Note : Les descriptions sont présentées dans la langue officielle dans laquelle elles ont été soumises.
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HYDROCARBON CONVERSION PROCESS AND
CATALYSTS USED THEREIN
BACKGROUND OF THE INVENTION
Hydrocracking continues to be an important refinery process for producing
modern fuels, and much attention has been devoted to the development of
hydrocracking catalysts. In particular, hydrocracking is useful in the
production of
middle distillate fractions. those petroleum fractions boiling in the range of
about
250°-700°F. ( 121 °-371 °C.) as determine by the
appropriate ASTM test procedure.
These middle distillate fractions are particularly desirable as valuable fuels
products.
The term "middle distillate" is intended to include the diesel, jet fuel and
kerosene
boiling range fractions. The kerosene or jet fuel boiling point range is
intended to
refer to a temperature range of about 280°-525°F. ( 138°-
274°C.) and the term "diesel
boiling range" is intended to refer to hydrocarbon boiling points of about
250°-700°F.
(121°-371°C.). Gasoline or naphtha is normally the C5 to
400°F. (204°C.) boiling
point fraction of available hydrocarbons. The boiling point ranges of the
various
product fractions recovered in any particular refinery will vary with such
factors as the
characteristics of the crude oil source, refinery local markets, product
prices, etc.
Reference is made to ASTM standards D-975 and D-3699-83 for further details on
kerosene and diesel fuel properties.
Processes for hydrocracking heavy petroleum feedstocks to middle distillates
generally involve successive hydrocracking steps or "stages". Successive
hydrocracking stages are usually performed in separate reaction vessels,
though this is
not required: Initial stages are operated at conditions and with catalysts
which are
active for sulfur and nitrogen removal, for aromatic saturation, and for other
hydrotreating reactions. First stage hydrocracking catalysts are designed for
high
conversion activity in the presence of feeds having high sulfur and nitrogen
contents,
and often with high aromatic contents. Typical first stage hydrocracking
catalysts
comprise a cracking component. a binder and a non-noble metal hydrogenation
component such as nickel. molybdenum, cobalt. and/or tungsten. The cracking
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component is generally an amorphous material such as silica-alumina. though a
zeolite may be included.
Hydrocracking stages after the first stage are operated at hydrocracking
conditions. using catalysts while are intended for high cracking rates. Such
second
stage catalysts may contain either base metal hydrogenation components. such
as
nickel, molybdenum, cobalt. and/or tungsten, or noble metal hydrogenation
components. such as platinum or palladium. While catalysts comprising a noble
metal
hydrogenation component are more active in certain situations than those
having a
base metal as the hydrogenation component. they are also susceptible to sulfur
contamination, which deactivates the catalysts and shortens their useful
lives. For
these reasons noble metal catalysts have been typically used in second-stage
hydrocracking units, or units wherein the feed comprises relatively small
amounts of
sulfur and nitrogen.
Traditionally, palladium has been the popular choice as the hydrogenation
component of second-stage hydrocracking catalysts because of its relatively
high
selectivity and substantially lower cost than platinum. However, continual
development of refinery processes has generated new requirements for second
stage
hydroprocessing. The common practice of feeding high sulfur feeds to the two-
stage
hvdrocracker has increased the amount of sulfur fed into the second-stake.
Upsets in
the sulfur control processes between the first and second stage also
occasionally result
in high sulfur loads in the second stage feed. Under these conditions. noble
metal
catalysts containing palladium are unacceptable, since the palladium catalysts
are
particularly sensitive to sulfur compounds. even relatively small amounts
contained in
some first stage effluents. Processes wherein substantially all the sulfur is
not
removed from the feed before introduction into the second-stage hydrotreater
readily
fouls the palladium catalyst decreasing both its selectivity and activiy.
A number of noble metal catalysts have been proposed. U.S. Patent No.
x.393.408 discloses a process for stabilizing a lubricating oil base stock.
Catalysts
taught for this process contain a hydrogenation component independently
selected
from the group consisting of the platinum group metals. nickel. cobalt.
chromium,
molybdenum. tungsten and tin. A combination of metals containing both platinum
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and palladium are also taught. but there is no suggestion of any possible
benefit for
using a particular range of platinum and palladium.
