EPA-650/2-74-062-0
NOVEMBER 1974
Environmental Protection Technology Series
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EPA-650/2-74-082-0
REFINERY CATALYTIC
CRACKER REGENERATOR
SOX CONTROL -
STEAM STRIPPER
LABORATORY TEST
by
T. Ctvrtnicek, T. Hughes,
C. Moscowitz, and D. Zanders
Monsanto Research Corporation
Dayton Laboratory
Dayton, Ohio 45407
Contract No. 68-02-1320, Task 1
Phase II
ROAP No. 21ADC-031
Program Element No. 1AB013
EPA Project Officer: Kenneth Baker
Control Systems Laboratory
National Environmental Research Center
Research Triangle Park, North Carolina 27711
Prepared for
OFFICE OF RESEARCH AND DEVELOPMENT
U.S. ENVIRONMENTAL PROTECTION AGENCY
WASHINGTON, D.C. 20460
November 1974
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This report has been reviewed by the Environmental Protection Agency
and approved for publication. Approval does not signify that the
contents necessarily reflect the views and policies of the Ageitcy,
nor does mention of trade names or commercial products constitute
endorsement or recommendation for use.
ii
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ABSTRACT
The report summarizes experimental results from steam
contacting of spent catalyst used in petroleum refinery
fluid catalytic crackers. This concept has been identified
as a potentially effective means of sulfur emission control
for fluid catalytic cracker regenerators. Correlations be-
tween sulfur removal efficiency from the catalyst and the
product of steam residence time in the stripper with the
steam stripping rate are presented for several stripper
designs. The extent of by-product formation, a discussion
of pertinent equipment design, and recommendations for further
investigation and development of this concept are also includ-
ed. Additionally, the economics are presented as a function
of steam stripping rate and fluid catalytic cracker unit
size.
ill
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TABLE OF CONTENTS
Page
1. CONCLUSIONS 1
2. RECOMMENDATIONS 8
3. INTRODUCTION 10
4. EXPERIMENTAL WORK 12
4.1 EXPERIMENTAL EQUIPMENT 12
4.1.1 Catalyst Test Unit 12
4.1.2 Spent Catalyst Procurement 14
4.1.3 Catalyst Handling 19
4.1.4 Steam Stripping Experiments and Catalyst 22
Characterization
4.1.4.1 Catalyst Steam Stripping . 22
4.1.4.2 Determination of the Sulfur Content 25
of Coke
4.1.4.3 Determination of the H20 Content of 25
Spent FCC Catalyst
4.1.4.4 Determination of the Coke Content of 26
Spent FCC Catalyst
4.1.5 Other Experiments 26
4.1.6 Analysis 27
4.1.6.1 Sulfur Compound and Hydrocarbon 27
Determination
4.1.6.2 Analysis for the Total Organic Carbon 28
Content of Stripper Condensate
4.1.6.3 Volatility of Coke on Spent Catalyst 28
4.2 EXPERIMENTAL RESULTS 28
4.2.1 Catalyst Characterization 28
4.2.2 Catalyst Steam Stripping 31
4.2.3 Hydrocarbon Volatilization During Steam 33
Stripping
4.2.4 Volatility of Coke on Spent Catalyst 69
4.2.5 By-Product Formation During Steam 73
Stripping
4.2.6 Effect of Steam Stripping on Catalyst 75
Activity
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TABLE OP CONTENTS (Continued)
4.3 DISCUSSION OP RESULTS 8l
4.4 DATA REGRESSION ANALYSIS 88
5. STEAM STRIPPING PROCESS DESIGN 102
5.1 CATALYST STEAM STRIPPER DESIGN 102
5.1.1 Semi-Batch Fluidized Bed Reactor 103
5.1.2 Rate Controlling Factors 110
5.1.3 Other Stripper Designs 112
5.1.3.1 Continuous Fluldlzed Bed Reactor 113
5.1.3.2 Plug-Flow Stripper 116
5.1.3.3 Counter-Current Stagewlse Contacting 120
5.2 CONDENSER DESIGN 124
5.3 ACIDIFIER/PHASE SEPARATOR DESIGN 126
5.3-1 Equilibrium Relationship 126
5.3.2 Combined Condenser/Acidifier/Phase 138
Separator Design (-Alternative Sour Water
Treatment System)
6. ECONOMICS 141
APPENDIX A SPENT FCC CATALYST STEAM STRIPPING 154
DATA SHEET
APPENDIX B DETAILED COST ESTIMATES OF THE STEAM 164
STRIPPING PROCESS
vi
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LIST OP FIGURES
Figure
1 Control panel, reactor, and analysis system 13
2 Catalyst test unit, schematic flow diagram 15
3 Motionless mixer used in the catalyst reactor 16
4 Water deaeration system 17
5 Photograph of catalyst storage container 21
with charging flask attached
6 Weighing flask mounted on catalyst charging 23
bomb
7 Charging of the catalyst into a hot reactor 24
8 Results of steam stripping experiments on 35
B-Series catalyst
9 Results of steam stripping experiments on 38
C-Series catalyst
10 Results of steam stripping experiments on 40
D-Series catalyst
11 Results of steam stripping experiments on 43
E-Series catalyst with motionless mixer in
stripper
12 Results of steam stripping experiments on 45
E-Series catalyst without motionless mixer
used in the stripper
13 Results of steam stripping experiments on 49
F-Series catalyst
14 Results of steam stripping experiments on 53
H-Series catalyst
15 Results of steam stripping experiments of 55
H-Series catalyst
vii
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LIST OF FIGURES (Continued)
Figure Page
16 Results of steam stripping experiments on 58
I-Series catalyst
17 Effect of steam stripping rate upon total 70
organic carbon content of stripper condensate
18 Summary of results of the regression analysis 93
performed on B-Serles catalyst
19 Summary of results of the regression analysis 94
performed on C-Series catalyst
20 Summary of results of the regression analysis 95
performed on D-Series catalyst
21 Summary of results of the regression analysis 96
performed on E-Series catalyst
22 Summary of results of the regression analysis 97
performed on F-Series catalyst
23 Summary of results of the regression analysis 98
performed on H-Series catalyst
24 Summary of results of the regression analysis 99
performed on H-Series catalyst
25 Summary of results of the regression analysis 100
performed on I-Series catalyst
26 Schematic of spent catalyst steam stripper 104
27 Schematic of continuous fluidized bed stripper 114
28 Schematic of plug flow steam stripper 117
29 Counter-current, stagewise contactor 121
30 Distribution diagram for hydrogen sulfide 130
(Note that [S=] has appreciable concentration
only in strongly basic solutions
31 Effect of temperature upon ionization constant 132
for H2S
viii
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LIST OF FIGURES (Continued)
Figure Page
32 H2S concentration in vapor and liquid stream 136
33 Combined condenser, acidifier, phase separator 139
clarifier, and scum oil removal system
34 FCC catalyst steam stripping, total invest- 1^5
ment cost
35 Summary of operating costs 146
36 Summary of total investment costs, typical
case
37 Summary of operating costs, typical case
38 Summary of total investment costs, worst case
39 Summary of operating costs, worst case 150
ix
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LIST OF TABLES
Table Page
1 Results of Spent FCC Catalyst Sample 30
Characterization Tests
2 Results of Steam Stripping Experiments 34
B-Series Catalyst, Without Motionless Mixer
3 Results of Steam Stripping Experiments 36
C-Series Catalyst, Without Motionless Mixer
4 Results of Steam Stripping Experiments 39
D-Series Catalyst, Without Motionless Mixer
5 Results of Steam Stripping Experiments 41
E-Series Catalyst, With Motionless Mixer
6 Results of Steam Stripping Experiments 44
E-Series Catalyst, Without Motionless Mixer
7 Results of Steam Stripping Experiments 46
F-Series Catalyst, With Motionless Mixer
8 Results of Steam Stripping Experiments 50
H-Series Catalyst, Without Motionless Mixer
9 Results of Steam Stripping Experiments 54
H-Series Catalyst, Without Motionless Mixer
10 Results of Steam Stripping Experiments 56
I-Series Catalyst, Without Motionless Mixer
11 Results of Experiments Performed to Determine 59
the Effect of Steam Stripping on Coke Volatility,
B-Series Catalyst
12 Results of Experiments Performed to Determine 6l
the Effect of Steam Stripping on Coke Volatility,
C-Series Catalyst
13 Results of Experiments Performed to Determine 63
the Effect of Steam Stripping on Coke Volatility,
E-Series Catalyst
14 Results of Experiments Performed to Determine 64
the Effect of Steam Stripping on Coke Volatility,
F-Series Catalyst
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LIST OF TABLES (Continued)
Table Pa,
15 Results of Experiments Performed to Determine 65
the Effect of Steam Stripping on Coke Volatility ,
H-Series Catalyst
16 Results of Experiments Performed to Determine 68
the Effect of Steam Stripping on Coke Volatility,
H-Series Catalyst
17 Results of Coke Volatility Experiments on 71
H-Series Catalyst
18 Results of Coke Volatility Experiments on 72
E-Series Catalyst
19 Summary of By-Product Formation Experiments , 7^
Sulfur and Hydrocarbon Compounds
20 Calculated Stripper Off-Gas Analysis 76
C-Series Catalyst
21 Calculated Stripper Off-Gas Analysis 77
H-Series Catalyst
22 Summary of Results of By-Products Formation 78
Experiments, Ammonia Formation
23 Catalyst Activity Tests 79
24 Analysis of Stripper Off-Gas Condensate 86
25 Refinery Wastewater Loadings for Typical 87
Refining Technology
26 Steam Stripping Data Regression Analysis 91
27 Steam Stripping Requirement for Sulfur 101
Reduction to 200 vppm
28 Recommended Limits of Solids in Boiler 125
Feedwater
29 lonization Constants for the H2S-Water 129
System at Various Temperatures
xi
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LIST OP TABLES (Continued)
Table Page
30 Henry's Law Constant for H2S Versus
Temperature
31 Capital and Operating Cost Summary for Steam
Stripping Concept
xii
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1. CONCLUSIONS
1. The steam stripping concept identified in Phase I of
this program has been tested experimentally in a semi-
batch fluidized bed reactor system. Catalysts tested
were obtained from existing petroleum refineries in
the United States. Altogether, nine catalyst samples
were acquired, two appeared to have high water content
indicating that they had been partially exposed to
air and were no longer representative of an FCC spent
catalyst. Consequently, they were eliminated from
further experimentation.
2. The spent catalyst samples were exposed to steams at
temperatures between 755 and 8ll K (900-1000°F),
pressures from 1.082 x 105 to 3.^27 x 105 Pa (1-35
psig), and steam stripping rates of 1 to 200 kg
H20/100 kg catalyst. Catalyst steam exposure times
ranged from 0 to 3600 seconds.
3. Both air-saturated and oxygen-free distilled waters
were used for superheated steam preparation. The
presence of oxygen appeared to have no significant
effect on sulfur removal efficiency from spent catalyst
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4. The experimental work did not reveal any information
that would contradict the conclusions drawn in Phase I
of this program. The steam stripping rate of 4 kg
H20/100 kg of catalyst assumed to evaluate the steam
stripping concept economics in Phase I report was
further expanded to cover the range between 4 to
100 kg H20/100 kg of catalyst. The results are
included in this report to allow economics comparisons
at those rates.
5. The experimental data demonstrate that reduction of
PCC regenerator SOX emissions is possible via spent
catalyst contacting with steam.
The weight percent sulfur content of coke on spent
catalyst samples acquired from petroleum refineries
ranged from 0.3 to 1.76$. With no'reduction of this
sulfur concentration, these catalysts would produce
regenerator off-gas containing between 2.43 x 10"1*
and 14.49 x lO"4 mole fraction (243-1449 vppm) S02.
Reduction of sulfur on spent FCC catalysts to levels
that are equivalent to 2 x 10~4 mole fraction (200
vppm) S02 in FCC regenerator off-gas was demonstrated
with four catalyst samples. Even though in the case
of three samples the equivalent levels of 2 x 10~4
mole fraction (200 vppm) S02 were not obtained, a
substantial reduction (40-50$) of sulfur on spent
catalyst was observed after their exposure to steam.
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6. The steam stripping rates needed to reduce sulfur
concentrations on spent catalysts to the equivalent
levels of 2 x IQ~k mole fraction (200 vppm) S02 in
FCC regenerator off-gas ranged from 2 to 100 kg
of steam/100 kg of catalyst.
7. The sulfur removed from spent catalyst appeared in
steam in the form of hydrogen sulfide which is readily
handled by refineries.
8. Catalyst steam contacting also removed volatilized
hydrocarbons from the spent catalyst in addition to
sulfur. The gaseous hydrocarbons were identified and
included methane, ethane, and propane, with methane
prevailing.
Heavier hydrocarbons were detected as TOC (total
organic carbon) condensate in the steam. A linear
log-log correlation was found between the steam
condensate TOC concentration and steam stripping rate
for all catalysts used in this study. This seems to
indicate that most of the hydrocarbons are removed
from the catalyst in the very initial period of steam
contacting and thus the steam stripping may also be
used to improve hydrocarbon recoveries in the petroleum
refineries. The effect of hydrocarbon removal on
catalyst regenerator operation and heat balance will
have to be further evaluated for each specific refinery
and steam stripping application.
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9. In cases of some catalysts, carbonyl sulfide has also
been formed in concentrations of about 2 x 10~6 mole
fraction (2 vppm) a thousand times lower than those
of H2S. These concentrations are not expected to
cause any problems in further processing of H^S-rich
streams.
10. Some formation of ammonia was also observed in steam
stripping experiments in concentrations between
3.Ml x 10-u and 5.20 x IQ~k mole fraction (3^1 to
520 vppm). The formation of ammonia apparently occurs
by the same mechanism as that for H2S. NOX emissions
in the regenerator off-gas may be reduced due to this
formation.
11. Exposure to steam at 755-797 K (900-975°P) and
2.39 x 105 Pa (20 psig) for 900 seconds (15 minutes)
caused no change in PCC catalyst activity. This is an
important observation to assure and maintain minimum
interference of steam stripping with present FCC
operations in existing refineries.
12. Composition of steam condensate seems to differ very
little from compositions of waste waters which refin-
eries presently handle. This suggests that no
additional waste water problems, except for increased
waste water volume are created as a result of appli-
cation of steam stripping to sulfur reduction on
spent catalyst.
13. A regression analysis of the experimental results from
all catalysts tested produced a correlation in which
the sulfur removal efficiency is proportional to the
product of steam catalyst contact time and steam
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stripping rate. The proportionality constant seems to
be a function of the catalyst type, form of sulfur on
the catalyst, contracting temperature, and design of
steam contacting equipment, and has to be experi-
mentally determined. The correlation equation has
been modified to describe various designs of steam
contacting reactors including semi-fluidized bed,
continuous fluidized bed, plug-flow reactor, and
counter-current stagewise contacting.
14. The economics of the steam stripping concept were
determined as a function of FCC unit size, steam
stripping rate, and catalyst attrition rate. The
analyses were performed to make their results
comparable with the costs for PCC feed desulfurization
and add-on processes presented in the Phase I final
report. In terms of capital investment costs, (see
Figure 34, page 145), steam stripping is competitive
with FCC feed desulfurization if the steam stripping
rates stay below 70 kg of steam/100 kg of catalyst
with an attrition rate of 0.57 kg of catalyst/m3 of
oil (0.2 Ib of catalyst/barrel) and below 140 kg
steam/100 kg of catalyst in the case of an attrition
rate of 0.29 kg/m3 (0.1 Ib/barrel). Operating costs
for a catalyst attrition rate of 0.57 kg/m3 (0.2 Ib/
barrel) are comparable to those for FCC feedstock
desulfurization in the range of 54-81 kg of steam/100
kg of catalyst for 1.84 x 1Q-2 m3/s (10,000 barrels
per stream day) FCC capacity, 34-49 kg of steam/100
kg of catalyst for 9.20 x 10~2 m3/s (50,000 barrels
per stream day) FCC capacity, and 31-48 kg of steam/100
kg of catalyst for 27.6 x 10~2 m3/s (150,000 barrels
per stream day) FCC capacity. Should the attrition
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rate be 0.29 kg/m3 of oil (0.1 Ib/barrel), the
operating costs of steam stripping are comparable to
those for FCC feedstock desulfurization even if the
ranges of steam stripping rates above are doubled
(see Figure 35, page 1^6).
15. Further reduction of the steam stripping costs may
result from the use of reduced stripping velocities.
The cost estimates presented in this report were
calculated based upon 0.61 m/s (2 ft/s) steam super-
ficial linear velocity through the stripper. Reduction
of this velocity will proportionally reduce steam
consumption. Steam catalyst contact time, however,
will have to be increased. This can be easily done
by a proper design of the stripper. These cost
savings are discussed and demonstrated in Section 6.
Additionally9 our cost analysis did not include the
additional benefits which may result from improved
hydrocarbon recovery and recovery of energy from
stripping steam condensation. The effects of hydro-
carbon removal on FCC regenerator operation, heat
balance, and economics have also not been included.
16. Sour operation of the stripping steam condenser may
produce streams containing as high as 25% volume H2S
with the condensate not exceeding present water
pollution standards for sulfide concentration.
17. The refineries are presently handling sour water
produced from several operations (crude distillation,
FCC fractionator, coker unit, HDS unit, sulfur plant,
etc.). Operations of sour water facilities are
similar to the operation of a sour stripping steam
condenser. Combining the sour water produced from
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steam stripping with sour waters from other operations
and treating these waters in one integrated system may
further increase the applicability of the steam
stripping concept to the refineries.
18. Further substantial improvements in sulfur removal
efficiencies may be expected if the steam stripping
concept is applied commercially. This statement is
based upon observations made in going from a pilot
scale to a commercial scale for similar stripping
operations. Commercially, the same effects on FCC
catalyst were produced by 2 to 5 times lower steam
rates than those measured in a pilot scale. Since
our experimental results were obtained on a much
smaller than pilot scale unit in a semi-batch
manner (not Identical to commercial operations), even
more significant improvements in steam stripping
effectiveness to reduce sulfur levels on a spent
catalyst should be expected.
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2. RECOMMENDATIONS
Based upon the demonstrated ability of the steam stripping
concept to reduce sulfur concentrations on spent FCC
catalyst with no evident effects on existing FCC unit
operation and additional favorable factors which may be
realized upon steam stripping operation scale-up, it is
highly recommended that this concept be carried through
pilot scale development. The pilot scale program should be
performed in one of several FCC pilot plants owned by
petroleum refinery research centers. This would substan-
tially reduce the cost of such effort.
We also recommend that the pilot scale program should
determine the effects of catalyst type, feedstock type, and
feedstock pretreatment upon maximum sulfur removal efficien-
cies via steam stripping. The operating conditions which
should be tested upon each of the above combinations include
temperature range between 728 and 8ll K (850-1000°F) ,
pressure range between 1.082 x ICr and 3.427 x 1CK Pa (1-35
psig), and catalyst residence time in stripper between 30
and 900 seconds. Several stripper designs should be tested,
including continuous fluidized bed, continuous counter-
current stage-wise contactor, and continuous co-current
plug-flow contacting. The complete steam analysis for each
process stream should be performed to yield complete material
and energy balances. The regenerator flue gas should
be analyzed to determine the extent of SO , NO , hydrocarbon,
X. X
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and particulate reduction. The stripper off-gas should
be analyzed to determine the concentration of products as
a function of operating conditions. Stripper condensate
should be characterized to determine the treatability and
compatibility of these wastes with other refinery wastes.
Finally, based on the results of these tests, new economics
should be determined for the steam stripping concept and
compared with those for other sulfur reduction techniques.
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3. INTRODUCTION
Upon completion of Phase I of EPA Contract No. 68-02-1320,
Task 1, Monsanto Research Corporation (MRC) has recommended
that several processes be considered as potential candidates
for refinery catalytic cracker regenerator SO control.
X
The rank ordering of promising processing techniques was
established during Phase I and is presented below.
1. Process Modification - steam stripping of
spent FCC catalyst
2. Dry Sorption (Westvaco Process and Shell Flue
Gas Desulfurization Process)
3. Sodium Sulfite Scrubbing (Wellman-Lord Process)
A detailed discussion and evaluation of each of these
techniques was presented in the final report for Phase I.
It was determined that steam stripping of FCC catalyst may
result in substantial reduction of sulfur compounds deposited
on the spent catalyst. This technique would then prevent the
sulfur compounds from entering the catalyst regenerator and
after their oxidation to sulfur oxides they would be emitted
in the regenerator flue gas to the atmosphere.
10
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Based upon the findings in Phase I, Monsanto Research
Corporation conducted a laboratory development program and
investigated the steam stripping of spent FCC catalyst con-
cept (the processing technique identified in Phase I as the
currently most feasible for reducing SO emissions from PCC
A.
regenerators). Additionally, information was to be acquired
to determine economic and environmental aspects of this
concept and to establish the needs for further investigation
on a pilot scale.
This report summarizes the results and evaluations of
experimental work performed during Phase II. After the
report conclusions, recommendations, and introduction in
Sections 1, 2, and 3, the next section describes the exper-
imental work with experimental apparatus, catalyst samples,
and analytical procedures utilized during the program in
Section M.I, and the experimental results and their inter-
pretations in Section 4.2. Process design considerations
appear in Section 5. The economic analysis is presented in
Section 6.
11
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4. EXPERIMENTAL WORK
During the Phase II experimental program the spent FCC
catalyst samples were exposed to various amounts of steam.
The experiments were carried out in a semi-batch fluidized
catalyst bed reactor. Catalyst samples investigated in the
program were obtained from three petroleum refining companies
in the United States. The effluent gases from the catalyst
testing chamber were analyzed for sulfur and other compounds
removed from the catalyst. The description of the experi-
mental equipment, spent catalyst samples, and the analytical
techniques and facilities utilized on this program are
presented below.
4.1 EXPERIMENTAL EQUIPMENT
4.1.1 Catalyst Test Unit
The test unit, Figure 1, used during this program was
designed and fabricated for the purpose of testing catalysts
and consisted of the following functional sections:
Reactor
Preparation and metering of simulated process or
combustion gases
Effluent gas analysis system
Catalyst handling system
12
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Figure 1. Control panel, reactor, and analysis system
13
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Figure 2 is a schematic flow diagram of the test apparatus.
It consisted of a heated reactor chamber mounted in an
insulated enclosure. The reactor was designed to operate at
temperatures up to 922 K (1200°P) and pressures up to 5.15
x 105 Pa (60 psig).
During the experimental work performed on this program, air,
nitrogen, and water converted into superheated steam were
fed into the reactor charged with PCC catalyst. The gases
leaving the reactor were analyzed for compounds stripped
and burned off from the catalyst.
For several steam stripping experiments, a motionless mixer
was placed inside the reactor chamber (see Figure 3). The
mixer was designed to improve gas-catalyst contacting and
investigate its effects on catalyst stripping efficiency.
Distilled water was used for superheated steam preparation.
For some experiments, the water was deaerated by bubbling
prepurified nitrogen through in order to reduce dissolved
oxygen content below 2 x 10"-3 kg/nr (0.02 ppm). The effects
of deaerated steam on catalyst stripping efficiency were
investigated. The deaeration system with the dissolved
oxygen analyzer is shown in Figure 4.
4.1.2 Spent Catalyst Procurement
The primary objective of the catalyst sample procurement
was to obtain samples that would represent the catalyst
conditions after the catalyst passed'through the FCC reactor
(spent catalyst) but prior to its regeneration. In addition,
samples containing a range of coke concentrations on spent
catalyst and sulfur concentrations were needed to establish
the effects of these variables on effectiveness of steam
stripping.
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Plant Air *-
*2-
H20-
* Instrumentation Air,
-75°F Dew Point
I
11
Control Panel
Off-Gas Analysis
Vent
Catalyst Addition
I
Removal
Vent
Reactor Assembly
Figure 2. Catalyst test unit, schematic flow diagram
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Figure 3. Motionless mixer used in the catalyst reactor
16
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Figure 4. Water deaeration system
17
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The spent FCC catalyst samples used in this program were
obtained from various existing petroleum refineries in the
U.S. Some refineries asked to supply spent catalyst samples
were unable to do so for various reasons; e.g., insufficient
manpower for sample collection, lack of appropriate sampling
ports, or physical Impossibility to collect the sample.
Catalyst samples investigated in this study were supplied
by three American oil companies.
The catalyst samples were delivered in the following amounts,
Catalyst nr Gallons
"A" 1.89 x 10~2 5
"B" 0.76 x 10"2 2
"C" 1.89 x 10~2 5
"D" 0.38 x 10"2 1
"E" 1.89 x 10~2 5
"P" 1.89 x 10~2 5
"G" 20.80 x 10~2 55
"H" 20.80 x 10~2 55
"I" 0.57 x 10"2 1.5
The "A" sample was shipped in an open top can inside a
disposable plastic garbage bag. The "G" sample was shipped
in an open top drum. Upon arrival, the lid on both contain-
ers was not securely fastened.
The samples were characterized according to the procedures
described in Section 4.1.4. The results of these analyses
are presented in Table 1 (Section 4.2) and indicate that the
samples contained about 2.5$ and 5%+ moisture, respectively.
18
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Evidently the samples absorbed moisture from the air, lost
their integrity and were not representative of spent FCC
catalyst. Consequently, they were eliminated from further
experimentation. Samples "B", "C", "D", "E", "F", "H", and
"I" were shipped in airtight containers and were used in
our experiment.
4.1.3 Catalyst Handling
As discussed in the Phase I final report, the cracking
catalysts in use today are primarily of the zeolitic type
and have an affinity for water. In the FCC unit operations
the catalyst is exposed to elevated temperatures, 811-922 K
(1000-1200°F). At these temperatures the catalyst is
essentially dry. Any exposure of the spent catalyst to
water vapor or oxidation atmosphere of ambient air could
result in water absorption and also yield slow oxidation of
hydrocarbons deposited on the spent catalyst. Both of these
phenomena would make the sample non-representative of spent
FCC catalyst. Furthermore, the stripping reactions could
occur under these conditions at a slow rate, removing some
sulfur and hydrocarbons before the steam stripping experi-
ments. Presence of air could also cause some side reactions
with the hydrocarbons and change the nature of the coke
originally present on the spent catalyst. In order to
maintain the spent catalyst sample integrity it was impera-
tive to prevent catalyst exposure to water vapor or air
during the sample collection, shipment, and handling in the
laboratory.
We advised the petroleum refineries who supplied the spent
catalyst that the samples should be collected in an inert
atmosphere (such as nitrogen) and shipped in an airtight
container under a nitrogen blanket.
19
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Determination of whether or not the samples were shipped to
MRC in airtight containers was made at the time the samples
arrived at our location. Each shipping container was
inspected for leaks and the catalyst from any container not
securely sealed was not used during our research program.
In all of our efforts, we tried to use only those catalyst
samples which were representative of the spent catalyst in
an FCC unit. The characterization analyses were an addi-
tional measure of the spent catalyst sample representative-
ness. It should be noted, however, that we do not know
whether or not each sample maintained integrity because of
the thermal cycling which occurred during the sample
collection, shipment, and experimentation. Each sample had
to be collected at an elevated temperature of 755 to 811 K
(900 to 1000°P), cooled down for shipment, and reheated
during our experimentation program.
In order to maintain catalyst integrity during our experi-
mentation a system was devised to handle the spent catalyst
samples without contamination. A flow capability of the
spent FCC catalyst was utilized in this system. The spent
catalyst sample container (a can or drum) was pressurized
with nitrogen to about 4.754 x 107 Pa (0.5 psig). This
pressure was maintained at all times to prevent air or
moisture leakage into the container. A vent pipe made of
9.53 x 10"-^ m (3/8 inch) stainless steel tubing was inserted
into the container for transferring the catalyst into a
preweighed nitrogen purged flask (see Figure 5). A stainless
steel ball valve with Teflon seals was used to control
catalyst flow out of the catalyst sample container. The
weighing flask was fitted with pinch clamps to prevent air
and moisture leakage to the catalyst contained inside the
flask. Typically, the flask was charged with 0.4 kg of
spent catalyst.
