0130
8.0230
8*0330
8*0430
8.0530
8.0630
8*0730
8*0830
8*0930
8 . 1130
8*1230
8.1330
8.1430
8.2030
8.2130
8.2230
8*2330
9.0030
VELOCITY
FT/SEC
3*98
4.00
3.98
4.01
3.94
3.89
3.84
3.74
3.79
3.82
3.85
3.86
4.38
3.86
3.86
3.80
-
3.84
3.85
3.71
3*69
3.70
3.73
3.72
3*68
3.76
3-77
3.73
3.75
3.69
3.76
3.79
3.70
3.70
3.76
3*67
4.30
4.14
4.27
3*83
3.98
AIR/FUEL
PCT STOIC
21 .2
21 .2
21.2
210
21.2
20.7
20.8
20.3
.20.3
20.3
20.3
20.4
20.3
20.2
19.1
19.4
18.7
19.5
19.3
18.5
18.5
18.5
19.0
18.9
18.9
19.5
19.6
19.5
19.5
19.5
24.9
25.0
25.3
25.3
26.7
25.7
22.2
22.7
22.9
20.5
20.6
CA/S RATIO
MOL/MOL
• 35
.35
.35
.89
• 89
.89
.89
.16
.16
.16
.16
-16
.16
.90
.90
.90
.90
.90
.90
.90
.90
.51
.51
.51
.51
•51
2.05
2.05
2.05
.90
5
5
5
5
5
.64
• 39
• 39
.39
.39
.81
REGENERATOR SULPHUR
LB/HR PCT OF FED
6*21
6*38
6.40
7.02
6.37
6.29
5.81
7.18
.16
.92
.17
.65
6.
6<
6.
7.
7.91
8.95
• 61
45
1 .
7,
6.56
6.95
5.24
6*48
4.90
6.07
77
31
86
52
4.
6<
5.
5.
5.61
6.47
7.08
5.93
6*12
6.60
6.75
6.81
.44
7.32
.00
6.27
6.03
4.76
6.05
66.6
66*8
73.3
66.5
65.7
60.7
75.0
64.4
72.2
64.4
7<9 . Q
82.
93.
16.
77.
68 <
72.
54.
67,
51.
63.
49.
65.
61
57.
58
67,
73.
61
f,
5
«
8
5
6
7
7
2
4
8
9
2
6
6
6
9
9
81.7
88.1
90.1
90.9
5.9
97.8
.0
71.1
68.3
53.9
68*6
-------
- 284 -
TABLE J-XI
Sheet 5 of 5
CAFB Heed Ratios and Regenerator Sulphur
Run 3
DAY.HOUR
9,
9.
9
9.
9.
9
9
9
9
9
9
0130
0230
0330
0430
0530
0630
0730
0830
0930
1030
1 130
9.1530
9-1630
1730
1830
1930
2030
2130
2230
2330
0030
0130
0230
0330
0430
10.0530
10.0630
10.0730
10.0830
10.0930
10.1030
1 130
1230
1330
1430
9
9
9
o
9
9
9
10
10
10
10
10
10
10
10
10
10*1530
10.1630
VELOCITY
FT/SEC
3.99
3.94
4. 12
4.05
4. 1 1
3.93
4.06
4. 19
4.09
4.27
3.95
3.74
_
3.91
3.88
3.91
3.78
3.94
3. -99
3.74
3.89
3.88
3.89
3-89
3.88
3.92
4.03
3.91
3.68
3.88
3.80
3.77
3.72
3.72
3.73
3.28
AIR/FUEL
PCT STOIC
22. 4
21.7
22.3
21.8
22.2
21.1
20.9
21.7
21.7
-
21.0
20.4
21 .4
21 .6
21 .4
21 .8
21 .2
22.0
22.2
20.4
21 .4
21 .3
21.4
21.3
21 .2
21 .2
22.4
22.2
22.2
23.3
23.1
23*0
22.5
22.5
22.5
17.6
CA/S RATIO
MOL/MOL
• 81
• 81
.81
.81
.04
.04
.04
• 04
.04
.04
.05
.91
.91
.91
.91
.91
.91
.91
.91
.91
.98
• 98
.98
.98
.98
.98
• 98
• 53
• 58
• 58
• 58
• 58
• 58
• 58
• 58
• 58
REGENERATOR SULPHUR
LB/HR PCX OF FED
5.95
5.93
5.52
5.74
5.60
5.98
5.76
6.06
5.87
5.43
.07
5.84
6.18
6.61
6.88
6.52
6.69
6.99
6.
6
7.
.80
.58
.34
6.94
6.94
7.26
6.97
7.05
6.98
6.38
5.74
5.
6.
5.
60
11
38
5.73
6.30
5.50
67.4
67.2
62.5
65.0
63.4
67.7
65.2
68.7
66.5
61.5
.8
66.7
70.6
75.5
78.6
74.4
76.3
79.7
77.7
75.1
83.8
79.2
79.2
82.9
79.5
80.5
79.7
80.3
72.3
70.5
76.9
67.8
72.1
79.3
69.2
-------
- 285 -
TABLE J-XII
Sheet 1 of :•-
CAFB Pilot Plant Solids Circulation Rates Computed
from Heat Balance
TIME
DAY.HOUR
FLOW OF HOT SOLIDS IN POUNDS PER HOUR
CAO CIRCULATION
TOTAL SOLIDS FLOW
TO FROM
REGENERATOR REGENERATOR
0230
0330
0430
0530
• 0630
• 0730
• 0830
• 0930
• 1030
• I 130
• 1230
• 1330
• 1430
• 1530
• 1630
• 1730
• 1830
• 1930
• 2030
• 2130
• 2230
• 2330
2.0030
2.0130
2*0230
2.0330
2.0430
2.0530
2.0630
2.0730
2.0830
2.0930
2.1030
2.1130
2.1230
2.1330
2.1430
2.1530
2.1630
2.1730
2*1830
6321
1927
1764
933
673
577
485
574
861
1015
618
648
746
639
564
463
516
687
656
710
469
551
508
640
549
546
571
757
643
597
388
532
440
345
498
442
350
498
395
474
6336
1949
1781
947
695
596
502
594
873
1035
639
665
764
660
586
490
541
715
684
732
494
577
535
666
576
571
595
779
666
617
408
558
465
364
524
459
366
522
415
497
6350
1965
I 79'.