U.S. Patent No. 4.387.258 teaches a selective hydrogenation process to convert
alkynes and dienes tg-alkene containing an olefinic double bond using a
catalyst
comprising platinum or palladium on a crystalline silica polymorph.
U.S. Patent No. 3,852.207 teaches a two stage process for producing a
lubricating aid having good UV stability. Catalysts taught for the
hydrogenation step
contain a hydrogenation component, including platinum or palladium. The silica-
alumina support of the preferred hydrogenation catalyst has an aluminum
content of
from 40 to 9~ wt % and or alumina/silica weight ratio between 40160 and 95/5.
T. Lin. et al.. Ind. Eng. Chem. Res. 1995, 3.1. 4284-4289 and J. Chiou, et
al.,
Ind. Eng. Chem. Res. 1995. 34, 4277-4283 disclose methods of improving the
sulfur
tolerance of supported platinum catalysts by adding a second metal, such as
palladium, to the supported catalyst.
U.S. Patent No. 3,962,071 to Itoh et al., claims a process which includes
hydrocracking, fractionation. and hydrogenation for enhancing lubricating oil
photostability. In Itoh et al., the hydrogenation catalyst includes palladium
on a silica
containing refractory inorganic oxide carrier having ~-40 weight percent
silica, a
surface area of 100-500 m'/g, a pore volume of 0.5-1.2 ml/g, an average pore
diameter
of 30-120 Angstroms, and a bulk density of 0.5-0.7 g/ml.
U.S. Patent No. 3.637,484, U.S. Patent No. 3.637,878, and U.S. Patent No.
3,703,461, all to Hansford, disclose a process for hydrogenating an aromatic
hydrocarbon feedstock with a catalyst having a support composed of a silica-
alumina
cogel in a large pore alumina gel matrix and containing a platinum group
metal. The
catalyst composition of Hansford has a pore volume of 0.8-2.0 mUg with about
0.3-1.0
ml/g of the pore volume in pores of diameter greater than about X00 Angstroms.
In
Hansford, the support employs a relatively high ratio of alumina to silica.
always in
excess of 60 weight percent of alumina as compared to silica. Conversely,
supports
having high ratios of silica to alumina have been disclosed in U.S. Patent
Nos.
4,139,493; 4.325,80; and 4.601.996. However, catalysts containing higher
ratios of
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silica to ahunina generally contain a hydrogenation component other than a
platinum
group metal, commonly nickel. tin. molybdenum, or cobalt.
Clark, et al., in U.S. Patent No. x.346.874 disclose a hydrogenation process
and a catalyst comprising from about 0.1 % to about 2.0% be weight of
palladium and
from about 0.1 % to about 2.0% be weight of platinum on a support comprising
borosilicate.
Noble metal catalysts. primarily comprising platinum and/or palladium, are
knowm for hydrogenation reactions such as olef n or aromatic hydrogenation. To
minimize yield losses, such hydrogenation catalysts contain little or no
cracking
activity, generally comprising alumina supports with no amorphous or
crystalline (e.g.
zeolitic) silica-alumina components. Conversions in these hydrogenation
processes
are low. While noble metal catalysts are also desirable for cracking
reactions, with
larger amounts of cracking leading to molecular weight reduction. they tend to
be
vulnerable to sulfur components in the feedstock, and can only be used for
cracking
very low sulfur feeds. Even upsets in an otherwise low sulfur process, which
may
dump large quantities of sulfur on a hydrocracking catalyst, fouls the
conventional
noble metal catalyst. It is desirable to have a process and a catalyst for
cracking a
hydrocarbonaceous oil with the performance advantages of a noble metal
catalyst but
with improved sulfur tolerance.
0 SUMMARY OF THE INVENTION
It is an object of the present invention to provide a process for
hydrocracking
low sulfur containing feeds while reducing the deleteriously effect of sulfur
on
catalyst performance and life. Another object of the present invention is to
provide a
process for increased middle distillate production. Still another object of
the present
2~ invention is to decrease hydrogen consumption.
