U.S. patent number 3,655,551 [Application Number 05/042,053] was granted by the patent office on 1972-04-11 for hydrocracking-hydrogenation process.
This patent grant is currently assigned to Union Oil Company of California. Invention is credited to Robert H. Hass, Paul F. Helfrey, Nicholas L. Kay.
United States Patent |
3,655,551 |
Hass , et al. |
April 11, 1972 |
HYDROCRACKING-HYDROGENATION PROCESS
Abstract
A cyclic hydrocracking-hydrogenating process comprising a
gasoline-producing cycle, passing a mineral oil feedstock through a
first and second catalyst contact zone to effect a substantial
synthesis of gasoline; and in a middle distillate producing cycle,
passing the feedstock through the first zone to effect a
substantial synthesis to middle distillate and passing unconverted
oil and middle distillate from the first contacting zone in the
substantial absence of hydrogen sulfide through the second
contacting zone to substantially hydrogenate without substantially
hydrocracking and unconverted oil and middle distillate; the
catalyst in the second contacting zone comprising a Group VIII
noble metal hydrogenation component support on an active zeolitic
cracking base.
Inventors: |
Hass; Robert H. (Fullerton,
CA), Helfrey; Paul F. (Whittier, CA), Kay; Nicholas
L. (Fullerton, CA) |
Assignee: |
Union Oil Company of California
(Los Angeles, CA)
|
Family
ID: |
21919813 |
Appl.
No.: |
05/042,053 |
Filed: |
June 1, 1970 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
Issue Date |
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792619 |
Jan 21, 1969 |
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592482 |
Nov 7, 1966 |
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Current U.S.
Class: |
208/59; 208/58;
208/89; 208/143; 208/144; 208/217; 208/254H; 208/111.15;
208/111.35 |
Current CPC
Class: |
C10G
65/10 (20130101); C10G 65/12 (20130101); C10G
47/18 (20130101) |
Current International
Class: |
C10G
65/12 (20060101); C10G 65/10 (20060101); C10G
47/18 (20060101); C10G 65/00 (20060101); C10G
47/00 (20060101); C10g 013/02 (); C07c 005/02 ();
B01j 011/40 () |
Field of
Search: |
;208/59,143,144 |
References Cited
[Referenced By]
U.S. Patent Documents
Primary Examiner: Gantz; Delbert E.
Assistant Examiner: Schmitkons; G. E.
Parent Case Text
RELATED APPLICATIONS
This application is a continuation-in-part of copending
application, Ser. No. 792,619, filed Jan. 21, 1969, now abandoned
which in turn is a continuation-in-part of Ser. No. 592,482, filed
Nov. 7, 1966, now abandoned.
Claims
We claim:
1. In a catalytic hydrocracking-hydrogenation system, a process for
converting a mineral oil feedstock containing aromatic hydrocarbons
and boiling above the gasoline range alternately to a relatively
aromatic gasoline product, and to a plurality of products including
an aromatic gasoline and a substantial proportion of a relatively
non-aromatic middle-distillate product boiling at least partially
above the gasoline range, said feedstock also comprising a heavy
fraction boiling above the end-boiling-point of said middle
distillate product, which comprises:
1. in a periodic gasoline-producing cycle, passing said feedstock
in admixture with hydrogen through a first catalyst contacting zone
in contact with an active hydrocracking catalyst at an elevated
pressure and a temperature above about 625.degree. F. to effect a
substantial synthesis of gasoline, separating the resulting
effluent into an aromatic gasoline product and an unconverted
higher boiling fraction, passing said higher boiling fraction
through a second catalyst contacting zone at an elevated pressure
and a temperature above about 500.degree. F. in the presence of
hydrogen, hydrogen sulfide and a hydrocracking catalyst comprising
a Group VIII noble metal hydrogenation component supported on an
active zeolite cracking base wherein at least half of the original
zeolitic sodium ions have been replaced by hydrogen ions,
polyvalent metal ions, decationized sites, or a combination
thereof, to thereby effect a substantial synthesis of gasoline,
separating the resulting second-stage effluent into an aromatic
gasoline fraction and an unconverted recycle oil, and recycling at
least a portion of said recycle oil to said first and/or second
contacting zones;
2. in an alternating periodic middle distillate-producing cycle,
passing said feedstock in admixture with hydrogen through said
first catalyst contacting zone as defined in step (1), separating
the resulting effluent into an aromatic gasoline product and a
relatively aromatic bottoms fraction comprising middle distillate
and a heavier fraction, passing at least the middle distillate
portion of said bottoms fraction plus added hydrogen through said
second catalyst contacting zone at an elevated pressure
substantially in the absence of hydrogen sulfide, and at a
temperature correlated with space velocity so as to saturate at
least about one-third of the aromatic hydrocarbons in said middle
distillate portion but to substantially avoid cracking reactions,
recovering a relatively non-aromatic middle distillate product from
the resulting effluent, and recycling to said first catalytic zone
a heavier fraction recovered from the effluent from said first
and/or second catalytic contacting zones.
2. A process as defined in claim 1 wherein the catalyst employed in
said second contacting zone is a molecular sieve zeolite having a
SiO.sub.2 /Al.sub.2 O.sub.3 mole-ratio above about 3, at least
about 20% of the ion-exchange capacity thereof being satisfied by
hydrogen ions, and deposited thereon between about 0.1% and 3% by
weight of palladium or platinum.
3. A process as defined in claim 1 wherein the entire bottoms
fraction recovered from the effluent from said first contacting
zone in step (2) is passed through said second contacting zone, and
wherein the effluent from said second contacting zone in step (2)
is fractionated to recover said non-aromatic middle distillate
product and said heavier fraction, and wherein said heavier
fraction is recycled to said first contacting zone.
4. A process as defined in claim 1 wherein said bottoms fraction
recovered from the effluent from said first contacting zone in step
(2) is further fractionated to separate said middle distillate from
said heavier fraction, and wherein the separated heavier fraction
is recycled to said first contacting zone.
5. A process as defined in claim 1 wherein said relatively
non-aromatic middle distillate product is a jet fuel boiling in a
range of about 350.degree. -530.degree. F.
6. A process as defined in claim 1 wherein a temperature between
about 200.degree. and 500.degree. F. is maintained during step (2)
in said second contacting zone.
