U.S. patent number 3,928,172 [Application Number 05/376,091] was granted by the patent office on 1975-12-23 for catalytic cracking of fcc gasoline and virgin naphtha.
This patent grant is currently assigned to Mobil Oil Corporation. Invention is credited to Francis E. Davis, Jr., Richard G. Graven, Wooyoung Lee, Robert A. Sailor.
United States Patent |
3,928,172 |
Davis, Jr. , et al. |
December 23, 1975 |
Catalytic cracking of FCC gasoline and virgin naphtha
Abstract
A combination process is described for improving the quality and
volatility of a refinery gasoline pool comprising the recracking of
gasoline product of gas oil cracking and separate product recovery
thereof, cracking of virgin naphtha and alkylating olefins formed
in the combination process for blending with pool gasoline.
Inventors: |
Davis, Jr.; Francis E.
(Woodbury, NJ), Graven; Richard G. (Westmont, NJ), Lee;
Wooyoung (Westmont, NJ), Sailor; Robert A. (Riverton,
NJ) |
Assignee: |
Mobil Oil Corporation (New
York, NY)
|
Family
ID: |
23483675 |
Appl.
No.: |
05/376,091 |
Filed: |
July 2, 1973 |
Current U.S.
Class: |
208/77; 208/16;
208/78; 208/164; 585/476; 208/120.01; 208/80; 502/67; 208/72 |
Current CPC
Class: |
C10G
11/18 (20130101) |
Current International
Class: |
C10G
11/18 (20060101); C10G 11/00 (20060101); B01J
029/28 (); C10G 011/04 (); C10G 037/02 () |
Field of
Search: |
;208/72,73,77,120 |
References Cited
[Referenced By]
U.S. Patent Documents
Primary Examiner: Gantz; Delbert E.
Assistant Examiner: Schmitkons; G. E.
Attorney, Agent or Firm: Huggett; Charles Gilman; Michael G.
Farnsworth; Carl D.
Claims
We claim:
1. A method for upgrading hydrocarbons which comprises
cracking a gas oil boiling range feed in a first cracking zone in
the presence of a crystalline zeolite cracking catalyst at an
elevated temperature of at least 850.degree.F to obtain conversion
of the gas oil feed to materials including a heavy cycle oil, a
light cycle oil, a heavy naphtha fraction and materials lower
boiling than said heavy naphtha fraction,
cracking the heavy naphtha fraction above obtained in a second
cracking zone in the presence of freshly regenerated crystalline
zeolite cracking catalyst at an elevated temperature within the
range of 850.degree.F to about 1000.degree.F, separating the
product of said heavy naphtha fraction cracking operation in a zone
separate from the gas oil product separation under conditions to
recover gasoline boiling range material rich in aromatics and
lighter hydrocarbon material from a light cycle oil product
material, combining said light cycle oil recovered products,
combining materials lower boiling than said heavy naphtha product
of gas oil cracking with said gasoline boiling and lighter
materials separated from the product of said heavy naphtha cracking
operation and
separating the combined materials lower boiling than light cycle
oil into a gasoline boiling range product fraction and a lower
boiling uncondensed vaporous product fraction.
2. The method of claim 1 wherein a freshly regenerated crystalline
aluminosilicate containing cracking catalyst is sequentially passed
through said heavy naphtha cracking step, said gas oil cracking
step and a virgin naphtha cracking step before stripping of the
catalyst and passing stripped catalyst to catalyst
regeneration.
3. The method of claim 2 wherein the cracking catalyst comprises a
faujasite type of crystalline zeolite in combination with a
crystalline aluminosilicate of the ZSM-5 type.
4. The method of claim 1 wherein freshly regenerated catalyst is
used for recracking the heavy cracked naphtha material and the gas
oil feed and catalyst used in each cracking operation is passed to
a common catalyst stripping operation before being regenerated.
5. The method of claim 1 wherein recracking of said heavy naphtha
fraction is accomplished in a high temperature riser cracking zone
at a temperature of at least 950.degree.F.
6. The method of claim 1 wherein the recracking of said heavy
naphtha fraction is accomplished in a riser cracking zone.
7. The method of claim 1 wherein recracking of said heavy naphtha
fraction is accomplished in a high temperature riser cracking zone
terminating in a dense fluid bed cracking zone.
8. The method of claim 1 wherein catalyst separated from said gas
oil riser cracking step is used to convert a virgin naphtha at an
elevated cracking temperature to improve its octane rating and the
products of said virgin naphtha cracking step are separated with
products of said gas oil cracking step.
9. The method of claim 1 wherein the heavy naphtha fraction
subjected to recracking contains hydrocarbon components boiling up
to about 500.degree.F.
10. The method of claim 1 wherein a virgin naphtha is cracked in
the presence of said heavy cracked naphtha.
Description
BACKGROUND OF THE INVENTION
The fluid catalyst system as we know it today embodies the
technique of utilizing finely divided solid particulate material in
a fluid or a freely moving state such that a mass of solids can be
circulated much in the same way as a liquid. Thus in catalytic
cracking operations, the fluid catalyst is caused to flow between
and through a hydrocarbon conversion zone and a catalyst
regeneration zone. The catalyst has been used in dispersed and/or
dense phase condition or combinations thereof to permit contact of
reactant materials therewith followed by separation of catalyst
particles from products of reaction. Thus the catalyst has been
used either as a dense fluid bed of catalyst in the reaction and
regeneration zones or passed in a dispersed phase condition
upwardly through one or both zones.
During World War II, the demands for large volumes of high octane
gasoline suitable for military use prompted the recracking of
gasoline products of thermal and catalytic gas oil cracking
operations over conventional amorphous silica-alumina catalysts in
order to achieve high octane materials. Throughout the history of
catalytic cracking, hydrocarbon materials insufficiently cracked to
produce gasoline boiling range material have been recycled and
subjected to further catalytic cracking. To improve the
crackability of heavy recycle fractions, they may be subjected to a
prehydrogenation treatment before further cracking thereof. Some
early patents directed to the techniques above expressed are:
Nelson U.S. Pat. No. 2,376,501 issued May 22, 1945; Newton U.S.