The present invention provides a process for converting a hydrocarbonaceous
oil into a product of lower average molecular weight and lower average boiling
point
comprising contacting a hydrocarbonaceous oil, under hydrocracking conditions
sufficient to convert at least about 30% of the hydrocarbonaceous oil which
boils
30 above X50°F to middle distillate components having a boiling point
of less than
X50°F_ with a catalyst comprising a cracking component. a binder.
and a
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hydrogenation component comprising platinum and palladium in a molar ratio of
between 10:1 and 1:10.
The present invention is directed to a hydrocracking process for hydrocracking
a high sulfur feed. including a feed containing greater than 20 ppm sulfur, up
to sulfur
levels of 100 ppm or higher. The present invention is further directed to a
hydrocracking process, such that at least about 30 vol%. preferably at least
about 50
vol%. of a hydrocarbonaceous oil which boils above 550°F is converted
to a cracked
product having a boiling point of less than S50°F. At least about 75
vol% of the
preferred hydrocarbonaceous oil of this process has a normal boiling point
above
about 550°F. More preferably, at least about 75 vol% of the
hydrocarbonaceous oil
has a normal boiling point above about 700°F. The preferred catalyst of
this process
comprises a cracking component, such as an amorphous and/or zeolitic silica-
alttmina
component. and a hydrogenation component comprising platinum and palladium in
a
molar ratio of between 10:1 and 1:10, preferably in the molar ratio of between
4:1 and
1:4.
Among other factors, the present invention is based on the discovery of a
hydrocracking catalyst, containing both platinum and palladium in a particular
ratio on
a single catalyst particle, which shows high tolerance to sulfur contained in
a
hydrocracking feed. Furthermore, while the catalyst of this invention has
considerably
more activity than the base metal catalysts, even in the presence of sulfur,
the catalyst
is also significantly and surprisingly more selective for producing middle
distillate
fuels than is either conventional base metal catalysts or conventional
catalysts which
contain palladium alone.
DETAILED DESCRIPTION OF THE INVENTION
The present process is directed to hydrocracking reactions for the production
of middle distillate fuels. Such hydrocracking reactions include removing
sulfur and
nitrogen from hydrocarbonaceous oils and convening components of the oils to
products of an average lower molecular weight and an average lower boiling
point.
While cracking of the petroleum components is desirable, it is preferable to
minimize
cracking to light products boiling below 250°F.
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The catalyst useful in the present process comprises a cracking component. a
binder, and a palladium-platinum hydrogenation component having a palladium to
platinum molar ratio of between i 0:1 and 1:10, preferably between about 4.0:1
to
1:4.0 and most preferably between about 2:1 and 1:2.
The cracking component includes an amorphous cracking component. which
generally is the base material for the catalyst. The preferred amorphous
cracking
component is silica-alumina, containing typically between 10 and 90 weight
percent
silica. preferably between 15 and 65 weight percent silica. and more
preferably
between about 20 and 60 weight percent silica, the remainder being alumina.
Base
materials suitable for preparing the catalysts used in~he process of this
invention are
commercially available, for example. from Condea Chemie. GmbH of Hamburg,
Germany. and base materials from Condea Chemie designated as "Siral 40" have
been
found to be particularly suitable to prepare catalysts employed in the present
invention. Alternatively, the silica-alumina base materials are prepared using
known
coprecipitation, cogelation, andlor comull procedures. Crystalline silica-
alumina
materials such as zeoiites may also be included as part of the support of the
hydrogenation catalysts to increase catalyst acidity.
The base material may further comprise in the range of 0-85 w % of a zeolite
cracking component, preferably a Y-type zeolite. The preferred catalyst
contains up to
25% zeolite and preferably in the range of 1-20% zeolite. One of the zeolites
which is
considered to be a good starting material for the manufacture of hydrocracking
catalysts is the well-known synthetic zeolite Y as described in U.S. Patent
No.
3,130,007, issued April 21, 1964. A number of modifications to this material
have
been reported, one of which is ultrastable Y zeolite as described in U.S.
Patent No.