20
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Figure 5- Photograph of catalyst storage container
with charging flash attached
21
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The catalyst was then transferred to a nitrogen purged
charging bomb (see Figure 6). This bomb was used to transfer
a spent FCC catalyst sample to a preheated reactor. Actual
charging was done under a nitrogen blanket and 4.8l x 10 Pa
(55 psig) pressure according to the diagram shown in Figure
7.
When the catalyst charging was completed, the charging bomb
was remounted onto the weighing flask to remove all the
residual catalyst retained in the bomb. The difference in
weight of the flask before and after the catalyst charging
operation was a measure of the amount of the catalyst
sample charged into the reactor.
The fluid nature of the spent FCC catalyst was utilized for
the quantitative removal of catalyst from the catalyst unit.
_Q -i
The catalyst was pneumatically transported to a 10 m
(1000 ml) Erlenmeyer collection flask using nitrogen as the
carrier gas. The catalyst was discharged through the
catalyst removal tubing (see Figure 2).
4.1.4 Steam Stripping Experiments and Catalyst
Characterization
The procedures used to perform the spent FCC catalyst steam
stripping experiments and catalyst characterization tests
are presented below.
4.1.4.1 Catalyst Steam Stripping -
A known weight of catalyst sample (0.35-0.40 kg) was placed
into the catalyst reactor. The catalyst was then heated to
the predetermined steam stripping temperature. After the
reactor reached the desired temperature, a known quantity of
superheated steam was passed through the catalyst bed at a
22
-------
Figure 6. Weighing flask mounted on catalyst
charging bomb
23
-------
Toggle Valve
Catalyst
Charging
Bomb
Ball Valves
Return Line
I Purge
Catalyst Reactor-
Chamber
r
^ To Gas
Analysis System
— Porous Stainless Steel Filter
Figure 7- Charging of the catalyst into a
hot reactor
-------
controlled rate which maintaned the catalyst bed in a
fluidized state. The entire off-gas stream was sent through
the HpS analysis system where the sulfur content of the
stripping steam was determined. After the steam stripping
experiments, the H2S analysis system was replaced with S0?/
SCU/H2SOjj analysis system and the same catalyst charge
exposed to air to determine the residual sulfur remaining on
the catalyst according to the procedure described in the
following section.
4.1.4.2 Determination of the Sulfur Content of Coke -
A known weight of catalyst was charged into the catalyst
reactor (0.35-0.40 kg) and oxidized with air at 922 K
(1200°P). The temperature of the catalyst bed was controlled
by adjusting the flow rate of air through the reactor. The
entire gas stream leaving the reactor was analyzed for
oxidation products of sulfur, S02/SO-/H2SOjj. Prom the
results of this analysis the total sulfur of spent catalyst
coke was calculated.
4,1.4.3 Determination of the H^O Content of Spent
FCC Catalyst -
A known weight of each spent catalyst type (0.35-0.40 kg)
was charged to the catalyst reactor in an inert nitrogen
atmosphere as described in Section 4.1.3. A slow, 1.6? x
C O
10"^ nr/s (1 liter/minute) nitrogen purge through the
catalyst at bed was used to remove water vapor desorbed
from the catalyst 783 K (950°P). This test was performed
for a standard period of 300 s (5 minutes) to prevent an
excessive volatilization of other compounds from the
catalyst. The off-gases were collected in a preweighed
drying tube filled with silica gel. The weight difference
of this tube before and after the test was a measure of a
catalyst moisture content.
25
-------
4.1.4.4 Determination of the Coke Content of Spent
FCC Catalyst -
A known weight of each spent catalyst type (0.35 - 0.40 kg)
was charged in the catalyst reactor. The coke deposits were
oxidized with air at 922 K (1200°P). Upon completion of the
coke combustion the catalyst was removed from the test unit.
The difference in weight of catalyst before and after the
test was the measure of the coke and the moisture content
of the catalyst. Subtracting the moisture content determined
according to the procedure described in the previous para-
graph produced the value for the catalyst coke content. This
value would, of course, be representative of all compounds
forming the coke (carbon, hydrogen, sulfur, nitrogen, oxygen,
etc.) .
4.1.5 Other Experiments
Analytical procedures were developed to determine the various
by-products formed during steam stripping of PCC catalyst.
Specifically, we analyzed the off-gases from the steam strip-
ping experiments for carbonyl sulfide, carbon disulfide,
mercaptan sulfur, ammonia, and hydrocarbons including methane
ethylene, propane, propylene, and benzene. The analytical
procedures used to analyze all these compounds will be
described in the following section. The purpose of these
tests was to determine the maximum concentration of products
formed by pyrolyzing the coke on spent catalyst. The tests
identified the types of compounds to be analyzed for in
subsequent experiments and determined the maximum possible
decrease in coke content caused by heating.
26
-------
4.1.6 Analysis
Wherever possible, standard and well established analytical
techniques were utilized. This was the case with SO
HpSO^, H2S, and NHg analyses. EPA Method #8 sampling train
and procedure for "Determination of Sulfuric Acid Mist and
Sulfur Dioxide Emissions from Stationary Sources" was used
for determination of oxidized sulfur compounds (Federal
Register, Vol. 36, No. 247, December 23, 1971, pp. 24893-
24895). An EPA Method #11 sampling train for "Determination
of Hydrogen Sulfide Emissions from Stationary Sources"
(Federal Register, Vol. 39, No. 47, March 8, 1974, pp. 9321-
9323) was used for hydrogen sulfide determination. This
method was modified in order to quantitatively condense
and collect superheated steam. The normal procedure requires
_-3 O
that midget impingers with 0.05 x 10~° itr (50 ml) capacity
_o o
be utilized. For this program, standard 0.5 x 10 mj
(500 ml) Greenburg-Smith impingers were used.
The ammonia analysis of the stripping steam condensate was
made according to the procedures outlined in Standard Methods
for the Examination of Water and Wastewater (New York,
American Public Health Association, 13th Edition, 1971,
Procedure #132).
4.1.6.1 Sulfur Compound and Hydrocarbon Determination -
The analysis for sulfur containing compounds and hydrocarbons
was performed on steam condensate samples as well as gas
samples collected into the Tedlar bag during the steam
stripping experiments. All samples were analyzed in the
F&M Model 720 chromatograph utilizing a dual column and
detection system. A Porapak-Q gas chromatographic column
was used to separate the various components of the sample
analyzed. The chromatographic system employed a temperature
27
-------
programming feature to aid in separation of hydrocarbons
and sulfur compounds. The detector system consldted of a
Tracer 1 sulfur selective flame photometric detector and a
hydrocarbon flame ionization detector. Both detectors were
used simultaneously to detect hydrocarbons and sulfur
compounds.
4.1.6.2 Analysis for the Total Organic Carbon Content of
Stripper Condensate -
In order to determine the extent of hydrocarbon contaminatio
of the stripping steam condensate the entire stripper off-
gas stream was condensed and quantitatively analyzed for
carbon. The samples were analyzed according t'o Procedure
#138 in the Standard Methods for the Examination of Water
and Wastewater using the Beckman TOG analyzer.
4.1.6.3 Volatility of Coke on Spent Catalyst -
In this test, the catalyst samples were placed in an
evacuated quartz tube sealed on one end and heated from
room temperature to 8ll K (1000°F). The other end of the
tube was connected to the mass spectrometer (Du Pont CEC
Model 21-103 C). Gases evolved during the test were
introduced into the instrument for-the analysis.
4.2 EXPERIMENTAL RESULTS
4.2.1 Catalyst Characterization
Each catalyst sample was characterized immediately before it
was submitted to steam stripping experiments in order to
minimize potential changes in catalyst samples due to
aging over long periods of time. Each test was performed
28
-------
at least three times and an average value was then calculated,
except for sample "D" which had only one characterization
analysis performed because of the sample limited amount.
The results for all catalysts obtained from the petroleum
refineries are presented in Table 1. The values in Table 1
are presented as averages with plus/minus percent devia-
tions observed when the multiple characterization tests
were performed. The sulfur content of coke deposited on
"as-received" catal'yst was calculated based on an average
value of catalyst coke content.
The data in Table 1 were used to calculate the equivalent
regenerator S02 emissions from the FCC regenerator. A
math model was developed during Phase I of this program
(see Appendix G in the Phase I final report) to predict
the FCC regenerator S02 emissions after calculating the
carbon, hydrogen, and sulfur contents of the coke and the
C02/C0 ratio of the gases leaving the regenerator. This
model was applied to convert the sulfur content of coke to
the equivalent regenerator S02 concentrations. The
following assumptions were made in using this conversion
method:
Hydrogen content of coke, H = 0.1 (weight fraction)
C02/C0 ratio, R«l (mole ratio)
According to /the math model, the weight fraction of sulfur
in coke before combustion would have to be 0.002^3 in order
to obtain a concentration of 2 x 10~l+ mole fraction (200
vppm) S02 in the regenerator off-gas with R = 1 and H = 0.1.
(This can also be determined from the diagram in Figure 5,
page 27 in the Phase I final report, which was produced
based on the equation in Appendix G). The calculated
29
-------
Table 1. RESULTS OP SPENT FCC CATALYST
SAMPLE CHARACTERIZATION TESTS
Sample
A
B
C
D
E
F
G
H
I
Moisture
(*)
2.488±0.763
1.260±3
0.617+8.4
0.363
0.0
0.0628+13
5+
0.492±4.4
0.680±24
Coke
(*)
0.489±4.3
0.663±4.7
1.202+3.2
1.889
1.016±3.3
0.854+1.55
N.A.
1.304±2.1
1.529±5-52
Sulfur
(50
0.451±2.0
l.?6±3.0
1.013±0.6
0.624
0.520+2.5
0.595±17
N.A.
0.295±2.3
0.748±6.45
S02 concentration
(mole fraction) _
3.71 x 10-*
14.49 x 10"*
8.34 x 10-*
5.14 x 10"*
4.28 x ID'*
4.90 x 10"*
2.43 x 10"*
6.16 x 10-*
30
-------
initial equivalent regenerator S02 concentrations are also
included in Table 1 for each of the catalyst samples used
in this program. An identical procedure was used to pre-
dict the equivalent regenerator S02 concentrations after
the steam stripping experiments.
As shown in Table 1, samples "A", "B", and "G" contained
2.1*9%, 1.26%t and 5+% (by weight) of moisture, respectively.
It was felt that samples "A" and "G" had lost their
integrity due to use of improper containers for their
collection and shipment (refer to Section 4.1.2). They
were thus eliminated from any further experimentation.
Although sample "B" contained 1.26$ (by weight) of moisture,
it was believed that this moisture content was marginal and
that this sample should undergo experimentation.
4.2.2 Catalyst Steam Stripping
The steam stripping experiments were performed according to
the procedure described in Section 4.1.4. Various spent
catalyst samples received from petroleum refineries have
been exposed to steam at different temperatures, pressures,
steam stripping rates, and catalyst residence times in the
stripper reactor. The residence times were expressed by
two variables, steam-catalyst contact time and catalyst-
steam exposure time.
The steam-catalyst contact time is defined as the length of
time that the steam is in contact with the catalyst. It is
calculated by dividing the fluidized catalyst height by the
steam superficial linear velocity and correcting for bed
porosity.
31
-------
Catalyst-steam exposure time is defined as the length of
time which the catalyst is located in a steam environment.
More accurately, it is the catalyst residence time in the
stripper.
The following are the ranges of the variables studied
during this program:
Stripping temperature.
K
op
755-811
900-1000
Stripping pressure,
Steam stripping rate,
Pa
psig
kg H20/100 kg
catalyst
1-35
1-200
Steam-catalyst contact time, s
Catalyst-steam exposure time, s
0.5-2.0
0-3600
After a catalyst sample steam stripping experiment and
after sulfur analysis of the steam streanij the catalyst
sample was oxidized in air and the effluent stream analyzed
for sulfur oxidation compounds. Thus, knowing the original
concentration of sulfur on the catalyst charged to the
reactor, a good sulfur balance check could be made by adding
total sulfur stripped from the catalyst with the steam and
total sulfur remaining on the catalyst and removed after
the oxidation with air.
32
-------
All the experimental and analytical data were recorded on
specially prepared blank forms. An example of this form
Is presented in Appendix A. It includes calculation pro-
cedures used to determine final results.
The final results of all steam stripping experiments have
been summarized in tabular form and are presented in Table 2
through 10. Each table identifies clearly the conditions
at which the experiments were carried out: temperature,
pressure, catalyst source, and other measured variables
including steam stripping rate, steam-catalyst contact time,
superficial steam catalyst contact time, catalyst-steam
exposure time, stripper superficial velocity, and oxygen
content of superheater feed water. The tables also list
the calculated results of percent sulfur removal by steam
stripping and equivalent regenerator S02 concentrations.
For convenience and to better observe the effect of steam
stripping rate on sulfur removal, the data are also pre-
sented in graphic form (Figures 8 through 16). Here, steam
stripping rate has been plotted against the equivalent
regenerator SC>2 concentration. The experiments identified
as not shown on graphs were obtained at steam stripping
rates that are higher than those presented on graphs. For
results of these experiments refer to corresponding experi-
mental data tables.
4.2.3 Hydrocarbon Volatilization During Steam Stripping
A number of tests were performed to determine the amount of
hydrocarbons that would volatilize during steam stripping
experiments and contaminate the steam condensate. The pro-
cedure followed in these tests was described in Section
4.1.6.3. The results are summarized in Tables 11 through 16.
For convenience, the experiment operating conditions are
also included in the data tables.
33
-------
Table 2. RESULTS OP STEAM STRIPPING EXPERIMENTS
B-Series Catalyst, Without Motionless Mixer
U)
_tr
Catalyst bed tenperature,(K)
Steam stripping rate (SSR) ,
(kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts> (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
f eedwater, (kg/m3 )
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, S0)
(mole fraction)
(vppm)
Experiment
B-3
811
1000
5.31
1.08xlOs
1.0
0.123
1.23
180
0.276
0.907
2.0xlO~5
32.4
9.79x10-"
979
B-5
811
1000
32.8
1. 08x10 5
1.0
0.837
0.791
720
0.408
1.31
2.0xlO"5
49-5
7.32x10-"
732
B-7
811
1000
17-6
2. 39x10 5
20
0.145
1.36
300
0.155
0.509
2.0xlO~5
34.1
9.51x10-*
951
B-8
811
1000
39.8
2. 39x10 5
20
0.156
1.48
738
0.197
0.646
2.0xlO~5
50.2
7.21x10-*
721
B-9
811
1000
9.01
3.08xlOs
30
0.212
2.05
180
0.134
0.438
2.0xlO-5
44.4
8.05x10-"
805
B-ll
811
1000
21.9
3. 43xlOs
35
0.192
1.81
360
0.158
0.517
2.0xlC~5
46.9
7. 69x10- u
769
B-13
811
1000
1372
3.43xl05
35
0.0164
-0.155
1920
0.140
0.458
2.0xlO~5
76.3
3.44x10""
344
Remarks: Equivalent regenerator S02 concentration before stripping, 14.49X10"1*mole fraction S02
1449 vppm
-------
IJL>
U1
"evj
O
en
E
Q-
l€
c — '
CD s^
O ~
0 O^
o in
0*0
(/) '4=
. o
^y ^O
^5 v
"(0 *^"
V- CD
c 2
o> ^E
ct:
"c
CD
.1
•3"
<1UUU
1800
1600
1400
1200
1000
800
600
400
200
n
Temperature, 811 k(1000°F)
, B-13 Not Shown on Graph
L • .
\
N.
• \^ B;7
B'3 ^^x! B-ll
' B-9 ^ ^Ji5 Bi8
^^^^^^^^^^^^
•^
-
-
1 1 1 1 1 i I 1 1 1 1 1 1 1 1
Steam Stripping Rate (kgH20/100 kg Catalyst)
Figure 8. Results of steam stripping experiments on B-Series catalyst
-------
u>
o\
Table 3. RESULTS OP STEAM STRIPPING EXPERIMENTS
C-Serles Catalyst, without Motionless Mixer
Experiment
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
(kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/m3)
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 Concentration, SQ,
(mole fraction)
(vppm)
C-6
811
1000
7-31
1. 08x10 5
1.0
0. 0939
0.888
180
0.396
1.30
2.0xlO-5
29.4
' 592
C-7
811
1000
14.1
l.OSxlO5
1.0
0.0971
0.924
360
0.408
1.34
2.0xlO-5
34.2
5.49X10-1*
549
C-8
811
1000
32.6
1.08x10 5
1.0
0.0842
0.798
720
0.469
1.54
2.0xlO-5
35-2
5.40X10-1*
540
C-9
811
1000
8.08
2.39X105
20
0.187
1.78
180
0.211
0.692
2.0x10-5
36.0
534
C-10
811
1000
16.0
2. 39x10 5
20
0.191
1.80
360
0.212
0.695
2.0x10-5
40.9
4.93X1Q-1*
493
C-ll
811
1000
5.98
2.39xl05
20
0.134
1.27
95
0.258
0.845
2.0x10-5
26.8
e.ioxio-1*
610
C-12
811
1000
37.2
2.39xl05
20
0.164
1.54
720
0.199
0.652
2.0x10-5
35-1
5.41X10"1*
541
Remarks: Equivalent regenerator S02 concentration before stripping, 8.34x10 ** mole fraction S02
834 vppm S
-------
Table 3 (continued). RESUIIPS OP STEAM STRIPPING EXPERIMENTS
C-Series Catalyst, without Motionless Mixer
C-14
0-15
0-17
C-18
C-19
C-20
C-21
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
(kg H2OAOO kg catalyst)
Stripper pressure, (Pa)
(psis)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Tg, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/m )
Sulfur removal (% by wt)
811
1000
29-5
3. 43x10 5
35
0.286
2.69
720
0.135
0.441
811
1000
7-59
3.43xl05
35
0.185
1-75
120
0.215
0.704
811
1000
219
2.39xl05
20
0.176
1.48
4050
0.220
0.722
811
1000
109
3.43xl05
35
0.258
2.43
2400
0.154
0.505
811
1000
467
1. 08x10 5
1.0
0.0295
0.279
3600
0.421
1.38
811
1000
738
2.39xl05
20
0.0404
0.389
3600
0.205
0.673
811
1000
631
3.43xl05
35
0.0657
0.631
3600
0.143
0.469
2.0xlO~5 2.0xlO~5 2.0xlO-5 2.0xlO~5
36.0 33-0 44.9 42.1
2.0xlO~5 2.0xlCT5 2.0xlO~5
60.7
56.0
Residual equivalent regenerator
S02 concentration, SQ,
(mole fraction) 5.34xlO-4 5.59X10-1* 4.60x10-^ 4.83X1Q-1*
(vpprn) 534 559 460 483
Remarks: Equivalent regenerator S02 concentration before stripping, 8.34x10
834 vppm
3.27x10-" 3-67X10-1*
327 367
h mole fraction S02
S02 (3j)
43.5
4.71x10- **
471
-------
uo
CO
CVJ
O
CO
0.
^
c
I
sv
I
CN]
O
o>
CT»
CD
•t-*
C
CD
(O
_>
'3
UJ
1000
900
800
_ 700
&
x 600
^ 500
2 40°
^j
o 300
^E
200
m f\f\
100
n
—
Temperature, 811 K (1000°F)
-
X-ll
- •'^'C~6 /"8
- A •"•* v^^-°-12 ^^
- \\ 14 ~~C~l
C-9 C-10
-
Not Shown on Graph
C-17.C-19, C-20.C-21,
"
iiiiiiiiiiiiiii
u 20 40 60 80 100 120 140 16
Steam Stripping Rate (kg H20/100kg Catalyst)
Figure 9. Results of steam stripping experiments on C-Series catalyst
-------
(JO
vo
Table 4. RESULTS OF STEAM STRIPPING EXPERIMENTS
D-Series Catalyst, without Motionless Mixer
Experiment
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
(kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/m )
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, So
(mole fraction)
(vppm)
D-2
811
1000
17-7
2.39xl05
20
0.173
1.63
360
0.201
0.659
:2.0xlO~5
25-0
3.85x10-"
385
D-4
811
1000
41.1
2.39xl05
20
0.149
1.40
720
0.201
0.661
<2.0xlO~5
64.3
1.83X10-1*
183
D-5
811
1000
90.5
2. 39x10 5
20
0.140
1.32
.500
0.201
0.661
<2.0xlO~5
57.8
2.17X10-1*
217
D-7
811
1000
16.6
2. 39x10 5
20
0.184
1-73
3600
0.201
0.661
<2.0xlO~5
67.3
1.68X10-1*
168
Remarks: Equivalent regenerator S02
concentration before stripping, 5.13x10" ** mole fraction S02
513 vppm S02(Si)
-------
1000
^ 800
o
*-•
CO
600
O X
CM CVJ
o o
to
-------
Table 5. RESULTS OF STEAM STRIPPING EXPERIMENTS
E-Series Catalyst, with Motionless Mixer
Experiment
Catalyst bed temperature , (K)
Steam stripping rate (SSR),
(kg H2OAOO kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
Oa content of superheater
feedwater, (kg/m3)
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, SQ,
(mole fraction)
(vppm)
E-6
853
1075
*97.8
l.OSxlO5
1.0
0.095
0.88
2400
0.411
1.35
<9.1xlO~5
62.5
1. 62x10- •*•
162
E-8
811
1000
148
1. 08x10 5
1.0
0.0639
0.584
2400
0.442
1.45
<9.1xlO~5
38.8
2.27x10-'*
227
E-15
811
1000
*6.32
l.OSxlO5
1.0
0.180
1.630
260
0.218
0.716
<2.0xlO-5
15.4
3. 62x10-"*
362
E-16
811
1000
13-3
l.OSxlO5
1.0
0.162
1.48
546
0.256
0.840
<2.0xlO"5
28.7
3.05X1Q-1*
305
E-17
811
1000
5-2
l.OSxlO5
1.0
0.189
1.84
264
0.197
0.645
<2.0xlO-5
15.4
3.61X10-1*
361
Remarks: Equivalent regenerator S02 concentration before stripping, 4.28xlO"1+ mole fraction on S02
* Estimated values 428 vppm S02
-------
-Cr
ro
Table 5 continued. RESULTS OP STEAM STRIPPING EXPERIMENTS
E-Serles Catalyst, with Motionless Mixer
Experiment
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
(kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
f eedwater , (kg/m3 )
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, SQ,
(mole fraction)
(vppm)
E-18
811
1000
2.85
1. 08x10 5
1.0
0.113
1.06
84
0.347
2.0xlO~5
20.8
3.62x10-"*
362
E-21
811
1000
11.6
1. 08x10 5
1.0
0.0989
0.937
300
0.399
2.0xlO~5
20.8
3. 39x10- *
339
E-22
8n
1000
1.34
1. 08x10 5
1.0
0.0708
0.642
24
0.555
2.0xKr5
8.17
3.93X1Q-1*
393
E-23
811
1000
6.16
1.08x10 5
1.0
0.0942
0.879
150
0.418
2.0xlO-5
14.0
3-68x10-^
368
E-24
811
1000
11.8
1. 08x10 5
1.0
0.0871
0.826
270
0.445
2. OxlO~ 5
14.6
3.66x10-"
366
Remarks: Equivalent regenerator S02 concentration before stripping, 4.28X10"1* mole fraction on S02
428 vppm S02
-------
U)
500
CM
O
CO
.S
400
3oo
<-> X
So
O (D
100
E-18
Temperature,811K(1000°F)
0
20
40 60 80 100 120
Steam Stripping Rate (kg F^O/lOO kg Catalyst)
E-8
140
Figure 11. Results of steam stripping experiments on E-Series
catalyst with motionless mixer in stripper
-------
Table 6. RESULTS OF STEAM STRIPPING EXPERIMENTS
E-Series Catalyst, vrLthout Motionless Mixer
Experiment
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
hg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
SteanHcatalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, '(s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/ra3)
Sulfur removal (% by wt)
Residual equivalent regenerator
SO2 concentration, SQ,
(mole fraction)
(vppm)
E-26
811
1000
8.96
1. 08x10 5
1.0
0.122
1.084
270
0.341
1.12
E-28
811
1000
7.9
1. 08x10 5
1.0
0.326
3-075
675
0.121
0.396
E-31
811
1000
9.6H
1. 08x10 5
1.0
0.0713
0.674
l&O
0.482
1.58
E-33
811
1000
185
l.OSxlO5
1.0
0.0741
0.702
3600
0.445
1.46
E-36
8U
1000
9.20
1. 08x10 5
1.0
0.0748
0.707
180
0.451
1.48
E-37
811
1000
33-8
1. 08x10 5
1.0
0.0813
0.769
720
0.445
1.46
<2.0xlO~5 <2.0xlO~s <2.0xlO~s <2.0xlO~5 <2.0xlO~5 <2.0xlO~5
29.8
30.6
9.04
23.1
13-3
25.0
3.00x10"* 2.97x10"" 3.89x10"* 3- 29x10-**
300 297 389 329
3.71X10-1* 3-21X10-1*
371 321
Remarks: Equivalent regenerator S02 concentration before stripping, 4.28xlO~1+ mole fraction on SO,
428 vppm S02 (S^
-------
Jr-
VI
500
CVJ
O
t/>
E
o.
Q.
>• 400
c
o
3 ~ 300
o •"•
O csi
CO O
og 200
CD "G
Ccu
•
21 100
(D
0
•E-31
Temperature,811 K (1000°F)
E-33 Not Shown on Graph
I ' I J 1 L
i i » L
J L
0
10
20 30 40 50 60
Steam Stripping Rate (kg H20/100kg Catalyst)
70
Figure 12. Results of steam stripping experiments on E-Series
catalyst without motionless mixer used in the stripper
-------
Table
Catalyst bed temperature, (K)
<°F)
Steam stripping rate (SSI),
kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts> (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/m3)
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, SQ,
(nole fraction)
(vppm)
7. RESULTS OP STEAM STRIPPING EXPERIMENTS
F-Series Catalyst, with Motionless Mixer
Experiment
F-9
811
1000
5.54
1. 08x10 5
1.0
0.080?
0.782
120
0.433
1.42
<2.0xlO~5
10.6
4.38X10-*
438
F-10
811
1000
10.5
1. 08x10 5
1.0
0.100
0.927
270
0.369
1.21
<2.0xlO~5
16.80
4.o8xlO-4
408
F-ll
811
1000
2.17
1. 08x10 5
1.0
0.0787
0.748
45
0.445
1.46
<2.0xKT5
2.69
4.77x10-^
477
P-12
811
1000
6.99
2.39xl05
20.0
0.145
1.37
120
0.241
0.791
<2.0xlO~5
15.1
4. 16x10- •»
416
F-13
811
1000
7.41
2. 39x10 5
20.0
0.135
1.29
120
0.258
0.847
<2.0xlO~5
10.3
4.40X10-1*
440
Remarks: Equivalent regenerator S02 concentration before stripping, 4.90x10- ** mole fraction on S02
490 vppm S02
-------
Table 7 continued. RESULTS OP STEAM STRIPPING EXPERIMENTS
F-Series Catalyst, with Motionless Mixer
Catalyst bed tenperature, (K)
Steam stripping rate (SSR),
kg H2C/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, '(s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/m3)
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, S,-,,
(mole fraction)
(vppm)
F-14
811
1000
11.1
2. 39x10 5
20.0
0.18?