7. In a catalytic hydrofining-hydrocracking-hydrogenation system, a
process for converting a mineral oil feedstock boiling above the
gasoline range alternately to a relatively aromatic gasoline
product, and to a plurality of products including an aromatic
gasoline and a relatively non-aromatic middle distillate product
boiling at least partially above the gasoline range, said feedstock
also comprising a heavy fraction boiling above the
end-boiling-point of said middle distillate product, which
comprises:
1. passing said feedstock plus added hydrogen through a catalytic
hydrofining zone at an elevated temperature and pressure to effect
decomposition of organic sulfur and/or nitrogen compounds contained
therein without substantial cracking of hydrocarbons;
2. in a periodic gasoline-producing cycle:
a. passing effluent from said hydrofining zone, without intervening
purification to remove ammonia and hydrogen sulfide, through a
second catalytic contacting zone at an elevated pressure and a
temperature between about 625.degree. and 850.degree. F., in
contact with a hydrocracking catalyst comprising a minor proportion
of a Group VIII metal and/or an oxide or sulfide thereof supported
on a crystalline zeolitic cracking base wherein the zeolitic
cations are predominately hydrogen ions and/or polyvalent metal
ions, to effect about 30-80 volume-percent conversion per pass to
gasoline-boiling range material,
b. separating the effluent therefrom into an aromatic gasoline
product, and a bottoms fraction comprising middle distillate and a
heavier fraction,
c. passing said bottoms fraction through a third catalytic
contacting zone in admixture with hydrogen and hydrogen sulfide at
an elevated pressure and a temperature between about 500.degree.
and 850.degree. F. in the presence of a hydrocracking catalyst as
defined in step 2(a), wherein the Group VIII metal is a noble
metal,
d. separating effluent from said third contacting zone into an
aromatic gasoline product and an unconverted recycle oil, and
e. recycling said unconverted recycle oil to said second and/or
third contacting zone; and
3. in an alternating periodic middle distillate producing
cycle:
a. continuing the operation of said hydrofining step (1) and said
second contacting zone as defined in steps (2)(a) and (2)(b),
b. passing at least the middle distillate portion of said bottoms
fraction through said third catalytic contacting zone in admixture
with hydrogen, but substantially in the absence of hydrogen
sulfide, and at a temperature correlated with space velocity so as
to saturate at least about one-third of the aromatic hydrocarbons
in said middle distillate portion, but to substantially avoid
cracking reactions,
c. recovering said relatively non-aromatic middle distillate
product from the effluent of said third contacting zone; and
d. recycling to said second contacting zone the heavier fraction
boiling above said middle-distillate product, said heavier fraction
being recovered from the effluent from said second and/or third
contacting zones.
8. A process as defined in claim 7 wherein the catalyst employed in
said third contacting zone is a molecular sieve zeolite having a
SiO.sub.2 /Al.sub.2 O.sub.2 mole-ratio above about 3, at least
about 20% of the ion-exchange capacity thereof being satisfied by
hydrogen ions, and deposited thereon between about 0.1 and 3% by
weight of palladium or platinum.
9. A process as defined in claim 7 wherein the entire bottoms
fraction recovered from the effluent from said second contacting
zone in step (3)(a) is passed through said third contacting zone,
and wherein the effluent from said third contacting zone in step
(3) is fractionated to recover said non-aromatic middle distillate
product and said heavier fraction, and wherein said heavier
fraction is recycled to said second contacting zone.
10. A process as defined in claim 7 wherein said relatively
non-aromatic middle distillate product is a jet fuel boiling in the
range of about 350.degree. -530.degree. F.
11. A process as defined in claim 7 wherein the bottoms fraction
recovered from the effluent from said second contacting zone in
step (3)(a) is further fractionated to separate said middle
distillate from said heavier fraction, and wherein the separated
heavier fraction is recycled to said second contacting zone.
12. A process as defined in claim 7 wherein a temperature between
about 200.degree. and 500.degree. F. is maintained during step
(3)(b) in said third contacting zone.
13. In a catalytic hydrocracking-hydrogenation system, a process
for converting a mineral oil feedstock containing aromatic
hydrocarbons and boiling above the gasoline range alternately to a
relatively aromatic gasoline product, and to a plurality of
products including an aromatic gasoline and a substantial
proportion of a relatively non-aromatic middle-distillate product
boiling at least partially above the gasoline range, which
comprises:
1. In a periodic gasoline-producing cycle, passing said feedstock
in admixture with hydrogen through a first catalyst contacting zone
in contact with an active hydrocracking catalyst at an elevated
pressure and a temperature above about 625.degree. F. to effect a
substantial synthesis of gasoline, separating the resulting
effluent into an aromatic gasoline product and an unconverted
higher boiling fraction, passing said higher boiling fraction
through a second catalyst contacting zone at an elevated pressure
and a temperature above about 500.degree. F. in the presence of
hydrogen, hydrogen sulfide and a hydrocracking catalyst comprising
a Group VIII noble metal hydrogenation component supported on an
active zeolite cracking base wherein at least half of the original
zeolitic sodium ions have been replaced by hydrogen ions,
polyvalent metal ions, decationized sites, or a combination
thereof, to thereby effect a substantial synthesis of gasoline,
separating the resulting second-stage effluent into an aromatic
gasoline fraction and an unconverted recycle oil, and recycling at
least a portion of said recycle oil to said first and/or second
contacting zones;
2. In an alternating periodic middle distillate-producing cycle,
passing said feedstock in admixture with hydrogen through said
first catalyst contacting zone as defined in step (1), separating
the resulting effluent into an aromatic gasoline product and a
relatively aromatic unconverted higher-boiling fraction, passing
said higher-boiling fraction plus added hydrogen through said
second catalyst contacting zone at an elevated pressure and
substantially in the absence of hydrogen sulfide and at a
substantially non-cracking temperature below about 450.degree. F.,
correlated with space velocity so as to saturate at least about
one-third of the aromatic hydrocarbons in said unconverted oil
which synthesizing less than about 5 volume-percent of C.sub.4
-400.degree. F. gasoline by hydrocracking, and recovering a
relatively non-aromatic middle distillate product from the
resulting effluent.
14. A process as defined in claim 13 wherein the catalyst employed
in said second contacting zone is a molecular sieve zeolite having
a SiO.sub.2 /Al.sub.2 O.sub.3 mole-ratio above about 3, at least
about 20% of the ion-exchange capacity thereof being satisfied by
hydrogen ions, and deposited thereon between about 0.1 and 3
percent by weight of palladium or platinum.
15. A process as defined in claim 13 wherein the effluent from said
second contacting zone in step (2) is fractionated to recover said
non-aromatic middle distillate product and a higher boiling recycle
fraction, and wherein said higher boiling recycle fraction is
recycled to said first contacting zone.
16. A process as defined in claim 13 wherein said relatively
non-aromatic middle distillate product is a jet fuel boiling in a
range of about 350.degree.-530.degree. F.
17. A process as defined in claim 13 wherein a temperature between
about 250.degree. and 450.degree. F. is maintained during step (2)
in said second contacting zone.