959
700
608
504
600
878
1034
632
662
759
654
576
473
527
703
673
734
484
564
520
655
563
562
590
782
661
607
396
543
449
354
509
451
356
507
404
485
-------
- 286 -
TABLE J-XII
Sheet 2 of 5
CAFB Pilot Plant Solids Circulation Rates
Computed from Heat Balance
TIME
DAY.HOUR
FLOW OF HOT SOLIDS IN POUNDS PER HOUR
CAO CIRCULATION
TOTAL SOLIDS FLOW
TO FROM
REGENERATOR REGENERATOR
2.1930
2.2030
2.2130
2.2230
.2330
.0030
.0130
.0230
.0330
.0430
.0530
.0630
.0730
.0830
.0930
J.1030
3.1130
T.1230
. 1330
. 1430
> 1630
. 1730
.2305
.0005
.0130
0230
130
.0430
.0530
.0630
.0730
.0830
.0930
. 1030
1 130
1230
1330
1430
4.1530
4.1630
4.1730
2.
3-
3.
3.
3.
3-
3.
3-
3.
3.
3.
3.
3.
3.
3-
3-
4<
4.
4
4
4.
4<
4<
4.
4.
4.
4.
4.
4.
4.
4.
383
430
449
366
438
319
385
414
419
472
377
327
449
512
501
431
424
434
402
536
386
592
695
634
1028
1095
856
1275
1078
61 I
749
858
467
828
703
584
562
599
614
664
402
450
469
384
460
338
406
436
441
495
495
399
345
471
535
524
451
444
456
422
556
402
616
720
654
1051
1116
874
1301
1 105
637
771
888
486
849
728
607
584
623
635
689
392
441
459
373
446
327
395
424
428
484
487
384
333
461
528
517
444
435
444
41 1
555
394
612
722
661
1054
1 123
880
1306
1 105
631
768
879
482
850
720
599
577
614
629
680
-------
- 287 -
TABLE J-XII
Sheet 3 of f.>
CAFB Pilot Plant Solids Circulation Rates
Computed from Heat Balance
Run 3
TIME
DAY.HOUR
FLOW OF HOT SOLIDS IN POJNDS PER HOUR
CAO CIRCULATION
TOTAL SOLIDS FLOW
TO FROM
REGENERATOR REGENERATOR
5.
5.
5.
5.
4.1830
A.1930
5.0130
5.0230
5.0330
5.0430
5.0530
5.0630
5.0730
.0930
.1030
• 1130
.1230
5.1330
5.1430
5.1530
5.1630
5.1730
5.2200
5.2230
5.2330
6.0030
6.0130
6*0230
6.0330
6.0430
6.0530
6.0630
6.1130
6.2030
6.2130
6.2230
6.2330
7.0030
7.0130
573
479
654
684
743
600
631
497
677
719
684
652
684
585
289
881
909
1492
552
478
413
448
551
541
221
661
518
543
551
523
428
597
499
675
706
766
623
654
516
701
740
707
675
707
605
306
903
931
1513
574
500
434
467
576
565
222
682
539
565
573
545
448
5SB
671
709
772
619
648
512
695
736
704
671
702
602
304
910
934
1545
571
490
427
463
569
559
221
689
541
563
568
540
44!
-------
- 288 -
TABl£ J-XII
Sheet 4 of 5
CAFB Pilot Plant Solids Circulation Rates
Computed from Heat Balance
TIME
DAY.HOUR
FLOW OF HOT SOLIDS IN POUNDS PER HOUR
CAO CIRCULATION
TOTAL SOLIDS FLOW
TO FROM
REGENERATOR REGENERATOR
7
7,
7,
7.
7.
7.
7,
7.
7,
7.
7.
7.
7.
7.
7.
7.
7,
7,
7.
7.
8.
8<
8<
8<
8.
8.
8-
8-
8<
8<
8.
8<
8.
8-
8<
8.
8.
8.
0230
0330
0430
0530
0630
0730
0830
0930
1030
1 130
1230
1330
1430
1530
1630
7.I 730
1930
2030
2130
2230
2330
0030
0130
0230
0330
0430
0530
0630
0730
0830
0930
1 130
1230
1330
1430
2030
2130
2230
2330
9.0030
468
441
422
438
426
433
441
354
408
362
458
336
398
362
514
449
452
458
41 7
424
449
445
406
481
498
401
420
460
421
427
479
505
462
495
146
626
570
514
471
489
462
442
458
447
454
462
374
428
382
478
356
420
384
531
471
474
478
437
444
471
466
427
503
521
422
441
482
440
447
499
526
482
518
146
648
591
534
491
483
454
434
448
439
447
457
362
421
371
471
342
406
367
538
459
463
473
428
440
465
465
421
498
519
417
433
472
432
439
490
517
473
508
146
642
585
532
484
-------
- 289 -
TABLE J-XII
Sheet 5
CAFB Pilot Plant Solids Circulation Rates
Computed from Heat Balanee
TIME
DAY*HOUR
FLOW OF HOT SOLIDS IN POUNDS PER HOUR
CAO CIRCULATION TOTAL SOLIDS FLOW
TO FROM
REGENERATOR REGENERATOR
9.0130
9*0230
9.0330
9.0430
9.0530
9.0630
9.0730
9.0830
9*0930
9.1030
9.1130
9.1530
9.1730
9.1830
9.1930
9.2030
9.2130
9*2230
9.2330
10.0030
10.0130
10.0230
10.0330
10.0430
10.0530
10.0630
10.0730
10.0830
10.0930
10.1030
10.1130
10.1230
10.1330
10.1430
10.1530
10.1630
514
536
448
782
734
639
649
652
640
590
175
624
610
710
519
406
520
455
435
41 1
421
459
919
479
51 1
558
488
488
521
541
457
510
480
476
445
534
557
467
806
759
662
672
675
664
612
173
647
634
734
542
427
541
476
454
431
442
481
937
501
532
580
510
509
542
562
477
530
500
496
465
528
551
460
803
758
657
669
671
660
609
171
642
629
728
535
419
532
466
445
420
432
472
965
491
523
571
500
502
537
557
470
526
494
489
461
-------
- 290 -
H. Figures from Continuous Pilot Plant Runs
-------
5
>,
u
CO
fc,
a>
M
r;
n
K
w
CJ
CAFB Pilot Plant Operation log (Run 1)
Sulphur Removal Efficiency and Lime Addition Hate
100
Sulphur Removal Efficiency
iiLuL-LiJL
w ..... i » .V .
i — H-*-r-
r 1 - •;::: f" • '"nTrn 1: r rqi
8 13 18 23
Data, August
12
15
-------
20
r.'FR "lint Plant Operation Log (Run 1)
Air Puel Ratio and Superficial Gas Velocity
o
*-!
t,
J - I ,-. . II-. -^ --. .-. - I - t
, 4 I- .-; --L.I • .-,-;. - .. ...i-1 Li. . I . ,-: i
.-4-. .-!__;-. I.L; ! .,-. -J.;-,-, ; ' •
-
18
'.:{(.,
16
o
o
l-\
to
K
K
Oi
a
. Air Fuel Ratio .-i- ..j ^^^
. 1.. .1 J. I i. i_! ; j
-H-(- I t-H : ;
Superficial Gas Velocity i~ \\~-. '. 1-
11
12
16 21
7 12 1?