Pat. No. 2,406,394 issued Aug. 27, 1946; Sweeney U.S. Pat. No.
2,426,903 issued Sept. 2, 1947; Nelson U.S. Pat. No. 2,425,555
issued Aug. 12, 1947 and Marshall U.S. Pat. No. 2,921,014 issued
Jan. 12, 1960.
During the last decade cracking catalyst compositions comprising
crystalline aluminosilicate zeolites have been widely adopted in
both fluid and moving bed cracking operations. The use of these
zeolite catalysts is so widespread that they represent over 90% of
catalyst used in fluid operations today. They are used in dispersed
phase riser reaction zones alone or in combination with dense
catalyst phase reaction zones. Most of these composite catalysts
consist of a small percentage (5 to 20%) of the crystalline zeolite
in a larger percentage of an amorphous cracking catalyst such as
silica-alumina.
Currently the extreme difficulties encountered throughout the
United States because of atmospheric pollution have imposed many
requirements or potential requirements on motor fuels. Every effort
is being made to reduce atmospheric pollution by products of
combustion from automobiles. These efforts, in turn, result in
specific detailed objectives with motor fuels. For example, reduced
use of tetraethyllead is sought. This necessitates gasolines of
high intrinsic octane number. Higher volatility is being considered
to ensure good driveability of new engines. For instance, it has
been recently proposed by General Motors to set the 50%, 90% and
end boiling points of quality gasoline in the range of
190.degree.-210.degree.F., 280.degree.-320.degree.F. and
330.degree.-380.degree.F. respectively. California has a bromine
number specification for gasoline because olefins in gasoline
contribute to pollution. Since sulfur compounds also contribute to
pollution, low sulfur levels are sought.
To meet higher volatility requirements, it is necessary to reduce
the boiling ranges of certain blending components of the gasoline
pool such as the gasoline product of catalytic cracking and
reforming.
One simple route for lowering the 50% boiling point of gasoline is
by undercutting the cracked stock or by lowering the end point of a
reforming charge stock. However, this is a most expensive route for
producing high volatility gasoline and in times of crude oil
shortage, operates to oppose the conservation of energy fuels.
SUMMARY OF THE INVENTION
It has been discovered that heavy cracked gasoline can be recracked
over a zeolite-containing catalyst at particular operating
conditions to provide an improved and modern approach to the
production of gasoline satisfying stringent new quality
specifications as now dictated by the need to reduce emissions from
internal combustion engines.
The zeolite catalyst is able to effect a degree of octane
improvement necessary for today's needs and not heretofore possible
with amorphous silica-alumina catalysts. Moreover the zeolite
catalyst effects significant volatility improvement and reduces
olefins in the liquid cracked product to a level not practically
attainable with a silica-alumina catalyst. Recracking removes the
bulk of the sulfur from the heavy cracked gasoline. It produces
high yields of propylene, butenes and isobutane which are valuable
as alkylation charge stocks.
Cracking of a virgin heavy naphtha over zeolite catalyst effects
large improvements in octane number and volatility, not practical
over silica-alumina. It produces gaseous olefins and isobutane,
valuable as alkylation feed stocks.
There are a number of ways of accomplishing these benefits. As one
alternative, cracked naphtha is recycled and/or virgin naphtha is
added to a riser cracking zone along with a gas oil feed. A second
alternative may involve a multi-stage operation in which a cracked
naphtha is recracked or a virgin naphtha is cracked in a separate
riser reactor or a dense fluid catalyst bed reactor. As a third
alternative, the choice of conditions for cracking heavy cracked
gasoline and heavy virgin naphtha can be met in a combination
process in which heavy cracked gasoline is recracked at high
temperature in a dense bed, gas oil is cracked in dispersed phase
in a riser, and virgin naphtha is cracked at a lower temperature in
a dense fluid catalyst bed.
DISCUSSION OF SPECIFIC EMBODIMENTS
It has been found when following the concepts of this invention
that the cracking of selected hydrocarbon fractions under
particular operating conditions offers a greater portential in
various refinery applications for producing high volatility
gasoline, high octane blending stock, light olefins for alkylation
reactions, aromatic concentrates for use in petrochemicals,
improved isobutane production and a reduction in olefin and sulfur
in the gasoline product.
Exploratory studies of the concepts of the present invention were
made in laboratory bench scale equipment which involved passing a
heavy gasoline charge material over 180 to 300 gms. of catalyst in
fixed fluidized bed at preselected process conditions. Time on
stream for each cycle of run was 5 minutes and after each cycle the
reactor was purged and the catalyst regenerated in-situ. The
products were cooled and weathered at 120.degree.F. The resulting
liquid product was found to have 40-50 volume percent of the total
C.sub.5's produced. Both liquid and gaseous products were fully
analyzed. The ranges of operating conditions selected were: reactor
temperature from 850.degree. to 1100.degree.F; catalyst to oil
ratio from 2 to 10 w/w; the liquid hourly space velocity from 1 to
6 and a vapor residence time of several seconds.
In another group of examples cracked and virgin naphthas were
cracked in a short-contact-time semi-commercial riser pilot unit.
The riser cracking zone discharge temperature was varied from about
900.degree. to about 1200.degree.F. and the catalyst-to-oil ratio
was varied from about 8 to 60 w/w. Oil and catalyst contact times
were of the order of 4 to 7 seconds.
Most of the examples relate to one identified cracked naphtha and
to one identified virgin naphtha. Properties of these naphthas are
reported in Table I. The one cracked gasoline or naphtha fraction
was produced by cracking over a zeolite catalyst. A few fractions
of similar materials but of different boiling ranges were
examined.