3,536,605, issued Oct. 27, 1970. To further enhance the utility of synthetic Y
zeolite,
additional components can be added. For example, U.S. Patent No. 3,835,027,
issued
on Sept. 10, 1974 to Ward, et al.. describes a hydrocracking catalyst
containing at least
one amorphous refractory oxide. a crystalline zeolitic aluminosilicate and a
hydrogenation component selected from the Group VI and Group VIII metals and
their sulfides and their oxides. Zeolites having small unit cell sizes are
described in
U.S. Patent Nos. 5.059.567 and 5.246.677, the disclosures of which are
incorporated
herein by reference for all purposes. The zeolite-containing catalyst
particles may be
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prepared using conventional methods. One such method is described in U.S.
Application Serial No. 07/870,011. filed by M.M. Habib et al. on April 15,
1992, and
now abandoned. the disclosure of which is incorporated herein by reference for
all
purposes. Also. so-called x-ray amorphous zeolites (i.e., zeolites having
crystallite
sizes too small to be detected by standard x-ray techniques) can be suitably
applied as
cracking components.
The preferred catalysts employed in the present invention contain a catalyst
support that is generally prepared from these base materials. The distribution
of silica
and alumina in the support may be either homogeneous or heterogeneous. but is
preferably heterogeneous. A homogeneous distribution is ordinarily obtained
when the
silica-alumina ratio is uniform throughout the support, resulting for example
from
conventional coprecipitation or cogelation techniques. These homogeneous
supports,
wherein the necessary silica content is uniformly distributed. are difficult
to prepare in
the large-pore forms required herein. Pure alumina, on the other hand. can
readily be
prepared in these forms, preferably using acid or base peptization methods. A
preferred form of the support consists of an alumina gel in which is dispersed
the
silica-alumina base material, which form is referred to herein as the
"heterogeneous"
support. The alumina gel is also referred to herein as the "oxide binder." The
support
may also contain refractory materials other than alumina or silica. such as
for example
other inorganic oxides or clay particles, provided that such material does not
adversely
affect the hydrogenation activity of the final catalyst. Other inorganic
oxides that may
be present in the support may include. but are not necessarily limited to,
titania,
magnesia and zirconia or combinations thereof. Generally, silica-alumina will
make
up at least 90 weight percent of the entire support. and most preferably the
support
will be substantially ail silica-alumina.
As stated above, the support is preferably prepared by mixing the base
material
with a peptized oxide binder, such as alumina, which has been treated with an
acid,
generally a strong acid such as nitric. acetic. or hydrochloric acids.
Generally, the
weight ratio of base material to oxide binder is in the range from 95/5 to
30/70 and
preferably 65 base material/3~ binder. Pore size is in part controlled in
supports
prepared as described herein by the length of time that the oxide binder is
exposed to
the acid during the peptizing step. and by the amount of acid used. To prepare
the
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particulate refractory inorganic support for the catalyst. the base material
is mixed
with an oxide binder which has been treated with acid, preferably with less
than about
3 weight percent of~ 100% pure acid (based on the weight of total calcined
solids). A
4-12% acid is used to make a mesoporous catalyst. The proportion of acid will
vary
depending on the n~pe of acid. the reactivity of the raw material powders, the
type of
mixing equipment. and the mixing time. temperature, etc.
The support used in the practice of the present invention is a particulate
support. The exact size and shape of the catalyst support particles will vary
depending
on the particular method that will be used to hydrogenate the lubricating oil
base
stock. The effective diameter of the zeolite catalyst particles are in the
range of from
about 1/32 inch to about 1/4 inch, preferably from about 1/20 inch to about
1/8 inch.
The catalyst particles may have any shape known to be useful for catalytic
materials,
including spheres. cylinders. fluted cylinders, prills, granules and the like.
For
non-spherical shapes, the effective diameter can be taken as the diameter of a
representative cross section of the catalyst particles. The zeolite catalyst
particles will
further have a surface area in the range of from about 50 to about 500
m2/gram.
When the support is to be extruded in the preparation of the catalyst
materials, mixing
the base material with the peptized oxide binder enhances the extrusion
process and
improves the strength of the completed catalyst pellets. The extrudate is
usually dried
and calcined in an oven to produce the support. After calcining, the support
is ready
for the addition of the hydrogenation component of the catalyst.