Time, Hour
13
Date, August
22
8 13 18
15
Fig J-2
-------
CAFB Pilot Plant Operation log (Run 1)
' i Gasifler Pressure
f T
13 18 PJ
12
Time, Hour
15
Date, August
15
-------
CAFB Pilot Plant Operation Log (Run 1)
and Gaalfier Bed Temperatures
£
*J
o
E
a
;-. 1.! .:,} Regenerator Bed Temperature I • - • •
i ' ———• j-*——>
,. .!' * ";
. . - - - I..L... ;.(.
•• -i : • : |i M-r h- • r
_Ll^4l ililidi 1 -*
,. I j _,i, .,
: \ ;--; !'- ' il - ' 111 '•' h" i rOaslfler Bed Temperature •; ','•-'•-•'• '..:..'•-. i-
- — i . i • — ^ — LJ — ; • ' ' ' li - ' ! • ' • _ , __ " .'. I :'*""" j"" : :'!' ": • :
- ! ' -'• H ; • ' J •-• •- - •• ' ' ••» i ^>s •-• : i-: ' - ! : _• _ «! • - • I I ..... i •
; • j r- • ;•-•" i•• I-'-1 i :
i . I... , .. | | , - .1. . J.^ _,.._:.
-- -
900
800 .._'.
12
Time, Hour..
15
Date, August
15
Fig J-k
-------
CAFB Pilot; Plant Operation IX»g CRu»
Regenerator Oas SO,, Concentration and Regenerator Air Rate
^
K
12
4)
+>
£
iu
<.
i~
0
-------
CAFB Pilot Plant Operation Log (Run 1)
Regenerator Gas Oxygen and COp Concentrations
r' hagenerator Gas Oxygen Concentration
LM: i .
! ' i ' ' " . ' ' ' : ••••-; rn-r
Regenerator Gas C02 Concentration
o\
11 lo 2J 2
12
12 17 "'?. ^
Time, Hour
Date.
35,
Fig J-6
-------
Pilot Plant Operation Log (Run 1)
Selectivity of Oxidation of CaS to CaO
o
a)
o
o
-p
CO
-------
C.'FB Pilot Plant Operation Log (Run 2)
20
5
o
•o
03
10
o U.
in
•O
•p
&
100
Solids Addition and Removal Rates
J - ri
•ar
• •t '
,-:
. ,- i
r-:Tr
• '. - 1 -
• L
;iii
. . i.i
•nit
fli;.
-| -r *~-
i ' t
- i - 1
I ! T '
. -1 - ,-
. I.I.I
-, (- '< 1
i •; t • *
!?!:.
lT-n:
-i ::t •
.-I ,.
1 . .. .
1 1 '-'
i j !
I : i .
-f • :
-U^ -, -.
"i ri "
-f H-
lit-
(- ;- + -.
TJ i"
1 .
:iri.--
.. - i—
»• - L « •
r '" '
,' •
1
: • f
: -t
li-ii-
IM"
H-Hrr' : " • - '• '""ttr
tT;4-f-~T Boiler Solids Removed f.p
1 i - . 4 ... 1-
. : ...i | . i j ;
1 • • !_[:• ;r( i
i--^^;
i!H
j . v.
. 1- :
1 -i . -;
; ! ri
1 ; !1_
— t-r}
"5ii
-- -i -'-< -
-Til
!
-*•- i
. , .. j
Trtr
*-. ., -^ -
_i P
i ;:; .::
•-'-T t
•-• J -j
i 14 i
1 • '
• » *-;-
. i r j
' * i
• T -i -'
-MJ4-
^!f
Stt-
r •
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.... 1 1.
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r- 1 ' --
laes Li
i . :
i - . * ',-
i. . . .
i ^n.::;.
i-iUVr;
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-i -i
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t-i-
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Iflf
t^rr
fiJ-t '^;
JWjMn
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if-
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Regenerator Cyclone J-< 4- -
Bollds Removed J.l.j |
'=•-: 1 Hi '-IH-tlir-
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:&
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Ei:
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ir
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11
'. J-8
50 Uii_i±iz:4j:
-------
CAPB Pilot Plant Operation Log
2)
Oil Peed and Calcium/Sulphur Feed Ratio -
1900 0500 1500 0100 1100 2100
78 9 October
Time, Hour/bate
0700
1700
10
0300
1JOO
11
Fig J-9
-------
CAPE Pilot Plant Operation Log
2)
"asifler Air/Fuel Ration and Gas Velocity
% 8
K O
O
o
10
a
o o
a
H Ul
!-i-i J-J.iii-i-i-i—i.i-^-VI-i • ;.,
Alr/Puel Ratio /-J^ffr
i i ' ;-r-r -,--• i ;•——' -! H't
E^II-M ^-imid
j i ; Casifier Superficial Oaa V*locityi
' 'i^J_t.:j.::
.I i if u -i i
-H-fr I-H-
—1 t-i.t r i .
1900
0500
1500
0100
.7
1100 2100
9 October
Time, Hour/Date
0700
1700
10
0300
U
Pig J-10
-------
CAPB Pilot Plant Operation Log
2)
Gaaifier Pressure and Bed Depth
4:--•*-A* *i--i i ;j -1
t- - J - ..-. 4 ; -:-i-j- i i.i -1 . i (.-^i 1 i ...