1.72
240
0.250
0.819
F-15
8n
1000
9.57
2. 39x10 5
20.0
0.166
1.50
180
0.250
0.819
Experiment
F-17
811
1000
9.87
3. 08x10 5
30.0
0.194
1.88
180.
0.203
0.665
P-18
811
1000
9.07
3. 08x10 5
30.0
0.144
1.36
120
0.288
0.946
F-22
811
1000
8.20
3.08xl05
30.0
0.272
2.63
210
0.144
0.474
<2.0xlO~5
30.1
<2.0xlO~5
23.0
14.8
3.42x10-** 3.77X1Q-1* 4.17x10-^
40.8
2.90x10-
<2.0xlO~5
35.5
3.16x10-**
Remarks: Equivalent regenerator S02 concentration before stripping, 4.90x10-^ mole fraction S02
490 vppm S02
-------
Table 7 continued. RESULTS
F-Series
CO
Catalyst bed temperature, (K)
Steam stripping rate (SSR) ,
kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(ffl/B)
(ft/s)
02 content of superheater
feedwater, (kg/m3)
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, S0,
(mole fraction)
(vppm)
OP STEAM STRIPPING EXPERIMENTS
Catalyst, with Motionless Mixer
Experiment
F-25
811
1000
11.9
3. 08x10 5
30.0
0.255
2.41
280
0.163
0.534
<2.0xlO"5
47-7
2.56X10-1*
256
F-27
811
1000
2.3
1. 08x10 5
1.0
0.360
3-29
210
0.116
0.379
<2.DxlO~5
13-8
4.22xlO"lf
422
F-29
811
1000
192
1. 08x10 5
1.0
0.0714
0.677
3600
0.460
1.51
<2.0xlO~5
49-9
2.45x10-^
245
F-30
811
1000
192
2. 39x10 5
20.0
0.160
1.50
3600
0.194
0.638
<2.0xlO~5
49.4
2.48x10-**
248
F-31
811
1000
186
3-43xl05
35.0
0.226
2.14 •
3600
0.140
0.460
<2.0xlO"5
37.3
3.07x10-^
307
Remarks: Equivalent regenerator S02 concentration before stripping, 4.90x10-** mole fraction S02
490 vppm S02 (Si)
-------
4=-
VO
CVJ
o
to
E
CL
9-
c.
f€ 300
"•
O> CM
1°
o
CO to
200
C.
CD
0
0
Pressures
1.0BxloFpa(lpsicj)
2.39xl05Pa(20psig)
Temperature,811 K
F-29. F-30, & F-31 Not Shown on Graph
I i I
I I
I I I
4 6 8 10 12
Steam Stripping Rate ( kg h^O/ 100 kg Catalyst)
14
Figure 13- Results of steam stripping experiments on F-Series catalyst
-------
Table 8. RESULTS OP STEAM STRIPPING EXPERIMENTS
H-Series Catalyst, without Motionless Mixer
Experiment
ui
o
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, '(s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/m3)
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, So,
(mole fraction)
(vppm)
H-5
811
1000
4.01
1. 08x10 5
1.0
0.0571
0.539
60
0.689
2.26
<2.0xlO~5
28.2
1.75x10-^
175
H-6
811
1000
7-68
1. 08x10 5
1.0
0.0884
0.847
180
0.454
1.49
<2.0xlO~5
41.1
1. 43xlQ-'t
143
H-7
811
1000
12.9
1. 08x10 5
1.0
0.107
1.01
360
0.384
1.26
<2,OxlO-5
44.0
1.36xlO-4
136
H-8
811
1000
32.1
1.08xl05
1.0
0.0858
0.810
720
0.475
1.56
<2. OxlO-5
42.9
1. 39x10-"*
139
H-ll
811
1000
11.5
1.08xl05
1.0
0.0800
0.755
2400
0.445
1.46
<2. OxlO-5
70.0
0.73X10-1*
73
H-12
811
1000
4.37
2.39xl05
20.0
0.116
1.10
60
0.314
1.03
<2. OxlO-5
54.5
l.lOxlO-1*
110
Remarks: Equivalent regenerator S02 concentration before stripping, 2.43x10-"* mole fraction S02
243 vppm S02 (SO
-------
Ul
Table 8 continued. RESULTS OF STEAM STRIPPING EXPERIMENTS
H-Series Catalyst, without Motionless Mixer
Experiment
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/m3)
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, SQ,
(mole fraction)
(vppm)
H-13
8n
1000
5.38
2.39xl05
20.0
0.189
1.78
120
0.204
0.669
<2.0xl(T5
43-9
1.36x10-"
136
H-14
811
1000
8.95
2.39xl05
20.0
0.169
1.60
180
0.204
0.670
<2.0xlO~5
47.8
1.27x10-"
127
H-15
811
1000
17.1
2.39xl05
20.0
0.177
1.68
360
0.202
6.662
<2.0xlO-5
40.6
1.44x10-"
144
H-16
8n
1000
31.6
2.39xl05
20.0
0.192
1.82
720
0.185
0.607
<2.0xlO-5
51.1
1.19x10-"
119
IM.7
811
1000
64.0
2. 39x10 5
20.0
0.195
1.84
1500
0.187
0.613
<2.0xlO~5
60.8
0.95x10-"
95
H-18
811
1000
185
2.39xl05
20.0
0.164
1.55
3600
0.201
0.659
<2.0xlO~5
65-5
0.84x10-"
84
Remarks: Equivalent regenerator S02 concentration before stripping, 2.43x10"" mole fraction S02
243 vppm S02
-------
Table 8 continued. RESULTS OP STEAM STRIPPING EXPERIMENTS
H-Serf.es Catalyst, without Jfotlonless Mixer
ru
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg HZ0/100 kg catalyst)
Stripper pressure, (Pa)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts> (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/m3)
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, So,
(mole fraction)
(vppm)
Experiment
H-19
811
1000
4.67
3-43xl05
35.0
0.151
1.42
60
0.225
0.737
<2.0xlO-
32.6
1.64x10-*
164
H-21
811
1000
164
3-43xl05
35.0
0.256
2.42
3600
0.136
0.445
H-22
811
1000
6.89
3.43xl05
35.0
0.204
1.92
120
0.163
0.535
5<2.0xlO"5<2.0xlO-5
73-9
0.63X10-4
63
39.2
1. 48x10-'*
148
H-23
811
1000
9-78
3.43x105
35.0
0.216
2.03
180
0.148
0.4^5
2.0xlO-5
42.2
126x10-"
126
H-24
811
1000
15.6
3. 43x10 5
35.0
0.271
2.54
360
0.134
0.439
2.0xlO-5
56.3
1.06x10-"
106
H-25
811
1000
41.8
3.43xl05
35-0
0.201
1.90
720
0.159
0.521
2.0xlO-5
59.1
0.99x10-"
99
H-26
811
1000
65.6
3.43xl05
35-0
0.268
2.53
1500
0.134
0.439
2.0x10-
55-0
1.80x10-"
109
Remarks: Equivalent regenerator S02 concentration before stripping, 2.43x10"^ mole fraction SO2
vppm SOz (Si)
-------
U)
O -7UU
CO
E
o.
~
o 400
.»— *
to
"c
o>
0 ^
3^ 300
CM >,
O ^,
*c. ° <
•§ c
If 200
§>£
c _.
-------
Table 9. RESULTS OF STEM STRIPPING EXPERIMENTS
H-Series Catalyst, without Motionless Mixer
VJ1
Catalyst bed tenperature, (K)
Steam stripping rate (SSR),
kg H2Q/1QO kg catalyst)
Stripper pressure, (Pa)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, '(s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/m3)
Sulfur removal (% by wt)
Residual equivalent regenerator
SOj concentration, SQ,
(mole fraction)
(vppm)
Experiment
H-27
755
900
11,6
1. 08x10 5
1.0
0.10
0.953
360
0.3^7
.1.14
H-28
755
900
72.0
l.OSxlO5
1.0
0.0851
0.803
1500
0.387
1.27
H-29
755
900
179
1. 08x10 5
1.0
0.0825
0.777
3600
0.415
1.36
H-30
755
900
18.2
1.08K105
1.0
0.8080
0.765
360
0.418
1-37
H-31
755
900
10.5
1. 08x10 5
1.0
.0705
0.663
180
0.482
1.58
H-32
755
900
3.81
1.08x105
1.0
0.0557
0.525
120
0.631
2.07
<2.0xlO~5 <2,OxHT5 <2.0xlO~5 <2.0xlO~5 <2.0xlO-s <2.0x10-5
57.6
57.6
65
54.2
58.2
26.0
1.03x10-^ 1.03x10-^ 0.85x10-^ 1-llxlO-4 1,02x10-'*
103 103 85 111 102 180
Remarks: Equivalent regenerator S02 concentration before stripping, 2.43x10~^ mole fraction
-------
Ul
VJ1
500
CM
O
E
Q.
Q.
>
400
O>
&
O i
CM
300
CM
.1
CD «*-
SP,CO
11
03
to
200
100
0
H-27 H-30
H-28
Temperature 755 K (900°F)
H-29
0
20
40 60 80 100 120
Steam Stripping Rate (kgH20/100 kg Catalyst)
140
160
180
Figure 15. Results of steam stripping experiments of H-Series catalyst
-------
Table 10. RESULTS OP STEAM STRIPPING EXPERIMENTS
I-Series Catalyst, without Motionless Mixer
VJl
cr\
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg H20/100 kg catalyst)
1-5
811
1000
4.19
Stripper pressure, (Pa) 1.08x10 5
(psig) 1.0
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
Cm/a)
(ft/s)
02 content of superheater
feedwater, (kg/m3) <2
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, S0,
(mole fraction) 4
(vppm)
0.0546
0.517
60
0.524
1.72
.OxlO-5
22.58
.77xlO~4
477
1-8
811
1000
29.1
1. 08x10 5
1.0
0,0631
0.596
480
0.475
1.56
<2..0xlO-5
37-27
3.86X10-1*
386
Experiment
1-9
811
1000
53.4
1. 08x10 5
1.0
0.690
0.648
960
0.433
1.42
<2.0xlO-5
41.48
*36l
1-10
811
1000
194
1. 08x10 5
1.0
0.0710
0.669
3600
0.436
1.43
<2. 0x10-5
50.64
3. 04xlO~ 4
304
1-11
811
1000
13.4
1. 08x10 5
1.0
0.0668
0.648
240
0.451
1.48
<2.0xlO-5
31-52
4.22X1Q-1*
422
Remarks: Rjuivalent regenerator SO2 concentration before stripping, 6.16x10 "* mole fraction on
vppm. SO^ CS^
-------
Table 10 continued. RESULTS
I-Series
VJ1
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg H2OAOO kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
02 content of superheater
feedwater, (kg/m3)
Sulfur removal (% by wt)
Residual equivalent regenerator
S02 concentration, SQ,
(mole fraction)
(vppm)
OF STEAM STRIPPING EXPERIMENTS
Catalyst, without Motionless Mixer
Experiment
1-12
811
1000
3-99
2.39xl05
20
0.254
2. HO
120
0.120
0.395
<2.0xlO~5
25-31
4. 60x10-**
460
1-13
811
1000
31.6
2. 39x10 5
20
0.128
1.21
480
0.232
0.761
<2.0xlO~5
35.02
4.00X10-1*
400
1-14
811
1000
97-1
2. 39x10 5
20
0.156
1.48
1800
0.191
0.628
<2.0xlO~5
50.67
3.04x10-^
304
1-15
811
1000
182
2.39xl05
20
0.167
1.58
3600
0.215
0.704
<2.0xlO~5
55.45
2.74x10-'*
274
1-16
811
1000
49.1
2. 39x10 5
20
0.155
1.46
900
0.201
0.659
<2.0xlO~5
51.31
S.OOxlO-1*
300
Remarks: Equivalent regenerator S02 concentration before stripping, 6.16X10"1* mole fraction on S02
616 vppm S02 (Si)
-------
00
o*
Q_
f- 800
c
^^
•^" /fat
i;
8^ 600
O _)
~^500
^S
£ c 400
1o .£
|| 300
iPji>
QC o /QO
^•~' ^~
c —
o>
| 1UO
cy
-U
»
m
Temperature, 811 K(1000°F)
•1.08xlt?Pa( Ipsig)
Pressure
"2. 39 x 105 Pa (20psig)
:"5
l-l^*» '"^ i-i?
" ^^**^^»A^ 1 -9
| jj *
1-16 1-14" — ' —ft
i -f5
1-10 Not Shown on Graph
iiitiiiiitiiiiiiii
0 20 40 60 80 100 120 140 160 180
Steam Stripping Rate ( kg H20 / 100 kg Catalyst)
. "Results of steam stripping, experiments on X-Series catalyst
-------
Table 11. RESULTS OF EXPERIMENTS PERFORMED TO DETERMINE THE
EFFECT OF STEAM STRIPPING ON COKE VOLATILITY
\o
B-Series Catalyst
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
Carbon content of stripper
condensate (kg/m3)
Carbon removal
(% by weight of coke)
Experiment
B-3
811
1000
5.31
1.08xl05
1.0
0.123
1.23
180
0.276
0.90?
1.016
0.817
B-5
811
1000
9.14
2. 39x10 5
20.0
0.166
1.57
180
0.197
0.646
0.714
0.989
B-7
811
1000
17.6
2.39xl05
20.0
0.145
1.36
300
0.155
0.509
0.518
1.45
B-8
811
1000
39-8
2.39xl05
20.0
0.156
1.48
738
0.197
0.646
0.168
1.01
-------
Table 11 continued.
RESULTS OF EXPERIMENTS PERFORMED TO DETERMINE THE
EFFECT OF STEAM STRIPPING ON COKE VOLATILITY
B-Series Catalyst
Experiment
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg fl20/100 Kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
Carbon content of stripper
condensate (kg/m3)
Carbon removal
(% by weight of coke)
B-9
811
1000
9.01
3.08xl05
30.0
0.212
2.05
180
0.13^1
0.438
0.739
1.01.
B-ll
811
1000
21.9
3.43xl05
35.0
0.192
1.81
360
0.158
0.517
0.215
0.713
B-12
811
1000
99.6
3.43xl05
35.0
0.212
2.00
1800
0.1*11
0.460
0.0721
1.09
B-13
811
1000.
1372
3.43xl05
35-0
0.0164
0.155
1920
0.140
0.458
0.0340
7.07
-------
Table 12. RESULTS OP EXPERIMENTS PERFORMED TO DETERMINE THE
EFFECT OF STEAM STRIPPING ON COKE VOLATILITY
C-Series Catalyst
Experiment
Catalyst bed temperature, (K)
(°P)
Steam stripping rate (SSR),
kg flzOAOO kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(*/s)
(ft/a)
Carbon content of stripper
condensate (kg/in3)
Carbon removal
(% by weight of coke)
C-6
811
1000
7.31
1. 08x10 5
1.0
0.0939
0.888
) 180
0.396
1.30
1.498
0.911
C-7
811
1000
14.1
l.OSxlO5
1.0
0.0971
0.924
360
0.408
1.34
0.325
0.793
C-10
811
1000
16.0
2.39xl05
20.0
0.191
1.80
360
0.212
0.695
0.427
1.18
C-ll
811
1000
5-98
2.39xl05
20.0
0.134
1.27
95
0.258
0.845
0.863
0.896
C-12
811
1000
37.2
2. 39x10 5
20.0
0.164
1.54
720
0.199
0.652
0.217
1.40
C-13
811
1000
109
3-43xl05
35.0
0.258
2.43
2400
0.156
0.475
0.008
1.52
C-14
811
1000
29.5
3-43xl05
35-0
0.286
2.69
720
0.135
0.444
0.797
4.09
-------
Table 12 continued. RESULTS OP EXPERIMENTS PERFORMED TO DETERMINE THE
EFFECT OF STEAM STRIPPING ON COKE VOLATILITY
C-Series Catalyst
Experiment
ro
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg .H20/100 kg catalyst)
Stripper pressure , (Pa)
(pslg)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity ,
(m/s)
(ft/s)
Carbon content of stripper
condensate (kg/m3)
Carbon removal
(% by weight of coke)
C-15
811
1000
7.59
3.43xl05
35.0
0.185
1.75
120
0.215
0.704
1.521
2.01
C-16*
811
1000
115
1.08xl05
1.0
0.0799
0.756
2400
0.445
1.46
1.355
27.0
C-18*
811
1000
109
3. 43x10 5
35.0
0.258
2.43
2400
0.154
0.505
1.899
36.0
C-20
811
1000
738
2. 39x10 5
20.0
0.0404
0.389
3600
0.205
0.673
0.0214
2.75
C-21
811
1000
631
3-43xl05
35.0
0.0657
0.631
3600
0.143
0.469
0.0406
4.45
C-24
811
1000
177
1.08xl05
1.0
0.0730
0.733
3600
0.460
1.51
0.0464
1.43
* Samples contaminated with acetone from bottle
-------
Table 13. RESULTS OP EXPERIMENTS PERFORMED TO DETERMINE THE
EFFECT OF STEAM STRIPPING ON COKE VOLATILITY
E-Series Catalyst
U)
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg JI2Q/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
Carbon content of stripper
condensate (kg/m3)
Carbon removal
(% by weight of coke)
E-31
811
1000
9.64
1.08x105
1.0
0.713
0.674
180
0.482
1.58
E-33
811
1000
185
1.08x105
1.0
0.0741
0.702
3600
0.445
1.46
Experiment
E-35
811
1000
70.5
1.08x105
1.0
0.0814
0.768
1500
0.485
1.49
E-36
811
1000
9.2
1.08x105
1.0
0.748
0.707
180
0.482
1.48
E-37
811
1000
33.8
1.08x105
1.0
0.0813
0.769
720
0.445
1.46
0.246
0.232
0.0532
0.966
0.0554
0.383
3.026
2.73
0.531
1.76
-------
Table 14. RESULTS OP EXPERIMENTS PERFORMED TO DETERMINE THE
EFFECT OF STEAM STRIPPING ON COKE VOLATILITY
F-Series Catalyst
Catalyst bed temperature, (K)
TO
Steam stripping rate (SSR),
kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
Carbon content of stripper
condensate (kg/m3)
Carbon removal
(% by weight of coke)
Experiment
F-29
811
1000
192
1. 08x10 5
1.0
0.0714
0.677
3600
0.460
1.51
F-30
811
1000
192
2.39xl05
20.0
0.160
1.50
3600
0.194
0.638
F-31
811
1000
186
3.43xl05
35-0
0.226
2.14
3600
0.140
0.460
0.126
2.83
0.0879
1.97
0.110
2.40
-------
Table 15. RESULTS OP EXPERIMENTS PERFORMED TO DETERMINE THE
EFFECT OP STEAM STRIPPING ON COKE VOLATILITY
H-Series Catalyst
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg ,H20/100 kg catalyst)
Stripper pressure, (Pa)
(pslg)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
Carbon content of stripper
condensate (kg/m3)
Carbon removal
(% by weight of coke)
Experiment
H-5
811
1000
4.01
1. 08x10 5
1.0
0.0571
0.539
60
0.689
2.26
H-6
811
1000
7.68
1. 08x10 5
1.0
0.0884
0.847
180
0.454
1.49
H-7
811
1000
12.9
1. 08x10 5
1.0
0.107
1.0
360
0.384
1.26
H-8
811
1000
32.1
1. 08x10 5
1.0
0.0858
0.810
720
0.475
1.56
H-9
811
1000
69.0
1.08xl05
1.0
0.0831
0.784
1500
0.463
1.52
H-ll
811
1000
115
1. 08x10 5
1.0
0.0800
0.755
2400
0.445
1.46
H-12
811
1000
4.37
2. 39x10 5
20.0
0.116
1.10
60
0.314
1.03
1.419
0.437
1.541
0.907
1.957
1.94
0.289
0.712
0.165
0.871
0.0300
.0.261
1.987
0.255
-------
Table 15 continued. RESULTS OF EXPERIMENT PERFORMED TO DETERMINE THE
EFFECT OP STEAM STRIPPING ON. COKE VOLATILITY
H-Series Catalyst
ON
Catalyst bed tenperature, (K)
Steam stripping rate (SSR),
kg .H20/100 kg catalyst)
Stripper pressure, (Pa)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/a}
Carbon content of stripper
condensate (kg/hi3)
Carbon removal
(% by weight of coke)
H-13
811
1000
5-38
2. 39x10 5
20.0
0.189
1.78
120
0.204
0.669
H-14
811
1000
8.95
2. 39x10 5
20.0
0.169
1.60
180
0.204
0.670
H-15
811
1000
17-1
2. 39x10 5
20.0
0.177
1.68
360
0.202
0.662
Experiment
H-16
811
1000
31-6
2. 39x10 5
20.0
0.192
1.82
720
0.185
0.607
H-17
811
1000
65.0
2. 39x10 5
20.0
0.195
1.84
1500
0.187
0.613
H-18
811
1000
185
2. 39x10 5
20.0
0.164
1-55.
3600
0.201
0.659
H-19
811
1000
4.67
3. 43x10 5
35-0
0.151
1.42
60
0.225
0.737
0.379
0.157
1.176
0.808
0.332
0.435
0.0581
0.141
0.173
0.864
0.0097
0.803
1.742
0.624
-------
Table 15 continued. RESULTS OF EXPERIMENTS PERFORMED TO DETERMINE THE
EFFECT OF STEAM STRIPPING ON COKE VOLATILITY
H-Series Catalyst
Experiment
Catalyst bed temperature, (K)
Steam stripping rate (SSR) ,
kg ,H20/100 kg catalyst)
Stripper pressure, (Pa)
(psig)
Steamr-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ft/s)
Carbon content of stripper
condensate (kg/m3)
Carbon removal
(% by weight of cote)
H-21
811
1000
164
3.43x105
35-0
0.256
2.42
3600
0.136
0.445
0.0545
0.686
H-22
811
1000
6.89
3-43x105
35-0
0.204
1.92
120
0.163
0.535
0.222
0.117
H-23
811
1000
9.78
3.43x105
35-0
0.216
2.03
180
0.148
0.485
0.177
0.134
H-24
811
1000
15.6
3.43x105
35.0
0.271
2.54
360
0.134
0.439
0.446
0.535
H-25
811
1000
41.8
3.43x105
35-0
0.201
1.90
720
0.159
0.521
0.0548
0.176
H-26
811
1000
65.6
3-43x105
35.0
0.268
2.53
1500
0.134
0.439
0.208
1.05
-------
Table 16. RESULTS OF EXPERIMENTS PERFORMED TO DETERMINE THE
EFFECT OF STEAM STRIPPING ON COKE VOLATILTTY
H-Series Catalyst
Experiment
CO
Catalyst bed temperature, (K)
Steam stripping rate (SSR),
kg H20/100 kg catalyst)
Stripper pressure, (Pa)
(pslg)
Steam-catalyst contact time, (s)
Superficial steam-catalyst
contact time, Ts, (s)
Catalyst-steam exposure time, (s)
Stripper superficial velocity,
(m/s)
(ffc/8)
Carbon content of stripper
condensate (kg/m3)
Carbon removal
(% by weight of coke)
H-27
755
900
14.6
1. 08x10 5
1.0
0.101
0.953
360
0.347
1.14
0.627
0.701
H-28
755
900
72.0
l.OSxlO5
1.0
0.0851
0.803
1500
0.387
1.27
0.0536
0.285
H-29
755
900
179
l.OSxlO5
1.0
0.0825
0.777
3600
0.415
1.36
0.0651
0.900
H-30
755
900
18.2
1. 08x10 5
1.0
0.0808
0.765
360
0.418
1.37
0.655
0.911
H-31
755
900
10.5
l.OSxlO5
1.0
0.0705
0.663
180
0.482
1.58
2.292
1.78
H-32
755
900
8.81
l.OSxlO5
1.0
0.0557
0.525
120
0.631
2.07
0.237
0.165
-------
Regardless of the catalyst type, the carbon content of stripper
condensate was plotted against the steam stripping on a
logarithmic paper, Figure 17, and a rather uniform relation-
ship between the two variables was observed. Regression
analysis of the data produced the following equations.
log TOG = -0.998 log (SSR) + 3.909 (1)
log (SSR) = -1.002 log TOC + 3-918 (2)
4.2.4 • Volatility of Coke on Spent Catalyst
In order to obtain information on types of compounds that
may volatilize from the catalyst at elevated temperatures
we have performed the catalyst coke volatility test. The
procedure followed to perform this test was described in
Section 4.1.6.3. As indicated by the procedure, the test
was carried out under vacuum and therefore is not an exact
simulation of conditions that would exist in the catalyst
stripper where some pressure is present and the coke is
exposed to continuous flow of steam. Also, during steam
stripping the steam may react with the hydrocarbons and
produce other hydrocarbon product mix. However, we feel
the test together with other experiments made in this
study may provide some Information as to the type of com-
pounds that can volatilize during the catalyst steam
stripping.
Two catalyst samples, "H" - and "E"-series, were used for
this test. The results are summarized in Tables 17 and 18.
Other catalyst samples were not tested because the purpose
of the test was to demonstrate that the presence of
69
-------
10
CT»
J£
o
o
0.1
log (SSR) = -1.002 log TOO 3.918
log TOC =-0.998 log (SSR)+3.909
O.OL
i i il
10
SSR(kg H20/100 kg Catalyst)
100
1000
Figure 17,
Effect of steam stripping rate upon
total organic carbon content of
stripper condensate
70
-------
a
b
Table 17. RESULTS OP COKE VOLATILITY
EXPERIMENTS ON H-SERIES CATALYST
Determination of Coke Volatility
Not Due to Steam Stripping
Temperature Range
293 - 423 K
68 - 302°P
423 - 573 K
302 - 572°F
573 - 673 K
572 - 752°P
673 - 8ll K
752 - 1000°P
Components3" (wt. % of catalyst)
4 36 6H6
0.001
0.002
0.001
0.001
0.01
0.02
0.01
0.01
ND
0.0092
0.0003
0.0008
totals13 (C + H)
(C)
0.005
0.0038
0.05
0.043
0.0103
0.0088
These were the only compounds detected in the gases evolved.
C + H = weight percent of catalyst volatilized as hydrocarbon,
C = weight percent of catalyst volatilized as carbon
(obtained by calculating carbon content of hydro-
carbon volatilized).
71
-------
Table 18. RESULTS OF COKE VOLATILITY
EXPERIMENTS ON E-SERIES CATALYST
Determination of Coke Volatility
Not Due to Steam Stripping
Componentsa (wt. % of catalyst)
Temperature Range CH4 C2H4 °3H6 C4H8
293 - 373 K
68 - 212°P 0.0004 0.006 0.005 ND
373 - 573 K
212 - 572°P 0.001 0.01 0.01 0.009
573 - 811 K
572 - 1000°P 0.001 0.003 0.004 ND
totals13 (C + H) 0.0024 0.019 0.019 0.009
(C) 0.0018 0.016 0.016 0.0077
?These were the only compounds detected in the gases evolved.
C + H = weight percent of catalyst volatilized as hydrocarbon.
C = weight percent of catalyst volatilized as carbon
(obtained by calculating carbon content of hydrocarbon
volatilized).
72
-------
hydrocarbons in the stripper off-gases may not be entirely
due to the action of steam upon the coke, but also hydro-
carbon volatilization at elevated temperatures of up to
752K (1000°F).
4.2.5 By-Product Formation During Steam Stripping
In order to determine the extent of by-product formation
during steam stripping, several tests were made. These
tests were performed to identify the types and concen-
trations of compounds formed during catalyst steam stripping.