The amount of palladium-platinum hydrogenation component on the catalyst
must be sufficient to act as an effective catalyst to hydrocrack the petroleum
feed.
Generally, the amount of alloy on the support used to catalyze a hydrogenation
process within the scope of the present invention will be within the range of
from
about 0.01 weight percent to about 5 weight percent. preferably the range is
from
about 0.1 weight percent to about 1 weight percent. Generally, adding greater
than
about 1 weight percent of the alloy does not significantly improve on the
activity of
the catalyst and is therefore economically disadvantageous. However, amounts
in
excess of 1 weight percent are usually not harmful to the performance of the
catalyst.
The preferred catalyst comprises from about 0.1 to 0.5 wt% palladium and from
about
0. > to-0.4 w-t% platinum. and more preferably from about 0.1 to 0.25 wt%
palladium
_g_
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and from about 0.1 to 0.2~ w% platinum. The ratio of palladium to platinum
useful as
the hydrogenation component is 10:1 to 1:10. A more preferred ratio is 2.5:1
to 1:2.5.
Most preferably, the palladium to platinum molar ratio is 2:1 to 1:2.
A number of methods are known in the art to deposit platinum and palladium
~ metal or their compounds onto the support. such as, for example. by ion
exchange,
impregnation, coprecipitation. etc. It has been found that depositing platinum
and
palladium on the supports used in the catalyst of the present invention is
particularly
advantageous when using a contacting solution containing active compounds of
both
platinum and palladium under a controlled pH. The contacting solution
preferably will
be buffered to maintain a pH within the range of from about 9 to about 10.
Values
outside of this pH range may be used to deposit platinum and palladium jointly
on the w
support. but the final distribution of the alloy on the support may not be as
favorable
as those obtained within this pH range.
When depositing platinum and palladium by impregnation. the metals are
usually added to the impregnating solution as a metal salt, generally as an
organic
amine complex salt of a mineral acid. Ammonium salts have been found to be
particularly useful in preparing the impregnating solution. Representative of
the
ammonium salts that may be used are nitrates, carbonates, bicarbonates and
lower
carboxylic acid salts such as acetates and formates. In the case of palladium,
an
ammonium nitrate salt or an ammonium chloride salt have been found to give
satisfactory results. However, other salts of the metals are also operable and
could be
used to impregnate the support. In such case, it may be useful to determine
the
optimal pH to use during impregnation for the particular salt selected in
order to
obtain the best distribution of metals on the support. It has been found that
excellent
2~ distribution of palladium will be obtained using the present support if an
impregnating
solution containing tetraamine palladium nitrate is buffered to a pH of from
between
about 9.6 and about 10.
Following impregnation, the impregnated support should be allowed to stand
before drying for a period of time sufficient for it to attain equilibration
with the
impregnating solution. For an extrudate, this period usually is at least 2
hours, and
periods of up to 24 hours are not detrimental to the finished catalyst. A
suitable time
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for a given support may be readily determined by one skilled in the art having
regard
to this disclosure by, for example, drying at various times after impregnation
and
measuring the metal distribution. Following impregnation and standing, the
catalyst is
again dried and/or calcined. The prepared catalyst may be reduced with
hydrogen as is
conventional in the art and placed into service.
The hydrocracking reaction takes place in the presence of hydrogen, preferably
at hydrogen pressures in the range of between about 500 psia and 3000 psia,
more
preferably in the range of about 900 psia to about 3000 Asia. The feed rate to
the
hydrogenation catalyst system is in the range of from about 0.2 to about 1.5
LHSV,
preferably in the range of about 0.2 to about 1.0 LHSV. The hydrogen supply
(makeup
and recycle) is in the range of from about SOO to about 20,000 standard cubic
feet per
barrel of lubricating oil base stock, preferably in the range of from about
2000 to
about 10,000 standard cubic feet per barrel.
The hvdrocarbonaceous feed to the hydrocracker is generally a distillate
stream
from a vacuum distillation process. The feed may be derived from a crude
stream or
from a refinery process such as an FCC process. a coker process, a residuum
demetallization and/or hydrotreating process, a synthetic fuel process, a
deasphalting
process, a hydrocracking or hydrotreating process or the like. The feed may
have been
processed, e.g. by hydrotreating, prior to the present hydrocracking process
to reduce
or substantially eliminate its heteroatom content. The preferred feed has a
boiling
point range starting at a temperature above X00°F. (260°C)., and
more preferably a
boiling point range between 500°-1050°F. (260-566°C.)