TABLE I
__________________________________________________________________________
CHARGE STOCKS Virgin Straight FCC Cracked Gasoline Run Naphtha,
260-380.degree.F. Fraction 260-380.degree.F. Fraction
__________________________________________________________________________
Properties Gravity, .degree.API 51.1 39.3 Sulfur, wt.% <0.001
0.10 Bromine No. <0.1 18.8 ASTM Distillation, .degree.F IBT 278
280 10% 288 293 30% 296 301 50% 303 309 70% 315 322 90% 334 343 FBP
352 361 Research Octane Number, clear 40.8 93.5 PONA Analysis,
vol.% Paraffins 44.7 12.7 Naphthenes 37.4 8.7 Olefins 0.4 13.6
Aromatics 17.5 66.0
__________________________________________________________________________
Four catalysts have been used. Catalyst A was a commercial
equilibrium fluid catalyst and consisted of about 10% REy zeolite
in a silica-alumina-zirconia-clay matrix. Catalyst B was also a
commercial equilibrium fluid catalyst and consisted of about 15%
REY zeolite in a similar matrix. Catalyst C was a laboratory
steamed catalyst and consisted of 5% of zeolite ZSM-5 on catalyst
B. Thus it contained about 15% REY and 5% ZSM-5. Catalyst D was a
laboratory steamed catalyst and consisted of 10% ZSM-5 on the
matrix material of catalyst A.
RECRACKING OF HEAVY FCC GASOLINE
Approximately 23 weight percent (wt.%) of the gasoline charge stock
was cracked to C.sub.4 or lighter products when a heavy
(260.degree.-380.degree.F) FCC gasoline was contacted in a dense
bed with commercial equilibrium Catalyst A at nominal gas oil
cracking conditions (e.g. 1050.degree.F. reactor temperature and 6
C/O ratio). Conversion, defined as the weight percent of charge
stock converted to C.sub.4 or lighter products and to coke
increased with reactor temperature and catalyst to oil ratio (C/O).
The cracking activity of Catalyst C was approximately the same as
that of Catalyst A.
Conversion to C.sub.4 and lighter, plus coke, is here defined as
"conversion" because these products are generally outside of the
gasoline boiling range and therefore not directly usable in motor
gasoline. However, isobutane and the light olefins are convertible
to alkylate and butanes can be used in the finished gasoline to
give desired vapor pressure.
Tables II and III report results of Examples I to XI involving
cracking of heavy FCC gasoline.
TABLE II
__________________________________________________________________________
Cracking of 260-380 FCC Naphtha in dense bed bench unit Example I
II III IV V Catalyst A A A C D
__________________________________________________________________________
Operating Conditions Avg.Reactor Temp. .degree.F. 1050 1050 950 950
1050 Cat.to Oil Ratio, wt. 2.1 6.0 6.1 6.2 6.3 WHSV 5.7 2.0 2.0 1.9
1.9 Conversion to C.sub.4 -, wt.% charge* 16.5 22.9 17.2 18.8 17.8
Product Yields, % Charge C.sub.5 + Liquid, vol.% 82.6 75.5 81.9
80.7 80.7 Total C.sub.5 's, vol.% 5.7 6.0 7.3 7.7 3.7 Total C.sub.4
's, vol.% 9.8 12.0 11.5 11.9 8.4 Total Dry Gas (wt.%) 8.4 11.5 7.1
7.9 11.3 Coke (wt.%) 1.07 3.04 2.03 2.46 0.53 Isobutane, vol.% 4.4
6.0 7.0 7.4 2.0 Butene, vol.% 4.2 4.0 2.4 2.4 5.9 Propene, vol.%
5.6 5.1 3.2 3.6 9.8 Product Properties Sulfur, wt.% 0.050 0.042
0.027 0.026 0.056 Research O.N. clear 99.5 102.0 100.0 101.1
>9.7 PONA analysis, C.sub.6 +, vol.% Paraffins 12.8 10.4 12.4
9.5 14.1 Naphthenes 5.2 4.4 5.3 3.4 7.9 Olefins 1.8 2.3 1.9 1.1 2.8
Aromatics 80.1 82.9 80.4 85.9 75.2
__________________________________________________________________________
*Includes conversion to coke.
TABLE III
__________________________________________________________________________
Cracking of FCC Naphtha in Riser Pilot Plant Example VI VII VIII IX
X XI Catalyst B B B B B A
__________________________________________________________________________
Boiling Range Charge 260-380 260-450 260-380 Operating Conditions
Riser Top Temp..degree.F. 986 1005 1028 1173 1010 1013 Cat.to Oil
Ratio,w/w 7.7 19.6 35.3 62.9 20.4 10.1 Oil Contact Time, sec. 5.5
5.1 4.8 5.2 5.0 5.6 Cat.Residence Time, sec. 7.1 6.8 6.7 7.4 6.7
7.2 Conversion to C.sub.4 -, wt.% charge* 15.9 24.0 32.9 55.2 25.9
13.2 Product Yields, % charge C.sub.5 + Liquid, vol.% 84.0 75.8
66.4 42.5 74.3 86.7 Total C.sub.3 's, vol.% 8.1 9.5 9.5 4.1 9.3 6.6
Total C.sub.4 's, vol.% 11.8 15.5 18.4 14.1 15.4 10.1 Total Dry
Gas, wt.% 6.2 9.3 12.7 29.1 10.2 5.3 Coke, wt.% 15.1 4.06 7.49
16.27 5.28 0.82 Isobutane, vol.% 6.5 9.5 11.8 5.8 9.5 4.6 Butene,
vol.% 3.2 2.2 1.5 2.6 2.0 4.2 Propene, vol.% 5.3 5.0 4.2 5.1 4.7
4.7 Product Properties Sulfur, wt.% 0.042 0.028 0.023 0.014 0.077
0.042 Bromine No. 5.5 3.1 2.1 1.7 3.1 8.7 Research O.N., clear
100.7 104.3 105.4 110.6 103.9 99.8 PONA Analysis, C.sub.6 +, vol.%
Paraffins 12.6 8.7 3.2 0.5 6.5 14.3 Naphthene 4.0 1.5 0.1 0.1 1.1
5.9 Olefins 1.9 1.6 1.1 0.1 1.0 3.0 Aromatics 81.6 88.2 95.6 99.3
91.3 76.7
__________________________________________________________________________
*Includes conversion to coke
Gasoline octane number is of prime importance and the octane number
in the absence of lead is of special importance because of problems
with air pollution. Example I, Table II, illustrates the fact that
recracking heavy FCC gasoline to about 17% conversion raises the
clear octane number by six units (from 93.5 to 99.5). More
extensive conversion increases octane number still further and 33%
conversion gives about 12 units (93.5 to 105.4, Example VIII, Table
III) while 55% conversion gives about 17 units (93.5 to 110.6,
Example IX). Recracking at 33-55% conversion converted the cracked
gasoline to a liquid product containing over 95% aromatics
(Examples VIII and IX). Conversions of 16% up to 55% were obtained
in dense bed and riser units. Conversion increased with temperature
and with catalyst-to-oil ratio; and octane number (and aromatic
content) increased with conversion. Catalysts A, B and C were
roughly similar in performance. In the riser cracking unit, heavy
FCC gasoline of 260.degree.-450.degree.F. boiling range cracked
very much like the 260.degree.-380.degree.F. fraction. (Compare
Examples VII and X). Similarly, cracked fractions boiling up to
500.degree.F. are desirable charge stocks. Amorphous silica-alumina
cracking catalysts are not able to produce these high
conversions.