18. In a catalystic hydrofining-hydrocracking-hydrogenation system,
a process for converting a mineral oil feedstock boiling above the
gasoline range alternately to a relatively aromatic gasoline
product, and to a plurality of products including an aromatic
gasoline and a relatively non-aromatic middle distillate product
boiling at least partially above the gasoline range, which
comprises:
1. passing said feedstock plus added hydrogen through a catalytic
hydrofining zone at an elevated temperature and pressure to effect
decomposition of organic sulfur and/or nitrogen compounds contained
therein without substantial cracking of hydrocarbons;
2. in a periodic gasoline-producing cycle:
a. passing effluent from said hydrofining zone, without intervening
purification to remove ammonia and hydrogen sulfide, through a
second catalytic contacting zone at an elevated pressure and a
temperature between about 625.degree. and 850.degree. F., in
contact with a hydrocracking catalyst comprising a minor proportion
of a Group VIII metal and/or an oxide or sulfide thereof supported
on a crystalline zeolitic cracking base wherein at least half of
the original zeolitic sodium ions have been replaced by hydrogen
ions, polyvalent metal ions, decationized sites, or a combination
thereof, to thereby effect about 30-80 volume-percent conversion
per pass to gasoline-boiling range material,
b. separating the effluent therefrom into an aromatic gasoline
product, and an unconverted higher boiling fraction,
c. passing said unconverted higher boiling fraction through a third
catalytic contacting zone in admixture with hydrogen and hydrogen
sulfide at an elevated pressure and a temperature between about
500.degree. and 850.degree. F. in the presence of a hydrocracking
catalyst as defined in step 2-(a), wherein the Group VIII metal is
a noble metal,
d. separating effluent from said third contacting zone into an
aromatic gasoline product and an unconverted recycle oil, and
e. recycling said unconverted recycle oil to said second and/or
third contacting zone; and
3. in an alternating periodic middle distillate producing
cycle:
a. continuing the operating of said hydrofining step (1) and said
second contacting zone as defined in steps (2)-(a) and (2)-(b),
b. passing said unconverted higher boiling fraction through said
third catalytic contacting zone in admixture with hydrogen, but
substantially in the absence of hydrogen sulfide, and at a
substantially non-cracking temperature between about 250.degree.
and 450.degree. F., correlated with space velocities above about
1.0 so as to saturate at least about one-third of the aromatic
hydrocarbons in said unconverted oil while synthesizing less than
about 5 volume-percent of C.sub.4 -400.degree. F. gasoline by
hydrocracking, the conversion to hydrocarbons boiling below
530.degree. F. being less than about 15 volume-percent; and
c. recovering said relatively non-aromatic middle distillate
product from the effluent of said third contacting zone.
19. A process as defined in claim 18 wherein the catalyst employed
in said third contacting zone is a molecular sieve zeolite having a
SiO.sub.2 /Al.sub.2 O.sub.3 mole-ratio above about 3, at least
about 20% of the ion-exchange capacity thereof being satisfied by
hydrogen ions, and deposited thereon between about 0.1 and 3% by
weight of palladium or platinum.
20. A process as defined in claim 18 wherein the effluent from said
third contacting zone in step (3) is fractionated to recover said
non-aromatic middle distillate product, and a higher boiling
recycle fraction, and wherein said higher boiling fraction is
recycled to said second contacting zone.
21. A process as defined in claim 18 wherein said relatively
non-aromatic middle distillate product is a jet fuel boiling in a
range of about 350.degree. - 530.degree. F.
22. A process as defined in claim 18 wherein the aromatic content
of the unconverted oil fed to said third contacting zone is above
30 volume-percent and the aromatic content of the middle distillate
product is below 20 volume percent.
23. A method for the hydrogenation of aromatic hydrocarbons in a
mineral oil feedstock, which comprises contacting said feedstock
plus added hydrogen, but substantially in the absence of hydrogen
sulfide, with a hydrocracking catalyst at an elevated pressure and
at a substantially non-cracking temperature between about
250.degree. and 425.degree. F., said temperature being correlated
with space velocity so as to saturate at least about one-third of
the aromatic hydrocarbons in said feedstock while synthesizing less
than about 5 volume-percent of C.sub.4 -400.degree. F. gasoline by
hydrocracking, said hydrocracking catalyst comprising a minor
proportion of a Group VIII noble metal supported on a crystalline
zeolite cracking base wherein at least half of the original
zeolitic sodium ions have been replaced by hydrogen ions,
polyvalent metal ions, decationized sites, or a combination
thereof.
24. A process as defined in claim 23 wherein said Group VIII noble
metal is palladium.
25. A process as defined in claim 23 wherein the conversion of
middle distillate is less than 10 volume percent.
Description
BACKGROUND AND SUMMARY OF INVENTION
This invention relates to catalytic hydrocracking, and more
particularly is concerned with a two-stage process wherein the
first stage converts fresh feed at relatively high temperatures to
high-octane gasoline, and unconverted oil from the first stage is
treated in the second stage which is operated alternately as either
(1) a hydrocracker for producing additional gasoline, or (2) a
non-cracking hydrogenator for producing a relatively saturated
middle distillate product such as jet fuel and/or diesel fuel.
Basic novel features of the process consist in operating the second
stage during the non-cracking hydrogenation cycle substantially in
the absence of hydrogen sulfide, and at very low temperatures such
that the average molecular weight of the feed thereto is not
substantially reduced, and in recycling any hydrocarbon effluent
therefrom boiling above the desired middle distillate product range
to the first stage for additional hydrocracking. During this cycle
of operation, there is substantially no production of gasoline by
hydrocracking in the second stage, which is advantageous in that
the gasoline produced under low-temperature hydrocracking
conditions is of very poor quality.
To convert the process to maximum gasoline production, process
conditions, principally temperature, are adjusted in the second
stage whereby the cracking activity of the catalyst becomes
operative and there is a resultant substantial synthesis of
gasoline therein, with unconverted oil therefrom being recycled
either to the first stage or the second stage. Further, when the
second stage is operated under these cracking conditions, it is
preferred to reduce the hydrogenation activity of the catalyst as
e.g., by adding a reversible poison to the feed such as hydrogen
sulfide, and/or to operate at even higher hydrocracking
temperatures while avoiding over-cracking by adjusting space
velocity upwardly and/or reducing the cracking activity of the
catalyst by adding reversible poison such as ammonia to the feed.
Reducing hydrogenation activity and/or raising hydrocracking
temperatures is required in order to obtain a high-octane, aromatic
gasoline from the second stage. It will be understood that the
first stage of the process operates throughout under these optimum
gasoline-producing conditions.
A principal object of the invention is to provide an integrated
hydrocracking process designed mainly for the production of
gasoline, but which can be easily regulated to produce substantial
yields, as the seasonal demands of the market may require, of high
quality jet fuel boiling for example in the 350.degree.
-530.degree. F. range, and/or a high quality diesel fuel boiling
for example in the 400.degree. - 650.degree. F. range. A further
objective is to minimize the total reforming capacity required to
produce the desired quantity of high-octane gasoline. A specific
object of the invention is to provide an integrated
hydrocracking-hydrogenation process of the nature described wherein
maximum gasoline quality is achieved by substantially eliminating
gasoline production under jet fuel hydrogenation conditions, and
wherein the aromaticity of the jet fuel product is also readily
controllable. Other objects will be apparent from the more detailed
description which follows.