The tests were performed following the same procedure as
that used for steam stripping experiments. The effluent
gas collection system consisted of an ice water sampling
train and a Tedlar bag. Thus, no constitutents could
leave the system. After the condensation of stripping
steam the fluidized bed reactor and the ice water sampling
train were flushed with nitrogen an-d all gases collected in
the Tedlar bag. Both the condensate and the contents of
the Tedlar bag were analyzed gas chromatographically. In
addition, the condensate was analyzed for ammonia. The
compounds formed include CHU, C2Hi», C3H6, COS, CS2, H2S,
S02, and NH3.
Table 19 presents the actual results of the tests. The
operation of the ice water sampling train requires some
water present in the train to absorb sulfur dioxide or
hydrogen sulfide. Consequently, the results obtained by
a direct analysis of the ice water train contents involve
some dilution of compounds originally present in the
stripping steam. Only compounds detected appear in Table 19.
73
-------
Table 19. SUMMARY OF BY-PRODUCT FORMATION EXPERIMENTS
Sulfur and Hydrocarbon Compounds
Sample
No.
D-4
C-25
C-26
H-35
Medium
Gas
Liquid
Gas
Liquid
Gas
Liquid
Gas
Liquid
Compounds
Present
H2S
N2
COS
CH,
H2S
N2
None
H2S
N2
COS
COS
H2S
N2
COS
Concentration
of Compounds,
(gas - mole fraction)
(liq - kg/m3)
630 x 10~6
13 x 10~6
443 x 1CT6
Balance
1.5 x 10~3
1400 x 10-6
99 x 1CT6
492 x 10~6
Balance
ND
2600 x 10~6
273 x 10~$
42 x 10~6
713 x 10~6
Balance
0.54 x 10~3
137 x 10~6
1.47 x 10~6
130 x 10 6
Balance
0.188 x 10~3
N2-free Gas
Compositions
(dry basis)
(vol. g)
58.0
1.2
40.8
70.3
4.9
24.8
71.6
7.5
1.2
19-7
51.2
0.3
48.5
Fraction of Feed
Sulfur & Carbon
in Sample
(wt %)
0.214 carbon
0.0132 carbon
64.3 sulfur
1.21 sulfur
0.445 carbon
0.063 carbon
41.2 sulfur
0.968
0.203
67.4
0.258
carbon
carbon
sulfur
sulfur
0.0468 carbon
0.453 sulfur
40.1 sulfur
0.25 sulfur
ND = not detected
-------
Since the amount of. stripping steam used in these experiments
was known, we have recalculated the results in Table 19 to
obtain apparent average concentrations of the detected
compounds in stripping steam. The results of these calcu-
lations appear in Table 20, 21, and 22.
4.2.6 Effect of Steam Stripping on Catalyst Activity
A serious concern was raised regarding the possible
deactivation of FCC catalyst during its extended exposure
to steam. Several tests were performed to determine the
effect of steam stripping on catalyst activity. The results
of these tests appear in Table 23.
Five catalyst samples identified by six-digit numbers and
included in Table 23 were analyzed by Davison Chemical
Division, W. R. Grace Company, Baltimore, Md. The analyses
included chemical analysis, physical analysis, and activity
determinations. More specific identification of the five
s ample s fo1lows.
Sample 165445 was an "as-received" catalyst. This sample
was neither steam stripped nor regenerated in our catalyst
test unit. It was used to establish the basis for com-
parison of samples received from the refinery and those
later exposed to steam stripping experiments.
Sample 165446 was not steam stripped but was regenerated
at 1.08 x 105 Pa (1 psig) and 866 K (1100°F) for 3600 s
(60 minutes) in our catalyst test unit. This sample was
analyzed to determine whether or not our catalyst regeneration
technique caused any catalyst deactivation.
75
-------
Table 20. CALCULATED STRIPPER OPP-GAS ANALYSIS
C-Series Catalyst (c-26)
Component
H2S
COS
CS2
NH3
H20
Concentration,
mole fraction
2520 x 10~6
1.78 x 10~6
N D
7200 x 10~6
50.5 x 10~6
N D
445 x 10~6
balance
ND = not detected
76
-------
Table 21. CALCULATED STRIPPER OFF-GAS ANALYSIS
H-Series Catalyst (H-35)
Component
H2S
COS
CS2
C3H6
NH3
H20
Concentration,
mole fraction
116 x 10~6
2.1 x 10~6
N D
35 x 10~6
N D
N D
341 x 10~6
balance
ND = not detected
77
-------
Table 22. SUMMARY OP RESULTS OP BY-PRODUCTS
FORMATION EXPERIMENTS
Ammonia Formation
NH3 Content of Condenser Off-Ga£
Experiment (mole fraction) _.
H - 34 3111 x 1Q-6
H - 35 346 x 10"6
C - 27 445 x 10~6
C - 28 520 x 10~6
-------
Table 23. CATALYST ACTIVITY TESTS
o
Chemical Analyses
A1203, wt %
Na20
SO,,
Pe
Re203
C
Ni
V
Cu
TV
Physical
SA ,
PV ,
ABD,
Part.
0-20
0-40
0-80
APS,
, wt %
, wt %
, wt %
, wt %
, wt %
, ppm
» ppm
, ppm
, wt %
Analyses5
mz/gm
cc/gm
gm/cc
Size Dlst.a
v
y
P
P
Davison Microactivity
CPP
GPP
where TV
SA
PV
ABD
M
CPP
GPP
APS
1654*5.
31.7
0.
0.
0.
2.
0.
654
123
11
2.
125
0.
0.
0
10
94
68
69.
0.
2.
91
06?
34
20
93
24
36
82
5
73
80
165446
32.2
0.
0.
0.
2.
0.
685
135
11
0.
121
0.
0.
0
H
64
72
69.
0.
4.
90
165448
21
0
050 o
35
17
01
85
35
78
3
87
65
« total volatiles wt %
** surface area, m2/gm
* pore volume, cc/gm
- apparent bulk density, gm/cc
* microns
* carbon production factor
= gas production factor
" average particle size, microns
0
2
0
660
145
11
2
130
0
0
0
1
64
73
67
0
3
a
.6
.92
.057
.33
.19
.02
.79
.36
• 79
.6
.87
.21
165449
31.1
0.
0.
0.
2.
0.
650
140
11
1.
122
0.
0.
0
4
82
67
71.
0.
4.
ppm =
m2/gm =
cc/gm =
gm/cc =
v =
90
051
33
14
02
63
36
82
4
76
52
165450
31.4
0
0
0
2
0
678
151
11
0
124
0
0
0
4
65
72
68
0
2
.87
.050
.34
.17
.01
.85
.3
.80
.8
.70
.59
1Q-11 % wt
103m2/kg
10~3m3/kg
103kg/m3
10" 6m
79
-------
Samples 165448, 165449 and 165450 were both steam stripped
and regenerated in our catalyst test unit. While the steam
stripping conditions varied for each sample, and were 797 K
(975°F), 783 K (950°P), and 761 K (910°P), respectively,
2.39 x 105 Pa at (20 psig) for 900 s (15 minutes), the
regeneration of all three samples was done at the same
conditions as those for sample 165446.
According to Mr. Warren Letzsch from Davison Chemical Div.»
W. R. Grace Co., "The samples appear to be representative
commercial products that had a higher than average nickel
level. This is reflected in the gas producing factor (GPF)
of the microactivity test which normally runs under 2.0."
Mr. Letzsch also concludes that "a comparison of the first
two samples shows that our regeneration technique removes
virtually all of the carbon (coke) without doing any damage
to the catalyst. The chemical and physical analyses are
virtual duplicates with the exceptions, of course, of the
carbon and TV [total volatiles] analyses. Stripping at the
L.S
relatively mild conditions shown for the last three catalySv
did not cause any significant deactivation. This is not to°
surprising when one considers that many commercial stripped5
run with almost 100$ steam at 783-811 K (950-1000°P) for up
to several minutes. As the data indicate, no real changes
occurred in either the chemical or physical analyses. TheTe
is no evidence of pore sintering, and sieve stability
appears to be excellent. Our microactivity test runs plus
or minus two numbers, so we would conclude that all of the
samples (even the 67.6 volume percent conversion) are
experimental error of the base catalyst.*
*Results of Davison microactivity test differing by ± 2.0
are within experimental error of the test and do not indi-
cate change In catalyst activity.
80
-------
Mr. Letzsch, who is a recognized authority on fluid catalytic
cracking catalysts and their production and application also
suggested that these initial tests appear to be very
encouraging and that further work is fully justified. He also
recommended that before a final design is set, steam tests
lasting several days should be undertaken.
4.3 DISCUSSION OF RESULTS
Over 160 FCC spent catalyst steam stripping experiments
were performed on a total of nine spent catalyst samples.
These samples were obtained from five U.S. refineries situated
in various geographical locations. As far as we could
determine at the time of sample collection, the refineries
operated on crude oils with sulfur contents ranging from
Q.25% to 1.0$ by weight. The sulfur content of the FCC
feedstocks for the catalyst samples ranged from 0.5$ to
1.9%. Some of the FCC feedstocks did receive pretreatment
(hydro-desulfurization) prior to processing in the catcracker.
The steam stripping conditions to which these catalyst
samples were exposed were outlined in Section 4.2.2.
The technical feasibility of steam stripping of spent FCC
catalyst -was demonstrated and it was shown that equivalent
regenerator SOX emissions of 2.0 x I0~k mole fraction (200
vppm) S02 or lower are feasible. Most of the catalysts
exposed to steam stripping rates of 1 to 100 kg of steam
per 100 kg of catalyst showed sulfur reduction that would
result in sulfur oxide emissions of 2 x 10"1* mole fraction
(200 vppm) in the regenerator off-gas. For three catalysts,
the experiments with steam stripping revealed sulfur
removal lower than that resulting in emissions of 2.0 x 10"14
mole fraction (200 vppm). However, even in these three
81
-------
cases (catalysts "B"-, "C"-, and "I"- series) substantial
reduction in sulfur concentration on coke (^0-50$) was
observed with steam stripping rates of about 50 kg per
100 kg of catalyst or lower.
Actual steam stripping rate requirements are a function of
many variables including catalyst type, type of feedstock,
temperature, pressure, catalyst contact time, and catalyst
residence time. Mathematical correlations were developed
for some of the variables and are presented in Section 4.4.
Because a large number of variables affect the steam
stripping process (many of which cannot be -quantitatively
described), only an experiment can determine whether or not
the contact with steam will result in a required sulfur
reduction for a specific catalyst.
Essentially no uniform trend between steam stripping and
sulfur reduction was found in the case of "C"- series
catalyst. The reasons for this phenomenon are unknown but
Dr. E. G. Wollaston of American Oil Company suggests that
this might be attributed to the metal sulfides content of
the coke deposits.
Examination of experimental results reveals that the steam
stripping rates to obtain an equivalent regenerator S02
concentration of 2.0 x 10~6 mole fraction (200 vppm) can
vary greatly for different types of catalysts. This fact
is not totally unexpected considering all possible catalys*
types, feedstock materials, feedstock pretreatments, and
process operating conditions in commerical FCC units.
However, the data presented by Conn and Brackin discussed
in the Phase I final report (page 52) indicate that con-
siderable steam savings can be obtained by increasing FCC
82
-------
capacity from pilot plant scale to commercial application.
Conn and Brackin demonstrated a substantial steam reduction
of 50 to 80%.
Our experiments were performed in a semi-batch fluidized bed
reactor which in no case is representative of commercial or
pilot scale FCC units. In addition to the reactor's small
size, its operation was not continuous which is contrary to
any FCC commercial unit. After placement in the reactor
the catalyst sample was exposed to steam for various lengths
of time. According to some theories of laminar and turbulent
conditions existing in fluidized beds* the conditions in our
reactor were laminar. To change these conditions we inser-
ted a motionless mixer in our reactor but found no signifi-
cant improvement in sulfur removal efficiencies. How the
motionless mixer changed the laminar conditions in the
experimental reactor was not determined due to the lack of
theories and empirical correlations applicable to such
systems.
Nevertheless, the semi-batch operation of our reactor,
significant reactor start-up and shut-down times, possible
wall effects, and reactor size should be considered as
important factors that would tend to decrease the efficiency
of steam-catalyst contacting. Since practical observations
were made in the past and showed significant improvements
in sulfur removal efficiencies by going from pilot to
commercial scale, the improvements of the efficiencies
observed in a laboratory scale reactor are even more likely.
*Bena, J., J. Ilavsky, E. Kossaczsky, and L. Neuzil. Changes
of the Flow Character in a Fluidized Bed. Collect. Czech.
Chem. Commun. (Prague). 28:293-308, 1963.
83
-------
The effect of steam stripping upon by-product formation
was also determined during this program. The data indicate
that there is no appreciable formation of sulfur compounds
other than the hydrogen sulfide. Some carbonyl sulfide
(COS) and carbon disulfide (CS2) were found in the stripper
off-gas but the maximum concentrations amounted to less
than 0.5% by volume of the total sulfur concentration in
the gas stream.
Other by-products, namely hydrocarbons, were also detected
in the stripper off-gas. These were present in relatively
low concentrations except for methane which often exceeded
the H2S concentration on molar basis.
Heavier hydrocarbons in stripping steam condensate were
detected as total organic carbon (TOG). A linear corre-
lation on logarithmic paper was observed between steam
stripping rate and TOG concentration as illustrated in
Figure 17- The concentration consistently decreases with
an increasing steam stripping rate. The absolute amount °*
hydrocarbons found in the stripper effluent condensate,
however, does not seem to vary. This suggests that only
limited and fixed amounts of hydrocarbons may be stripped
off the catalyst with additional amounts of steam diluting
the condensate stream.
This also indicates that essentially all strippable hydro"
carbons will leave the catalyst with first steam in the
very initial phase of catalyst steam contacting. Thus,
separating this first steam that carries most of the
hydrocarbons may reduce hydrocarbon concentrations in the
steam condensate.
84
-------
Table 24 presents some approximations of maximum concentrations
for the stripper off-gas condensate based on results obtained
in our experimental program. Table 25 summarizes waste-
water loadings for typical refinery operations. Comparing
the data in Tables 2 4 and 25, we can conclude that refineries
are currently treating effluents which have waste loadings
similar to or higher than those produced by steam stripping.
Hence, no new control technology will be needed to solve
expected water pollution problems. Expansion of the existing
wastewater treatment facilities may be required, however,
due to the increased wastewater flow. Some carbonyl
sulfide may be present in the steam condensate. It appears
that the amount of COS formed in steam stripping depends
upon the type of catalyst and probably its history.
Consequently, the COS concentration should be determined
experimentally. Also, the effect of acidity of steam
condensate on the amount of dissolved COS is not known.
This would be important if the steam condenser is operated
in the manner described in Section 5.3.
Presence of ammonia in the stripper off-gas was also ob-
served. Its formation occurs apparently in the same manner
as that of hydrogen sulfide (Phase I report, pages 58 & 59).
Actually, all of the ammonia present in the stripper off-
gas dissolved in the condensate.
Several tests were performed to determine the effect of
spent catalyst steam stripping upon catalyst deactivation.
Samples of spent catalyst which had been steam stripped and
then regenerated in our laboratory were sent for analysis
to Warren Letzsch of W. R. Grace Company, a manufacturer
of FCC catalyst. The results of the analyses performed
85
-------
Table 24. ANALYSIS OP STRIPPER OFF-GAS CONDENSATE
Stripper Operating Conditions
Temperature = 811 K (1000°F)
Pressure = 2.39 x 105 Pa (20 psig)
Steam stripping rate = 6 kg H20/100 kg catalyst
Condensate Composition (no sulfide separation)
Component Concentration
(kg/m3)
Total organic.carbon (TOG) 1.300
£»
Biological oxygen demand 2.600
Ammonia 0.420
aCalculated from TOG (BOD5 = 2 x TOG).
•L
No hydrocarbon recovery was assumed.
86
-------
Table 25. REFINERY WASTEWATER LOADINGS FOR TYPICAL REFINING TECHNOLOGY'
CD
Fundamental Process
Crude oil and product
storage
Crude desalting
Crude fractionation
Thermal cracking
Catalytic cracking
Reforming
Polymerization
Alkylation
Solvent refining
Dewaxing
Hydrotreating
Deasphalting
Drying and sweetening
Wax finishing
Grease manufacturing
Lube or finishing
Blending and packaging
DWL/5
(kg/m3)
0.30
1.20
0.0005
0.060
0.040
t
0.0026
0.0020
b
2.60
0.240
b
0.150
b
b
b
b
gal/bbl feed
0.4
0.2
50
2
30
6
140
60
8
23
1
b
40
b
b
b
b
m3/m3
0.0095
0.0048
1.19
0.048
0.714
0.143
3-33
1.43
0.190
0.548
0.024
b
0.952
b
b
b
b
(kg/m3)*
b
0.0060
0.0024
0.012
0.080
0.014
t°
0.0003
0.043
0.008
tc
b
0.030
b
b
b
b
OUJ.1 J.UCO ,
(kg/m3)
b
1.20
0.0024
0.060
0.012
0.020
0.0086
0.020
tC
tC
0.240
b
b
b
b
b
b
Jones, H. R., Pollution Control in the Petroleum Industry, Noyes Data Corporation,
Park Ridge, N. J., 1973-
Data not available for reasonable estimate.
tcTrace
-------
indicated that the exposure of catalyst to steam for 900 5
(15 minutes) at 2.39 x 105 Pa (20 psig) and temperatures
ranging from 755 to 797 K (900 to 975°F) did not cause any
catalyst deactivation.
In summation, the experimental work did not reveal any
information that would require changing conclusions drawn
in the Phase I final report (pages 39-43). The steam
stripping of spent catalyst is a technically feasible meth0
of reducing FCC regenerator SOX emissions. This method
does not form large quantities of undesirable compounds t>u*
forms H2S which can be converted to saleable grade sulfur
at existing refineries. Also, it appears that steam
stripping will cause no drastic catalyst deactivation as
suspected at several petroleum refineries.
4.4 DATA REGRESSION ANALYSIS
The experimental data presented in the previous section
were analyzed using regression analysis techniques. The
empirical correlation form which best fits the data obtai^6
from all catalysts is presented below. For some catalysts*
experimental data may seem to fit other correlations bettef
than that in Equation (3). Our intent, however, was to
obtain a general correlation that would represent best the
reaction phenomena for all catalysts.
ln(S02)out = ln(S02)ln - k TS(SSR) (3)
88
-------
where (S02)0ut = equivalent S02concentration (vppm)
after catalyst steam stripping
(equivalent to residual sulfur
content on coke)*»+
(S02)in = equivalent regeneration S02
concentration (vppm) before
catalyst steam stripping (equivalent
to initial sulfur content on coke)*»+
k = proportionality constant
Ts = steam residence in catalyst bed
in stripper (minutes)++
(SSR) = steam stripping rate (kg H20/100 kg
of catalyst)
If we define fractional sulfur removal efficiency as
(soz)ln-(so2)out
(S02) ,.
= 1 - > < °ut
^nr
Equation (3) becomes
In(l-X) = - k T (SSR) (6)
s
*Both concentrations were calculated according to the
procedure outlined in Section 4.2.
+vppm = 106 mole fraction
++1 min = 60 seconds
89
-------
The data used in regression analysis for each of the
catalysts and their manipulation to determine the constant3
of Equation (3) are summarized in Table 26. This table
includes additional calculated values according to the
following nomenclature:
Equation (3) can be simply transcribed
in the form
Y = A + BX
where Y = ln(S02)QUt
X = TS(SSR)
A = ln(S02)±n
B = -k
and X = mean of X
crx = standard deviation of X
TT = T - test
EE = standard error of estimate
R = simple correlation coefficient
At the end of this section a summary of correlation
equations obtained for each catalyst from regression
analysis is presented and their agreement with experi-
mental results demonstrated (Figures 18 through 25).
90
-------
Table 26. STEAM STRIPPING DATA REGRESSION ANALYSIS
Experiment
B 3
5
7
8
9
11
C 6
7
8
10
11
15
17
18
21
D 2
5
7
E 6
15
16
17
18
21
22
23
24
F 9
10
11
12
13
14
15
17
18
22
25
27
(SO?) out
1.23
0.791
1.36
1.48
2.05
1.81
0.888
0.924
0.798
1.80
1.27
1.75
1.48
2.43
0.631
1.63
1.32
1.73
0.880
1.63
1.48
1.84
1.06
0.937
0.642
0.879
0.826
0.782
0.927
0.748
1.37
1.29
1.72
1.50
1.88
1.36
2.63
2.41
3.29
5.31
32.8
17.6
39-8
9.01
21.9
7.31
14.1
32.6
16.0
5.98
7-59
219
109
631
17.7
90.5
166
97.8
6.32
13-3
5.20
2.85
11.6
1.34
6.16
11.8
5.54
10.5
2.17
6.99
7.41
11.1
9.57
9.87
9.07
8.20
11.9
2.30
979
732
951
721
805
769
592
549
540"
493
610
559
460
483
471
385
217
168
162
362
305
361
362
339
393
368
366
438
408
477
416
440
342
377
417
290
316
256
422
6.5313
25.945
23.936
58.904
18.471
39.639
6.4913
13.028
26.015
28.800
7.5946
13-283
324.12
264.87
398.16
28.851
119-46
287.18
86.064
10.302
19.684
9.5680
3.0210
10.869
0.8603
5.4146
9.7468
4.3323
9-7335
1.6232
9.5763
9.5589
19.092
14.355
18.556
12.335
21.566
28.679
7.5670
.6.8865
6.5958
6.8575
6.5806
6.6908
6.6451
6.3835
6.3081
6.2916
6.2005
6.IU35
6.3262
6.1312
6.1800
6.1549
5.9532
5.3799
5.1240
5.0876
5.8916
5-7203
5.8889
5.8916
5.8260
5.9738
5.9081
5.9026
6.0822
6.0113
6.1675
6.0307
6.0868
5.83^8
5.9323
6.0331
5.6699
5-7557
5.5452
6.0450
TT
R2
28.904 18.210 -2.105 0.5257 6.86 5.26xlO~3
120.26
160.28
-3.189 0.6791 6.33 5.26X10-1*
145.16
131.07
-2.446 0.8568 5-92 3.00xlO~3
17.281 26.355 -23-067 0.9870 5-96 l.OxlQ-2
13.081
7.732 -4.388 0.6582 6.19 2.0xlO-2
-------
Table 26 continued. STEAM STRIPPING DATA REGRESSION ANALYSIS
Experiment
•H
ro
5
6
7
8
11
12
13
14
15
16
17
18
19
21
22
23
24
25
26
28
29
30
31
32
5
8
9
10
11
12
13
14
15
16
0.539
0.847
1.01
0.81
0.755
.10
.78
.60
.68
.82
.84
.55
1.42
2.42
1.92
2.03
2.54
1.90
2.53
0.803
0.777
0.765
0.663
0.525
0.517
0.596
0.648
0.669
0.648
.40
.21
.48
.58
4.01
7.68
12.9
32.1
115
4.37
5-38
8.95
17.1
31.6
65.0
185
4.67
164
89
78
1.46
15.6
41.8
65.6
72.0
179
18.2
10.5
8.81
4.19
29.1
53-"
194
13.4
3-99
31.6
97.1
182
49.1
175
T43
136
139
73
110
136
127
144
119
95
84
164
63
148
126
106
99
109
103
85
111
102
180
477
386
361
304
422
460
400
304
274
300
X
2.1614
6.5050
13.029
26.001
86.825
4.8070
9-5764
14.320
28.728
57.512
119.60
286.75
6.6314
396.88
13.229
19.853
39.624
79-420
165.97
57-816
139.08
13-923
6.9615
4.6253
2.1662
17.344
34.603
129-79
8.6832
9.5760
38.236
143.71
287-56
71.686
y x
5.1648 72.496
4.9628
4.9127
4.9345
4.2905
4.7005
4.9127
4.8442
4.9698
4.7791
4.5539
4.4308
5-0999
4.1431
4.9972
4.8363
4.6634
4.5951
4.6914
4.6347 44.482
4.4427
4.7095
4.6250
5.1930
6.1675 74.335
5.9558
5.8889
5-7170
6.0450
6.1312
5.9915
5-7170
5.6131
5-7038
ux
106.23
57.134
90.186
TT
R2 A
-4.388 0.6357 4.91 2.02xlO-3
-5.447 0.4331 4.87 3.24xlO~3
-4.362 0.7040 6.03 1.82xlCr3
This experiment was performed at 900°F; all others were performed at 1000°F.
-------
7.0
c
"ro
i_
•4-»
c
o>
o
CSJ
O
v.
o
•*->
ro
CD
C
O)
6.0
| 5.5
Temperature, 811 K(1000°F)
mole fraction 2xlO"4 (200 vppm)
ro
'5
o-
"o
E
J L
| ' ' 1 1 1—-I 1 L
J L
0 4 8 12 16 20 24 28 32 36 40 44 48 52 56 60
TS(SSR)
Figure 18. Summary of results of the regression
analysis performed on B-Series catalyst
-------
7.0
o
co
"c
-------
a
o
'-t-f
ro
CD
0 i C
§ 6'5
O
CM
O
LO
6.0
0)
c
CD
5.5
•5 5.0
(O
S1
Temperature, 811 K(1000°F)
mole fraction 2xlO"4 (200 vppm)
i i 1—I—I—I—I—I i
4 5 ''«i 1—i 1 1 1 1 1 1 1 1—i i
' 0 20 40 60 80 100 120 140 160 180 200 220 240 260 280 300
TS(SSR)
Figure 20. Summary of results of the regression
analysis performed on D-Series catalyst
95
-------
7.0
o
U-»
co
o
o
CM
O
CO
6.5
Temperature,811 K (1000°F)
co
1_
-------
f.u
c
o
ro
OS
g 6.5
0
O
C\J
0
CO
I 6-0
CD
C
-------
c
o
"03
s—
•4—'
c
o
o
o
o
CO
^
CO
i_
o>
c
-------
c 7.0
0
ro
"c
o>
o
g 6.5
o
(XI
o
v_
o
I 6-°
C
CD
CP
Oi
1 5.5
CO
._
cr
LU
«*—
Q
f^J
0
Temperature, 755 K{900<*F)
•»
, mole fraction 2xlO"4 (200 vppm)
•
MM
*, 11' I J ^J^L^T ••_ | • ,
10 20 30 40 50 60 70 80 90 100 110 120 130 14
TS(SSR)
10
Figure 24. Summary of results of the regression
analysis performed on H-Series catalyst
99
-------
7.0
c
o
H—»
ro
L_
•4—•
C
O)
/• r
6.5
o
o
CVJ
O
o 6.0
03
O>
O)
en
c
CD
5.5
o
E
5.0
03
SM.5
Temperature, 811 K(1000°F)
0
^mole fraction 2xlO"4 (200vppm)
50
100
150
Ts(SSR)
200
250
300
350
Figure 25. Summary of results of the regression
analysis performed on I-Series catalyst
100
-------
Summary of Correlation Equations
Catalyst
B
C
D
E
F
Ha
ln(S02)out
ln(S02)out
ln(S02)0ut
ln(S02)out
ln(S02)out
ln(S02)out
ln(S02)out
ln(S02)out
\J m
6.
5.
5.
6.
4.
4.
6.
33 -
92 -
96 -
19 -
91 -
87 -
03 -
^ . c u A -L u ig v,'->l->riy - J> • u
5.262xlO-4 Ts (SSR)± 1.9
3.
1.
2.
2.
3.
1.
00x10-
0 xlO-
0 xlO-
02x10-
24x10-
82x10-
3
2
2
3
3
3
T
T
T
T
T
T
s
s
s
s
s
s
(SSR)
(SSR)
(SSR)
(SSR)
(SSR)
(SSR)
± 8.2
± 0.1
± 3.9
± 7.0
± 10.
± 3.8
a = 1000°P
b = 900°P
The percent error for each equation was calculated at the
mean X for interval of 2 times EE, which should include
95% of all data in regression analysis. Table 27 summarizes
the term Ts (SSR) calculated from the correlation equations
to obtain 200 vppm equivalent regenerator S02 concentrations.