Olefin content and bromine number of liquid products thus obtained
is also of interest because of the effect of olefins on air
pollution. Whereas the original heavy FCC gasoline contained 13.6%
olefins, 20% conversion of this fraction by recracking reduced
olefins to about 2%; 33% conversion dropped olefins to near 1%
while 55% conversion removed essentially all olefin. Similarly, 16%
conversion reduced bromine number from 18.8 to about 6 and higher
conversion reduced bromine number still more. With these zeolite
catalysts, a degree of olefin removal is attained which is not
practical with amorphous silicaalumina.
Light olefins (C.sub.2.sup.= .about.C.sub.4.sup.=) produced could
be further processed to produce gasoline blending stocks or
separated for certain petrochemical manufacturing. Gasoline
cracking over ZSM-5 catalyst showed exceptionally high yields of
light olefins as illustrated by Examples V and XIX. For instance,
70-75 wt.% of the C.sub.4 + lighter products in these examples were
light olefins. Therefore, if light olefins are desired, heavy
naphtha or FCC gasoline could be processed over ZSM-5 catalyst.
Recracking of heavy FCC gasoline improves the volatility thereof
and this contributes to improved volatility of the whole gasoline
pool. Table IV reports simulated distillation data obtained by gas
chromatography for the products from a dense fluid catalyst bed
example.
TABLE IV
__________________________________________________________________________
THE EFFECT OF RECRACKING HEAVY FCC GASOLINE ON VOLATILITY DENSE BED
__________________________________________________________________________
UNIT Example Charge I II III IV V Conversion 0 16.5 22.9 17.2 18.8
17.8 Simulated Distillation, .degree.F. IBP 239 113 98 77 17 30 10%
273 231 221 222 227 231 30% 296 279 273 279 274 278 50% 328 300 291
296 286 292 70% 351 335 330 335 322 328 90% 382 376 375 379 354 359
95% 404 414 421 419 375 371 EP 553 606 596 613 527 494
__________________________________________________________________________
The 10% point is reduced 40.degree.-50.degree.F. and the 50% point
is reduced about 30.degree.-40.degree.F. Table IV indicates some
reduction in 90% point but an increase in end-point. At the dense
fluid catalyst bed conditions there are some polymerization and
condensation reactions to form a small amount of heavy ends. These
can be removed as a very few percent of byproduct by
distillation.
Table V reports ASTM distillation data for charge and products of a
riser pilot plant. t2 TABLE V-THE EFFECT OF RECRACKING HEAVY FCC
GASOLINE ON VOLATILITY? ? -RISER PILOT PLANT? -Example Charge VI
VII VIII IX X XI ? - -Conversion 0 15.9 24.0 32.9 55.2 25.9 13.2
-ASTM Distillation, .degree.F. - IBP 280 102 88 103 135 95 102 -
10% 293 202 189 202 220 188 217 - 30% 301 283 272 265 244 274 291 -
50% 309 306 292 283 259 294 313 - 70% 322 329 318 312 286 325 335 -
90% 343 391 422 441 456 471 413 - 95% --? 499 544 --? 520 574 580 -
E.P. 361 539 552 591 623 607 605? -
Again, the 10% point is reduced by a large amount. The 50% point is
reduced 25.degree. to 60.degree. at the 33-55% conversion. The
riser system produces more heavy ends and, in this case, the 90%
point was increased. The 90% point of the charge is retained by
removal of 9-11 vol.% of heavy ends by redistillation.
Recracking of heavy FCC gasoline has the added beneficial effect of
reducing sulfur content. Examples I-IV, Table IV, indicate 50-75%
sulfur removal at about 16-23% conversion. Similarly, Examples
VI-IX and XI, Table V, resulted in 58-85% sulfur removal at 16-55%
conversion.
Cracking of heavy FCC gasoline, unlike reforming, produces olefins
which can be alkylated to increase gasoline yield. During the
recracking of heavy FCC gasoline significant yields of propene and
butene are produced. Examples I to XI, Tables IV and V, indicate
that the amounts produced are greater at higher cracking
temperatures and are greatest at moderate conversions, decreasing
at higher conversion levels. At the same time, significant amounts
of isobutane are produced which can be alkylated with these
olefins. The amount of isobutane produced is highest at low
cracking temperatures and at high conversion levels.
In summary, the recracking of a gasoline product of gas oil
cracking experiments has demonstrated that the heavy gasoline
fraction can be recracked to produce lighter, cleaner and higher
octane gasoline at the expense of the gasoline volume. However, if
C.sub.3 =, C.sub.4 = alkylate material is included, the loss in
total recoverable liquid per octane boost is much less.