Pressure Above Oaaifler Bed
1900
0500
Time, Hour/Date
-------
CAPS Pilot Plant Operation Lop (Run 2)
Regenerator and Gasifier Bed Temperatures
•o
s
to
tt
2
1200
1100
1000
Regenerator Temperature
r li
rr-TTTTi-r-bh;:
Gasifier Temperature
iLbrLl/ '-l^u-.^-1 '• : '- LLlL.iJ;:--1
800
1900
0500
1500
0100
1100 2100
9 October
Time, Hour/Date
10
to
I
Pig J-12
-------
Pilot Plant Operation Lo
2)
Regenerator Air Rate and SO- Concentration
o
85
'
25
20 ~ IT
15
10
5
. Air Rate to Regenerator
' I ! f? I I •- :
!••-!• ! I I I I • •
i • : • • ! : ; • i i : i
SO? in Regenerator Gas
1900
0500
1100 2100
9 October
Time, Hour/Date
-------
CAFB Pilot Plant Operation Log (Run 2)
Regenerator Gas Oxygen and CO^ Concentration
to
3
s
«>*.
C •
§}
01
g
S
Oxygen in Regenerator Gas
"^'-£:£-*--H
COg In Regenerator Gas
1900
0500
1500
0100
1100 2100
9 October
Time, Hour/date
0700 1700
10
OJOO
1JOO
11
Pig J-lfr
-------
CAFB Pilot Plant Operation Log (Run 2)
Selectivity of Ca3 Oxidation and 30C Output of Regenerator
Selectivity of CaS Oxidation
302 Output of Regenerator
1900
0300
1500
1100 2100
9 October
Tine, Hoi:;/Date
-------
CONTINUOUS PILOT PLANT OPERATION LOG(RUN.3.)
SULPHUR REMOVAL
100,-
90
2 80
LJ
70
O
i
z
ce
o
I
-------
CONTINUOUS PILOT PLANT OPERATION UOG(RUN.3.)
OIL RATE
45
40
35
ct
o
o
UJ
30
o
5
CD
1.0230 1.1230 2.0030
2.1230 3.0030 3.1230 4.0030 4.1230 5.0030 5.1230
DAY: HOUR
< 45,-
40
35
30
8
?
5l
5.2230 6.1230 7.0030 7.1230 8.0030 8.1230 9.0030 9.1230 10.0030 10.1230
DAY:HOUR
g
-I
RG.J-17
-------
CAFB PILOT PLANT OPERATIONS LOG i RUN. 3.)
CUMULATIVE LIME ADD'TION AND REGENERATION SOLIDS REMOVAL
co
CO
1200,-
I 100
1000
< 900
g 800
cc 700
< 600
oc
£ 500
s 400
300
200
100
REGENERATOR TOTAL
REMOVAL
///UNIT NOT
GASIFING
5000
CO
CD 4000
Q- 3000
£ 2000
iu 1000
TOTAL LIME ADDITION
I
I
I
IO 30 5O 70 90 110 I3O 150 170 190 210 230
HOURS FROM START OF GAS I F 1C AT I ON
FIG. J-18
-------
CAFB PILOT PLANT OPERATIONS LOG (RUN. 3.)
CUMULATIVE SOLIDS REMOVAL FROM BOILER AND FLUE
1600 -
1500
1400
1300
1200
co MOO
CD
-i 1000
o
Id
<
oc
e
900
800
700
600
500
40O
300
2OO
100
0
MATERIAL TOTAL COLLECTION
IN THE BOILER AND FLUE
/// UNIT NOT GASIFYING
\/
O
/
1
*
-------
CAFB PILOT PLANT OPERATIONS LOG(RUN.3.)
RATES OF LIME FFED AND REGENERATOR SOLIDS REMOVAL
16
i .4
-------
CAFB PILOT PLANT OPERATIONS LOG (RUN.3.)
REGENERATOR CYCLONE AND BOILER FLUE SOLIDS REMOVAL RATES
20
16
12
o:
^ 8
CO
03 4
_j ^
•
1 f\
< °
>
O
K 20
16
12
8
4
••
_
„
-TL
-
-
_
-
- r~"
10
i— 1_
._.
L
-^
i
30
1 *— i
I I
LJ
L_^
I 1 I
50
;/
•/
~L
REGENERATOR CYCLONE
r-n COLLECTION RATE
1
-
_
1 — T7 /// UNIT NOT GASIFYING -1
; 1_n
/ TH
Hjl i
BOILER FLUE MATERIAL
COLLECTION RATE
[-1
rrf1 n H-i f
^ LJ~^ i
/
i—i /
n_ /
\
'
\ ., i '
4 \ '' &
70 90 110 130 150 170 190
HOURS FROM START OF GASIFICATION
-
•
-I
1
1
_ L- J 1
1, , ,-
210 230
u
FIG.J-21
-------
CONTINUOUS PILOT PLANT OPERATION LOG(RUN.3.)
BED REPLACEMENT RATE
O
o
O
CO
rf 0
l-
<
QL
a!
UJ
o
I
I
I
I
o
I
z
cr
o
O)
I
I
1.0230 1.1230 2.0030 2.1230
3.0030
DAY
3.1230
HOUR
4.0030 4.1230 5.0030 5.1230
i
&
IM
4r
(t
(D
o
I
cr
CD
I
I
5.2230
6.I23O -7.OO3O 7.I23O
8.OO3O
DAY :
81230
HOUR
9.OO3O
9.123O • IO.OO3O IO. I23O
-------
CONTINUOUS PILOT PLANT OPERATION LOG (RUN. 3.)
AIR/FUEL RATIO
30 r
25
o
* 20
i is
o
z
tr
ac.
13
ffl
V)
u_
O
* 30
o
uj 25
20
1.0230 1.1230 2.0030
2.1230 3.0030 3.1230 4.0030
DAY: HOUR
4.1230 5.030 5.1230
•
01
OJ
I
15
z
o:
K
CD
J
5.2230
6.1230 7.0030
7.1230 8.0030 8.1230 9.0030
DAY : HOUR
9.1230 10.030 10.1230
FIG J-23
-------
CONTINUOUS PILOT PLANT OPERATION LOG (RUN. 3.)
SUPERFICIAL GAS VELOCITY
4.5
4.0
3.5
ft 3.0
O
3
Ul
.2230
6.I23O 7.OO3O
7.I23O 8.0030
DAY : HOUR
J.
J.
8.I23O 9.OO3O 9.I23O IO.OO3O IO.I23O
-------
CONTIMUOUS PILOT PLANT OPERATION LOG (RUN. 3.)