Table 27
Catalyst
STEAM STRIPPING REQUIREMENT FOR
SULFUR REDUCTION TO 200 vppm
Tg (SSR)
B
C
D
E
F
H
I
297
1961
207
66.2
45.6
136 vppm initially
402
101
-------
5. STEAM STRIPPING PROCESS DESIGN
Applying the steam stripping process for refinery FCC u
regenerator SOX control requires several processing steps-
In the Phase I final report several process alternatives
were proposed and one of these (Option 1) was evaluated in
detail. The laboratory development program (Phase II) did
not reveal any evidence that would require a modification
of the Option 1 alternative. However, in many cases new
technical information was obtained or generated which
enables a better understanding of the individual process
steps for optimization of equipment design. In this
section, discussions on individual processing steps and
technical background information are presented and appli6
to proper processing equipment design.
5.1 CATALYST STEAM STRIPPER DESIGN
In applying the Option 1 process alternative in the Phase
final report, we indicated that several types of equipmefl
may be used to contact the spent catalyst with steam.
Namely, these are:
Pluidized bed catalyst stripper (similar to existing
FCC unit reactors and regenerators)
Counter-current, stagewise contacting (similar
strippers presently used in refineries and
in Figure 3, Phase I final report, p. 17)
102
-------
Co-current, plug-flow contacting (similar to riser
reactor concept applied to FCC hydrocarbon cracking;
this concept would require a catalyst disengagement
step following the stripper)
These concepts are discussed below.
Computer analysis of the experimental data obtained in the
laboratory development program revealed that the removal
efficiency of sulfur from the coke on spent catalyst is a
function of several variables. Specifically, the mathe-
matical correlations containing the steam stripping rate
and time during which the catalyst is exposed to steam were
developed for all catalysts tested. The correlations were
presented in Section 4.2.3 and appeared to have the general
form of Equation (6) for all catalysts.
The constant in the Equation (6) includes factors such as
temperature, pressure, type of catalyst, type and concen-
tration of sulfur compounds on coke, and type of equipment
in which the catalyst steam contacting takes place. The
equation may become very useful in further development of
the steam stripping process. Its further extrapolation to
large scale units, however, will have to be verified.
5.1.1 Semi-Batch Fluidized Bed Reactor
The following theoretical discussions will further clarify
the meaning and interpretation of the constant in Equation
(6). As a starting point, we will assume that the catalyst
steam stripper can schematically be depicted in the manner
shown in Figure 26. The operation of this stripper is
similar to our experimental reactor.
103
-------
Spent FCC Catalyst
sulfur concentration [SI
Steam +H2$
Steam (no Sulfur)
Figure 26. Schematic' of spent catalyst steam stripper
-------
Assuming that the rate of sulfur removal from spent FCC
catalyst is a function of the sulfur content of coke and
amount of steam to which the catalyst is exposed, the
following correlation may be used to describe this relation-
ship.
- = k [S]* [H20]b (7)
where [S] = concentration of removable sulfur on
spent catalyst
[H20] = concentration of steam
a»b _ constants expressing the order of
stripping reactions
k = proportionality constant in any of the
equations below, this constant must be
experimentally determined
t - time
Previous work related to catalyst coke composition revealed
that the sulfur on the coke can be in various forms. This
was also discussed in the Phase I final report, Section
5.2.2. Mathematically we can describe the sulfur forms on
the catalyst as follows:
ST = SR + S (8)
where Sip = total sulfur on spent catalyst
SR - residual sulfur
S = sulfur removed by steam stripping
Thus, in Equation (7) above, [S] represents the sulfur that
can be removed by contacting the catalyst with steam.
105
-------
In order to integrate Equation (7), certain boundary
conditions must be defined. These are:
[S] = Sj at t = 0
[S] = S2 at t = T_ (9)
O
[S] = SR at t = -
where TC = catalyst residence time in the stripper, or
time during which the catalyst is exposed to steal11
In operation with excess steam we can assume [H20] =4= f
conversion). Integrating Equation (7) and applying the
above boundary conditions we obtain
= -k [n2o]
b T
1-a L""^' - v«i/ j ». ^2^ ic
for a-f* 1, and
In [S2] -In [SJ = -k [H20]b T
C
for a = 1.
if*1
In Equations (10) and (11), Sj represents the initial suJ-*
concentration on coke and S2 the sulfur concentration on
coke after the time T .
To better compare the experimental results with these
correlations we will define a new variable, X, represent^
if,1
the fraction of sulfur removed from the catalyst in time
-------
This variable is identical to the one defined by Equation
, Section 4.4.
Substituting Equation (12) into Equations (10) and (11) we
obtain
--le [H20]bTc (13)
1-a
and
In (1-X) = -k [H2o]b Tc (14)
Further modification will be made with the right side of
Equations (13) and (14). First, we will multiply the terms
on the right side by TS/TS, where Ts is the steam residence
time in catalyst stripper.
-k [H2o]b Tc = -k [H20]b Tc |a (15)
Examination of the Equation (15) will reveal that new terms
which were actually measured in our experimental studies
can be introduced in Equations (13) and (14).
The steam residence time Ts can also be expressed as a
function of reactor volume VR:
107
-------
where Vg = volume rate of steam (m3/s)
VD = reactor volume (m3)
rt
ps = steam density (kg/rn3); in excess conditions
equal to steam concentration
M0 = steam mass rate (kg/s)
Similarly, we can express catalyst mass in the reaction ^
C = VR pc (17)
where C = mass of catalyst used for an experiment (Kg/
pc = catalyst density in the reactor (kg/m3)
Substituting TS and using Equation (17) we will introduce
a steam-to-catalyst ratio term in Equation (15).
-k [H20]b Tc = -k[H2o]b n s P
-------
= -k [H20]b |H Ts (SSR) (21).
L—a J
for a -f 1, and
In (1-X) = -k [H20]b ^ T_ (SSR) (22)
Pg S
for a = 1.
Equations (21) and (22) are in a general form which can be
used to determine the effect of steam stripping rate and
steam residence time on sulfur removal efficiency. The
right side can be further simplified according to the
following assumptions. If the effect of steam on the sulfur
removal is of the first order, the constant b = 1 (our
regression analysis of experimental data showed that this is
a reasonable assumption as presented in Section 4.4) and
the steam concentration will cancel out with actual steam
density. This is possible because of high excess of steam.
Also, if the catalyst reaction density pc is considered
constant over a narrow range of operating conditions and
its change is expressed in terms of steam residence time
T (normally T ,x p = constant), the term p may be
o S C C
included into the proportionality constant k. Thus, our
final simplified equations will become
"a
= -k Ts (SSR) (23)
for a + 1, and
In (1-X) = -k T (SSR) (24)
S
109
-------
for a = 1, or
X = l-e-k T (SSR) (25)
s
The regression analysis of experimental data produced a
correlation that very well satisfies Equation (24), which
suggests that the effect of sulfur on the sulfur removal
rate is also of the first order.
Together, Equations (23) and (24) can be used to predict
the sulfur removal efficiency as a function of steam
stripping rate and steam residence time. Assuming that
proportionality constant k will represent the effects of
temperature, pressure, type of catalyst, and type of
catalyst contacting device, the applicability of the
equations is further expanded.
5.1.2 Rate Controlling Factors
Several groups of data obtained for the same catalyst at
equal catalyst residence times but various stripper steam
velocities confirmed that the same sulfur removal can be
obtained with lesser amounts of steam at lower steam
velocities as long as the catalyst residence time remain0
the same. Several steps might be involved in bringing ^e
sulfur on the catalyst to the form in which it is removed-
Some of the steps are suggested below even though it is ^°
known which of these steps are the controlling ones.
(1) Rate of diffusion of steam through the pores in
catalyst particle.
110
-------
(2) Rate of reaction of steam and sulfur compounds on
catalyst surface.
(3) Rate of diffusion of product H2S through the pores in
catalyst.
(*1) Rate of diffusion of product H2S through laminar boundary
layer surrounding each catalyst particle.
Should the first or third factor be controlling, our
experimental data would have shown essentially no effect of
pressure on sulfur removal efficiency. Increased pressures,
however, resulted In higher removal efficiency (see Figure 13),
Should the second factor be controlling, a significant
difference would be expected between the data measured at
different temperatures. This effect, however, can be
partially compensated for or enhanced by the other effects,
the fourth one in particular. Nevertheless, essentially
very minimal temperature effect was observed.
A single effect of the fourth factor should indicate a
significant improvement in sulfur removal with an increased
turbulence in the reactor. Insertion of static mixer in the
reactor should enhance the turbulence of the fluidized bed
but it did not result in a remarkable improvement and
consequently does not support the significance of laminar
layer diffusion. However, the theory of the laminar and
turbulent conditions in fluidized beds is not well defined
and, as mentioned earlier, the effect of static mixer on
turbulent conditions in the fluidized bed is not easy to
quantify. Therefore, it is difficult to objectively
evaluate the improvement of turbulent conditions in the
fluidized reactor by the use of a static mixer.
Ill
-------
In conclusion, the minimal effect of temperature and rather
significant effect of pressure on sulfur removal efficiency
seem to indicate that kinetics (Factor 2) is the controlling
step, with elevated pressures resulting in easier sulfur
removal. As a result, potential improvement in sulfur
removal kinetics may be sought through an investigation of
catalytic effects of trace elements normally contained in
petroleum feedstocks and deposited on FCC catalyst during
the cracking process. These effects were not evaluated
in this program.
Better sulfur removal efficiencies were observed by going
from pilot to commercial scale stripper (Phase I final
report, Section 5.2.1). This seems to suggust that Factor
4 has some significance. This can be partially explained
by the more uniform conditions and the minimization of wall
and start-up effects in commercial units. Considering this
observation, our experiments performed in a fluidized bed
reactor in a rather semi-batchwise manner should not be
viewed as representative of FCC commercial units since
substantial wall, start-up, and mixing effects were
probably present. Consequently, we feel that the experi-
mental results observed on the laboratory scale can be
significantly improved in operations of commercial size.
5.1.3 Other Stripper Designs
The petroleum industry has spent roughly 40 years develop-
ing various catalytic cracking processes. Research performed
in this area has resulted in improvements in FCC reactor
design, spent catalyst steam stripping for hydrocarbon
removal, and spent catalyst regeneration. The experience
112
-------
gained in these extensive R&D efforts can be used to aid
in designing spent catalyst steam strippers for the pur-
pose of sulfur removal. Our semi-batch experimental
reactor was described previously. Three other design
alternatives are currently envisioned, including continuous
fluidized bed, counter-current, and co-current reactors.
The theoretical math model developed previously for semi-
batch bed is extended below to cover the other two
stripper design alternatives.
5.1.3.1 Continuous Fluidized Bed Reactor -
The same type of kinetic model development used to
determine the behavior of a batch fluidized bed catalyst
stripper can be used for a continuous fluidized bed
catalyst stripper if we assume that the stripper behaves
as a constant stirred tank reactor (CSTR), Figure 27.
Basically, we can assume that [S] = S2 = constant and
no concentration change occurs in the reactor, or -rr- = 0.
The sulfur concentration entering the reactor must be equal
to the concentration leaving the reactor plus the amount
of sulfur reacted, or
Sj -S2 -k S2a [H2o]b Tc = 0 (26)
Further modification of the equation is possible:
o Q
- " * [H2o]b T (27)
113
-------
H20 + H2S
Catalyst [s] - S2
\
Catalyst Steam
Figure 27. Schematic of continuous fluidized bed stripper
-------
for a + I, and
S2 _ 1 (28)
Sj - 1 + k [H20] T
for a = 1.
Considering Equation (12) and assuming a = 1, and b = 1,
the sulfur removal can also be expressed as
1 - X «
+ k [H20] T(
or
v _ k [H20] TG
using Equation (16) and the following Equation (32)
,„
s/C
M , - (SSR)
M
115
(30)
A - 1 + k [H20] Tc
Expressing T in the following form,
T = ^A1 (3D
-------
we can include steam stripping rate into Equation (30)
k[H20] £& TS (SSR)
1 + k[H20] -°- Ts (SSR)
" S
Applying the assumptions listed on page 109} we obtain
equation in its final form
k TR
1 + k T (SSR)
5.1.3.2 Plug-Flow Stripper -
A spent catalyst steam stripper can be operated such
j
the steam and catalyst pass through the stripper in the $®
direction, or co-currently . Co-current steam stripping ^
thus performed in a transfer line, as in a plug-flow or ^
riser reactor, all three terms implying the type of opera
shown in Figure 28.
For the plug-flow reactor Equation (7) can be expressed
as sulfur concentration change with the distance
- ft • k ts]a [H2o]b (35)
Due to large excess of steam, it can be assumed that [H2°
+ f(x), and after separation of variables, integration °
Equation (35) is possible.
e
l
116
-------
Steam
+H2S * I
Catalyst t-Tc=Ts
= So
Steam
tt
Catalyst t=0
LSure 28. Schematic of plug flow steam stripper
117
-------
= -k[H20]b dx (36)
_k[H20]b L (3T)
for a 4=1, and
In (S2) -In (S!) = -k[H20]b L
for a = 1
where L = length of the plug reactor
Since T = T , the reaction length can also be expressed
c s
in terms of time as follows :
L=vxT =vxT
c s
where v = velocity of steam-catalyst mixture through
the reactor
Using Equations (16) and (31) the ratio of T /T can
c s
determined:
T " c P
S
from which
T = T
c C s ps
118
-------
Expressing steam-to-catalyst ratio by steam stripping rate
we can write
C 100
Now by substituting Equations (39), (4l), and (42) into
Equations (37) and (38) we obtain
- -k|K20]b v T <> (1,3)
for a + 1 , and
in (S2) -In (SO = -k[H20]b v TB I|ggl
for a » 1.
Applying the simplification assumptions of a = 1 and b = 1,
cancelling ps with [H20], and including pc/100 into k, we
obtain
In g^- = k v Tg (SSR) (45)
Using Equation (12) we can write
In (1-X) = -k v Tg (SSR) (46)
or
_e -k v Ts (SSR) (||?)
119
-------
5.1,3.3 Counter-Current Stagewlse Contacting -
The design of a counter-current, stagewise catalyst steam
stripper is depicted schematically in Figure 29. For
Jt ffP*
of the stages in the contactor, the behavior of the fluid*2
bed is the same as in the continuous f luidized bed strip?6
In this design, catalyst is flowing downward by gravity
from stage to stage with each stage performing as an
size fluidized bed. Each stage uses the off-gases from
next lower stage for f luidization.
It is assumed in this model that the sulfur concentration
in the vapor phase does not affect the sulfur removal
efficiency. The sulfur material balance can be determin6
by considering the desorption of sulfur from the spent
catalyst while disregarding the re-adsorption of sulfur W
the catalyst. This assumption can only be verified by
experimentation. Our experiments have supported the fac*
that the sulfur removal is controlled kinetically rather
..41$
than by equilibrium. Consequently it is Justified to a»p
that re-adsorption has a minimal effect. With this in
the counter-current reactor becomes a backmix reactor.
Each stage is equal in size and can then be described
a continuous fluidized bed reactor and the following
relationships for each stage of the contactor exist:
S -S! -k Sla [H20]b T „ 0
= k [H20]b
s a <• ci
120
-------
catalyst in
porous distribution
plates
steam in
catalyst out
Figure 29. Counter-current, stagewise contactor
121
-------
s — s
1 ? = k [H20]b T (50)
S — S
n" n= k [H20]b T
S
n
where n = number of reactor stages
Assuming a=l, and applying Equation (12) in the form
x =
n
n-!
the set of Equations (49) through (51) will become
Si
So 1 + k [HaO]1
82
__
(53)
Sn-l n 1 + k [H20]b T
Multiplying Equations (53) through (55) and assuming
T = T = = T (5
ci C2 en
122
-------
we obtain
S
=
'o (T + k [H20]b Tcn]n
(57)
Defining sulfur removal efficiency for the whole reactor as
S -S S
X - -4-^- 1- ^ (58)
So So
and catalyst contact time in the reactor as
Tc = n Tcn (59)
we can modify Equation (57) as follows:
1-X = ± E =- (60)
Fl + k [H2or Tc/nln
or
X". 1 - { ± E f (61)
k [H20]D Tc/n
Using Equations (16), (3D, and (32), assuming b=l, applying
assumptions from page 109, and including n In k, Equation
(61) will yield
(62)
123
-------
5.2 CONDENSER DESIGN
The catalyst stripper effluent steam may contain H2S,
and hydrocarbons (see Tables 20 through 21). This steam ^
condensed to recover heat and to achieve a separation of
H2S from the water. A common shell and tube heat exchang61"
may be used with the stripper off-gas being passed through
the tube side of the heat exchanger. Cooling water, on tb
heat exchanger shell side, is used as process water for
subsequent steam stripping. The stripping steam is cooled
from 811 K (1000°F), condensed, and subcooled while the
process water is vaporized to produce saturated steam at
9.63xl05 Pa (125 psig).
In order to minimize corrosion in the condenser, the
materials of construction should be at least a low grade
stainless steel, such as 5% Cr plus 1/2$ Mo alloy.* How^6
eU^
the current trend is to more expensive stainless steels B"
as types 321 and 3^7 after stabilized annealing.**
The composition of the process water should meet or surpaS
the boiler feedwater specifications before it enters the
steam superheater. These specifications are presented i*1
Table 28.
*Pontana, M. G., and N. D. Greene. Corrosion Engineer!^'
New York, McGraw-Hill Book Co., 1967.
**Evans, P. L. Refiners Pace Corrosion Pacts.
Processing. £3:109-112, April 197^.
124
-------
Table 28. RECOMMENDED LIMITS OP SOLIDS IN BOILER PEEDWATER
. Below 600 to 1000 to Over 2000
Drum pressure 600 psi 1000 psi 2000 psl psl
a
Total solids, ppm
Total hardness as
ppm CaCOs 0
Iron, ppm
Copper, ppm
Oxygen, ppm
PH
Organic
0
0
0
8
0
.1
.05
.007
.0-9.5
0
0
0
0
8
0
.05
.03
.007
.0-9.5
0
0
0
0
0
8
0
.15
.01
.005
.007
.5-9.5
0.
0
0.
0.
0.
8.
0
05
01
002
007
5-9.
5
aSteam/its generation and use. New York, Babcock and
Wilcox, 1972
bl psl - 6.895 x 103 Pa
In our cost analysis, the design of the condenser was
based upon the normal heat exchanger design equation (Phase I
final report, Appendix E)
" Q - UAATm (63)
where U was assumed to be 3975 W/m2K (700 Btu/ft2'°P-hr).
The AT was calculated from the heat exchanger terminal
m
temperatures according to the following example:
811 K (1000°P) • stripper off-gas inlet to heat exchanger
311 K (100°P) * stripper off-gas outlet from heat exchanger
300 K (80°P) = process water inlet
451 K (353°F) " 9.63xl05 Pa (125 psig) steam, saturated
125
-------
AT = (B11-ljj IK, 1:L-300) = 100.1 K (6«)
m , oll-*ol
in 311-300
Based on the above assumptions and knowing the amount and
conditions of steam used for catalyst steam stripping, tfte
heat transfer area for the condenser can be calculated
from Equation (63).
-i ,/(.}
A/Q'= 3975x100 1 = 2.5l4xl06m2/W (7.92 sq ft/106Btu/hr) W
5.3 ACIDIPIER/PHASE SEPARATOR DESIGN
5.3.1 Equilibrium Relationship
The design of the acidifier/phase separator system will
be dictated by several factors. These include the hydrog6
sulfide content of the steam stripper off-gas, the temp6*"3
ture, the system pressure, the' hydrogen ion concentration
(pH) of the condensate, and the allowable H2S concentratl0
of the effluent wastewater. The H2S content of the strlpP
gas will be dictated by the sulfur content of the spent
catalyst and the efficiency of the stripper. However, #2
\
concentrations of up to 2.0xl03 mole fraction (2000
can be expected in the stripper overhead vapors. The
of the condensate and its sulfide content may also be
affected by the ammonia content of the stripper off-gas-
The allowable H2S content of the phase separator conden- .
sate will be dictated by either federal, state, or local
126
-------
standards for refinery effluents. However, a sulfide
concentration 10~3kg/m3 will probably be the highest
tolerable level.*
It should be noted that as long as the condensate contain-
ing dissolved and unreacted H2S (g) is exposed to ambient
air after leaving the condenser, the H2S gas will diffuse
out of the liquid and leave zero H2S concentration level.
The rate of this diffusion will depend on the H2S concen-
tration, the temperature of the condensate, and the
effectiveness of liquid-air contacting downstream of the
condenser. Consequently, the final residual sulfide
concentration in the condensate effluent will be a function
of the diffusion rate, the amount of mercaptans condensed,
and the amount of compounds that can tie with H2S and form
sulfide salts.
Since essentially pure steam is used in the steam stripping,
the compounds that can react with H2S to form sulfides must
be formed in the steam stripping process. Our experiments
showed that only one such compound is formed, ammonia.
Thus, the total sulfide concentration [TSS] in the conden-
ser water effluent may be expressed as follows
[TSS] = [H2S] + [MSH] + [(NHlt)2S] (66)
*Topical Law Reports, Pollution Control Guide. New York,
Commerce Clearing House, Inc., Vol. 2, Part 419, pp. 9627-
9627-19.
127
-------
where [H2S] = dissolved hydrogen sulfide
[MSH] = mercaptan sulfide
[(NHif)2S] = ammonium sulfide
Data from our experiments indicated that practically no for-
mation of mercaptans occurs. Hence, the second term of
Equation (66) may be neglected.
[TSS] = [H2S] + C(NH4)2S] (67)
As shown in Table 22, we have observed ammonia formation
in concentrations ranging from 3. ^xlCT^-S. 2xlO~'tmole fraction
to 520 ppm).
Table 29 summarizes the first step dissociation constants
for hydrogen sulfide in the temperature range between 5 and
60°C.* Assuming a temperature of 25°C we can calculate the
relative contents of species produced from hydrogen sulfide
as a function of pH, Figure 30. The second step dissociation
constant at this temperature is 1.0 x 10~15.* Both constants
can be written as follows
Kl =
[HS-]
* Gmelins Handbuch Der Anorganischen Chemie (Gmelins Hand-
book of Inorganic Chemistry), 9th Edition, Number 9,
Sulfur, Section Bl. Weinheim/Bergstrasse, Germany: Verlag
Chemie, GMBH; 1953-
128
-------
Table 29. IONIZATION CONSTANTS FOR THE
H2S-WATER SYSTEM AT VARIOUS TEMPERATURES
Temperature, 7
on K, x 10'
o .1
5 0.471
10 0.574
15 0.747
18 0.910
20 0.853
25 1.08
25 1.15
30 1.26
40 1.64
50 2.03
60 2.39
at
k = tc + 273.15.
129
-------
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w p ct
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O i— 13
D CQ
0"9 ii &
MI_J H-
ft)
cr
pj CO P3
M 3
H- p
O T3 H)
•O O
« 4 4
O fT>
M o tr
C H-«<
rt p fj.
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3 Q 0«3
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o
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o c
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ct
1.0
0.9
0.8
0.7
0.6
0.5
0.4
0.3
0.2
0.1
0
HS"
25°C
3
10 11
12
13
14
-------
For our purposes, K^ values of up to 100°C are needed. They
were obtained by fitting the data presented to an empirical
equation below. The equation shown also in Figure 31, pre-
dicts the ionization constant at any temperature in the
range of 0 to 60°C with greater than 99.9% confidence.
K! = [(0.0356655)T + 0.2288326] x 10~7 (70)
where T = temperature (°C)
In the absence of data for temperatures above 60°C, an
extrapolation of data found in the literature was made.
The degree of reliability of such an approach, however,
should be verified by ascertaining experimental data at
these temperatures since significant deviations from
extrapolated values may occur. The above equation may be
used to obtain appropriate K! values at various tempera-
tures over the range of 0 to 199°C.
As indicated in Figure 30 the [S=] concentration becomes
significant only at high pH (pH>13) or in strongly basic
solutions. In our case where there is an excess of hydro-
gen sulfide present the pH of the condensate will never
approach high pH values and formation of [S=] may be
neglected. Consequently, the species present in the
condensate will include HS" and H2S. All ammonia will also
be in the form of NHi»HS.
Increasing temperatures will allow higher ratios of ionized
species at lower pH values. This change, however, may be
considered insignificant since a 35°C increase will move the
curves in Figure 30 by only 0.3 pH unit to the range of
lower pH values.
131
-------
CD
f-H
X
c
CO
•4—1
CO
C
o
o
c
o
*-•
03
'o
o
Dissociation Constant for
H2S-H20 System
Kr
0
20 40 60 80
Temperature, (C)
100
120
Figure 31. Effect of temperature upon ionization constant
for H2S
132
-------
In case some hydrogen sulfide is present in the gaseous
phase above the condensate, the concentration of total
sulfur can be determined from Henry's law.
[H2S]g = H CH2S]aq (71)
where H = Henry's law constant for hydrogen sulfide
(atm/mole fraction)
[H2S] = partial pressure of H2S in gaseous phase (atm)
O
fraction of hydrogen sulfide in solution
aq
The values of Henry's law constant for temperatures between
0 and 100°C are presented in Table 30.
The condenser material balance can now be determined by
means of Equation (71) and Equation (72) if an assumption
of steam condensation ratio is made.
V (1-y) [ 2 ]g + Vy [H2S]aq = Vx (72)
where V = volume of steam (moles)
y = fraction of steam condensed
x = mole fraction of H2S in steam entering the
condenser
•n = condenser pressure (atm)
Volume of steam occurs in each term of Equation (72) and
may be cancelled. [H2S] may be expressed from Henry's law,
Equation (71), and we obtain
H |H2S|
m + y [H2S]aq = x (73)
133
-------
TABLE 30. HENRY'S LAW CONSTANT FOR H?S VERSUS TEMPERATURE
Temperature, Henry's Law Constant
C atm/mole fr. HzS in solj
o
0 20,300
10 36,700
20 48,300
30 60,900
40 74,500
50 88,400
60 103,000
70 119,000
80 135,000
90 144,000
100 148,000
aPerry, J.H., Chemical Engineers' Handbook, 4th
New York, McGraw-Hill Book Co., 1963.
blatm = 1.01325xl05 Pa.
134
-------
Knowing the concentration of H2S in the stripper effluent
x, and Henry's law constant at various temperatures and degrees
of steam condensation y at the operating temperature of the
condenser, we can determine condenser pressure from steam
tables and subsequently calculate [H2S] . Substituting
[H2S] in Equation (71) with proper H we can determine
aq
concentration of H2S in vapor phase.
Results obtained for a stream containing 2.0xlO~3 mole
fraction (2000 ppm) H2S at three different temperatures
and various fractions of steam condensation are presented
in Figure 32.
As can be seen in Figure 32, the condensation of stripping
steam containing 2.0xlO~3 mole fraction (2000 ppm) H2S may
produce a large range of concentrations in vapor and liquid
phase depending on the operating conditions maintained in
the condenser. Specifically, 99% condensation at tempera-
tures lower than about 60°C will not result in liquid
concentrations that would exceed 1.0xlO~3kg/m3 (1 ppm).
This is valid, of course, only if no ammonia is present.
Condensation at higher temperatures or higher condensation
ratios would produce solution with H2S concentrations
higher than 1.0xlO~3kg/m3 (1 ppm).
As pointed out earlier, once the solution of H2S is exposed
to the gas phase containing no H2S, the diffusion of H2S
from liquid phase to gaseous phase would occur until a new
equilibrium was reached. If the liquid were exposed to
ambient air the H2S concentration would eventually go to zero.