The process of this invention, in its basic objectives and in some
of its physical features, bears a substantial similarity to the
process described in U.S. Pat. No. 3,132,090. The process of the
patent also embraces a two-stage hydrocracking system wherein the
first stage is operated at relatively high temperatures for the
production of gasoline, and wherein the second stage is operated
alternately (A) with added sulfur at relatively high temperatures
for gasoline production, or (B) at relatively low temperatures in
the absence of sulfur for jet fuel production. However, the
patented process always envisages substantial hydrocracking in the
second stage, both in the (A) and (B) cycles of operation, as is
clearly evident from the fact that any unconverted oil from the
second stage which is not desired as product is always recycled
back to the second stage. Such an operation obviously could not be
maintained unless substantial hydrocracking, i.e., molecular weight
reduction, is taking place in the second stage.
In the process of the above patent, it is not possible to maintain
a significant conversion of high boiling hydrocarbons to lower
molecular weight jet fuel hydrocarbons in the (B) cycle of
operation without also synthesizing substantial amounts of
gasoline, which, under the disclosed conditions, is highly
saturated and has a very poor octane value. Moreover, under the (B)
cycle hydrocracking conditions it is difficult or impossible to
control the degree of aromaticity of the jet fuel product, such
product always being substantially completed saturated. Complete
saturation of aromatic hydrocarbons in the jet fuel product is
seldom required in order to meet commercial specifications, and
represents an unnecessary and wasteful consumption of hydrogen. In
the process of this invention the first stage is operated at
relatively high temperatures in the presence of hydrogen sulfide
whereby high-octane aromatic gasoline is produced with a minimum of
hydrogen consumption. And in the non-cracking, hydrogenation cycle
designed for jet and/or diesel fuel production, the second stage
functions merely as a saturator for like-boiling-range material
derived from the first stage. Hydrogenation without cracking is
achieved by the conjoint effect of a sulfur-free environment and
the use of low temperatures, correlated with the space velocity to
avoid any substantial cracking. At the same time, these conditions
can be further adjusted to control kinetically the degree of
hydrogenation of the product. As noted, such kinetic control over
the extent of hydrogenation is substantially impossible under the
hydrocracking conditions of the patent.
Furthermore, in the process of the invention herein, the
unconverted oil from the second stage is saturated without being
substantially hydrocracked to gasoline and/or middle distillate
during the hydrogenation cycle and this oil may be recycled to the
first stage for conversion to jet fuel and/or diesel fuel. It has
been found that middle distillate yields and quality are improved
when the feed is largely saturated and therefore the process of the
invention allows the improvement in middle distillate yield and/or
quality of the recycling of the saturated unconverted oil to the
first stage. Prior art processes have always recycled unconverted
oil from the second stage back to the second stage with substantial
conversion occurring therein. In the process herein, unconverted
oil from the second stage is recycled to the first stage wherein
conversion to gasoline and/or middle distillate occurs. Hence,
substantial conversion to products need not occur in the second
stage and the second stage is utilized principally for saturating
the middle distillate and/or unconverted oil. The process of the
invention thereby attains the above-described advantages of higher
gasoline octane and improved middle distillate quality.
A critical feature of the invention resides in the nature of the
catalyst employed in the second hydrocracking zone. For maximum
efficiency in the hydrocracking cycle, it is desirable to employ a
catalyst comprising a highly active cracking base which will
effectively crack hydrocarbons at temperatures below about
700.degree. F. Such cracking bases include primarily the
crystalline zeolites, e.g., of the X, Y or L crystal types, wherein
the zeolitic cations are predominately hydrogen ions and/or
polyvalent metal ions. Yet it is precisely this type of cracking
base which would appear to be of most doubtful operability in the
hydrogenation cycle of the process where hydrogenation activity
must be maintained at temperatures below effective cracking
temperatures.
It has now been discovered that the Group VIII noble metals,
particularly palladium and platinum, when supported upon such
zeolite cracking bases, exhibit extraordinarily high hydrogenation
activity such that, if a substantially sulfur-free atmosphere is
maintained, effective hydrogenation of aromatic hydrocarbons can
occur at temperatures of 200.degree. - 500.degree. F., pressures of
500-3,000 pgis, and at space velocities of 0.5-20 and that those
conditions may be correlated such that little or no cracking of
hydrocarbons occurs. Further, it has been discovered that in the
hydrocracking cycle of the second stage, any one or more of several
process variables can be altered so as to effect substantial
hydrocracking with minimal hydrogen consumption and the production
of a high quality aromatic gasoline. It would be a simple matter to
initiate hydrocracking simply by raising the temperature, but this
alone has been found to result in substantially complete saturation
of aromatics and consumes much hydrogen. To avoid the latter
consequences, any one or more of the following procedures are
adopted:
1. Sufficient sulfur, as for example in the form of hydrogen
sulfide, is mixed with the second-stage feed to provide at least
about 0.01, and preferably at least 0.5, millimoles of hydrogen
sulfide per mole of hydrogen. Concomitantly, temperatures are
raised by about 50.degree.-150.degree. F. to achieve the desired
crack per pass. The added hydrogen sulfide appears to repress the
hydrogenation activity of the catalyst, and the temperature
elevation activates the cracking centers of the catalyst.
2. A basic nitrogen compound such as ammonia is added to the feed
in amounts between about 50 and 2,000 parts per million of nitrogen
based on feed, and temperatures are concomitantly raised about
100.degree.-350.degree. F. The ammonia appears to repress cracking
activity to such an extent that temperatures can be elevated to a
level at which hydrogenation is not thermodynamically favored.
3. Space velocity can be elevated to a level of e.g., 5 to 20, and
temperatures concomitantly raised about 50.degree.-200.degree. F.
Under these conditions, cracking rates tend to outstrip
hydrogenation rates with the result that an aromatic gasoline can
be produced with minimal hydrogen consumption.
Obviously, any combination of two or more of the foregoing
alternates may be adopted, with the resultant hydrocracking
temperature being in all cases adjusted to maintain a predetermined
conversion per pass to gasoline, normally about 30-70 volume
percent based on feed. Alternates (1) and (3), or a combination of
both are normally preferred.
Another critical feature of the process resides in the use of an
initial feedstock which is substantially aromatic in character.
Suitable feeds include coker distillate gas oils, cycle oils from
catalystic or thermal cracking operations, as well as aromatic
straight-run gas oils. These feedstocks may be derived from
petroleum crude oils, shale oils, tar sand oils, coal hydrogenation
products and the like. Specifically, it is preferred to employ
feedstocks boiling between about 400.degree.-1,000.degree. F.
having an API gravity of about 20.degree.-35.degree., and
containing at least about 20% by volume of aromatic hydrocarbons.
Such oils may also contain up to about 5% by weight of sulfur and
up to about 2% by weight of nitrogen. Aromatic feedstocks of this
character are required inasmuch as the relatively low hydrocracking
temperatures and relatively high pressures employed do not
thermodynamically favor the synthesis of aromatics from
non-aromatics, and hence the aromatics appearing in the gasoline
product are primarily unhydrogenated fragments of high boiling
aromatics initially present in the feed. If non-aromatic feedstocks
were employed, the products obtained herein would all be
substantially parafinnic and/or naphthenic, and the non-cracking
hydrogenation cycle in the second stage would have little or no
utilitarian value.