Preferred feedstocks to the
reaction zone therefore include gas oils having at least 70 vol% of their
components
boiling above 650°F. (343°C.) and at least 90 vol% of their
components boiling below
950°F. (510°C.). The hydrocracking feedstock may contain
nitrogen, usually present
as organonitroeen compounds in amounts greater than 1 ppm. It is a feature of
the
present invention that high nitrogen feeds, e.g. containing up to 100 ppm of
organonitrogen may be treated in the present process. The feed will normally
also
contain sulfur containing compounds sufficient to provide a sulfur content of
greater
than 1 ppm. and up to 250 ppm.
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The preferred process is a dow~nflowing reaction process. The hydrocracking
reaction zone is operated at hydrocracking reaction conditions. including a
reaction
temperature in the range of from about 250°C to about 500°C,
pressures up to about
300 bar (30.5 MPa) and a feed rate (vol oil/vol cat hr) from about 0.1 to
about 10 hr's.
3 Hydrogen circulation rates are generally in the range from about 350 std
liters H2/kg
oil to 1780 std liters HZ/kg oil. Preferred reaction temperatures range from
about
340°C to about 455°C. Preferred total reaction pressures range
from about S00
pounds per square inch absolute (psia) to about 3,500 Asia (about 3.5 MPa -
about
24.2 MPa), preferably from about 1,000 psia to about 3,000 psia (about 7.0 MPa
-
about 20.8 MPa). With the preferred catalyst system described above, it has
been
found that preferred process conditions include contacting a hydrocarbonaceous
feed
with hydrogen in the presence of the layered catalyst system under
hydrocracking
conditions comprising a pressure of about 16.0 MPa (2,300 psia), a gas to oil
ratio at
from about 606-908 std liters H~/kg oil (4,000 scf/bbl to about 6,000
scf/bb)1, a LHSV
I S of about 1.0 hr's, and a temperature in the range of 360°C. to
427°C (680°F. - 800°F.).
The hydrocracked oil product exiting the hydrocracking reaction zone includes
normally liquid phase components, which comprise reaction products and
unreacted
components of the VGO, and normally gaseous phase components, which comprise
gaseous reaction products and unreacted hydrogen. The reaction products
include
cracked products having a boiling point below that of the feed to the
hydrocracking
process. such that at least 5%. more preferably at least about 10% by volume
and still
more preferably at least about 30% by volume of the components in the VGO
which
boil above 550°F are converted in the reaction zone to components which
boil below
about 550°F. The hydrocracker effluent is further decreased in nitrogen
and sulfur
content. Preferably the normally liquid products present in the hydrocracker
reaction
zone effluent contain less than about 20 ppm sulfur and less than about 10 ppm
nitrogen. more preferably less than about 10 ppm sulfur and about ~ ppm
nitrogen.
EXAMPLES
Example 1
Several base metal catalysts, a palladium catalyst and the platinum-palladium
catalyst of this invention were tested in a pilot plant for middle distillate
selectivity.
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Catalyst A is a catalyst of this invention, containing 0. I 6 Wt% Palladium,
0.2
Wt% Platinum.80% silica-alumina, 16% alumina with 4% ultra-stable Y zeolite.
Catalyst A was prepared by impregnating the SiO~-A1~03/A1z03/zeolite base with
a
solution containing platinum and palladium salts, drying the impregnated
catalyst and
calcining at 842°F in air for 30 minutes.
Catalyst B and Catalyst C were from two catalyst preparations containing 10
wt% Ni0 and 24 wt% W03 on a silica-alumina support.