135
-------
u>
uo
N>
O
O
3
O
0)
O
3
f»
en
ct-
•^
0)
p
10*
X
I 1
o
TO
Q. O
a
>
103
0.001
30°C
=0.99
0.95
=0.9
X=2xlO
-3
0.01
0.1
1.0
10
-------
Analyzing the hydrogen sulfide dissociation constant, Equation
(68) or Figure 30, we see that an increase of hydrogen ion
concentration would shift the equilibrium to non-dissociated
H2S. Since only non-dissociated H2S can diffuse out of
solution this would increase the driving force of the H2S
diffusion to gaseous phase and enhance the rate of H2S depletion,
A pH value of about 5 would convert essentially all dissolved
sulfide species to non-dissociated H2S and result in the
maximum rate of H2S depletion.
Essentially the same principles apply in the presence of
ammonia. Some of the acid, however, will react with the
ammonia and increase the acid consumption.
Applying the above principles, we may conclude that an
increase of hydrogen ion concentration by injection of an
acid into the condenser would further lower the H2S concen-
trations in the condensate and increase H2S vapor contents..
pH values around 5 would cause the H2S to stay in vapor
phase and prevent it from going into the condensate. Thus,
very rich hydrogen sulfide vapor stream and sulfide free
condensate would result.
In the design of the steam condenser a trade-off will have
to be evaluated between efficiency of the H2S recovery by
treating the condensate, or handling acidic streams at
elevated temperatures that may range from 300 to 373K (80-
212°F). In one case, the acid treatment may be done in a
condensate tank. However, the depletion of H2S from the
bulk of the condensate will be diffusion controlled. In
the second case, the condenser interior has to be acid
resistant at condenser operating conditions and no diffusion
factors apply since the H2S will never enter the condensed.
phase.
137
-------
5.3.2 Combined Condenser/Acidifier/Phase Separator Desig£
(Alternative Sour Water Treatment System)
63'1
In the process outlined in the Phase I final report (page
condensation, phase separation, acidification, and clarif**
cation occur in four separate vessels. Retrofit of such a
system to existing FCC units may be difficult because of
severe space limitations which exist at petroleum refiner1
A system which accomplishes all four processing steps
separation of oil from the water in an integrated space-
saving unit is desirable. One approach to design of such
+"1.0
system is proposed and presented in Figure 33. Its a
is described below.
Acidification of the condensate is achieved by injecting
sulfuric acid into the stripper off-gas vapor through a
spray nozzle. Two or more nozzles should be available t°
prevent shut-down precipitated by possible corrosion of *
nozzle. The H2SOit injected into the 811 K (1000°F) stre^
will decompose and form sulfur trioxide. To ensure homo"
geneous distribution of I^SOij and S03 in the vapor and ^e
condensate, the vapor stream should be properly mixed.
motionless mixer is visualized to fulfill this function-
Vapors containing H2S, NH3, S03, hydrocarbons and c.ataiy8
fines are condensed in a vertical floating head heat
exchanger. The E2SOt+ injected into the vapor stream i5
used to control the pH of the condensate.
f
A vapor-liquid separating cone is located in the bottom °
the condenser. Liquids formed in the condenser are sent
a liquid surge drum. The vapors pass through a demist6*"
pad and are sent to an existing refinery Glaus unit f°r
sulfur recovery.
138
-------
Water Inlet
To Slop
Vapor Liquid Separating Cone
Conden sate Outlet
S04 Spray Nozzles
Floating Head Cover
pHC - pH Control
LC - Level Control
TC -Temperature Control
To Water Trsatment Plant
33. Combined condenser, acidifier, phase separator
clarifier, and scum oil removal system
139
-------
The liquid surge drum is used to achieve several process*116
steps simultaneously - oil and water separation, cataly8*
fines removal (clarification), and residual H2S removal-
Water leaving the liquid surge drum is to be sent to the
refinery wastewater treatment for disposal or recycle.
140
-------
6. ECONOMICS
In the Phase I final report we proposed two possible con-
ceptual methods of applying steam stripping to existing
FCC units:
1. An increase of the present steam stripping rates in
existing equipment to the levels needed to achieve an
adequate SO reduction in the generator flue gas.
Jt
2. The use of a secondary open (add-on) stripping system
in which the spent catalyst Is removed from the exist-
ing equipment and transferred to a secondary stripper.
Sulfur free catalyst is then returned to the regenerator.
Both these methods were evaluated in Section 5.3 of the Phase
I final report. The first method was found economically
infeasible due to limited size of existing stripper, reactor,
overhead vapor line and FCC fractlonator, and the inability
of this equipment to handle increased flow rates resulting
from supplemental stripping steam.
The second method was also evaluated in detail and several
options were proposed for SO emission control by steam
X
stripping. Option 1 was believed to be the most unfavorable
economically and was analyzed further in two cases, the worst
and the typical case. Both cases represented the cost of
overall system presented in Figure 13, p. 63, Phase I final
report including the waste water treatment facilities.
11*1
-------
A detailed description and analysis of the Option 1 and
definitions of the worst and typical cases were presented
in Section 5.3.3 and Appendix E of the same report.
In evaluating the worst and the typical cases of Option 1
several assumptions were made and are repeated below:
Worst Case
Catalyst-to-oil ratio 12 6
Catalyst attrition rate,kg/m3 0.57 O-2
Ib/barrel 0.2 0.1
Steam stripping rate,
kg H20/100 kg of catalyst 4 4
The data generated during the laboratory development
program revealed that a variety of steam stripping rates
may be required for different catalysts. Specifically*
rates as high as lOOkg/lOOkg of catalyst were required t
obtain 2xlO~1*mole fraction (200 vppm) of sulfur oxides i
the FCC regenerator off-gas.
Change of steam stripping rate will require a change i*
equipment capacity for each processing step and result
V"*
in different process economics. Consequently, we have
panded the economic evaluations performed in Phase I to
higher stripping rates such as 6, 12, 40, and 100kg of
steam per 100kg of catalyst. All the assumptions used *
preparation of these evaluations were the same as those
applied in the Phase I report, which makes the results
of both reports (Phase I and Phase II) comparable. Als°' j
by using the results from the estimates prepared in Pha&
for the steam stripping rate of 4kg/100kg of catalyst } °
-------
overall range of steam stripping rate is from 4 to 100kg/
100/kg catalyst. Using this range of steam stripping rates
allowed presentation of the capital investment and operating
costs as a function of both steam stripping rate and re-
finery size.
As in Phase I, estimates were prepared for both the typical
and worst cases. The results of these evaluations are
summarized in Table 31 and Figures 34 through 39. Detailed
cost estimates and calculations are presented in Appendix B.
It should be noted that one of the very important assumptions
used in the economic analysis was the steam superficial
velocity in the stripper. As indicated previously [Equation
(3)], the sulfur removal efficiency appears to be a function
of the product steam residence time in the stripper and
steam stripping rate. The steam residence time is inversely
proportional to steam superficial linear velocity which is
a function of steam stripper design. With current stripper
designs, the steam residence time can be varied over a wide
range of values. Specifically, superficial linear velocity
in the stripper can vary from 1.52xlO~27.62m/s (0.05 to 25
ft/s) a factor of 500.
Equation (3) suggests that an increase in steam residence
time in the stripper can substantially reduce steam
stripping rate with no change in sulfur removal efficiency.
This can be easily done by decreasing the stripping steam
superficial velocity.. It will be shown that the reduction
in steam stripping rate will have an influence on the
economics of the steam stripping concept because the steam
superheater capacity will be reduced, the capacities and size
of equipment downstream the stripper will be reduced, and
condensate treatment will be minimized.
-------
-t
-tr
Table 31. CAPITAL AND OPERATING COST SUMMARY FOR STEAM STRIPPING CONCEPT
FCC Unit Capacity
10,000 bpsd 50,000 pbsd
150,000 pbsd
Operating Conditions
Typical stripping
operation
s/cu
s/cb
s/cb
s/cb
4
6
12
40
S/Cb «• 100
Worst stripping
operation
s/c°
s/cb
s/cb
s/cb
s/cb
4
= 6
= 12
= 40
= 100
0.72
0.92
1.47
3-32
6.17
19.82
22.09
29.34
57.00
114.7
2.08
2-73
4.37
9.90
18.51
13-73
14. ?7
20.43
45-20
94.50
4.51
5.76
9.21
19-^7
-•••.45
10.99
12.64
17.84
40.77
87-01
1 bpsd = 1.84xlO~s mVs
3S/C = Steam to catalyst ratio
sl bbl = 0.15899 m3
Capital Investment
Cost
($xlO-*)
2.82
3-59
5.75
13.05
24.46
4.51
5.76
9.21
19.^7
-•••.45
Operating Cost0
U/bbl)
5-96
6.83
9.64
21.63
45-75
10.99
12.64
17.84
40.77
87-01
-------
100
10
tq
"o
O
ons
J
03
c.
0.1
l50,000bpsd
50,000bpsd
lO,OOObpsd
FCC Feed Desulfuriration Cost Range
• 10,000bpsd
. 50,000bpsd
* 150,000bpsd
Typical Case
Worst Case
lbpsd*1.84xlO~6m3/S
' i i i 11
10 100
Stream Stripping Rate (kg H20/100kg Catalyst)
Figure 3^1. FCC catalyst steam stripping, total investment cost
-------
-t
O\
120
100
I/I
a
oJ
CX
O
80
60
40
20
0
FCC Feed Oesulfurization Cost Range
• lO.OOObpsd
. 50,000bpsd
*150,000bpsd
— Typical Case
Worst Case
lbbl=0.15899m3
Ibpsd=1.84xl0"6m3/s
10
20 30 40 50 60
Steam Stripping Ratefog H20/100 kg Catalyst)
70
80
90
100
Figure 35. Summary of operating costs
-------
100
= 10
o
(/I
C
O
CO
-------
1000
100
1*
I/I
o
s10
CO
o>
Q.
O
SSR = Steam Stripping Rate
(kg H20/100kg Catalyst)
Ibb! - 0.15899m3
Ibpsd=1.84xl0-6m3/s
•SSR= 12
SSR= 6
SSR= 4
1 10 100
FCC Capacity (Thousands of Barrels Per Day)
Figure 37- Summary of operating costs, typical case
148
-------
100
10
I/I
1_
ro
~O
O
c
o
o
o
•t-*
c
o
i
0.1
SSR= 12
SSR= 6
SSR= 4
SSR = Steam Stripping Rate
(kg H20/10Qkg Catalyst}
Ibpsd-1.84xl0"6m3/s
J.
10 100
FCC Capacity (Thousands of Barrels Per Day)
Figure 38. Summary of total investment costs, worst case
-------
-O
O
O
on
CD
a.
O
SSR= Steam Stripping Rate
{kg H20/100 kg Catalyst)
100
10
Ibbl = 0.15890m3
SSR=100
SSR=40
SSR=12
SSR=6
SSR=4
10 100
FCC Capacity (Thousands of Barrels Per Day)
1000
Figure 39. Summary of operating costs, worst case
150
-------
In our cost estimate we have assumed a steam superficial
velocity in the stripper of 0.6lm/s (2 ft/s). In the following
paragraphs we will discuss the effect of this velocity on
the overall economics of the steam stripping concept. We
will use an example of decreasing the velocity from 0.61 to
0.30m/s (2 to 1 ft/s). Essentially the same technique can be
used to obtain costs of steam stripping at any velocity.
The reduction of steam superficial velocity from 0.6l to
0.30m/s (2 to 1 ft/s) will decrease the steam stripping
rate by 50$. Since the size of most of the equipment used
in the process is based upon steam capacity (all equipment
except the stripper, which is based on catalyst residence
time), the investment and operating costs will also de-
crease. The investment cost will decrease according to the
0.67 power of steam capacity. The raw materials and utili-
ties costs will be cut in half.
Using the data in Table 31, the investment and operating
costs for 9.20xlO-2m3/s (50,000 barrels/day) FCC unit
capacity with steam stripping rate of 6 and 12kg of steam
per 100kg of catalyst, and other typical stripping con-
ditions, we can summarize
Steam stripping rate, Investment cost, Operating costs,a
kg H?0/100 kg catalyst millions $ (fr/bbl
6 1.71 8.34
12 2.73 11.51
al bbl = 0.15899 m3.
151
-------
Reduction of the superficial linear velocity in the stripper
from 0.61 to 0.30m/s (2 to 1 ft/s) for the steam stripping
rate of 12kg/100kg would decrease the investment and opera-
ting costs to $1,746,200 and 9.54
-------
The economic analyses presented in this report do not include
the additional economic benefits of steam stripping. The
costs include only items Incurred in performing the steam
stripping operation. No by-product credit is taken for
increased sulfur production, Increased hydrocarbon recovery
and heat recovery from stripping steam upon its condensa-
tion. Each of these three factors would further improve the
economics of the steam stripping concept.
153
-------
APPENDIX A
SPENT FCC CATALYST STEAM STRIPPING DATA SHEET
I. EXPERIMENT NUMBER -
II. DATE -
III. EXPERIMENT PERFORMED BY -
IV. SPENT FCC CATALYST IDENTIFICATION
A. COMPANY -
B. REFINERY LOCATION -
C. FCC UNIT -
D. CATALYST TYPE -
E. CATALYST HISTORY -
V. PURPOSE OF THIS EXPERIMENT -
VI. VARIABLES TO BE TESTED -
VII. CATALYST ANALYSIS
A. C = COKE ON SPENT CATALYST
DV,
GSC = % BY WEIGHT
B. SG = SULFUR CONTENT OF COKE
S = % BY WEIGHT
-------
VIII. EXPERIMENTAL DATA
A. CATALYST CHARGING DATA
1. CAT = CATALYST CHARGED TO REACTOR
r
CAT = GRAMS
r
2. S = SULFUR CHARGED TO REACTOR
r
Sr = (CATr) x (Csc) x (SQ) x (0.1)
S = MILLIGRAMS
r
B. STEAM STRIPPING OF SPENT FCC CATALYST
1 T ,_ * CATALYST BED TEMPERATURE
CD
T
cb
2. WATER FLOW RATE SETTING =
3. Wu n = WATER FLOW RATE
MILLILITERS/MIN
"
L T SUPERHEATED STEAM TEMPERATURE
ss
T
S3
5. P = STRIPPER PRESSURE
s
P = _ psig
s -- •
fi D SUPERHEATED STEAM DENSITY
ss
lb/ft3 (PROM STEAM TABLES)
7. V * CALCULATED STRIPPER LINEAR VELOCITY
3 _ i
V = (0,00263) x (WH Q) x (-— )
s "2 pss
ft/sec
8. AP = STRIPPER PRESSURE DROP
s
= _ INCHES OF WATER
s -- " "
q a STATIC CATALYST BED DENSITY
•? * " n oi~
^
pcat
155
-------
10. Fgx = FLUID BED VOLUME EXPANSION FACTOR
p
ex
11. H = (0.158) x (CAT ) x (-—) x Fpy
r pcat ex
Hfc = ft
12. VOLfb = VOLUME OF FLUIDIZED BED
VOLfb = (0.0140) x (Hfc)
VOLfb = ft3
13. VOLcat = VOLUME OF CATALYST IN FLUID BED
VOL , = (0.0022) x (-i—) x (CAT )
cat pcat r
VOLcat ' ft
14. VOL . , = VOLUME OF FLUID BED NOT OCCUPIED BY
void CATALYST
VOLvoid - VOLfb - VOLcat
15. T = STEAM/CATALYST CONTACT TIME
T r^^T
T = 27,240 (Pss) x ( void)
sc * w
WH00
T = SECONDS
so
16. Tg = TOTAL STRIPPING TIME
T = SEC.
s
17. Wu A = ESTIMATED WATER USAGE
WH n = MILLILITERS (GRAMS)
n2u
18. S/C = STEAM TO CATALYST RATIO
156
-------
100 H2°
q/P - ±uu ^
S/C CATr
S/C = lb R?0 / 100 Ib CATALYST
SYSTEM PURGING
1. NITROGEN PLOW RATE SETTING =
2. NITROGEN PLOW RATE = ____l/mln
3. PURGING TIME = min
4. PURGE VOLUME = 1
D. COKE COMBUSTION
1. CATALYST BED TEMPERATURE - SEE ATTACHED SHEET
2. T = PREHEATED AIR TEMPERATURE
a
Ta •
3. COMBUSTION AIR PLOW RATE SETTING =
1|. V = COMBUSTION AIR PLOW RATE
a
V = _ 1/min AT S.T.P.
a --
5. P = COMBUSTION CHAMBER PRESSURE
a
P = _ psig
a - --
6. AP = COMBUSTION CHAMBER PRESSURE DROP
CL
AP = __ _ INCHES OP WATER
a — •
7 p COMBUSTION AIR DENSITY
°a liqp 1^.7 + P
P = <0-0808) x
ea
= lb/ft3
8. Va = CALCULATED COMBUSTION CHAMBER LINEAR VELOCITY
460 + T
v = (0.00125) x (-^ ? + p ) x (V )
a
Va
ft/sec
157
-------
9. P = FLUID BED EXPANSION FACTOR
10. Hfc = DEPTH OF FLUIDIZED BED
Hfc - - -ft
11. VOL., = VOLUME OF FLUIDIZED BED
ID
VOLfb = _ ft3
12. VOL . = VOLUME OF CATALYST IN FLUID BED
C3.T/
VOLoat = - ft 3
13. VOL .„ = VOLUME OF FLUID BED NOT OCCUPIED BY
CATALYST
VOLvo±a = - ft 3
14. T = AIR/CATALYST CONTACT TIME
T = (56.871) x (VOLvold) x <"•? + V
cLC ' • " --.'--•-. ,
(Vn) (460 + TQ)
ct d
T = SECONDS
ac
15. T = TOTAL COMBUSTION TIME
3.
T = MIN.
3, "'" '""'<: ••-.--•-
16. V ~ VOLUME OF COMBUSTION AIR
C 3.
Vca - \ x Ta
vca - liters
E. SYSTEM PURGING
1. NITROGEN FLOW RATE SETTING =
2. NITROGEN FLOW RATE = 1/min (S.T.P.)
3. PURGING TIME = min
4. PURGE VOLUME = liters
S.T.P. = 32°F; 29.92 in Hg.
158
-------
IX. ANALYTICAL DATA
A. H2S ANALYSIS
1. Vt = VOLUME OF STANDARDS SODIUM THIOSULFATE TITRANT
V = mJL FOR BLANK
V = ml FOR SAMPLE
ts
2. N = NORMALITY OF STANDARD SODIUM THIOSULFATE
^ TITRANT
N, = g-eq/llter
3. VSOLN = TOTAL VOLUME OF SAMPLE + IODINE + ACID
VSOLN = Si
Jj V = VOLUME OF ALIQUOT TITRATED
a
V = ml
a •
5. V o ~ WEIGHT OF SULFUR COLLECTED AS H S
WH2S " t t BLANK t t
W . = MILLIGRAMS SULFUR
B. S02/S03 ANALYSIS
STTT.FUR COLLECTED AS ACID MIST AND SULFUR TRinYTTW
1 v = VOLUME OF BARIUM PERCHLORATE TITRANT USED
fc FOR SAMPLE
v - - ml
2 v VOLUME OF BARIUM PERCHLORATE TITRANT USED
tb FOR BLANK
Vtb
3. N - NORMALITY OF BARIUM PERCHLORATE
N - g - eq / liter
159
-------
4. V TOTAL SOLUTION OF SULFURIC ACID, (FIRST
soin + FILTER)
Vsoln - - ml
5. V = VOLUME OF SAMPLE ALIQUOT TITRATED
3.
va = _ ml
6. W = WEIGHT OF SULFUR COLLECTED AS ACID
SULPUR TRIOXIDE
,so, -^ v
c H $ a
WTJ en en = MILLIGRAMS SULFUR
xlpOUhjoU-
SULFUR COLLECTED AS SULFUR DIOXIDE
1. V, = VOLUME OF BARIUM PERCHLORATE TITRANT USED
r FOR SAMPLE
2. Vt = ml
2. V,. = VOLUME OF BARIUM PERCHLORATE TITRANT USED
FOR BLANK
Vtb = ml
3. N = NORMALITY OF BARIUM PERCHLORATE TITRANT
N = g - eq / liter
^' V^n1n = TOTAL SOLUTION VOLUME OF SULFUR DIOXIDE
(SECOND AND THIRD IMPINGERS)
soln ~
5. V = VOLUME OF SAMPLE ALIQUOT TITRATED
a
va =
6. WSQ = WEIGHT OF SULFUR COLLECTED AS SULFUR
2 . 16 (Vt-Vtb
2
MILLIGRAMS SULFUR
160
-------
SULFUR COLLECTED AS SULFUR DIOXIDE
1. V. = VOLUME OF BARIUM PERCHLORATE TITRANT
C FOR SAMPLE
V. = _ ml
Lt -""
2. V,K = VOLUME OP BARIUM PERCHLORATE TITRAWT
tb FOR BLANK ^HAWI
Vtb = - ml
3. N = NORMALITY OF BARIUM PERCHLORATE TITRANT
M = _ g - eq / liter
*• Vsoln = TOTAL SOLUTION VOLUME OF SULFUR
soln (SECOND AND THIRD IMPINGERS)
V n = _ ml
soln -
5. V = VOLUME OF SAMPLE ALIQUOT TITRATED
3l
v = mi
a -
6. WSQ = WEIGHT OF SULFUR COLLECTED AS SULFUR DIOXIDE
so2
W = __ MILLIGRAMS SULFUR
ibUo ~~~~~~~~ ~^~~~~ ~~
X. SULFUR BALANCE AROUND STRIPPER
A S = SULFUR CHARGED TO REACTOR
r
s = _ MILLIGRAMS SULFUR
r -- •
B. W = WEIGHT OF SULFUR COLLECTED AS H S
HpS • ^
w « MILLIGRAMS SULFUR
H S _ - ' --
+ Wso = FINAL WEIGHT OF SULFUR IN COKE
qo
* 3
3
Wu Qn qn + Wso = _ MILLIGRAMS SULFUR
W ' 2
161
-------
D. $WCT = PERCENT OP FEED SULFUR ACCOUNTED FOR
s
XI. H20 ANALYSIS - (TO BE DONE DURING STEAM STRIPPING)
A. TOTAL H20 VOLUME IN IMPINGERS
VT = FINAL VOLUME = _ ml (GRAMS)
1f
VT = INITIAL VOLUME = ml (GRAMS)
AV± = VOLUME OF WATER COLLECTED IN IMPINGERS
AV, = VT - VT = _ ml (GRAMS)
i if 11
B. WEIGHT OF H20 COLLECTED BY SILICA GEL
V = FINAL WEIGHT = _ GRAMS
Vsgi = INITIAL WEIGHT = _ GRAMS
AV = WEIGHT OF WATER COLLECTED BY SILICA GEL
AVsg - Vsgf - Vsgi - - : - GRAMS
C. Wu n = TOTAL WEIGHT OF H00 COLLECTED
- AV. + AV
i sg
grams water
XII. RESULTS AND CONCLUSIONS
A. CATALYST CHARGED TO REACTOR = _ GRAMS
B. WATER USED IN STRIPPING = GRAMS
C. STEAM TO CATALYST RATIO = S/C
GRAMS H20
100 GRAMS CATALYST
162
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D. SULFUR CONTENT OF COKE INITIALLY = % BY WEIGHT
E. SULFUR CONTENT OF COKE AFTER STRIPPING = % By
WEIGHT ~
F. PERCENT SULFUR CONTENT OF COKE REDUCTION = _^^^ %
G. DISCUSSION -
163
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APPENDIX B
DETAILED COST ESTIMATES OF THE STEAM STRIPPING PROCESS
A detailed capital cost estimate was prepared for the typical
case at a steam stripping rate of 6 kg HpO/100 kg catalyst and
50,000 bpsd nominal size FCC unit. This estimate was obtained
by cost estimating the major equipment necessary for the process,
designated as purchased equipment cost in the following section.
Fixed capital investment cost has been calculated by applying a
fixed capital investment factor* of 4.8 to the purchased equip-
ment cost (see Table B-13).
The capital investment cost for the worst case using the same
size FCC unit (50,000 bpsd) has been calculated by assuming a
scaling factor of 0.67 and applying it to the fixed capital in-
vestment for the typical case. The cost for the worst case is
summarized in Table B-43. Capital investment cost estimates
were made for both the typical and the worst case, and for 10,000
and 150,000 bpsd FCC unit nominal sizes at steam stripping rates
of 6, 12, 40, and 100 kg H20/100 kg catalyst. The estimates
were obtained by applying the scaling factor of 0.67 to the cor-
responding fixed capital investment cost figures determined for
the 50,000 bpsd unit. The operating costs were prepared on an
individual basis for each FCC unit size and steam stripping rate.
1 bpsd = 1.84 x 10~6 mVs.
*Peters, M. S., and K. D. Timmerhaus. Plant Design and Economics
for Chemical Engineers, 2nd Edition. New York, McGraw-Hill
Co., 1968. 850 pp.
-------
For convenience, we have repeated the list of assumptions that
were applied in the economic analyses in both the Phase I and
Phase II evaluations. Detailed capital and operating cost esti-
mates are presented in tabular form in the order of increasing
PCC unit capacity and increasing steam stripping rates in Tables
B-l through B-60.
Since the original estimate in Phase I was made for a 45,000 bpsd
catcracker we have revised it for the catcracker with 50,000 bpsd
capacity.
B-l. ASSUMPTIONS USED FOR CAPITAL INVESTMENT COST ESTIMATE AND
EQUIPMENT SIZE CALCULATIONS
. Catalyst-to-oil ratio (C/0) is 6 kg of catalyst per
kg of total feed for the typical case, and 12 kg of
catalyst per kg of total feed for the worst case
. Catalyst attrition rate is 3.33 x 10~" kg per kg
(0.1 Ib of catalyst per barrel) of feed for the typical
case and 6.66 x 10"* kg per kg (0.2 Ib of catalyst per
barrel)of feed for the worst case
. Steam line pressure, 9.63 x 105 Pa (125 psig)
Sulfur content of coke before steam stripping, 1.5
wt %
. Sulfur content of coke after steam stripping, 0.2^3
wt %
. Stripper operating temperature, 8ll K (1000°F)
1 barrel = 0.15899 m3.
1 bpsd = 1.84 x 10"6 nr.
165
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Stripper operating pressure, 3.^3 x 10"5 Pa (35 psig)
Velocity in stripper feed transfer lines, 12.2 m/s
(40 ft/s)
Vapor velocity in lines leaving the stripper, 30.5 m/s
(100 ft/s)
Velocity in stripper standpipe, 2.1 m/s (7 ft/s)
• Stripper bed density, 240 kg/m3 (15 Ib/cu ft)
• Catalyst bulk density, 801 kg/m3 (50 Ib/cu ft)
Catalyst density in the standpipe, 561 kg/m3 (35 lb/
cu ft)
Hydrogen sulfide produced in the steam stripper will
be fed into existing Glaus unit and no additional cost
was assumed to be needed for expansion of this facility
• Velocity in stripper, 0.61 m/s (2 ft/s)
Depth of fluidized bed in the stripper was assumed to
be 3.05 m (10 ft) with the fluid bed occupying 50$
of the total stripper volume
- Weight of FCC feed, 136.1 k ©/barrel (300 Ib/barrel)
Fixed capital investment factor = 4.8
• Start-up cost - 10% F.C.I. *
1 barrel = 0.15899 m3.