GASIFIERGAS SPACE PRESSURE
251-
20
15
10
1.0230 1.1230 2,0030
2.1230 3.0030 3.1230 4.0030 4.1230 5.0030 5.1230
DAY 1 HOUR
OJ
25
^ 20
15
10
ui
o
8
i
z
2
5.2230
6.1230 7.0030
7.1230 8.0030 8.1230 10.0030 9.1230 IO.OO30 10.1230
DAY : HOUR
FIG. J-25
-------
CAFB PILOT PLANT OPERATION LOG (RUN. 3.)
DISTRIBUTOR PRESSURE DROP
35
30
25
20
15
x
O
3 1.0230 1.1230 2.0030 2.1230 3.0030 3.1230 4.0030
* DAY:HOUR
35
CO
5 30
I 25
20
15
10
5
4.1230 5.0030 5.1230
w
O)
oc
i
5.2230 6.1230 7.0O3O 7.1230 8.030 8.1230 9.0030 9.1230 10.0030 IO.I23O
DAY : HOUR
FIG. J-26
-------
-317-
CAFB PILOT PLANT OPERATIONS LOG(RUN.3)
PRESSURE-DROPS DOWNSTREAM OF GASIFIER
10
(9
UJ
(T
UJ
PRESSURE DROP THROUGH CYCLONES
PRESSURE DROP THROUGH BIFURCATED DUCT
PRESSURE DROP THROUGH BOILER TUBES
PRESSURE DROP THROUGH BOILER FLUE
1.0130 1.1230
2.0030 2.1230 3.0030 3.1230 4.0030 4.1230 5.O030
DAY: HOUR
5.1230
PRESSURE DROPS DOWNSTREAM OF GASIFIER
T
T
q 10
to
-r
T
PRESSURE DROP THROUGH CYCLONES
PRESSURE DROP THROUGH BIFURCATED DUCT-
PRESSURE DROP THROUGH BOILER TUBES
PRESSURE DROP THROUGH BOILER FLUE
5.2130 61230 7.0030 71230
8.0030 8.1230
DAY, HOUR
9.0030 9.1230
10.0030 10.1230
FIG.J27
-------
CONTINUOUS PILOT PLANT OPERATION LOG (RUN.3.)
GASIFIER BED DEPTH
to
UJ
X
o
z
20
18
16
14
12
10
1.0230 1.1230 2.0030
2.1230 3.0030 3.I23O 4.OO3O
DAY : HOUR
4.1230 5.0030 5.I23O
< 20
03 18
ui
E 16
<7i
o (4
12
to
o
i
o
5.2230
6.1230 7.0O3O 7.1230
8.0030 8.1330
DAY : HOUR
9.0030 9.1230 IO.OO30 IO.I23O
FIG J-28.
-------
CONTINUOUS PILOT PLANT OPERATION LOG (RUN.3.)
BED TEMPERATURES
1100
1000
o
o
UJ
oc
a:
UJ
a.
900
830
1100
1000
900
830.
REGENERATOR TEMPERATURE
I-
o
GASIFIER TEMPERATURE
O
ac
1.0230 1.1230 2.0030 2.1230 3.0030 3.1230 4.0030 4.1230 5.0030 5.1230
DAY : HOUR
REGENERATOR TEMPERATURE
UJ
o
§
I
IK
GASIFIER TEMPERATURE
at
j
5.2230 6.1230 7.0030 7.1230 8.0030 8! 230 9.0030 9.1230 10.0030 101230
DAY : HOUR
FIG J-29
-------
CONTINUOUS PILOT PLANT OPERATION LOG(RUN.3.)
GASIFIER TO REGENERATOR SOLIDS TRANSFER RATE
o
o
O
m
UJ
at
UJ
u.
CO
lOOOr
800
600
400
200
_L
_L
_L
-L
_L
.0230 1.1230 2.0030 2.1230
3.0030 3.1230 4.0030
DAY : HOUR
4.1230 5.0030 5.1230
OJ
ro
O
lOOOr
800
600
400
200
5.2230 6.1230 7.003O
7.1230 8.0030
DAY
8.1230 9.0030
HOUR
9.1230 10.0030 10.1230
RG. J-3O
-------
CONTINUOUS PILOT PLANT OPERATION LOG (RUN.3.)
SO2 FROM REGENERATOR
12
10
8
6
T 4
> 2
GO
5
UJ
o
o
o
CM
8
? 1.0230 1.1230 2.0030 2.1230
3.030 3.1230
DAY : HOUR
4.0030 4.1230 5.0030 5.I23O
ro
i
I2
10
8
UJ
o
o
I
OJ
m
5.2230
6.1230 7.O030
7.1230 8.0030 8.1230 9.O030
DAY : HOUR
9.1230 10.0030 10.1230
FIG. J-31.
-------
CONTINUOUS PILOT PLANT OPERATION LOG(RUN.3.)
AND 02 FROM REGENERATOR
CO
QC
O
I Or
9 0
1.0230
S I0i-
1.1230 2.0030
2.1230 3.0030
DAY : HOUR
3.1230 4.0030
4.1230 5.0030
. 15%
5.I23O
M
I
8
2
O
UJ
O
o
o
C02
3.223O
6.1230 7.0O3O
7.1230 8.003O
DAY : HOUR
J J (*&£=! .
8.1230 9.0030 9.I23O IO.OO3O IOTT23O"
-------
CONTINUOUS PILOT PLANT OPERATION LOG (RUN.3.)
SULPHUR FROM REGENERATOR
12
10
8
i-
o
i
z
01
•3 ip
co '^
10
8
6
4
2
UJ
o
o
I
z
-------
CONTINUOUS PILOT PLANT OPERATION LOG (RUN.3.)
SULPHUR FROM REGENERATOR AS % OF SULPHUR FED
o
UJ
UJ
o
UJ
QC
o
-------
CONTINUOUS PILOT PLANT OPERATION LOG (RUN. 3.)
% CaS CONVERTED TO CoO
9O
60
30
o
g 1.0230 1.1230 2.0030
a:
to
x CALCULATED FROM BED SULPHUR CONTENT
O
i
DC
ffl
2.1230 3.0030 3.1230 4.0030 4.1230 5.0030
DAY : HOUR
5.1230
o
o
90
60
30
u
o
i
o
I
rvn
x CALCULATED FROM BED SULPHUR CONTENT
oe x
5.2230 6.1230 7.0030 7.1230 8.0030 8.1230 9.0030 9.1230 10.0030 10.1230
DAY : HOUR
FIG. J-35.