The process may be operated either with raw feedstock or with
preliminary hydrofining thereof. Normally, for feedstocks
containing substantial quantities of sulfur and/or notrigen
compounds, it is preferred to employ a prehydrofining step to
effect at least partial desulfurization, denitrogenation,
stabilization, etc. The hydrofining treatment may desirably be of
the "integral" type, i.e., wherein total hydrofiner effluent is
passed directly to the first hydrocracking stage without
intervening condensation or purification to remove ammonia and/or
hydrogen sulfide.
Reference is now made to the attached drawing, which is a flow
sheet illustrating a preferred mode of practicing the invention. It
will be understood that the drawing has been simplified by the
omission of certain conventional elements such as valves, pumps,
compressors, instrumentation and the like. The initial feedstock is
brought in via line 2, mixed with recycle and makeup hydrogen from
line 4, preheated to incipient hydrofining temperatures in heater 6
and then passed directly into hydrofiner 8, where catalystic
hydrofining proceeds under substantially conventional conditions.
Suitable hydrofining catalysts include for example mixtures of the
oxides and/or sulfides of cobalt and molybdenum, of nickel and
molybdenum, or of nickel and tungsten, preferably supported on a
substantially non-cracking carrier such as alumina, or alumina
containing a small amount of coprecipitated silica gel. Other
suitable catalysts include in general the oxides and/or sulfides of
the Group VIB and/or Group VIII metals, preferably supported on
adsorbent carriers such as alumina, silica, titania and the like.
Suitable hydrofining conditions are in general as follows:
HYDROFINING CONDITIONS
Broad Preferred Range Range Temperature, .degree.F. 550-850 650-750
Pressure, psig 500-5,000 800-2,500 LHSV, v./v./Hr. 0.5-10 1-5
H.sub.2 /oil ratio, MSCF/B 0.5-20 2-10
the above conditions are suitable adjusted so as to reduce the
organic nitrogen content of the feed to below about 100 parts per
million, preferably below about 25 parts per million.
The total hydrofined product from hydrofiner 8 is withdrawn via
line 10, blended with any of the hereinafter described recycle oils
in line 13, and transferred via heat exchanger 12 to first-stage
hydrocracker 14, preferably without intervening condensation of
separation of products. Heat exchanger 12 is for the purpose of
suitably adjusting the temperature of the total feed to
hydrocracker 14; this may require either cooling or heating,
depending upon the respective hydrofining and hydrocracking
temperatures employed and the relative volume of cool recycle oils
in line 13. Inasmuch as first-stage hydrocracker 14 and hydrofiner
8 are preferably operated at essentially the same pressure, it is
entirely feasible to enclose both contacting zones within a single
reactor, using appropriate temperature control means.
FIRST-STAGE HYDROCRACKING CATALYST
Suitable catalysts for use in hydrocracker 14 comprises in general
any refractory, solid cracking base having a cracking activity in
excess of that corresponding to a Cat-A Activity Index of about 40,
upon which is distributed a minor proportion of Group VIII metal or
metal sulfide hydrogenating components. Operative cracking bases
include for example mixtures or two or more refractory oxides such
as silica-alumina, silica-magnesia, silica-zirconia, alumina-boria,
silica-titania, silica-zirconia-titania, acid treated clays and the
like. Acidic metal phosphate gels such as aluminum phosphate may
also be used. The preferred cracking bases comprise crystalline,
siliceous zeolites, sometimes referred to in the art as molecular
sieves, composed usually of silica, alumina and one or more
exchangeable cations such as hydrogen, magnesium, rare earth
metals; or other polyvalent metal ions. These zeolites are further
characterized by crystal pores of relatively uniform diameter
between about 4 and 14 Angstroms. Suitable zeolites include for
example the synthetic molecular sieves, A, L, S, T, X and Y, and
natural zeolites such as chabazite, mordenite, etc. It is preferred
to employ zeolites having a relatively high SiO.sub.2 /Al.sub.2
O.sub.3 mole-ratio, between about 3.0 and 12, and even more
preferably between about 4 and 8. Specifically preferred zeolites
are those of the Y and L crystal types.
The naturally occurring molecular sieve zeolites are usually found
in a sodium form, an alkaline earth metal form, or mixed forms. The
synthetic molecular sieves normally are prepared first in the
sodium form. In any case, for use as a cracking base it is
preferred that most or all of the original zeolitic monovalent
metals be ion-exchanged out with a polyvalent metal, or with an
ammonium salt followed by heating to decompose the zeolitic
ammonium ions, leaving in their place hydrogen ions and/or exchange
sites which have actually been decationized by further removal of
water: ##SPC1##
In some cases, as in the case of synthetic mordenite, the hydrogen
forms can be prepared by direct acid treatment of the alkali metal
sieves. Hydrogen or "decationized" Y sieve zeolites of this nature
are more particularly described in U.S. Pat. No. 3,130,006. Mixed
polyvalent metal-hydrogen zeolites may be prepared by
ion-exchanging first with an ammonium salt, then partially
back-exchanging with a polyvalent metal salt, and then
calcining.
Both the hydrogen zeolites and the decationized zeolites described
above possess desirable catalytic activity. Both of these forms,
and the mixed forms are designated herein as being
"metal-cation-deficient." The preferred cracking bases are those
which are at least about 10%, and preferably at least 20%,
metal-cation-deficient, based on the initial ion-exchange capacity.
A specifically desirable and stable class of zeolites are those
wherein at least about 20% of the ion-exchange capacity is
satisfied by hydrogen ions, and at least about 10% by polyvalent
metal ions such as magnesium, calcium, zinc, rare earth metals,
etc.
The essential active metals employed herein as hydrogenation
components are those of Group VIII, i.e., iron, cobalt, nickel,
ruthenium, rhodium, palladium, osmium, iridium and platinum, or
mixtures thereof. The noble metals are preferred, and particularly
palladium and platinum. In addition to these metals, other
promoters may also be employed in conjunction therewith, including
the metals of Groups VIB and VIIB.
The amount of hydrogenating metal in the catalyst can vary within
wide ranges. Broadly speaking, any amount between about 0.05 and
20% by weight may be used. In the case of the noble metals, it is
normally preferred to use about 0.1-3% by weight. The preferred
method of adding the hydrogenating metal is by ion exchange. This
is accomplished by digesting the zeolite, preferably in its
ammonium form, with an aqueous solution of a suitable compound of
the desired metal wherein the metal is present in a cationic form,
as described for example in U.S. Pat. No. 3,200,083.
Following addition of the hydrogenating metal, the resulting
catalyst powder is then filtered off, dried, pelleted with added
lubricants, binders, or the like if desired, and calcined at
temperatures of e.g., 700-1,200.degree. F. in order to activate the
catalyst and decompose zeolitic ammonium ions. The foregoing
catalysts may be employed in undiluted form, or the powdered
catalyst may be mixed and copelleted with other relatively less
active adjuvants, diluents or binders such as activated alumina,
silica gel, coprecipitated silica-alumina cogel, magnesia,
activated clays and the like in proportions ranging between about 5
and 50% by weight. These adjuvants may be employed as such, or they
may contain a minor proportion of an added hydrogenating metal,
e.g., a Group VIB and/or Group VIII metal.