166
-------
• Working capital - 10.5% P.C.I. *
Interest on construction loan - construction period of
12 months; financed fixed capital at the rate of 8$/yr
for average of half of construction period assumed
• Does not include sulfur recovery plant capital cost
• Base period - February 1973
• Scaling factor - 0.67
* CE plant cost Index
1968 113.7
1969 H9.0
1970 125.7
1971 132.3
1972 137.2
Feb. 1973 1^0.4
• Other assumptions used will be presented at the time
of their use
* Fixed Capital Investment
167
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B-2. DETERMINATION OF PURCHASED EQUIPMENT COST FOR 50,000 BPSD
FCC UNIT, TYPICAL CASE
a. Catalyst Stripper
Catalyst Circulation Rate (CCR)
(catalyst to oil ratio) x (300 Ib/bbl of oil)
OCR Ob/hrO - _ x (bbl/day of feed oil)
CCR (Ib/hr; -- (24 hr/day)
CCR = 75 x (50,000) = 3.750 x 106 Ib/hr
4.725 x 102 kg/s
Steam Stripping Rate (SR)
SR (Ib/hr) = (Ib of steam per Ib of catalyst) x CCR
SR = 0.06 x 3.75 x 106 = 225,000 Ib/hr
28.3 kg/s
Volumetric Flow of Steam (VF)
SR x 359 cu ft/lb mole x (1000+460) x 1M.7
4°u ~50~
VTP = _ „ - J - 7 -
v* Ib lb/lb mole x 3600
VF = 5.17 x 103 x 225,000 = 1163 acfs
32.9 m3/s
Cross-Sectional Area (AL ) and Diameter (DL) of Feed
Transfer Lines
AT = linA, = °'025 x ll63 = 29-1 S9 ft
L ^O ft/s ^
D_ = / 4AT = 1.128 x /AT = 6.08 ft
L F"L L 1.85 m
1 barrel = 0.15899 m3.
168
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Cross-Sectional Area (Ag) and Diameter (D£) of the Stripper
ft
Dc =* 1.128 x /T7~ = 27.2 ft
s S 8.29 m
Cross-Sectional Area (A£) and Diameter (DE of Lines Leaving
the Stripper
VP
= 0.01 x 1163 = 11.63 sq ft
AE = 100 ft/s "•- --•-.- lt08V
D,, = 1.128 x / Av = 3.85 ft
E £ 1.17 m
Cross-Sectional Area (Ap) and Diameter (Dp) of Stripper
Standpipe
CCR
Ap (sq ft) - 3600 x (bed density) x (velocity)
^.750 x 106 = z, 25 sa ft
P = 3600 x 35 x Y 0;39 m*
D0 = 1.128 x /TIT = 2.33 ft
P p 0.71 m
Catalyst Inventory (CI) in the Stripper
CI (Ib) = A x 10 ft x (stripper bed density)
01 - o'oTS x 582 = 87,300 g
The cost of the steam stripper was estimated based on weight of
this equipment. The unit price for 5% Or, 1/2* Mo steel, which
was assumed to suit this application, was estimated at $l.i8/kg
(53.6
-------
b. Condenser
The area of the condenser was calculated according to the
following assumptions.
Stripping steam will be cooled from 811 K (1000°F) to
311 K (100°F) and condensed
The cooling water at 300 K (80°F) will enter the con-
denser and will be evaporated to produce saturated
steam at 452 K (353°F)
The overall heat transfer coefficient was assumed to
be 3973 W/m2 K (700 Btu/ft2 hr °F)
The cost of heat exchanger was assumed at $7.75/sq ft
(5% Cr, 1/2% Mo steel)
Calculation of Heat Transfer Area
Superheated Steam
2.048 x 10s Pa (15 psig)
811 K (1000°F)
3.56 x 106 J/kg
(1533 Btu/lb)
Saturated Steam
9.632 x 10s Pa
(125 psig)
452 K (353°F)
(1193 Btu/lb)
HX
Water
311 K (100°F)
1.58 x 105 J/kg
(68 Btu/lb)
•Process Water
300 K (80°F)
1.12 x 105 J/kg
Using the heat transfer equation
Q = UA AT
m
170
-------
Where AT = (1000-353) - (100-80) =
Where Alm - 1000-353
ln 100-80 ^55 K
Q = 225,000 x (153-68) = 329.6 x 106 Btu/hr
9.657 x 107 W
Q _ 329.6 x 106 f. f.
A = UKT = 700 x 180 " 2616 sq ft
^
Cost $ = 7.75 x 2616 = $20,275
329 6 x 106
Water requirement = ^Toi - ITS — = 2°?p9 x 10 3 Ib/hr
1193 - 4» = 0>363 kg/s
= 575 gpm
The unit will produce 7. 94 kg/s (63,000 Ib/hr) of 9.63 x
10s Pa (125 psig) steam
The pump delivering this amount of water through the
condenser and superheater operating at 9.63 x 10s Pa
(125 psig) will have 37,300 W (50 HP)
c. Steam Superheater
It was assumed that the saturated steam from the condenser will
be used as the feed for the superheater. The cost of the 2.63 x
10 7 W (90 x 106 Btu/hr) superheater was estimated at $200,000.*
Natural gas was considered as the fuel.
Heat input required for the superheater was calculated as follows
o = 225 000 Ib/hr x (1533 - 1193) = 76.5 x 106 Btu/hr
* 2.24 x 107 W
*Private Communication with Struther-Wells Corporation
171
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Superheater scale down:
$ = 200,000 ( gQ5) = $179,400
The amount of natural gas was approximated based on 155? heat loss
and 3.722 x 107 J/m3 (1000 Btu/cu ft) natural gas heating value:
Natural gas = ltl5 x 7?QQ0X 10—X-^- = 2.11 x 106 cu ft/day
0.692 m3/s
d. Phase Separator
A 2.83 m3 (750 gallon) stirred tank was assumed to be used for
phase separation. The tank is made of 5% Cr, 1/2% Mo steel. The
price of this tank was estimated at $6,600.
e. Acidifier
An epoxy-resin-lined, carbon steel, stirred, 2.83 m3 (750 gallon)
tank was assumed to suit this application. The cost of this tan^
was assumed to be $4,400. Acid sludge at 9.26 x 10"1* kg/s
(7.35 Ib/hr) is required to acidify the contents of the acidifiel>
to pH 3- This amount was calculated based on the assumption
the sludge will contain 90$ sulfuric acid.
f. Neutralizer
A 2.83 m3 (750 gallon) carbon steel, stirred tank was assumed
to suit this purpose. The cost of this tank was estimated to
be $2,500. The amount of lime needed to neutralize the acid
sludge was calculated to be 6.3 x 10"1* kg/s (5 Ib/hr).
g. Clarifler
The mass flow rate through the clarifier was assumed to be
172
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9.51 m3/m2 s (84 gal/sq ft hr). The area of the clarifier was
determined as follows:
«.«-^,wwv, Ib/hr _
«.3 Ib/gallon x 84 gallon/sq ft hr ~ *:£ s? ft
30 m
From this the diameter of the tank was calculated to be 6.19 m
(20.3 ft), or %20 ft. The depth of the clarifier was assumed to
be 3.05 m (10 ft). The cost of this vessel was assumed to be
$5,800.
h. Vacuum Filter
It was assumed that catalyst fines can be filtered by a vacuum
filter operating at the load of 2.03 x 10"2 kg cake/m2 s (15 lb
cake/sq ft hr) with 70$ moisture in the cake.
Assuming that 0.045 kg (0.1 lb) of catalyst per barrel of oil
feed will be carried out from the steam stripper, the weight of
filter cake and area of filter can be determined as follows:
Q'l x 50>000 =700 Ib/hour
d* x U>:S 8.82 x 10~2 kg/s
and - = 46 sq ft
15 4.27 m2
The cost of this equipment was assumed to be $14,500
B-3. ASSUMPTIONS USED FOR OPERATING COST ESTIMATES
The operating cost estimates were individually calculated for
each of the process operating conditions mentioned previously,
The assumptions used to arrive at the final operating cost
estimates are outlined as follows.
barrel = 0.15899 m3.
173
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Catalyst loss was assumed at 3-33 x 10 ** kg/kg (0.1 lb
of catalyst per barrel) of oil feed for the typical
case and 6.66 x 10"1* kg/kg (0.2 lb of catalyst per
barrel) of oil feed for the worst case
The cost of catalyst was assumed at $600/ton
Operating labor - 2 men per shift
The cost of raw materials and utilities was assumed
to change proportionally with the size of the FCC
unit
The cost of sulfuric acid was assumed at $^0.8/ton
(as 100$ H
The cost of lime was assumed to be $19.50/ton
Labor - $5.50/manhour
. Maintenance labor - 2% F.C.I.*
Maintenance materials - 2% F.C.I.
. Process water - 7-93^/m3 (30$/1000 gal)
Plant overhead - 80% total labor
. Taxes & insurance - 2% F.C.I.
. G&A, sales, research - 6% F.C.I.
*F.C.I. = Fixed Capital Investment
1 barrel = 0.15899 m3.
-------
Depreciation - 10% F.C.I.
Interest on working capital - 6% working capital
Return on investment - 20%
. Value of steam - 0.047
-------
Table B-l. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 10,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 4 kg H00/100 kg Catalyst
Fixed capital investment $360,100
Initial catalyst cost 4,400
Start-up cost 36,000
Working capital 37,800
Interest on construction loan 14,400
Total investment $452,700
1 bpsd = 1.84 x 106 mVs
176
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Table B-2. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 10,000 pbsd at 90% Capacity
Worst Stripping Operation
Fixed Capital Investment: $360,000
Steam Stripping Rate: 4 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 7,200
3 Control laboratory 19.300
4 Total labor 122,900
Materials
5 Raw and process - acid sludge 200
6 lime 100
7 catalyst replacement 99,000
8 Maintenance 7,200
9 Operating 9,600
10 Total materials 116,100
Utilities
11 Process water 11,000
12 Electricity 700
13 Fuel 31,700
14 Total utilities 43,000
15 Total direct operating costs (4, 10 & 14) $282,400
Indirect Operating Costs
16 Plant overhead 98,300
17 Taxes and insurance 7,200
18 Plant cost (15, 16 & 17) 387,900
19 General & administrative, sales, research 21,600
20 Cash expenditures (18 & 19) 409,500
21 Depreciation 36,000
22 Interest on working capital 2,300
23 Total operating costs* (20, 21 & 22) $447,800
24 Cost (cents/bbl) 13.84
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~6 mVs.
177
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Table B-3. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 10,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 6 kg H90/100 kg Catalyst
Fixed capital investment $459,900
Initial catalyst cost 5,200
Start-up cost 45,900
Working capital 48,300
Interest on construction loan 18,400
Total investment $577,700
1 bpsd = 1.84 x 106 mVs.
178
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Table B-4. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 10,000 bpsd at 90% Capacity
Worst Stripping Operation
Fixed Capital Investment: $459,900
Steam Stripping Rate: 6 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 9,200
3 Control laboratory 19,300
4 Total labor 124,900
Materials
5 Raw and process - acid sludge 300
6 lime 100
7 catalyst replacement 97,200
8 Maintenance 9,200
9 Operating 9,600
10 Total materials 116,400
Utilities
11 Process water 16,100
12 Electricity 900
13 Fuel 46,500
14 Total utilities 63,500
15 Total direct operating costs (4, 10 & 14) $304,800
Indirect Operating Costs
16 Plant overhead 98,300
17 Taxes and insurance 9,200
18 riant cost (15, 16 & 17) 413,900
19 General & administrative, sales, research 27.600
I® Cash expenditures (18 & 19) 441,500
21 Depreciation 46,000
22 Interest on working capital 2,900
23 Total operating costs* (20, 21 & 22) 490,400
^ Cost (cents/bbl) 15.14
Dopn not include by-product credit or recovery costs.
~ 1.84 x 10~6 mVs.
179
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Table B-5. SUMMARY OF CAPITAL INVESTMENT COSTS
PCC Unit Size: 10,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 12 kg H20/100 kg Catalyst
Fixed capital investment $731,700
Initial catalyst cost 10,500
Start-up cost 73,200
Working capital 76,800
Interest on construction loan 29,300
Total investment $921,500
1 bpsd = 1.84 x 106 mVs.
180
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Table B-6, SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 10,000 bpsd at 9Q% Capacity
Typical Stripping Operation
Fixed Capital Investment: $731,700
Steam Stripping Rate: 12 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 14,600
3 Control laboratory 19,300
4 Total labor 130,300
Materials
5 Raw and process - acid sludge 600
6 lime 200
7 catalyst replacement 97,200
8 Maintenance 14,600
9 Operating 7,700
10 Total materials 120,300
Utilities
H Process water 32,200
12 Electricity 1,900
13 Fuel 93,000
I** Total utilities
15 Total direct operating costs (4, 10 & 14) $377,700
Indirect Operating Costs
16 Plant overhead 104,300
1? Taxes and insurance 14,600
IU Plant cost (15, 16 & 17) 496,600
19 General & administrative, sales, research 43,900
^° Cash expenditures (18 & 19) 5^0,500
jl Depreciation ' 73,200
^2 Interest on working capital 4,600
23 Total operating costs* (20, 21 & 22) $618,300
^ Cost (cents/bbl) 19.08
not
include by-product credit or recovery costs
1.84 x 10~6 m3/s.
181
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Table B-7. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 10,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 40 kg H^O/lOO kg Catalyst
Fixed capital investment $1,639,000
Initial catalyst cost 35,000
Start-up cost 163,900
Working capital 172,100
Interest on construction loan 65,600
Total investment $2,075,600
1 bpsd = 1.84 x 106 mVs.
182
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Table B-8. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 10,000 bpsd at 90% Capacity
Typical Stripping Operation
Fixed Capital Investment: $1,639,000
Steam Stripping Rate: 40 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 32,800
3 Control laboratory 19,300
4 Total labor 148,500
Materials
5 Raw and process - acid sludge 2,100
6 lime 800
7 catalyst replacement 97j200
8 Maintenance 32,800
9 Operating 9,600
10 Total materials 142,500
Utilities
ll Process water 119,100
12 Electricity ' 6,400
13 Fuel 309>500
14 Total utilities 435,000
15 Total direct operating costs (4, 10 & 14) $ 726,000
Indirect Operating Costs
16 Plant overhead 118,800
17 Taxes and insurance 32,800
18 Plant cost (15, 16 & 17) 877,600
19 General & administrative, sales, research 98.300
20 cash expenditures (18 & 19) 975,000
21 Depreciation 163,900
22 interest on working capital 10,300
23 Total operating costs* (20, 21 & 22) $1,150,100
^ Cost (cents/bbl) 35.50
>oes not include by-product credit or recovery costs
bpsd = 1.84 x 10~6 mVs.
3psd"="l?8i x 10~"6 mVs.
183
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Table B-9. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 10,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 100 kg HpO/100 kg Catalyst
Fixed capital investment $3,029,000
Initial catalyst cost 87,600
Start-up cost 302,900
Working capital 318,000
Interest on construction loan 121,200
Total investment $3,858,700
1 bpsd = 1.84 x 106 raVs.
-------
Table B-10. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 10,000 bpsd at 90% Capacity
Typical Stripping Operation
Fixed Capital Investment: $3,029,000
Steam Stripping Rate: 100 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 60,600
3 Control laboratory 19,300
4 Total labor 176,300
Materials
5 Raw and process - acid sludge 5,600
6 lime 1,900
7 catalyst replacement 97,200
8 Maintenance 60,600
9 Operating 9,600
10 Total materials 174,900
Utilities
11 Process water 268,000
12 Electricity 16,000
13 Fuel 774,400
14 Total utilities 1,058,500
15 Total direct operating costs (4, 10 & 14) $1,409,700
Indirect Operating Costs
16 Plant overhead 141,000
17 Taxes and insurance 60,000
18 Plant cost (15, 16 & 17) 1,611,300
19 General & administrative, sales, research 181,700
20 Cash expenditures (18 & 19) 1,793,000
21 Depreciation 302,900
22 Interest on working capital 19,100
23 Total operating costs* (20, 21 & 22) $2,115,000
24 Cost (cents/bbl) 65.28
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~6 mVs.
185
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Table B-ll. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 50,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 4 kg H20/100 kg Catalyst
Fixed capital investment $1,030,400
Initial catalyst cost 17,500
Start-up cost 103,000
Working capital 108,200
Interest on construction loan 41
Total investment $1,300,300
1 bpsd = 1.84 x 106 mVs.
186
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Table B-12. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 50,000 bpsd at 90% Capacity
Typical Stripping Operation
Fixed Capital Investment: $1,030,400
Steam Stripping Rate: 4 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 20,600
3 Control laboratory 19*300
4 Total labor 136,300
Materials
5 Raw and process - acid sludge 1,100
6 lime 400
7 catalyst replacement 486,000
8 Maintenance 20,600
9 Operating _ 9,600
10 Total materials 517,700
Utilities
11 Process water
n
14 Total utilities 211,800
15 Total direct operating costs (4, 10 & 14) $ 865,800
Indirect Operating Costs
16 Plant overhead
X7 Taxes and insurance
}8 riant cost (15, 16 & 17)
19 General & administrative, sales, research
20 Cash expenditures (18 & 19) 1,057,200
21 Depreciation
Interest on working capital
23 Total operating costs* (20, 21 & 22) $1,166,700
^ Cost (cents/bbl) 7.20
uoes not include by-product credit or recovery costs
1 bpsd = 1.84 x 10"6 m3/s.
187
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1 bpsd = 1.84 x 106 m3/s.
Table B-13. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 50,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 6 kg H20/100 kg Catalyst
A. Catalyst stripper $ 48,250
B. Condenser 20,270
C. Steam superheater 179,400
D. Phase separator 6,600
E. Acidifier 4,400
F. Neutralizer 2,500
G. Clarifier 5,800
H. Vacuum filter l4,50Q
Total purchased equipment costs $ 281,720
Fixed capital investment $1,352,000
Initial catalyst cost 26,200
Start-up cost 135,200
Working capital 142,000
Interest on construction loan 54,10J3
Total investment $1,709,500
188
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Table B-14. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 50,000 bpsd at 90% Capacity
Typical Stripping Operation
Fixed Capital Investment: $1,352,000
Steam Stripping Rate: 6 kg H~20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 27,000
3 Control laboratory 19,300
4 Total labor 1*12,700
Materials
5 Raw and process - acid sludge 1,600
6 lime 600
7 catalyst replacement 486,000
8 Maintenance 27,000
9 Operating 9,600
10 Total materials 524,800
Utilities
ll Process water 80,600
12 Electricity 4,800
13 Fuel 232,400
14 Total utilities 317,800
15 Total direct operating costs (4, 10 & 14) $ 985,300
Indirect Operating Costs
16 Plant overhead 114,200
17 Taxes and insurance 27,000
18 Plant cost (15, 16 & 17) 1,126,500
19 General & administrative, sales, research 81,100
20 Cash expenditures (18 & 19) 1,207,600
21 Depreciation ' 135,200
22 Interest on working capital 8,500
23 Total operating costs* (20, 21 & 22) $1,351,300
^ Cost (cents/bbl) 8.34
oes not include by-product credit or recovery costs.
bpsd = 1.84 x 10~6 m3/s.
189
-------
Table B-15. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 50,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 12 kg H20/100 kg Catalyst
Fixed capital investment $2,151,000
Initial catalyst cost 52,400
Start-up cost 215,100
Working capital 225,900
Interest on construction loan 86,000
Total investment $2,730,400
1 bpsd = 1.84 x 106 m3/s.
190
-------
Table B-l6. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 50,000 bpsd at 90% Capacity
Typical Stripping Operation
Fixed Capital Investment: $2,151,000
Steam Stripping Rate: 12 kg HpO/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 43,000
3 Control laboratory 19,300
4 Total labor 158,700
Materials
5 Raw and process - acid sludge 3,200
6 lime 1,100
7 catalyst replacement 486,000
8 Maintenance 43,000
9 Operating 9,600
10 Total materials 542,940
Utilities
H Process water 161,200
12 Electricity 9,600
13 Fuel 464,900
14 Total utilities 635,700
15 Total direct operating costs (4, 10 & 14) $1,337,300
Indirect Operating Costs
16 Plant overhead 127,000
17 Taxes and insurance 43,000
18 Plant cost (15, 16 & 17) $1,507,300
19 General & administrative, sales, research 129,100
20 cash expenditures (18 & 19) 1,636,400
21 Depreciation 215,500
22 Interest on working capital 13,600
23 Total operating costs* (20, 21 & 22) $1,865,100
24 cost (cents/bbl) 11.51
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~* mVs.
191
-------
Table B-17. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 50,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 40 kg H20/100 kg Catalyst
Fixed capital investment $4,819,000
Initial catalyst cost 174,600
Start-up cost 481,900
Working capital 506,000
Interest on construction loan 192,800
Total investment $6,174,300
1 bpsd = 1.84 x 106 m3/s.
192
-------
Table B-18. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 50,000 bpsd at 905? Capacity
Typical Stripping Operation
Fixed Capital Investment: $4,819,000
Steam Stripping Rate: 40 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 96,400
3 Control laboratory 19,300
4 Total labor
Materials
5 Raw and process - acid sludge
6 lime
7 catalyst replacement
8 Maintenance
9 Operating
10 Total materials
Utilities
11 Process water
12 Electricity
13 Fuel
14 Total utilities
15 Total direct operating costs (4, 10 & 14)
Indirect Operating Costs
16 Plant overhead
17 Taxes and insurance
18 Plant cost (15, 16 & 17)
19 General & administrative, sales, research
20 Cash expenditures (18 & 19)
21 Depreciation
22 Interest on working capital
23 Total operating costs* (20, 21 & 22)
24 Cost (cents/bbl)
*Does not include by-product credit or recovery
1 bpsd = 1.84 x 10-1& m3/s.
212,100
10,600
3,800
486,000
96,400
9,600
606,400
537,500
32,100
1,553,300
2,122,900
$2,941,400
169,700
96,400
3,207,500
289,100
3,496,600
481,900
30,400
$4,008,900
24.75
costs.
193
-------
Table B-19. SUMMARY OP CAPITAL INVESTMENT COSTS
FCC Unit Size: 50,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 100 kg H20/100 kg Catalyst
Fixed capital investment $8,905,000
Initial catalyst cost 436,800
Start-up cost 890,500
Working capital 935,000
Interest on construction loan 356,200
Total investment $11,523,500
1 bpsd = 1.84 x 106 m'/s.
-------
Table B-20. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 50,000 bpsd at 90$ Capacity
Typical Stripping Operation
Fixed Capital Investment: $8,905,000
Steam Stripping Rate: 100 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 178,100
3 Control laboratory 19,300
4 Total labor 293,800
Materials
5 Raw and process - acid sludge 26,300
6 lime 9,500
7 catalyst replacement 486,000
8 Maintenance 178,100
9 Operating 9,600
10 Total materials 709,500
Utilities
11 Process water 1,343,700
12 Electricity 80,300
13 Fuel 3,888,600
14 Total utilities 5,312,600
15 Total direct operating costs (4, 10 & 14) $6,315,900
Indirect Operating Costs
16 Plant overhead 235,000
17 Taxes and insurance 178,100
18 Plant cost (15, 16 & 17) 6,729,000
19 General & administrative, sales, research 534,300
20 Cash expenditures (18 & 19) 7,263,300
21 Depreciation 890,500
22 Interest on working capital 56,100
23 Total operating costs* (20, 21 & 22) $8,209,900
24 Cost (cents/bbl) 50.68
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~* m3/s.
195
-------
Table B-21. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 150,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 4 kg H20/100 Ib Catalyst
Fixed capital investment $2,209,000
Initial catalyst cost 65,300
Start-up cost 221,000
Working capital 232,000
Interest on construction loan 88,400
Total investment $2,816,600
1 bpsd = 1.84 x 106 m3/s.
196
-------
Table B-22. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 150,000 bpsd at 90% Capacity
Typical Stripping Operation
Fixed Capital Investment: $2,209,900
Steam Stripping Rate: 4 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 44,200
3 Control laboratory 19,300
4 Total labor 159,900
Materials
5 Raw and process - acid sludge 3,300
6 lime 1,300
7 catalyst replacement 1,485,000
8 Maintenance 44,200
9 Operating 9,600
10 Total materials 1,543,400
Utilities
ll Process water 164,300
12 Electricity 10,700
13 Fuel 475,000
14 Total utilities 650,000
15 Total direct operating costs (4, 10 & 14) $2,353,300
Indirect Operating Costs
16 Plant overhead 127,900
17 Taxes and insurance 44,200
IS Plant cost (15, 16 & 17) 2,525,400
19 General & administrative, sales, research 132,600
2Q Cash expenditures (18 & 19) 2,658,000
21 Depreciation 221,000
22 Interest on working capital 13,900
23 Total operating costs* (20, 21 & 22) $2,892,900
^ Cost (cents/bbl) 5-96
*E>oes not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~* raVs.
197
-------
Table B-23. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 150,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 6 kg H^O/lOO kg Catalyst
Fixed capital investment $2,822,600
Initial catalyst cost 78,700
Start-up cost 282,300
Working capital 296,400
Interest on construction loan 112,900
Total investment $3,592,900
1 bpsd = 1.84 x 106 m3/s.
198
-------
Table B-24. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 150,000 bpsd at 9Q% Capacity
Typical Stripping Operation
Fixed Capital Investment: $2,822,600
Steam Stripping Rate: 6 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 56,400
3 Control laboratory 19,300
4 Total labor 172,100
Materials
5 Raw and process - acid sludge 4,700
6 lime 1,700
7 catalyst replacement 1,458,000
8 Maintenance 56,500
9 Operating 9,600
10 Total materials 1,530,500
Utilities
11 Process water 241,900
12 Electricity 14,500
13 Fuel 697,300
14 Total utilities 953,700
15 Total direct operating costs (4, 10 & 14) $2,656,300
Indirect Operating Costs
16 Plant overhead 137,700
1? Taxes and insurance 56,500
18 Plant cost (15, 16 & 17) 2,850,500
19 General & administrative, sales, research 169,400
20 Cash expenditures (18 & 19) 3,019,900
21 Depreciation 282,300
22 Interest on working capital 17,800
23 Total operating costs* (20, 21 & 22) $3,320,000
24 Cost (cents/bbl) 6.83
*£>oes not include by-product credit or recovery costs
1 bpsd = 1.84 x 10"6 mVs.
199
-------
Table B-25. SUMMARY OP CAPITAL INVESTMENT COSTS
PCC Unit Size: 150,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 12 kg H-0/100 kg Catalyst
Fixed capital investment $4,491,000
Initial catalyst cost 157,300
Start-up cost 449,100
Working capital 471,600
Interest on construction loan 1793600
Total investment $5,748,600
1 bpsd = 1.84 x 106 m3/s.
200
-------
Table B-26. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 150,000 bpsd at 90% Capacity
Typical Stripping Operation
Fixed Capital Investment: $4,491,000
Steam Stripping Rate: 12 kg HpO/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 89,800
3 Control laboratory 19,300
4 Total labor 205,500
Materials
5 Raw and process - acid sludge 9,500
6 lime 3,400
7 catalyst replacement 1,458,000
8 Maintenance 89,800
9 Operating 9,600
10 Total materials 1,570,300
Utilities
11 Process water 483,700
12 Electricity 28,900
13 Fuel 1,394,600
14 Total utilities 1,907,200
15 Total direct operating costs (4, 10 & 14) $3,683,000
Indirect Operating Costs
16 Plant overhead 164,400
17 Taxes and insurance 89,800
18 Plant cost (15, 16 & 17) 3,937,200
19 General & administrative, sales, research 269,500
20 Cash expenditures (18 & 19) 4,206,700
21 Depreciation 449,100
22 Interest on working capital 28,300
23 Total operating costs* (20, 21 & 22) $4,684,100
24 Cost (cents/bbl) 9.64
*Does not include by-product credit or recovery costs.
1 bpsd = 1.84 x 10~* mVs.