CRACKING OF HEAVY VIRGIN NAPHTHA
Heavy virgin naphtha is more easily cracked than a heavy gasoline
product of catalytic cracking and the cracking reaction can be
accomplished at lower temperatures. Tables VI and VII report the
results of Examples XII to XXV. Heavy virgin naphtha was cracked in
both a dense fluid catalyst bed test unit and in a riser pilot
plant test unit.
TABLE VI
__________________________________________________________________________
CRACKING OF 260-380.degree. VIRGIN NAPHTHA IN DENSE BED BENCH UNIT
__________________________________________________________________________
Example XII XIII XIV XV XVI XVII XVIII XIX Catalyst A A A A A B C D
Operating Conditions Avg.Reactor Temp..degree.F 850 950 1050 1050
1100 1100 950 1050 Cat.to Oil Ratio,wt. 10.0 6.2 6.1 2.0 10.1 10.1
6.1 6.0 WHSV 1.2 1.9 2.0 6.0 1.2 1.2 2.0 2.0 Conversion to C.sub.4
-, wt.% charge* 23.7 25.7 34.4 25.2 45.7 54.2 24.1 17.6 Product
Yields, % charge C.sub.5 + Liquid, vol.% 77.9 74.4 65.5 74.6 53.8
44.6 76.6 82.2 Total C.sub.5 's, vol.% 13.1 11.0 10.7 7.0 9.3 7.7
13.4 3.4 Total C.sub.4 's, vol.% 21.0 20.2 22.0 14.5 20.4 21.1 19.0
7.8 Total Dry Gas, wt.% 7.0 9.6 16.0 13.7 21.3 31.9 9.0 11.1 Coke,
wt.% 1.32 1.11 1.89 0.59 4.23 6.64 0.98 0.45 Isobutane, vol.% 15.2
13.4 12.8 8.3 11.3 11.6 12.2 2.0 Butene, vol.% 1.8 3.1 5.3 4.0 6.3
4.9 3.4 5.1 Propene, vol.% 2.6 4.4 7.0 9.0 9.8 9.2 4.6 8.5 Liquid
Product Properties Research ON clear 75.1 69.6 73.2 62.7 82.4 87.8
72.5 58.6 Gravity, .degree.API 49.6 48.2 46.1 49.3 41.5 38.8 48.8
50.7 Sulfur, wt.% .001 .003 .001 .005 .007 .003 .001 .004 PONA
Analysis, C.sub.6 +, vol.% Paraffins 38.2 41.3 37.8 42.5 30.6 24.6
39.0 42.9 Naphthenes 16.2 17.7 17.2 21.4 12.4 8.5 15.7 32.4 Olefins
1.1 1.8 2.5 2.7 2.4 2.3 1.6 2.6 Aromatics 44.3 39.1 42.5 33.4 54.6
64.4 43.8 22.1
__________________________________________________________________________
*Includes conversion to coke
TABLE VII
__________________________________________________________________________
CRACKING OF VIRGIN NAPHTHA IN RISER PILOT PLANT
__________________________________________________________________________
Example XX XXI XXII XXIII XXIV XXV Catalyst B B B B B A Boiling
Range, charge .fwdarw.260-380 .fwdarw. 380-500 260-380 Operating
Conditions Riser Top Temp..degree.F. 998 995 1036 914 977 985 Cat.
to Oil Ratio,w/w 13.3 17.3 26.8 37.7 15.1 9.4 Oil Contact Time,
sec. 5.3 4.4 4.3 5.0 4.2 4.9 Cat Residence Time,sec 6.9 5.8 5.8 7.1
5.6 6.3 Conversion to C.sub.4 - wt.% charge* 27.2 33.6 49.2 32.5
32.7 20.9 Product Yields,% charge C.sub.5 + Liquid, vol.% 75.0 68.1
51.8 69.5 72.0 79.9 Total C.sub.5 's, vol.% 15.3 16.0 17.7 17.8
17.5 11.0 Total C.sub.4 's, vol.% 20.8 26.1 33.8 26.1 24.6 16.6
Total Dry Gas, wt.% 10.1 11.7 18.9 9.5 11.5 7.7 Coke, wt.% 1.71
2.73 5.49 3.80 4.13 0.83 Isobutane, vol.% 12.4 16.3 21.8 17.5 14.5
8.7 Butene, vol.% 5.0 4.0 3.4 2.1 5.3 3.3 Propene, vol.% 7.8 7.4
7.7 4.3 8.5 6.9 Product Properties Research ON clear 79.6 87.7 97.2
90.0 90.6 77.0 Gravity, .degree.API 52.0 53.0 46.6 51.4 50.2 53.7
Sulfur, wt.% .004 .003 .006 .003 .12 <.001 PONA Analysis,
C.sub.6 +, vol.% Paraffins 37.7 30.5 16.8 28.4 27.0 40.6 Naphthene
13.2 6.9 1.9 5.7 6.7 17.5 Olefins 2.4 1.9 1.3 1.1 2.3 3.5 Aromatics
46.8 60.8 80.1 64.0 64.1 38.4
__________________________________________________________________________
*Includes conversion to coke
Conversion increases with temperature and with catalyst-to-oil
ration as shown by the above tables. Catalyst B was more active
than catalyst A. Catalyst C had an activity much like that of
Catalyst A.
Octane improvements are greater on cracking virgin naphtha than on
recracking gasoline product of gas oil cracking, but octane numbers
reached are not as high. At conversion levels of 25 to 55%, the
clear octane number was raised from 41 for the change to products
in the 70 to 90 octane number range.
TABLE VIII
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THE EFFECT OF CRACKING HEAVY VIRGIN NAPHTHA ON VOLATILITY DENSE BED
UNIT Example Charge XII XIII XIV XV XVI XVII XVIII XIV Conversion 0
23.7 25.7 34.4 25.2 45.7 54.2 24.1 17.6 Simulated Distillation,
.degree.F IBP 239 -- 52 78 77 -- -- 20 21 10% 264 -- 175 213 229 --
-- 163 232 30% 295 -- 257 258 267 -- -- 253 275 50% 317 -- 286 288
293 -- -- 281 298 70% 391 -- 319 317 322 -- -- 308 323 90% 367 --
354 353 354 -- -- 338 351 95% 384 -- 392 393 383 -- -- 353 364 E.P.