Catalyst D was prepared like Catalyst B but with 4% ultra-stable Y zeolite
Catalyst E contained 0.5 Wt% Pd on silica-alumina with 4% zeolite
A vacuum gas oil from commercial refinery praduction was fed through a two
stage pilot-plant unit. The first stage comprised a Ni-W-Ti on silica-alumina
catalyst
and was run under such conditions as to target approximately 40% conversion to
components boiling below 550° F. The effluent from the first-stage was
subjected to
vacuum distillation and the residuum was fed to the second stage. The second-
stage
contained the subject catalysts and targeted an approximately 60% per pass
conversion
to products boiling below 550° F. Feed to the second-stage contained
less than 6ppm
sulfur in each case. The results shown in Table 1 demonstrate the catalyst of
this
invention is more selective for middle distillate than the base metal and
palladium
catalysts. It also demonstrates that the catalyst of this invention is more
active than
the base metal catalysts. The results further demonstrate that less hydrogen
is
consumed using the catalyst of the present invention.
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Table 1
Second Stage CatalystB C D E A
Catalyst Temp. F 716 718 701 604 671
Overall,
Two-Stage
Yields,
Wt %
(zero
bleed
basis)
C4- 6.8 6.2 6.1 5.2 5.2
Naphtha (CS-250F) 25.0 25.0 26.6 25.6 22.8
-
Light Distillate (250-400F)35.8 35.4 36.5 37.7 37.3
Middle Distillate 34.8 35.5 33.0 34.t 36.7
(400-330F)
Total Distiliate (250-550F)70.6 70.9 69.5 71.8 74.0
H2 Consumption. SCFB 1610 1620 1600 1570 1530
Example 2
In this example, a palladium catalyst (catalyst E from example 1 ) and a
platinum-palladium catalyst of this invention (catalyst A from example 1 )
were tested
under high sulfur conditions. The conditions of this experiment were
substantially the
same as those in example 1 except the feed to the second stage of the pilot-
plant unit
was spiked «7th sulfur compounds to achieve a second-stage feed containing 20
ppm
sulfur. The results in Table 2 show that the platinum-palladium catalyst of
this
invention is more selective for middle distillate production than the
palladium
catalyst.
Table 2
Second-Stage Catalyst E A
Overall Wt%
Yield.
C4- 5.8 5.1
Naphtha (CS - 250F) 27.5 24.1
Middle distillate (250 69.3 72.8
- 5~0F)
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Example 3
Ln this example, a base metal catalyst (catalyst B of example 1 ) and a
platinum-palladium catalyst of this invention (catalyst A of example 1 ) were
tested '
under substantially the same conditions as in Example 1. The middle distillate
products were further analyzed for smoke point and freeze point and distilled
to
compare 5, 50 and 95% cuts. The results in Table 3 indicate that the platinum-
palladium catalyst of this invention is more selective for the middle
distillate products
than is the base metal catalyst. The results also show that less hydrogen is
consumed
in the process using the platinum-palladium catalyst of this invention. The
results
further indicate that the properties of the middle distillate made with a
platinum-
palladium catalyst of this invention are virtually the same as those made with
the base
metal catalyst.
Table 3
!,, Second-Stage Catalyst B A
Overall eld. Wt%
Yi
C4- 6.7 4.9
Naphtha (CS - 250F) 27.0 24.1
Middle distillate (250 69.2 ~ 73.2
- 550F)
H2 Consumption, SCFB 1560 1480
Middle Distillate Product
Quality
Smoke Point, mm 31 34
Freeze Point, C -61 -59
5% cut, F 259 ?59
50% cut, F 401 400
95% cut, F 543 5=t5
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Example 4
This example was run under substantially the same conditions as that of
example 3 except the feed to the second stage was spiked with sulfur compounds
to
achieve 30 ppm sulfur and a 25°F ascending temperature profile was
practiced in the
S first-stage and a 15°F ascending profile was practiced in the second
stage to model
typical commercial conditions. The results summarized in table 4 indicate that
the
even under the ascending profile and high sulfur conditions the platinum-
palladium
catalyst of this invention was more selective for the middle distillate
products than the
base metal catalyst and consumed less hydrogen. Comparison with the results
obtained in example 3 revealed that the base metal catalyst was more
susceptible to
catalyst aging (with a 3.2% decline in selectivity) than the platinum-
palladium catalyst
of this invention ( 1.1 % decline in selectivity). Analysis products indicates
that the
middle distillate product made with a platinum-palladium catalyst of this
invention
were virtually the same as those made with the base metal catalyst.