201
-------
Table B-27. SUMMARY OP CAPITAL INVESTMENT COSTS
FCC Unit Size: 150,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 40 kg H20/100 kg Catalyst
Fixed capital investment $10,062,000
Initial catalyst cost 524,400
Start-up cost 1,006,200
Working capital 1,056,500
Interest on construction loan 402,500.
Total investment $13,051,600
1 bpsd = 1.84 x 106 mVs.
202
-------
Table B-28. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 150,000 bpsd at 90% Capacity
Typical Stripping Operation
Fixed Capital Investment: $10,062,000
Steam Stripping Rate: 40 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 201,200
3 Control laboratory 19,300
4 Total labor 316,900
Materials
5 Raw and process - acid sludge 31,600
6 lime 11,400
7 catalyst replacement 1,458,000
8 Maintenance 201,200
9 Operating 9,600
10 Total materials 1,711,800
Utilities
11 Process water 1,609,600
12 Electricity 96,400
13 Fuel 4,648,800
14 Total utilities 6,354,800
15 Total direct operating costs (4, 10 & 14) $ 8,383,500
Indirect Operating Costs
16 Plant overhead 253,500
17 Taxes and insurance 201,200
18 plant cost (15, 16 & 17) 8,838,200
19 General & administrative, sales, research 603,700
20 Cash expenditures (18 & 19) 9,441,900
21 Depreciation ' 1,006,200
22 interest on working capital 63,400
23 Total operating costs* (20, 21 & 22) $10,511,500
24 cost (cents/bbl) 21.63
*E>oes not include by-product credit or recovery costs.
1 bpsd = 1.84 x 10"* mVs.
203
-------
Table B-29. SUMMARY OP CAPITAL INVESTMENT COSTS
FCC Unit Size: 150,000 bpsd
Typical Stripping Operation
Steam Stripping Rate: 100 kg H~20/100 kg Catalyst
Fixed capital investment $18,590,000
Initial catalyst cost 1,311,000
Start-up cost 1,859,000
Working capital 1,952,000
Interest on construction loan 743,60(3
Total investment $24,455,600
1 bpsd = 1.84 x 10
204
-------
Table B-30. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 150,000 bpsd at 90% Capacity
Typical Stripping Operation
Fixed Capital Investment: $18,590,000
Steam Stripping Rate: 100 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 371,800
3 Control laboratory 19,300
4 Total labor 487,500
Materials
5 Raw and process - acid sludge 79,100
6 lime 28,400
7 catalyst replacement 1,458,000
8 Maintenance 371,800
9 Operating 9,600
10 Total materials 1,946,900
Utilities
11 Process water 4,031,100
12 Electricity 241,000
13 Fuel 11,677,000
14 Total utilities 15,949,100
15 Total direct operating costs (4, 10 & 14) $18,383,500
Indirect Operating Costs
16 Plant overhead 390,000
17 Taxes and insurance 371,800
18 Plant cost (15, 16 & 17) 19,145,300
19 General & administrative, sales, research 1,115,400
20 Cash expenditures (18 & 19) 20,260,700
21 Depreciation 1,859,000
22 Interest on working capital 117,100
23 Total operating costs* (20, 21 & 22) $22,236,800
24 Cost (cents/bbl) 45.75
*Does not include by-product credit or recovery costs
1 bosd = 1.84 x 10"* m3/s.
bpsd
205
-------
Table B-31. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 10,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 4 kg H20/100 kg Catalyst
Fixed capital investment $572,900
Initial catalyst cost 8,700
Start-up cost 57,300
Working capital 60,200
Interest on construction loan 22,900^
Total investment $722,000
1 bpsd = 1.84 x 106 m3/s.
206
-------
Table B-32. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 10,000 bpsd at 90% Capacity
Worst Stripping Operation
Fixed Capital Investment: $572,900
Steam Stripping Rate: 4 kg H-0/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 11,500
3 Control laboratory 19,300
4 Total labor 127,200
Materials
5 Raw and process - acid sludge 400
6 lime 200
7 catalyst replacement 198,000
8 Maintenance 11,500
9 Operating 9,600
10 Total materials 219,700
Utilities
11 Process water 22,000
12 Electricity ' 1,400
13 Fuel 63,400
14 Total utilities 86,800
15 Total direct operating costs (4, 10 & 14) $433,700
Indirect Operating Costs
16 Plant overhead 101,800
17 Taxes and insurance 11,500
18 Plant cost (15, 16 & 17) 547,000
19 General & administrative, sales, research 34,400
20 Cash expenditures (18 & 19) 581,400
21 Depreciation 57,300
22 Interest on working capital 3,600
23 Total operating costs* (20, 21 & 22) $642,300
24 Cost (cents/bbl) 19.82
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10"® m3/s.
bpsd
207
-------
Table B-33. SUMMARY OP CAPITAL INVESTMENT COSTS
FCC Unit Size: 10,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 6 kg H20/100 kg Catalyst
Fixed capital investment $731,700
Initial catalyst cost 10,500
Start-up cost 73,200
Working capital 76,800
Interest on construction loan 29,300
Total investment $921,500
1 bpsd = 1.84 x 106 mVs.
208
-------
Table B-34. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 10,000 bpsd at 90% Capacity
Worst Stripping Operation
Fixed Capital Investment: $731,700
Steam Stripping Rate: 6 kg HpO/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 14,600
3 Control laboratory 19,300
4 Total labor 130,300
Materials
5 Raw and process - acid sludge 600
6 lime 200
7 catalyst replacement 194,400
8 Maintenance 14,600
9 Operating 7,700
10 Total materials 217,500
Utilities
11 Process water 32,200
12 Electricity 1,900
13 Fuel 93,000
14 Total utilities 127,100
15 Total direct operating costs (4, 10 & 14) $474,900
Indirect Operating Costs
16 Plant overhead 104,300
17 Taxes and insurance 14,600
18 Plant cost (15, 16 & 17) 593,800
19 General & administrative, sales, research 43,900
20 Cash expenditures (18 & 19) 673,700
21 Depreciation 73,200
22 Interest on working capital 4,600
23 Total operating costs* (20, 21 & 22) $715,500
24 Cost (cents/bbl) 22.08
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~6 m3/s.
bpsd
209
-------
Table B-35. SUMMARY OP CAPITAL INVESTMENT COSTS
FCC Unit Size: 100,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 12 kg H"20/100 kg Catalyst
Fixed capital investment $1,164,300
Initial catalyst cost 20,800
Start-up cost 116,400
Working capital 122,300
Interest on construction loan 46,600
Total investment $1,470,400
1 bpsd = 1.84 x 106 mVs.
210
-------
Table B-36. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 10,000 bpsd at 90$ Capacity
Worst Stripping Operation
Fixed Capital Investment: $1,164,300
Steam Stripping Rate: 12 kg HpO/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 23,300
3 Control laboratory 19,300
4 Total labor 139,000
Materials
5 Raw and process - acid sludge 1,300
6 lime 500
7 catalyst replacement 194,400
8 Maintenance 23,300
9 Operating 9,600
10 Total materials 229,100
Utilities
11 Process water 64,400
12 Electricity 3,900
13 Fuel 186,200
14 Total utilities 254,500
15 Total direct operating costs (4, 10 & 14) $622,600
Indirect Operating Costs
16 Plant overhead 111,200
17 Taxes and insurance 23,300
18 Plant cost (15, 16 & 17) 757,100
19 General & administrative, sales, research 69,900
20 Cash expenditures (18 & 19) 827,000
21 Depreciation 116,400
22 Interest on working capital 7,300
23 Total operating costs* (20, 21 & 22) $950,700
24 Cost (cents/bbl) 29.34
*Does not include by-product credit or recovery costs.
1 bpsd = 1.84 x 10~6 m3/s.
bpsd
211
-------
Table B-37. SUMMARY OP CAPITAL INVESTMENT COSTS
FCC Unit Size: 10,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 40 kg H20/100 kg Catalyst
Fixed capital investment $2,608,000
Initial catalyst cost 70,000
Start-up cost 260,800
Working capital 273,800
Interest on construction loan 10*J,300_
Total investment $3,316,900
1 bpsd = 1.84 x 106 mVs.
212
-------
Table B-38. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 10,000 bpsd at 9Q% Capacity
Worst Stripping Operation
Fixed Capital Investment: $2,608,000
Steam Stripping Rate: 40 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 52,200
3 Control laboratory 19,300
4 Total labor 167,900
Materials
5 Raw and process - acid sludge 8,500
6 lime 1,500
7 catalyst replacement 194,400
8 Maintenance 52,200
9 Operating 9,600
10 Total materials 266,200
Utilities
11 Process water 214,400
12 Electricity 12,900
13 Fuel 619,100
14 Total utilities 846,400
15 Total direct operating costs (4, 10 & 14) $1,280,500
Indirect Operating Costs
16 Plant overhead 134,400
17 Taxes and insurance 52,200
18 Plant cost (15, 16 & 17) 1,467,000
19 General & administrative, sales, research 156,500
20 Cash expenditures (1$ & 19) 1,623,500
21 Depreciation 206,800
22 Interest on working capital 16,400
23 Total operating costs* (20, 21 & 22) $1,846,700
24 Cost (cents/bbl) 57.00
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~* m3/s.
213
-------
Table B-39. SUMMARY OP CAPITAL INVESTMENT COSTS
FCC Unit Size: 10,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 100 kg H20/100 kg Catalyst
Fixed capital investment $1,819,000
Initial catalyst cost 171,600
Start-up cost 181,900
Working .capital 506,000
Interest on construction loan 192,800
Total investment $6,1?1,300
1 bpsd = 1.81 x 106 m3/s.
211
-------
Table B-40. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 10,000 bpsd at 90% Capacity
Worst Stripping Operation
Fixed Capital Investment: $4,819,000
Steam Stripping Rate: 100 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 96,400
3 Control laboratory 19,300
4 Total labor 212,100
Materials
5 Raw and process - acid sludge 10,600
6 lime 3,800
7 catalyst replacement 194,400
8 Maintenance 96,400
9 Operating 9,600
10 Total materials 314,800
Utilities
11 Process water 536,100
12 Electricity 32,100
13 Fuel 1,553,300
14 Total utilities 2,121,500
15 Total direct operating costs (4, 10 & 14) $2,648,400
Indirect Operating Costs
16 Plant overhead 169,700
17 Taxes and insurance 96,400
18 Plant cost (15, 16 & 17) 2,914,500
19 General & administrative, sales, research 289,100
20 Cash expenditures (19 & 19) 3,203,600
21 Depreciation 481,900
22 Interest on working capital 30,400
23 Total operating costs* (20, 21 & 22) $3,715,900
24 Cost (cents/bbl) 114.7
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10"* rnVs.
215
-------
Table B-4l. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 50,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: M kg H20/100 kg Catalyst
Fixed capital investment $1,639,^00
Initial catalyst cost 35,000
Start-up cost 163,900
Working capital 172,100
Interest on construction loan 65,600
Total investment $2,076,000
1 bpsd = 1.84 x 106 m3/s.
216
-------
Table B-42. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 50,000 bpsd at 90$ Capacity
Worst Stripping Operation
Fixed Capital Investment: $1,639,400
Steam Stripping Rate: 4 kg HpO/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 32,800
3 Control laboratory 19,300
4 Total labor 148,500
Materials
5 Raw and process - acid sludge 3,200
6 lime 1,100
7 catalyst replacement 972,000
8 Maintenance 32,800
9 Operating 9,600
10 Total materials 1,018,700
Utilities
11 Process water 161,200
12 Electricity 9,600
13 Fuel 464,800
14 Total utilities 635,600
15 Total direct operating costs (4, 10 & 14) $1,802,800
Indirect Operating Costs
16 Plant overhead 118,800
17 Taxes and insurance 32,800
18 Plant cost (15, 16 & 17) 1,954,400
19 General & administrative, sales, research 98,400
20 Cash expenditures (18 & 19) 2,052,800
21 Depreciation 163,900
22 Interest on working capital 6,900
23 Total operating costs* (20, 21 & 22) $2,223,600
24 Cost (cents/bbl) 13-73
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10"* m3/s.
bpsd
217
-------
Table B-43. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 50,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 6 kg H^O/100 kg Catalyst
Fixed capital investment $2,151,000
Initial catalyst cost 52,400
Start-up cost 215,100
Working,capital 225,900
Interest on construction loan 86,000.
Total investment $2,730,400
1 bpsd = 1.84 x 106 mVs.
218
-------
Table B-44. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 50,000 bpsd at 90% Capacity
Worst Stripping Operation
Fixed Capital Investment: $2,151,000
Steam Stripping Rate: 6 kg HpO/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 43,000
3 Control laboratory 19,300
4 Total labor 158,700
Materials
5 Raw and process - acid sludge 3,200
6 lime 1,100
7 catalyst replacement 972,000
8 Maintenance 43,000
9 Operating 93600
10 Total materials 1,028,900
Utilities
11 Process water 161,200
12 Electricity . 9,600
13 Fuel 464,900
14 Total utilities 635,700
15 Total direct operating costs (4, 10 & 14) $1,823,300
Indirect Operating Costs
16 Plant overhead 127,000
17 Taxes and insurance 43,000
18 Plant cost (15, 16 & 17) 1,993,300
19 General & administrative, sales, research 129,100
20 Cash expenditures (18 & 19) 2,122,400
21 Depreciation 215,100
22 Interest on working capital 13,600
23 Total operating costs* (20, 21 & 22) $2,351,100
24 Cost (cents/bbl) 14.51
*Does not include by-product credit or recovery costs,
1 bpsd = 1.84 x 10~* m3/s.
219
-------
Table B-45. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 50,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 12 kg H^O/lOO kg Catalyst
Fixed capital investment $3,422,600
Initial catalyst cost 104,900
Start-up cost 342,200
Working capital 359,400
Interest on construction loan 136,900.
Total investment $4,366,000
1 bpsd = 1.84 x 106 m3/s.
220
-------
Table B-46. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 50,000 bpsd at 90% Capacity
Worst Stripping Operation
Fixed Capital Investment: $3,422,600
Steam Stripping Rate: 12 kg H00/100 kg Catalyst
c.
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 68,400
3 Control laboratory 19,300
4 Total labor 184,100
Materials
5 Raw and process - acid sludge 6,300
6 lime 2,270
7 catalyst replacement 972,000
8 Maintenance 68,400
9 Operating 19,300
10 Total materials 1,068,300
Utilities
11 Process water 322,500
12 Electricity 19,300
13 Fuel 929,800
14 Total utilities 1,271,600
15 Total direct operating costs (4, 10 & 14) $2,524,000
Indirect Operating Costs
16 Plant overhead 147,300
17 Taxes and insurance 68,500
18 Plant cost (15, 16 & 17) 2,739,800
19 General & administrative, sales, research 205,400
20 Cash expenditures (18 & 19) 2,945,200
21 Depreciation 342,300
22 Interest on working capital 21,600
23 Total operating costs* (20, 21 & 22) $3,309,100
24 Cost (cents/bbl) 20.43
*Does not include by-product credit or recovery costs.
1 bpsd = 1.84 x 1Q~* mVs.
221
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Table B-47. SUMMARY OP CAPITAL INVESTMENT COSTS
FCC Unit Size: 50,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 40 kg H20/100 kg Catalyst
Fixed capital investment $7,668,000
Initial catalyst cost 349,200
Start-up cost 766,800
Working capital 805,100
Interest on construction loan 306,700
Total investment $9,895,800
-1 bpsd = 1.84 x 10e m3/s.
222
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Table B-48. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 50,000 bpsd at 90% Capacity
Worst Stripping Operation
Fixed Capital Investment: $7,668,000
Steam Stripping Rate: 40 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 153,400
3 Control laboratory 19,300
4 Total labor 269,100
Materials
5 Raw and process - acid sludge 21,100
6 lime 7,600
7 catalyst replacement 972,000
8 Maintenance 153,400
9 Operating 9,600
10 Total materials 1,163,700
Utilities
11 Process water 1,075,000
12 Electricity 64,200
13 Fuel 3,106,500
14 Total utilities 4,245,700
15 Total direct operating costs (4, 10 & 14) $5,678,500
Indirect Operating Costs
16 Plant overhead 215,300
17 Taxes and insurance 153,400
18 Plant cost (15, 16 & 17) 6,047,200
19 General & administrative, sales, research 460,000
20 Cash expenditures (18 & 19) 6,507,200
21 Depreciation 766,800
22 Interest on working capital 48,300
23 Total operating costs* (20, 21 & 22) $7,322,300
24 Cost (cents/bbl) ^5.20
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10"6 mVs.
223
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Table B-49. SUMMARY OP CAPITAL INVESTMENT COSTS
FCC Unit Size: 50,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 100 kg H00/100 kg Catalyst
c.
Fixed capital investment $14,168,000
Initial catalyst cost 873,000
Start-up cost 1,416,800
Working capital 1,487,600
Interest on construction loan 366,700
Total investment $18,512,100
1 bpsd = 1.84 x 106 mVs.
224
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Table B-50. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 50,000 bpsd at 90% Capacity
Worst Stripping Operation
Fixed Capital Investment: $14,168,000
Steam Stripping Rate: 100 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 283,400
3 Control laboratory 19,300
4 Total labor 399,100
Materials
5 Raw and process - acid sludge 52,900
6 lime 19,000
7 catalyst replacement 972,000
8 Maintenance 283,400
9 Operating 9,600
10 Total materials 1,336,900
Utilities
11 Process water 2,687,400
12 Electricity 160,600
13 Fuel 7,766,300
14 Total utilities 10,614,300
15 Total direct operating costs (4, 10 & 14) $12,350,300
Indirect Operating Costs
16 Plant overhead 319,300
17 Taxes and insurance 283,400
18 Plant cost (15, 16 & 17) 12,953,000
19 General & administrative, sales, research 850,000
20 Cash expenditures (18 & 19) 13,803,000
21 Depreciation • 1,416,800
22 Interest on working capital 89,300
23 Total operating costs* (20, 21 & 22) $15,309,100
24 Cost (cents/bbl) 94.50
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~6 mVs.
225
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Table B-51. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 150,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 4 kg H20/100 kg Catalyst
Fixed capital investment $3,516,100
Initial catalyst cost 130,700
Start-up cost 351,600
Working capital 369,200
Interest on construction loan 140,600
Total investment $4,508,200
1 bpsd = 1.84 x 106 mVs.
226
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Table B-52. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 150,000 bpsd at 90£ Capacity
Worst Stripping Operation
Fixed Capital Investment: $3,516,100
Steam Stripping Rate: 4 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 70,300
3 Control laboratory 19,300
4 Total labor 186,000
Materials
5 Raw and process - acid sludge 6,700
6 lime 2,700
7 catalyst replacement 2,970,000
8 Maintenance 70,300
9 Operating 9,600
10 Total materials 3,059,300
Utilities
11 Process water 328,700
12 Electricity 21,300
13 Fuel 950,000
14 Total utilities 1,300,000
15 Total direct operating costs (4, 10 & 14) $4,545,300
Indirect Operating Costs
16 Plant overhead 148,800
17 Taxes and insurance 70,300
18 Plant cost (15, 16 & 17) 4,764,400
19 General & administrative, sales, research 211,000
20 Cash expenditures (18 & 19) 4,975,400
21 Depreciation 351,600
22 Interest on working capital 22,200
23 Total operating costs* (20, 21 & 22) $5,349,200
24 Cost (cents/bbl) 10.99
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~s m3/s.
227
-------
Table B-53. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 150,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 6 kg H20/100 kg Catalyst
Fixed capital investment $4,491,000
Initial catalyst cost 157,300
Start-up cost 449,100
Working capital 4?2,600
Interest on construction loan 179»60.0.
Total investment $5,748,600
1 bpsd = 1.84 x 106 m3/s.
228
-------
Table B-54. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 150,000 bpsd at 90$ Capacity
Worst Stripping Operation
Fixed Capital Investment: $4,491,000
Steam Stripping Rate: 6 kg H00/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 89,800
3 Control laboratory 19,300
4 Total labor 205,500
Materials
5 Raw and process - acid sludge 9,500
6 lime 3,400
7 catalyst replacement 2,916,000
8 Maintenance 89,800
9 Operating 9,600
10 Total materials 3,028,300
Utilities
11 Process water 483,700
12 Electricity 28,900
13 Fuel 1,394,6000
14 Total utilities 1,907,200
15 Total direct operating costs (4, 10 & 14) $5,141,000
Indirect Operating Costs
16 Plant overhead 164,400
17 Taxes and insurance 89,800
18 Plant cost (15, 16 & 17) 5,395,200
19 General & administrative, sales, research 269,500
20 Cash expenditures (18 & 19) 5,664,700
21 Depreciation 449,100
22 Interest on working capital 28,300
23 Total operating costs* (20, 21 & 22) $6,142,100
24 Cost (cents/bbl) 12.64
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~* mVs.
229
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Table B-55. SUMMARY OP CAPITAL INVESTMENT COSTS
PCC Unit Size: 150,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 12 kg H20/100 kg Catalyst
Fixed capital investment $7,1^5,400
Initial catalyst cost 31^,600
Start-up cost 714,500
Working capital 750,300
Interest on construction loan 285,800
Total investment $9,210,600
1 bpsd = 1.84 x 106 mVs.
230
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Table B-56. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 150,000 bpsd at 90% Capacity
Worst Stripping Operation
Fixed Capital Investment: $7,145,400
Steam Stripping Rate: 12 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 1*12,900
3 Control laboratory 19,300
4 Total labor 258,600
Materials
5 Raw and process - acid sludge 19,000
6 lime 6,800
7 catalyst replacement 2,916,000
8 Maintenance 1*12,900
9 Operating 9,600
10 Total materials 3,094,300
Utilities
11 Process water 1,075,000
12 Electricity 57,900
13 Fuel 2,789,000
14 Total utilities 3,921,900
15 Total direct operating costs (4, 10 & 1*1) $7,274,800
Indirect Operating Costs
16 Plant overhead 206,800
17 Taxes and insurance 142,900
18 Plant cost (15, 16 & 17) 7,624,500
19 General & administrative, sales, research 428,700
20 Cash expenditures (18 & 19) 8,053,200
21 Depreciation 714,500
22 Interest on working capital 45,000
23 Total operating costs* (20, 21 & 22) $8,812,700
24 Cost (cents/bbl) 18.13
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10"6 m3/s.
231
-------
Table B-57 . SUMMARY OP CAPITAL INVESTMENT COSTS
FCC Unit Size: 150,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 40 kg I^O/lOO kg Catalyst
Fixed capital investment $16,009,000
Initial catalyst cost 1,048,800
Start-up cost 1,600,090
Working capital 1,680,900
Interest on construction loan
Total investment $20,980,000
1 bpsd = 1.84 x 10s mVs.
232
-------
Table B-58. SUMMARY OF ANNUAL OPERATING COSTS
FCC Unit Size: 150,000 bpsd at 90$ Capacity
Worst Stripping Operation
Fixed Capital Investment: $16,009,000
Steam Stripping Rate: 40 kg HpO/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 320,200
3 Control laboratory 19,300
4 Total labor 435,900
Materials
5 Raw and process - acid sludge 63,300
6 lime 22,700
7 catalyst replacement 2,916,000
8 Maintenance 320,200
9 Operating 9,600
10 Total materials 3,331,800
Utilities
11 Process water 3,224,900
12 Electricity 192,800
13 Fuel 9,297,500
14 Total utilities 12,715,200
15 Total direct operating costs (4, 10 & 14) $16,482,900
Indirect Operating Costs
16 Plant overhead 348,700
17 Taxes and insurance 320,200
18 Plant cost (15, 16 & 17) 17,151,800
19 General & administrative, sales, research 960,500
20 Cash expenditures (18 & 19) 18,112,300
21 Depreciation 1,600,900
22 Interest on working capital 100,900
23 Total operating costs* (20, 21 & 22) $19,814,100
24 Cost (cents/bbl) 40.77
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~6 m3/s.
233
-------
Table B-59. SUMMARY OF CAPITAL INVESTMENT COSTS
FCC Unit Size: 150,000 bpsd
Worst Stripping Operation
Steam Stripping Rate: 100 kg H20/100 kg Catalyst
Fixed capital investment $29,579,000
Initial catalyst cost 2,622,000
Start-up cost 2,957,900
Working capital 3,105,800
Interest on construction loan 1,183,200
Total investment $39,W,900
1 bpsd = 1.8*1 x 106 mVs.
234
-------
Table B-60. SUMMARY OP ANNUAL OPERATING COSTS
FCC Unit Size: 150,000 bpsd at 90$ Capacity
Worst Stripping Operation
Fixed Capital Investment: $29,579,000
Steam Stripping Rate: 100 kg H20/100 kg Catalyst
Direct Operating Costs
Labor
1 Operating $ 96,400
2 Maintenance 591,600
3 Control laboratory 19,300
4 Total labor 707,300
Materials
5 Raw and process - acid sludge 158,200
6 lime 56,700
7 catalyst replacement 2,916,000
8 Maintenance 591,600
9 Operating 9,600
10 Total materials 3,732,100
Utilities
11 Process water 8,062,200
12 Electricity 482,100
13 Fuel 23,243,800
14 Total utilities 31,788,100
15 Total direct operating costs (4, 10 & 14) $36,227,500
Indirect Operating Costs
16 Plant overhead 565,800
17 Taxes and insurance 591,600
18 Plant cost (15, 16 & 17) 37,384,900
19 General & administrative, sales, research 1,774.700
20 Cash expenditures (1.8 & 19) 39,159,600
21 Depreciation 2,957,900
22 Interest on working capital 186P3QO
23 Total operating costs* (20, 21 & 22) $42,303,800
24 Cost (cents/bbl) 87.04
*Does not include by-product credit or recovery costs
1 bpsd = 1.84 x 10~* mVs.
235
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TECHNICAL REPORT DATA
{Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-6 50/2-74-082-a
2.
3. RECIPIENT'S ACCESSION-NO.
4. TITLE AND SUBTITLE
Refinery Catalytic Cracker Regenerator SOX Control-
Steam Stripper Laboratory Test
5. REPORT DATE
November 1974
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
T. Ctvrtnicek, T. W.Hughes , C. M. Moscowitz, and
PL. Zanders
8. PERFORMING ORGANIZATION REPORT NO.
MRC-DA-446
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Monsanto Research Corporation
Dayton Laboratory
Dayton, Ohio 45407
10. PROGRAM ELEMENT NO.
1AB013: ROAP 21ADC-031
11. CONTRACT/GRANT NO.
68-02-1320
(Taskl, Phase II)
12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
NERC-RTP, Control Systems Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD COVERED
Phase II Final, 11/73-9/74
14. SPONSORING AGENCY CODE
15. SUPPLEMENTARY NOTES
kc The report summarizes experimental results from steam contacting of
spent catalyst used in petroleum refinery fluid catalytic crackers. This concept has
been identified as a potentially effective means of sulfur emission control for fluid
catalytic cracker regenerators. Correlations between sulfur removal efficiency
from the catalyst and the product of steam residence time in stripper and steam
stripping rate are presented for several stripper designs. The extent of by-product
formation, a discussion of pertinent equipment design, and recommendations for
further investigation and development of this concept are also included. Additionally:
the economics are presented as a function of steam stripping rate and fluid catalytic
cracker unit size.
16. ABJ
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Air Pollution
Petroleum Refining
Catalytic Cracking
Strippers
Tests
Regeneration
Sulfur Oxides
Desulfurization
Cost Effective-
ness
Air Pollution Control
Stationary Sources
Steam Contacting
13B, 07B
13H, 07D
07A
14A
14B
Unlimited
19. SECURITY CLASS (This Report)
Unclassified
21. NO. OF PAGES
20. SECURITY CLASS -(This page)
Unclassified
248
22. PRICE
EPA Form 222D-1 (9-73)
236
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