The process conditions in hydrocracker 14 are suitably adjusted so
as to provide about 20-70% conversion to gasoline per pass, while
at the same time permitting relatively long runs between
regenerations, i.e., from about 4 to 12 months or more. The
specific selection of operating conditions depends largely on the
nature of the feedstock, pressures in the high range normally being
used for highly aromatic feeds, or feeds with high end points. The
range of operative conditions contemplated for reactor 14 are as
follows, assuming the feed thereto is hydrofiner effluent
containing ammonia and hydrogen sulfide:
FIRST STAGE HYDROCRACKING CONDITIONS
Broad Preferred Range Range Temperature, .degree.F. 625-850 650-800
Pressure, psig 400-5,000 800-2,500 LHSV, v./v./Hr. 0.5-10 1-5
H.sub.2 /oil ratio, MSCF/B 0.5-20 2-10
the effluent from hydrocracker 14 is withdrawn via line 16,
condensed in heat exchanger 18, then mixed with wash water injected
via line 20 into line 22, and the entire mixture is then
transferred to high-pressure separator 24. Sour recycle hydrogen,
now substantially free of ammonia but still containing substantial
proportions of hydrogen sulfide, is withdrawn via line 26, and
aqueous wash water containing dissolved ammonia and some of the
hydrogen sulfide is withdrawn via line 28. The liquid hydrocarbon
phase in separator 24 is then flashed via line 30 into low-pressure
separator 32, from which flash gases comprising hydrogen, methane,
ethane, propane and the like are exhausted via line 34. The liquid
hydrocarbon phase in separator 32 is then transferred via line 36
to fractionating column 38.
Fractionating column 38 performs the dual function of separating
gasoline products synthesized in first-stage hydrocracker 14, and
of recovering unconverted higher boiling oils for treatment in
second-stage hydrocracker-hydrogenerator 62 and/or for recycle to
the first-stage hydrocracker 14. Light C.sub.4 -C.sub.6 gasoline is
normally taken off as overhead via line 40, while C7+ gasoline is
withdrawn as a light-side-cut via line 42. The remaining
unconverted oil may be treated according to two principal
alternates, as follows:
ALTERNATE A
In this alternate, which generally is preferred where maximum jet
and/or diesel fuel yields are desired, the entire bottoms fraction
from column 38 boiling above the gasoline range is sent to
second-stage hydrocracker-hydrogenator 62. To operate in this
manner side-cut line 41 is closed via valve 39, valve 43 is closed
and valve 45 opened, whereby the entire bottoms from column 38 is
transferred via lines 50, 47, 44, and preheater 60 to
hydrocracker-hydrogenator 62.
ALTERNATE B
In this alternate, which has the advantage of maximizing gasoline
quality, only the fraction from column 38 which boils substantially
within the range of the desired jet and/or diesel fuel product is
treated in hydrocracker-hydrogenator 62 during the non-cracking
hydrogenation cycle. Any higher boiling material is recycled to the
first stage for further cracking. To operate in this manner, valves
39 and 43 are opened and valve 45 closed, whereby the intermediate
side-cut fraction taken from column 38 via line 41 becomes the sole
feed to hydrocracker-hydrogenator 62, the remaining bottoms
fraction being transferred via lines 50, 49 and 13 to hydrocracker
14.
During the total gasoline production cycle, when
hydrocracker-hydrogenator 62 is being operated under cracking
conditions, side-cut line 41 is normally not utilized, all products
from column 38 boiling above the gasoline range being recycled to
hydrocracker-hydrogenator 62 via lines 50 and 47, or to
hydrocracker 14 via lines 50 and 49, or a portion may be recycled
to each zone. Normally it is preferred to recycle all or a
substantial portion of the bottoms fraction to the second-stage
reactor 62 in order to keep the cracking loads more evenly balanced
in the two reactors.
In any of the foregoing alternates, the second-stage feedstock in
line 47 is mixed with the recycle and makeup hydrogen from line 58,
preheated to incipient hydrocracking or hydrogenation temperatures
in heater 60 and passed into second-stage hydrocracker-hydrogenator
62. This second-stage feedstock differs considerably from the feed
of nitrogen and sulfur compounds. The choice is thus presented of
operating the second stage with or without significant amounts of
added sulfur.
In the modification illustrated, variations in sulfur
concentrations in hydrocracker-hydrogenator 62 are obtained by the
alternate use of separate and mixed recycle gas systems from
high-pressure separators 24 and 68. As previously noted, the
recycle gas from separator 24 normally contains a substantial
proportion of hydrogen sulfide which was not removed by the
water-washing operation. To operate hydrocracker-hydrogenator 62 in
the non-cracking cycle and substantially in the absence of sulfur
(separate recycle systems), valve 51 is opened and valves 52 and 54
closed, thus sending the sour recycle gas from line 26 through line
4 back to hydrocracker 14, and the sweet recycle gas from separator
68 back to hydrogenator 62 via lines 70 and 58. To operate reactor
62 as a hydrocracker with added sulfur, valve 51 is closed and
valves 52 and 54 opened, thereby diverting sour recycle gas from
line 26 into lines 55 and 70, where it mingles with sweet recycle
gas from separator 68. The mixed gases are then split, one portion
flowing to hydrofiner 8 via lines 56 and 4, and the other portion
flowing to reactor 62 via line 58.
While operating under Alternate A, the process variables are
suitably adjusted in hydrocracker 14 for maximum jet fuel and/or
diesel fuel synthesis (e.g., 30-70% conversion per pass). The
temperature and space velocity in hydrocracker-hydrogenator 62 are
correlated so as to effectively hydrogenate the unconverted oil and
middle distillate without substantial conversion of unconverted oil
to middle distillate and gasoline (e.g., less than about 15%,
preferably less than 10%, to 530.degree. F. minus and less than
about 20%, preferably less than 15%, to 650.degree. F. minus). For
purposes herein, conversion is defined as the volume percent of
unconverted oil in the hydrogenator feedstock minus the volume
percent of unconverted oil in the hydrogenator effluent divided by
the volume percent of unconverted oil in the feedstock, with
unconverted oil being defined as the oil boiling above the end
point of the particular middle distillate. For example, if the feed
to the hydrocracker-hydrogenator contains 70% oil boiling higher
than 530.degree. F. and the effluent contains 60% boiling above
530.degree. F., the conversion to 530.degree. F. jet fuel and below
is 70-60/70 .times. 100% or 14.3%. A substantial portion, if not
the total, of the conversion in the process herein is due to
boiling-point reduction from hydrogenation alone and not from
hydrocracking. As an illustration, the boiling point of naphthalene
is 424.degree. F. which when saturated yields decalin having a
boiling point of 382.degree. F., a 42.degree. F. boiling-point
reduction from hydrogenation alone. Hence, the fact that some of
the hydrocracker-hydrogenator feedstock is converted to middle
distillate and gasoline does not necessarily means that the feed is
hydrocracked.