485 -- 597 599 577 -- -- 492 516
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TABLE IX
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THE EFFECT OF CRACKING HEAVY VIRGIN NAPHTHA IN VOLATILITY RISER
PILOT PLANT
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Example Charge XX XXI XXII XIII XXIV XXV Conversion 0 27.2 33.6
49.2 32.5 32.7 20.9 ASTM Distillation,.degree.F. IBP 278 88 83 83
89 86 87 10% 288 149 128 130 133 122 156 30% 296 249 217 224 219
187 267 50% 303 284 281 273 277 276 288 70% 315 309 309 301 306 366
310 90% 334 363 421 499 369 553 351 95% -- 508 -- -- -- -- -- E.P.
352 522 528 576 522 640 455
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Cracking virgin naphtha improves volatility. The 50% point is
lowered 20.degree. to 40.degree.F. at conversions of 25 to 50%. The
90% point is reduced in the dense fluid catalyst bed examples
(Table VIII). In the riser pilot plant test unit (Table IX) there
was some formation of heavy ends and redistillation to remove about
5 to 8 weight percent of heavy ends is needed to maintain the 90%
point and the end point. There is a very minor formation of heavy
ends in the dense fluid catalyst bed examples which raised the
endpoint.
As obtained in recracking of cracked gasoline, liquid cracked
products from virgin naphtha have a very low olefin content.
Cracking of virgin naphtha produces very large yields of isobutane,
butene and propene, all valuable alkylation charge stocks. Cracking
is especially selective to isobutane over these zeolite catalysts
at low temperatures. At 850.degree.F. and 10 catalyst-to-oil ratio
there is a yield of 15% isobutane at 24% conversion. Lower
temperatures, such as 800.degree.F., and higher catalyst-to-oil
ratios, such as 20, will increase isobutane yield.
As for (FCC) fluid catalytic cracking gasoline recracking, ZSM-5
catalysts produce very large yields of light olefins (C.sub.2.sup.=
.about.C.sub.4.sup.=). These light olefins could be alkylated for
blending into gasoline pool or separated for petrochemical
manufacturing.
The amounts of potential alkylate produced are so large that, at
moderate conversions, combined yields of recracked gasoline and
alkylate, adjusted to gasoline vapor pressure, approach 100% of the
volume of the charge. Table X reports the combined yields of
recracked gasoline plus alkylate.
TABLE X
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RECOVERY OF TOTAL LIQUID PRODUCT FROM HEAVY FCC GASOLINE C.sub.5 +
Liquid Potential Outside Total Liq. Vol.% of Chg. Alkylate
i-C.sub.4 to be Recovery Vol.% of Chg. purchased Vol.% of Chg.
Vol.% of Chg.
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Ex. I 82.6 17.2 7.5 99.8 II 75.5 16.1 5.0 91.6 III 81.9 9.9 -.2
91.8 IV 80.7 10.6 -.1 91.3 V 80.7 27.5 17.1 108.2 VI 84.0 15.1 3.8
99.1 VII 75.8 12.8 -.6 88.6 VIII 66.4 10.2 -4.7 76.6 IX 42.5 13.7
3.6 56.2 X 74.3 12.0 -1.3 86.3 XI 86.7 15.7 6.1 102.4
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In summary, it has been demonstrated that a significant benefit can
be achieved by cracking heavy virgin naphtha over fluid catalytic
cracking catalysts since a relatively high octane gasoline product
was obtained which can be blended in a gasoline pool and high
yields of isobutane are obtained which can be alkylated to produce
a high octane blending stock.
BRIEF DESCRIPTION OF THE DRAWINGS
FIG. I is a diagrammatic sketch in elevation of one arrangement of
apparatus and system for practicing the separate riser recracking
of gasoline product of gas oil cracking in another separate riser
reactor with related product recovery equipment.
FIG. II is a diagrammatic sketch in elevation of another
arrangement for practicing the concepts identified with FIG. I
except a dense fluid bed of catalyst is relied upon to carry out
the recracking of gasoline product of gas oil cracking alone or
along with a virgin naphtha fraction.
FIG. III is a diagrammatic sketch of a further embodiment and
arrangement of reactor systems sequentially connected for
practicing the concepts of the invention wherein gasoline product
of gas oil cracking is recracked in a dense fluid catalyst bed
reaction zone with freshly regenerated catalyst, the catalyst used
for gasoline recracking is then used for gas oil cracking in a
riser cracking zone and catalyst separated from the riser cracking
operation is relied upon for effecting cracking of virgin
naphtha.