-------
-326-
Gasifier Distributor Nozzles After Run
Fig J-36
Empty Gasifier Interior After Run
Fig' J-37
•
•»XL
Nft&'H
%
-------
•327-
Gasifler Interiorf Bed in Place, After Run
Fig J-3&
Gaslfier Interior, Bed Partly Removed, After Run
Fig J-39
"X
>,
* , *
-------
-328-
Regenerator Distributor After Run
Fig
-------
-329-
Peposlt From Regenerator Bottom After Run-3.
Fig J-kl
-------
-330-
Regenerator Interior Prom Bottom After &vg» 3
Regenerator Top Gas Outlet After Run 3
-------
-331-
Rlght Hand (frsifler Cyclone From Top After Run !>
Fig J-kk
Left Hand Gaslfier Cyclone Prom Top After Run 3
Fig J-U5
-------
-332-
Outlets From Gasifier To Cyclone After Run 3. Fig J~k6
Local Deposits in Boiler Fire Tube After Run 3 Fig J-kJ
-------
-333-
Core of Solids from Boiler Tube After Run 3
Fig J-kQ
-------
-334-
Back End of Boiler After Run
Fig J-kg
-------
- 355 -
APPENDIX K
BURNING RATE MODEL FOR CARBON ON LIME IN CAFB GASIFIER
Tests in the CAFB batch unit have revealed an interaction between effectr
temperature and air/fuel ratio on the carbon concentration in the lime and on
loss in sulphur absorption efficiency.
It has been hypothesized that the concentration of carbon on lime is
determined by a dynamic equilibrium between the rate of carbon deposition
due to cracking of the fuel and the rate of carbon burning by the air supplie<
To test the reasonableness of this hypothesis, a model has been formulated
using a simplified geometry and contacting system together with literature
values for the rate of carbon oxidation.
Carbon levels computed with this model show order of magnitude corr-
espondence with values actually measured, and show the same type of dependency
on temperature and air fuel ratio as observed. The model also predicts an
increase in carbon content of lime with increasing superficial air rate at
constant temperature and air/fuel ratio. The rapid increase of predicted
carbon concentration below 800°C suggests that preheating the gasifier air
could improve operations at low air fuel ratio and high gas velocity. The
model also suggests that the height of oil injection to the gasifier could
be an important variable and warrants further investigation.
A. Physical Model
The physical model assumed is illustrated in Figure K-l.The gasification
reactor is assumed to consist of two zones, a carbon burning zone and a carbon
deposition zone. The burning zone is below the oil injector, and the deposition
zone is above it. It is assumed that oxygen which reaches the deposition zone
reacts with hydrocarbon gases and becomes unavailable for coke burning.
Each zone behaves as a well mixed reactor for solids but is plug flow in gas.
The temperature is constant throughout. Lime circulates between the two zones
carrying a layer of coke with it. Coke is burned only in the burning zone.
The coke content of the solids is low enough to permit thin layer burning
kinetics to apply.
Deposition Zone
Lime
t Circulation
J,- - -
Burning Zone
Top of Fluid Bed
Oil Injection
Air Distribution
Simplified View of Carbon Burning
in Gasifier
Figure K-l
-------
The rate of carbon burning is assumed to follow the first order kinetic
equation given by Weisz and GoodwinCS), modified by inclusion of a term for
average oxygen concentration. Lime circulation rate is estimated from an
empirical equation reported by Kunii and Levenspiel(9). The total coke
production rate, equal to the steady state coke burning rate, is taken as
the product of oil feed times fraction Conradson Oarbon. This assumption is
roughly in agreement with fluid coking experience. The average bed coke
level is assumed to be the weighted average of the levels in the two zones.
Actual equations employed are described in Section D.
B. Results
Fig.K-2 shows curves of coke concentration vs temperature calculated
with the model. The principal feature is the rapid rise of carbon content
as temperature falls below about 800°C. The calculations also show a steep
rise of carbon content with decreasing air/fuel ratio and with increasing
superficial gas velocity. These calculated results are all directionally in
agreement with observed behaviour of the CAFB batch reactor results. However
the experimental results generally show a somewhat higher increase of carton
with decreasing temperature than calculated.
-------
Computed Coke Concentration
on Gasifier Lime
o
12
10
8
6
Superficial
Velocity
ft/sec
780 820 860
Gasifier Temperature, °C
-------
- 258 -
The model contains several parameters which could be adjusted to bring
experimental results into better agreement with calculations. For example
the boundary between coke burning rnd deposition zones has been taken as
the level of the oil injector. Lowering this boundary would increase the
calculated carbon concentrations. The same result would be achieved by a
higher coke make than given by Conradson Carbon. It is recommended that
additional data on actual coke concentrations be compiled before attempts
are made to adjust parameters.
Table III lists values of carbon concentration and recycle rate computed
for various conditions. This table shows that differences between predi"V'J
coke contents in the two zones are not large in magnitude at the circulate
rates estimated. At the higher temperatures, however, the differences do
become large relative to the absolute values involved. The calculated
circulation rates gives turnover times of the order of a minute. With the
circulation rates listed, the average particle residence times in the zo"-^
are as follows:
Average Residence Time, Min
Superficial Velocity
ft/sec. Burning Zone Deposit inn Zone
4 0.51 1.J4
6 0.28 0.75
8 0.20 0.52
C. Implications
This model assumes a uniform temperature in the carbon burning zone.
Obviously the real case is a non uniform temperature distribution due to the
cooling effect of the entrance air and the heating effect of the coke
combustion. It is likely that the cooling effect reduces the effective
volume of the burning zone and increases the carbon concentration. Pre-
heating the gasification air supply should have a beneficial effect on ca.^">;
burning and permit operation at lower air/fuel ratios and higher gas velooi+V:
If the effective boundary between zones is determined by point of oil
injection, the location of this point could affect carbon burning. Obviously
if the oil were injected too low,the air would contact it before burning much
coke. On the other hand, if oil injection is too high there could be insuf-
ficient height for sulphur absorption above it. Determination of optimum
oil injector location should be made during Phase II studies in the continuous
pilot plant.
-------
- 339 -
D. Equations Used in Model
The equations used in the model and their basis are summarized below.
Table I lists the symbols and nomenclature used, and Table II is a listing oi
the computer prograj«B in Basic language. The equations are formulated on the
basis of a steady state operating period of the batch reactor.