The same variables are correlated for hydrogenation of the middle
distillate in Alternate B. For purposes of the operation wherein
the feed to the hydrocracker-hydrogenator comprises only middle
distillate, conversion is defined as that portion of the middle
distillate boiling above the temperature corresponding to the 50%
point of the feed which is converted to hydrocarbons boiling below
that temperature. The 50% point for jet fuel, for example, is about
400.degree. F. and the 50% point for diesel oil is about
530.degree. F. The conversion in middle distillate hydrogenation is
generally less than 10% and is preferably less than 5%.
In order to operate hydrocracker-hydrogenator 62 under
substantially non-cracking conditions, the reactor inlet
temperature is reduced to a level such that the cracking activity
of the catalyst becomes substantially inoperative, and the average
molecular weight of the C.sub.4 + effluent therefrom does not
differ by more than plus-or-minus 15%, preferably not more than
about plus-or-minus 10%, from the average molecular weight of the
feed thereto. Under these conditions there is substantially no
synthesis of gasoline or middle distillate by hydrocracking; the
effluent may, however, contain slightly larger proportions (e.g.,
1-10% preferably 1-5%) of these materials than the feed to
hydrogenator 62 by virtue of a slight reduction in boiling range
brought about principally by hydrogenation. Ordinarily, however,
the total gasoline synthesis is less than 10 volume-percent,
preferably less than 5%, based on feed.
To achieve the desired objective of effecting substantial
hydrogenation without cracking, the hydrogen sulfide concentration
in the hydrogenation zone should be maintained at a value below
about 0.2, and preferably below about 0.01 millimols hydrogen
sulfide per mole of hydrogen. This assures maximum hydrogenation
activity at temperatures sufficiently low to avoid cracking.
Hydrogen sulfide concentrations between about 0.01 and 0.2
millimols per mole of hydrogen may be preferred in order to control
the degree of hydrogenation and therefore avoid wasteful hydrogen
consumption while still producing a middle distillate meeting
product specification. Operative hydrogenation conditions fall
within the following general ranges:
NON-CRACKING SECOND-STAGE HYDROGENATION CONDITIONS
Broad Preferred Range Range Temperature, .degree.F. 200-500 250-450
Pressure, psig 400-5,000 800-2,500 LHSV, v.v./Hr. 0.2-20 1-8
H.sub.2 /Oil Ratio, MSCF/B 0.5-20 2-12 H.sub.2 S/H.sub.2 Ratio,
millimoles/mole <0.2 <0.01
As will be understood by those skilled in the art, the specific
selection of operating conditions within these ranges to
substantially saturate the feed (e.g. effect sufficient
hydrogenation to meet product quality specifications) without
substantial conversion will depend on several factors, mainly the
relative activity of the catalyst and the aromaticity of the feed.
The aromatic content of the feed to the hydrocracker-hydrogenator
is generally between about 20 to 70% and normally above 30%. For
satisfactory jet fuel quality, it is usually desirable to reduce
the aromatic content of the feed to below a maximum of 20% and
preferably to below 15%; and for satisfactory diesel fuel quality,
the aromatic content should be reduced to below 20%, and preferably
below 15%. When feedstocks containing at least 30% aromatics are
hydrogenated to products containing less than 20% aromatics, it
will be apparent that at least one-third of the aromatic
hydrocarbons are being hydrogenated. The degree of hydrogenation
desired can readily be controlled by simply varying temperature
and/or space velocity. For example, while operating under Alternate
A, space velocities between about 1.2 and 3.0 and temperatures
between about 325.degree. and 450.degree. F. generally give
satisfactory hydrogenation with little if any attendant
hydrocracking. At higher space velocities, higher temperatures are
maintained to achieve the necessary saturation yet not
substantially hydrocrack the unconverted oil. The operation at the
lower space velocities and temperatures is, however, preferred to
the operation at higher space velocities and temperatures.
While operating under Alternate B, lower temperatures (e.g.,
250.degree. - 425.degree. F.) are utilized to hydrogenate the jet
fuel and/or diesel oil and generally lower space velocities are
prevalent due to the decrease volume of hydrocarbons available for
hydrogenation in the hydrocracker-hydrogenator.
It is contemplated also that ammonia in amounts between about 1 and
2,000 parts per million by weight may be included with the feed
being hydrogenated in order to further repress cracking activity of
the catalyst. If sufficient ammonia is added, e.g., 200-2,000 ppm,
hydrogenation temperatures in excess of 500.degree. F., and up to
about 700.degree. F. may be attained without encountering cracking
and without significant reduction in hydrogenation activity of the
catalyst.
In order to convert hydrocracker-hydrogenator 62 to the
hydrocracking cycle, the principal operative requirement is to
raise the inlet temperature to a level which gives the desired
crack per pass. However, as previously noted, raising the
temperature is not alone sufficient to give a gasoline product of
desired aromaticity. Any one or more of the previously mentioned
expedients of (1) adding sulfur to the feed, (2) adding ammonia to
the feed, and (3) raising the space velocity may be employed to
reduce effective hydrogenation of the product. Suitable
hydrocracking conditions for the production of an aromatic gasoline
product fall within the following ranges:
SECOND-STAGE HYDROCRACKING CONDITIONS
Broad Preferred Range Range Temperature, .degree.F. 500-850 550-800
Pressure, psig 400-5,000 800-2,500 LHSV v.v./Hr. 0.5-30 1-10
H.sub.2 /Oil Ratio, MSCF/B 0.5-20 2-12 Sulfur, Millimoles/Mole
H.sub.2 >0.01 >0.2 Ammoniacal Nitrogen, ppm of feed 0-2000
2-200
SECOND-STAGE CATALYSTS
A critical feature of the invention resides in the nature of the
catalyst employed in hydrocracker-hydrogenater 62. To achieve the
desired flexibility of operating under either cracking or
non-cracking conditions, an active hydrogenating component
comprising one or more Group VIII noble metals in amounts of about
0.05-4% by weight is required. Specifically included are the
metals, ruthenium, rhodium, palladium, osmium, iridium, and
platinum, with palladium being preferred. These hydrogenating
metals may be supported on substantially any of the previously
described cracking bases used in the first stage of the process,
but is is distinctly preferred to employ a zeolite base in its
hydrogen or decationized form, i.e., one which is at least about
20%, and preferably at least about 35%, metal-cation-deficient,
which bases display maximum cracking activity. While cracking
activity is generally regarded as being substantially independent
of hydrogenation activity, it has now been found that the
hydrogenation activity of Group VIII moble metals is greatly
enhanced when supported upon one of the metal-cation-deficient
zeolites. This is dramatically illustrated by a series of
low-temperature hydrogenations carried out with several different
palladium-zeolite catalysts, wherein naphthalene and tetralin
feedstocks were subjected to hydrogenation in a stirred autoclave:
##SPC2##
Thus, palladium deposited on hydrogen zeolites appears to exhibit a
hydrogenation activity about 3 to 7 times that of palladium
deposited upon sodium zeolites or magnesium zeolites. Hence, a
preferred catalyst is one where a substantial amount (e.g., 50-90%)
of the total of the exchangeable cations in the zeolite are
satisfied by hydrogen ions and a most preferred catalyst is
palladium on a hydrogen Y zeolite.