Referring now to FIG. I there is shown a gas oil riser cracking
reactor with a product fractionation step in combination with a
separate heavy naphtha or gasoline product of gas oil cracking
operation in a separate riser reactor with its own independent
product recovery system. The recracking operation recovery system
is related to the primary fractionation system in a manner to
recover a common gasoline product stream of improved octane rating
and a common light cycle oil stream. However, it is also
contemplated using a single riser reactor system wherein the
gasoline or naphtha to be subjected to recracking initially
contacts the hot freshly regenerated catalyst introduced to the
riser and the gas oil charge is then introduced to a downstream
portion of the riser reactor for cracking thereof to gasoline
boiling product. In the arrangement of FIG. I, a regenerator 2 is
shown containing a bed of catalyst 4 which is contacted with
oxygen-containing regeneration gas such as air introduced by
conduit 6 to an air distributor grid 8. Cyclone separators 10
provided with diplegs 12 are located in the upper portion of the
regenerator for separating flue gases from entrained catalyst
particles. The separated catalyst particles are returned by the
diplegs 12 to the catalyst bed 4 and flue gases are removed as by
conduit 14. Regenerated catalyst is removed from bed 4 as by
withdrawal well 16 and conveyed to conduit 18 communicating with
the lower end of riser 20. A catalyst flow control valve 22 is
provided in conduit 18. Regenerated catalyst is also conveyed from
well 16 by conduit 24 to the bottom portion of riser 26. Catalyst
flow control valve 28 is provided in conduit 24. A gas oil feed
boiling in the range of 650.degree.F. to about 1000.degree.F. is
introduced by conduit 30 to the bottom portion or riser 26 for
admixture with hot regenerated catalyst introduced by conduit 24. A
catalyst-oil suspension is thus formed providing a temperature of
at least about 950.degree.F. and more usually in the range of
1000.degree.F. up to about 1100.degree.F. which is then passed
upwardly through the riser reactor 26 at a velocity to provide a
hydrocarbon residence time therein within the range of about 1
second up to about 10 seconds. During passage of the suspension
through the riser conversion of the gas oil feed to lower and
higher boiling products occurs. These products are separated after
removal of catalyst therefrom in a product fractionator as
discussed below. The catalyst-hydrocarbon suspension after
traversing the riser reactor is caused to flow directly into a
plurality of cyclonic separators 32 attached to the end of the
riser through a T-connection. Diplegs 34 attached to separators 32
pass separated catalyst to an annular stripping zone 36 provided
with baffles 38. Stripping gas such as steam is introduced to the
lower portion of the stripping zone by conduit 40. Stripped
catalyst is removed from the lower portion of the stripping zone by
conduit 42 and conveyed to the bed of catalyst 4 in the
regeneration zone. A flow control valve 44 is provided in conduit
42.
Stripping gas and stripped hydrocarbon material are removed from
the bed of catalyst in the stripping zone and enters cyclone
separator 46 wherein entrained catalyst particles are separated
from the stripping gas. Separated catalyst particles are returned
to the catalyst bed by dipleg 48. Stripping gas and hydrocarbon
material are then passed from separator 46 by a connecting conduit
to a plenum chamber 52. Hydrocarbon material separated from the
riser reactor 26 by separators 32 pass by connecting conduit 50 to
plenum chamber 52. Hydrocarbon material and stripping gas are
passed from chamber 52 by conduit 54 to a fractionator 56.
For the purpose of this discussion, fractionator 56 is relied upon
to separate a heavy cycle oil (HCO) withdrawn by conduit 58; a
light cycle oil (LCO) withdrawn by conduit 60; a heavy naphtha
withdrawn by conduit 62; material boiling below the heavy naphtha
withdrawn by conduit 64 and a bottoms fraction withdrawn by conduit
66. All or a portion of the bottoms fraction may be passed through
heater 66 and returned to the bottom of the tower 56 by conduit 68.
Generally the temperature of the bottom of the tower will be about
690.degree.F. The material boiling below the heavy naphtha fraction
and withdrawn from the fractionator by conduit 64 is passed through
cooler 70 and thence by conduit 72 to drum 74 maintained at a
temperature of about 100.degree.F. In drum 74 a liquid condensate
is recovered and recycled by conduit 76 to the upper portion of the
fractionator 56 as reflux. Uncondensed product is withdrawn from
drum 74 by conduit 78 and passed to cooler 80 and conduit 82 to
drum 84 maintained at a temperature of about 100.degree.F.
The heavy naphtha separated in fractionator 56 and withdrawn by
conduit may be passed all or in part by conduit 86 to the inlet of
riser reactor 20 where it is combined with hot regenerated catalyst
introduced by conduit 18 to form a suspension at a temperature
within the range of 950.degree.F. to about 1250.degree.F. When
virgin naphtha is used as the hydrocarbon feed to riser 20 instead
of the heavy naphtha fraction it may be introduced by conduit 88.
The suspension formed in the bottom of the riser 20 is passed
upwardly therethrough under conditions to provide a hydrocarbon
residence time in the range of 1 to about 10 seconds before
separating the suspension into a catalyst phase and a hydrocarbon
phase in cyclone separator 90. The catalyst phase separated in
cyclone 90 is passed by dipleg 92 to the bed of catalyst in the
stripping zone as above discussed. To complete the separation of
catalyst particles from hydrocarbon products of cracking the
hydrocarbon phase is removed from separator 90 by conduit 94 and
passed to a second separator 96. Catalyst separated in separator 96
is passed by dipleg 98 and 92 to the catalyst bed being stripped.
Hydrocarbon vapors are recovered from separator 96 and conveyed by
conduit 100 to cooler 102 wherein the vapors may or may not be
cooled. The vapors then pass by conduit 104 to tower 106 maintained
at a bottom temperature of about 550.degree.F. and a top
temperature of about 350.degree.F. A bottoms product boiling in the
light cycle oil boiling range is withdrawn from the bottom of the
tower by conduit 108 and combined with light cycle oil in conduit
60 withdrawn from fractionator 56. An overhead hydrocarbon portion
is withdrawn from tower 106 by conduit 110 and combined with
hydrocarbon material in conduit 78. Condensate material comprising
gasoline boiling range material is withdrawn from drum 84 by
conduit 112 and recycled in part by conduit 114 to tower 106 as
reflux. The remaining gasoline boiling condensate material is
recovered by conduit 116. Uncondensed vaporous material is
withdrawn from drum 84 by conduit 118 and sent to, for example, the
refinery gas plant.
The processing arrangement of the present invention contemplates
injecting the heavy naphtha to be recracked at the base of riser
reactor 26 and introducing gas oil to be cracked to a downstream
portion of the riser by either or both of inlet conduits 120 or
122. In such an arrangement it is contemplated cracking a virgin
naphtha in riser 20 in combination with cracking gas oil alone or
in combination with heavy naphtha cracking in riser 26.