Coke Input to System
All of the Conradson. Carbon of the feed oil is assumed to form coke.
Coke input = F * Cl (l)
Coke Burning Rate
Weisz and Goodwin ("'have reported kinetics for coke burning on catalysts
and other solids. It is assumed that the same kinetics apply to burning of coke
on lime in the gasifier.
The Weisz and Goodwin equation is
f= -KC (2)
whereK =4x10? exp(-376GO/R*Tl) rain'1 (3)
The Weisz and Goodwin work vias conducted with air at atmospheric pressure.
For the present work it is assumed that the ante is proportional to oxygen concentra-
tion, and a term for average oxygen concentration is included.
= -K* C * Al (4)
dt
Multiplication by the quantity of solids in the burning zone gives an
equation for mass rate of coke burning.
SI * dC = - K * C * Al # SI (5 )
dt
Coke Concentration in Burning Zone
Equating the rate of coke input to th^ rate of coke burning gives an
equation for coke concentration in the burning zone.
P # Cl = K * C * Al * SI (6)
C = P * C1/(K* Al * SI) (6a)
Air Consumption
The coke is assumed to contain hydrogen in addition to carbon. The
hydrogen also will consume oxygen. The products of carbon oxidation are CO
and C02- From experimental results it appears that the composition is kinetically
-------
conti'olled as there is much less CO present than.equilibrium would indicate.
The jramgof oxygen consumed for 100 Grams of oil feed is given by:
01 - Cl » (H3 * 8 + C2 * ((C3 * 16 + (1-C3) * 32)/12:) (7)
The quantity of oxygen supplied per 100 grams of oil is related to the
percent of stoichiometric air/fuel ratio by:
02 = 14 * S * 0.232 (8)
The fraction of oxygen consumed is:
03 = 01/02 (9)
The log mean average oxygen fraction in the burning zone is:
Al = 03/Log (1/(1-03)) (10)
Coke Concentration in Deposition Zone
A coke balance into and out of the carbon burning zone provides an
equation for the coke concentration in the deposition zone.
Coke transferred in = Coke Burned + Coke transferred out
C5 * R3 = P * Cl + C » R3 (11)
C5 = (F * C1/R3) + C (lla)
Coke Circulation Rate
The vertical circulation of coke across the boundary between zones is relatej
to gas velocity, particle size, and other fluidisation variables. The rate of coke
circulation has not been measured for the CAFB case. For this model, the equatic.t.
of Talmor and Benenati as reported by Kunii and Levenspiell9) is used. The
original equation is:
S = 0.785 (Uo-Umf) EXP (-66.3 dp) (12)
In the nomenclature of this model it becomes:
R3 = 60 * 0.785 * A * 30.48 (V-VL, EXP (-.00663D) (13)
Average Coke Concentration
The average coke concentration is determined from the relative heights
of the burning and deposition zones by:
C6 = (C * HI < C5 » (H2-H1))/H2 (l4j
-------
-341-
TABIJB K-I
Symbols and Nomenclature
A Bed Area sq. cm.
Al Log mean average oxygen concentration, fraction of that In air.
C Coke on lime in burning zone, percent-
Cl Conradson carbon of fuel oil, percent.
C2 Carbon In coke, wt. fraction.
C3 CO in carbon oxides from burning zone, mole fraction
C4 Line charge to bed, grams
C5 Coke on lime in deposition zone, percent
C6 . Average coke on lime in bed, percent.
dp Average lime particle diameter, cm ' '
D Average lime particle diameter, microns. .' '-
F Oil Feed rate, gm/mln. v !
HI Height of burning zone, cm '
H2 Height of bed, cm ;
H3 Hydrogen In coke, wt. fraction.
J Solids circulation rate between zones, gm/cm2-seo
K Oxidation rate constant, mln'1 (K=R2) ;
01 Oxygen consumption rate, gm/100 gm oil feed.
02 Oxygen feed rate, gm/100 gm oil feed.
03 Oxygen consumption, fraction of 02 fed
R Gas Constant, Cal/g Mole - Oeg kelvin.
Rl Gas density, Ib/cu-ft. . ' •
R2 Oxidation rate constant, mln-1 (R2WC).
R3 Solids circulation rate, gm/mln.
S Air fuel ratio, percent of stoichlometrte.
SI Solids in burning zone, g. :
t Time, minutes• '
T Temperature, Oeg C.
Tl Temperature, Deg K.
T2 Temperature, Deg R. - -
Uo Superficial Gas velocity, cm/sec.
Umf Minimum fluidisation velocity, cm/sec.
V Superficial gas velocity, ft/sec
VI Minimum fluidisation velocity, ft/sec.
-------
- 3*2 -
TABLB K-H
Computer Programme for Carbon Calculation
II")
0204
I30
149
150
160
I7fl
IPO
190
200
210
220
230
240
>>50
•260
270
280
290
3«0t
310
320
333
340
353
360
370
AIR/FUEL RATIOa****.* PERCENT STOlC-U
393
400
,11 e
420
430
440
450
460
470
490
500
519
520
530
540
550
563
570
OATA 4* 710* 15* 12.7, 46
OATA K2-7I3* 24R* 13., .9796*
DATA .5 * ?00* 1«5 . ,
READ V* T* S* HI* H2* C4* A* Cl* C2* H3* C3* D* VI
REK SUPERFICIAL VELOCITY* TEKP C, FRACTION STpXC AIR,
W P'JRN ZONE HEIGHT* CM. • "'
REf LIKE CHARGE GP* AREA SO CH* PCT CON CARSON,. FR C IN COKE
REPFR H IN COKE . ' "," . '
REM CO FRACTION* PARTICLE OIA MICRONS* KIN FLUID VELOCITY
PRINT"CALCULATION OF CARBON ON STONE" ..;.
PRINT" FOR"
PRINT 'JS.INC 6I0*H|*H2
PRINT MS ING 620* C4* A
PRINT USING 63Fl»Cl*C2
PRINT USING 640*H3*C3
PRINT. USING 650* D* VI
PRINT
FOR V»4 TO 8 STEP 2
FOR S*15 TO 30 STEP 5
PRINT USING 300*V*S
VELOCITY****.* FT/SEC.
PRINT -
PRINT USING 580 .-.-•,
PRINT USING 590 _ * . ^.-; ± ",,.