To complete the process description, effluent from
hydrocracker-hydrogenator 62 is withdrawn via line 64, condensed in
cooler 66 and transferred to high-pressure separator 68, from which
recycle hydrogen is withdrawn via line 70 and utilized as
previously described. The liquid hydrocarbons in separator 68 are
then flashed via line 72 into low-pressure separator 74 from which
C.sub.1 -C.sub.3 flash gases are withdrawn via line 76. The
remaining liquid hydrocarbon product in separator 74 is withdrawn
via line 78 and transferred to second-stage product fractionation
column 80, wherein it is fractionated into various gasoline, jet
fuel and diesel fuel fractions as may be desired.
During the gasoline-producing hydrocracking cycle in second-stage
reactor 62, light gasoline blending stock is withdrawn overhead
from column 80 via line 82, C.sub.7 -plus gasoline via line 84, and
the remaining unconverted oil is withdrawn as bottoms via line 98
and recycled to the second stage via lines 97, 47 and 44, or to
first-stage reactor 14 via lines 99 and 13, or a portion may be
recycled to each reactor. During this operation, side-cut stripping
column 90 is ordinarily not used.
During the non-cracking hydrogenation cycle in reactor 62, very
small amounts of gasoline fractions may be recovered via lines 82
and 84, and the major jet fuel product is withdrawn as a side-cut
via line 88 and transferred to stripping column 90, from which
overhead gasoline hydrocarbons are returned to column 80 via line
92, while the desired jet fuel product is withdrawn as bottoms via
line 94. Remaining bottoms fraction from column 80 may be withdrawn
via line 86 as diesel fuel, or recycled via lines 98, 99 and 13 to
first-stage hydrocracker 14. It will be understood however that
column 80 is normally operated in this manner only when
second-stage reactor 62 is operated according to Alternate (A)
described above, i.e., when the feed to reactor 62 comprises the
entire non-gasoline bottoms fraction from column 38. When reactor
62 is operated according to Alternate (B), i.e., when the feed
thereto comprises only the light side-cut withdrawn via line 41
from column 38, fractionator 80 can be by-passed with the entire
effluent stream in line 78 being sent directly to side-cut
stripping column 90 for separation of any minor gasoline fraction
from the jet fuel product. In some cases, the gasoline content of
the effluent from reactor 62 is so insignificant that the entire
product in line 78 may be sent to storage and blending facilities
without fractionation.
The following examples are presented to illustrate the operation
and results of the process as above described, but these examples
should not be construed as limiting in scope:
EXAMPLE 1.
A series of three two-stage operations as described in connection
with the drawing were carried out, with the second-stage being used
under non-cracking hydrogenation conditions according to Alternate
(A) described above, and with second-stage product boiling above
the jet fuel range being recycled to the first stage. The initial
feedstock was a blend of catalytic cracking cycle oils and
straight-run gas oils having a gravity of 23.9.degree. API, a
boiling range of 400.degree.-870.degree. F., containing 69
volume-percent aromatics, 0.184 weight-percent nitrogen, and 1.2
weight-percent sulfur. The feed to the hydrocracker-hydrogenator
contained about 38 volume-percent aromatics. The hydrofining
catalyst was composed of about 3% NiO and 15% MoO.sub.3 supported
on a carrier composed of 5% SiO.sub.2 coprecipitated with 95%
Al.sub.2 O.sub.3, the catalyst being sulfided before use. The
catalyst used in the first-stage hydrocracker, and in the
second-stage hydrogenator was a copelleted mixture of (a) 20% by
weight of activated alumina and (b) 80% by weight of a Y molecular
sieve zeolite containing 0.5 weight-percent palladium, and wherein
about 30% of the ion-exchange capacity was satisfied by magnesium
ions, 10% by sodium ions, and 60% by hydrogen ions. During runs 1
and 2, the second-stage hydrogenation was operated substantially in
the absence of sulfur, while in run number 3 the indicated
proportion of sulfur was added to the feed. Each contacting zone
was operated at a pressure of about 1,500 psig, with hydrogen/oil
ratios of 8,000-10,000 SCF/B of oil. The significant conditions and
results of the runs were as follows: ##SPC3##
The effect of sulfur in run 3 on product aromaticity is readily
apparent, and runs 1 and 2 illustrate the ease with which product
aromaticity can be controlled by varying temperature in a
sulfur-free environment.
If runs 1 and 2 are modified by raising the second-stage
temperature to around 510.degree. F. and recycling effluent boiling
above the jet fuel range back to the second-stage (as in U.S. Pat.
No. 3,132,090), a jet fuel product of substantially zero aromatic
content is obtained, with concomitant synthesis of about 20
volume-percent of C.sub.7 -400.degree. F. gasoline, which gasoline
has a leaded octane rating of about 72. Hydrogen consumption also
increases markedly.
Run 3 above can easily be modified for 100% gasoline production by
raising the second-stage temperature to about 600.degree.
-625.degree. F., and recycling thereto all material boiling above
gasoline. The second-stage gasoline produced under these conditions
is slightly inferior to the first-stage gasoline, having a leaded
octane number in the range of about 78-80, which is nevertheless
much superior to the second-stage gasoline produced at lower
temperatures in the absence of hydrogen sulfide.
EXAMPLE 2.
Another run was carried out under conditions described in Example
1, but operating according to Alternate (B), i.e., with the
525.degree. F..+-. fraction of effluent from the first-stage
hydrocracker being recycled directly to that stage, the sole feed
to the second-stage hydrogenation unit being the
400.degree.-525.degree. F. fraction of first-stage effluent. The
significant conditions and results of this run were as follows:
TABLE 3
Run No. 6 First-Stage Hydrocracking Avg. Bed Temp., .degree.F. 744
LHSV 2.94 Conv./Pass to 400.degree.F. minus 40
Yields, Vol. % of Fresh Feed C.sub.5 -C.sub.6 Gaso. 27.9 C.sub.7
--400.degree.F. Gaso. 57.6
Octane No's., F-1 + 3 ml TEL C.sub.5 -C.sub.6 Gaso. 99.5 C.sub.7
--400.degree.F. Gaso. 89.8
Second-stage Hydrogenation Avg. Bed Temp., .degree.F. 414 LHSV 0.29
H.sub.2 S/H.sub.2 Ratio, millimoles/mole 0.005
Property Feed Product ASTM Boiling Range, .degree.F. 406-554
400-537 Gravity 36.0 40.6 Aromatics 38 3 Saturates 60 97
Avg. Mol. Wt. Ratios, 2nd Stage Feed/C.sub.4 + Product 1.016
H.sub.2 Consumption, SCF/B 2,100
although lower jet fuel yields are obtained in the above run as
compared to Examples 1 and 2, the yield-octane value of the total
gasoline product is improved without sacrificing jet fuel
quality.
The following claims are believed to define the true scope of the
invention, which is not limited to the exemplary details described
above:
* * * * *