Referring now to FIG. II, there is shown diagrammatically in
elevation an embodiment of the reactor arrangement of FIG. I in
which the heavy naphtha fraction is recracked in a riser
discharging into the bottom of a dense fluid bed of catalyst and
deactivated catalyst separated from the dense fluid bed cracking
step is combined with catalyst separated from the gas oil cracking
operation and passed through the catalyst stripping zone. In the
arrangement of FIG. II, a heavy naphtha fraction separated from the
product of gas oil cracking as shown in FIG. I is introduced by
conduit 1 to the bottom of riser 3 for admixture with freshly
regenerated catalyst introduced by conduit 5 containing catalyst
flow control valve 7. A suspension is formed in the lower portion
of riser 3 at an elevated cracking temperature. The suspension
passes upwardly through riser 3 and is discharged into the bottom
of an enlarged zone 9 containing a dense fluid bed of catalyst 11.
Cracking of the gasoline fraction is accomplished in the dense
fluid catalyst bed 11. Hydrocarbon vapors comprising the recracked
gasoline vapors are passed through one or more cyclone separators
13 provided with catalyst dipleg 15. The cracked gasoline vapors
are withdrawn by conduit 100 and passed to product separation as
defined with respect to FIG. I.
In the arrangement of FIG. II, catalyst is withdrawn from the upper
surface of fluid bed 11 into a well 17 defined by baffle 19. The
catalyst is withdrawn from well 17 by conduit 21 provided with a
flow central valve 23. The riser reactor 27 provided for converting
a gas oil feed introduced by conduit 25 is intended to be operated
in the same manner as described with respect to FIG. I for riser
26. Thus in the arrangement of FIG. II, the hydrocarbon product of
gas oil cracking and stripping gas are withdrawn from the top of
the vessel by conduit 54 for separation in a manner similar to that
described with respect to FIG. I.
FIG. III departs from the arrangements of either FIG. I or II by
the combination of cascading regenerated catalyst first through a
dense fluid catalyst bed gasoline recracking zone, then a gas oil
riser cracking zone and the catalyst separated from the gas oil
cracking operation and collected as a dense fluid bed of catalyst
is then relied upon to crack virgin naphtha prior to the catalyst
passing to a stripping zone. The stripped catalyst is then passed
to a catalyst regeneration operation. The hydrocarbon products of
the gasoline recracking step and the separate gas oil cracking step
are recovered in a manner similar to that described with respect to
FIG. I. Products of virgin naphtha cracking are recovered in the
gas oil separation system. In the arrangement of FIG. III, a heavy
cracked gasoline fraction is introduced by conduit 31 for admixture
with hot regenerated catalyst withdrawn from a catalyst
regeneration zone by conduit 33 provided with a catalyst flow
control valve 35. A suspension is formed at an elevated temperature
of at least 1000.degree.F. which is conveyed by inlet conduit 37
into the bottom portion of a dense fluid bed of catalyst 39
confined in a cracking zone 41. Hydrocarbon product of gasoline
cracking is passed through cyclonic separation means 43 provided
with separated catalyst dipleg 45. Hydrocarbon vapors are withdrawn
by conduit 100 and passed to product separation similar to that
described in FIG. I. Catalyst is withdrawn into a well 47 from an
upper portion of bed 39 and conveyed therefrom by conduit 49
provided with valve 51 to the bottom portion of gas oil cracking
riser 53 to which a gas oil feed is introduced by conduit 55. A
suspension at an elevated gas oil cracking temperature is formed in
the lower portion of riser 53 and passes upwardly therethrough for
discharge into cyclonic separation zones 57 and 59. Hydrocarbon
vapors separated in zones 57 and 59 are withdrawn by conduits 61
and 63 communicating with chamber 65 and withdrawal conduit 54.
Catalyst particles separated by cyclonic means 57 and 59 are
conveyed to a dense fluid bed of catalyst 67 by diplegs 69 and 71.
A virgin naphtha fraction is introduced to the collected bed of
catalyst discharged from riser 53 by conduit 73 for conversion
thereof to higher octane product and olefin constituents as
described above. Catalyst bed 67 is a continuous downwardly moving
bed of catalyst which passes into a stripping zone beneath the
virgin naphtha inlet distributor grid. The catalyst thus
sequentially used as above identified passes downwardly through a
stripping zone 75 provided with stripping gas introduced by conduit
77. Stripped catalyst is then withdrawn by conduit 79 for transfer
to a catalyst regeneration operation. Hydrocarbon vapor product of
virgin naphtha cracking and stripping gas pass through cyclonic
separation zones 81 and 83 wherein entrained catalyst is separated
and returned to the catalyst bed 67 by diplegs 85 and 87.
Hydrocarbon vapors and stripping gas then pass by conduit 89 and 91
to chamber 65. Vaporous material withdrawn by conduit 54 is then
separated in the manner described with respect to FIG. I.
Catalyst withdrawn from the stripping zone by conduit 79 provided
with valve 97 is combined with regeneration gas 93 in the lower
portion of a riser regenerator 95. The suspension thus formed
passes upwardly through the riser regeneration zone and is
discharged into an enlarged separation-regeneration zone 99 and
above a dense fluid bed of catalyst 101 in the lower portion
thereof. Additional regeneration gas is introduced to a lower
portion of the regeneration zone by conduit 103. In the riser and
dense catalyst phase regeneration zones, carbonaceous material
deposited during hydrocarbon conversion is removed by burning with
the introduced oxygen containing regeneration gas. Gaseous products
of combustion pass through separators 103 and 105 wherein entrained
catalyst fines are removed from the flue gases. Separated catalyst
fines are returned to the catalyst bed 101 by diplegs 107 and 109.
Flue gases then pass by conduits 111 and 113 into chamber 115 from
which they are withdrawn by conduit 117.
In the arrangement of FIG. III, it is preferred that the catalyst
employed be a mixture of crystalline zeolite conversion catalyst of
small and large pore diameter crystalline materials and that the
small pore crystalline material be of the ZSM-5 type. The large
pore crystalline zeolite may be either of the X or Y type of
crystalline zeolite.
Having thus provided a general discussion of this invention and
provided specific embodiments going to the very essence thereof, it
is to be understood that no undue restrictions are to be imposed by
reason thereof except as defined by the following claims.
* * * * *