FOR T a 7C0 TO 900 STEP 20 - -: •• .;, .•
LET T1=T*273.16 . > . -
LET T2=1.B*TI
LET 01=Cl*tH3*P*C2*(CC3*16+(1-C3>*32)/12»
LET 02=I4*S*.23?
LET 03=01/02
LET Al =03/LOO<1/<1-03»
LET SI = (Ht/H2)*C4
LET AJ? = V*A*.001073
LET A3=A?*RI*60*454
LET r=A3/<.14*S)
LET R2=4£7*EXP<-37600/ **.#* **.#,»
FEED HEIGHT = ****.* CM
CHARGE(LI^E) = **#**.« GRAMS*
-H itatt.it Pr'<~T.'€.NJ
Ci;,\' .. f. Has a F.H.*
PARTICLE DIA = **»•». fIC«
-------
- 343 -
TABLE K-III
Summary of Computed Results
CONDITIONS
FEED HEIGHT
CHARGE ( LIME ) =
CON CARBON
COKE HYDROGEN =
PARTICLE DIA =
GAS VELOCITY STOICH
FT/SEC
4
"
n
ir
n
n
n
ti
n
•i
n
n
6
"
"
if
"
»
n
tr
"
n
n
n
8
"
"
»
»
»
M
It
It
If
II
II
PCT
15
n
if
20
n
n
25
n
"
30
if
15
M
"
20
"
M
25
M
II
30
"
15
IF
II
20
if
n
25
n
30
"
12.7
8200.0 GRAMS
10.0 PERCENT
.0204 FR. ,
800. MICRONS
AIR TEMP
C
700.
800.
900.
700.
800.
900.
700.
800.
900.
700.
800.
900.
700.
800.
900.
700.
800.
900.
700.
800.
900.
700.
BDO.
900.
700.
800.
900.
700.
800.
900.
700.
800.
900.
700.
800.
900.
BED HEIGHT
BED AREA =
COKE CARBON
CO FRACTION
, MIN FLUID VEL »
PERCENT COKE ON
B ZONE D ZONE
12.63 13.34
1.87 2.51
.38 .97
8.Y4 9-27
1.29 L77
.26 .70
6.69 7.11
.99 1.38
.20 .55
5.42 5.78
.80 l . 12
.16 .46
18.95 19.54
2.81 3.34
.57 1.06
13.10 15-55
1.94 2.34
• >9 .76
10.04 10.39
1.49 l.8l
.30 .60
8.14 8.43
1.21 1.47
•25 .49
25.26 25.81
3.74 4.24
.76 1.21
17.47 17.88
2.59 2.96
.53 .86
13.38 13.71
1.98 2.28
.40 .67
10.85 H.12
1.61 1.85
•33 -.55
46.0 CM
248. SQ.CM
.9796
.500
FR.
1.50 FT/SEC
LIME
TOTAL
13.14
2.34
.81
9.1?
1.64
• 58
7.00
1.27
.46
5.68
1.04
-38
19-37
3-19
.92
13.42
2.23
.66
10.29
1.72
• 51
8.35
1.40
.42
25.66
4.10
1.09
17-77
2.86
.77
13.62
2.20
.60
11.05
1.79
.49
RECYCLE
GM/MIN
4425
n
M
n
tl
M
tt
II
II
It
II
II
7965
II
II
II
If
II
It
II
II
II
II
II
11505
It
It
It
It
II
tf
It
It
II
II
II
-------
-II-
PHOTOGRAPH OF I-A BATCH REACTOR
FIG. 2.
-------
- 22 -
Figure 10 is': an interior view of the completed pilot plant
showing gasifier unit, boiler and mechanical equipment console.
'
CAFB PILOT PLANT INTERKDR
FIG. 10.
-------
Figure 11 shows the pilot plant exterior with oil feed storage
tank, rear end of boiler, cooling tower, and flue stack.
CAFB PILOT PLANT EXTERIOR
FIGJI.
-------
Assembly of the gasifier unit steel framework and casting of
the high density refractory concrete was performed with the gasifier in
place at the Esso Research Centre. After pouring each lift of concrete,
it was vibrated to provide maximum density.
Figure 15 is a photograph showing the gasifier at an early
stage of construction. The plastic patterns which form the regenerator
and solids transfer ducts are in place together with the wooden pattern
of the lower gasifier bed. Three of the pre-insulated mild steel
encasing walls also are erected.
GASIFIER DURING CONSTRUCTION
FIG. 13.
-------
-38-
AGGLOMERATES FROM BATCH REACTOR AFTER
CYCLE TEST WITH BCRI690 STONE
FIG. 20.
-------
- 185 -
The layout of the bed transfer ducts was largely dictated by the
geometry of the gasifying and regenerating compartments. It was not
possible, however, to produce a detailed design on theoretical grounds.
There was no way of predicting the dimensions and positions of the
catchment pockets which would provide satisfactory results for example,
and it woulc1 have required an excessive degree of confidence in the
outcome to have simply guessed the dimensions of the transfer ducts.
It was, therefore, decided at an early stage that the best procedure
would be to build a full scale cold model so that these finer points
could be settled experimentally. A view of the cold model is shown
in Figure Q* 4.
Fig.G-4
-------
- 191 -
PHOTOGRAPH OF GASIFIER WITH LOWER CONCRETE
LIFTS POURED
-------
- 192 -
PHOTOGRAPH OF COMPLETED GASIFTER
Fig G-9
-------
-326-
Gaslfier Distributor Nozzles After Run ~*>
Fig J-
Empty Gasifler Interior After Run
Fig J-37
-------
•327-
Gasifier Anterior, Bed in Place, After Run
Fig J-?8
Gasifier Interior, Bed Partly Removed, After Run
Fig J-39
\
-------
-328-
Recenerator Distributor After Run
Fig J-kO
-------
-329-
Deposit From Regenerator Bottom After Run
Fig J-
-------
Regenerator Interior From Bottom After Run
Fig J-
Regenerator Top Gas Outlet After Run 3
Fig J-l
-------
Right Hand Gasifier Cyclone From Top After Run 3
Fig J-H
Left Hand Gasifier Cyclone From Top After Run
Fig J-
-------
-332-
Outlets Prom Gasifier To Cyclone After Run J.
Fig J-U6
Local Deposits in Boiler Fire Tube After Run
Fig J-
-------
-333-
Core of Solids from Boiler Tube After Run
Fig J-48
-------
-334-
Back End of Boiler After Run
Fig J-
------- |