Table 4
Second-Stage Catalyst B A
Overall Wt%
Yield,
C4- 8.0 S.1
Naphtha (C; - 250F) 27.8 24.8
Middle distillate (250 67.0 72.4
- S50F)
HZ Consumption, scf/bbl 1580 1470
Middle Distillate Product
Quality
Smoke Point. mm 30 30
Freeze Point, C -62 -57
5% cut. F 259 259
50% cut. F 400 400
95% cut, F X43 550
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Example 5
The sulfur tolerance of a palladium catalyst, Catalyst F. (Catalyst F is
Catalyst
A without Platinum) and the platinum-palladium catalyst of this invention
(Catalyst A
of example 1 ) were tested. Each catalyst was tested over an extended period
with feed
containing high amounts of sulfur, and for short intervals with feed
containing at least
100 ppm sulfur. Test conditions were substantially the same as in Example 1.
Catalyst F was contacted for short intervals with feed containing first 50 ppm
sulfur
and then 100 ppm sulfur, followed by extended runs using feed containing 20
ppm
sulfur. Table 5 shows that Catalyst F suffered significant loss in performance
(measured by jet fuel yield loss) when contacted with the high sulfur feed. At
the end
of the test. the catalyst produced 10.7% less jet fuel yield than near the
beginning of
the test. By contrast, Catalyst A was contacted with 200 ppm sulfur feed.
followed
by an extended period with 20 ppm feed. In this test, Catalyst A produced 1.6%
less
jet fuel at the end of the test as near the beginning of the test. In a second
test,
Catalyst A was contacted with 700 ppm sulfur feed, followed by an extended
period
with 30 ppm feed. In this second test, Catalyst A produced 3.0% less jet fuel
at the
end of the test, compared to near the beginning of the test. These results
show that the
palladium containing catalyst was significantly more susceptible to sulfur
poisoning
that was the catalyst of the invention. which contained a platinum/paliadium
alloy.
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Table 5
Sulfur in Contact Jet Fuel Relative
CatalystTest Feed. m Time, Yield. Yield
~ ~ hrs wt% Loss,
2 800 71.3
50 180 66.2 7.2
100 132 64.3 9.8
CatalystTl 2 220 69.6 2.4
F
10 350
2 400 69 3.2
20 470 63.7 10.7
<6 300 74.0
T2 200 8
20 170 72.8 1.6
Catalyst <6 500 73.2
A
30 840 72.4 1.1
T3
700 4
30 3000 71.0 3.0
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WHAT IS CLAIMED IS:
1. A process for converting a hydrocarbonaceous oil into a product of lower
average
molecular weight and lower average boiling point comprising contacting a
hydrocarbonaceous oil. under hydrocracking conditions sufficient to convert at
least about 30% of the hydrocarbonaceous oil which boils above X50°F to
middle
distillate components having a boiling point of less than 550°F, with a
catalyst
comprising a cracking component. a binder. and a hydrogenation component
comprising palladium and platinum in a molar ratio of between 10:1 and 1:10.
2. The process according to claim 1 wherein the molar ratio of palladium to
platinum
in the alloy is between about 4.0:1 and about 1:4Ø
3. The process according to Claim 1 wherein the hydrogenation component
comprises from about 0.10 to 0.25 wt% palladium and from about 0.10 to 0.25
wt% platinum.
4. The process according to Claim 1 wherein the cracking component comprises
silica-alumina.
5. The process according to Claim 4, wherein the cracking component further
comprises an ultra-stable Y-type zeolite.
6. The process according to Claim 1 wherein the binder is alumina.
7. The process according to Claim 1 wherein the hydrocarbonaceous feed is a
denitrified VGO containing less than about 200 ppm nitrogen
8. The process according to Claim I wherein at least about 70 vol% of the
hydrocarbonaceous oil has a boiling point above about 650°F.
9. The process according to Claim 8 wherein at least about 75 vol% of the
hydrocarbonaceous oil has a normal boiling point above about 700°F.
10. The process according to Claim 1 conducted at hydrocracking conditions
sufficient to convert at least about 50% of the hydrocarbonaceous oil which
boils
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