Note: Descriptions are shown in the official language in which they were submitted.
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HYDROCRACKING CATALYST CONTAINING BETA AND Y ZEOLITES,
AND PROCESS FOR ITS USE TO MAKE DISTILLATE
FIELD OF THE INVENTION
[0001] The invention relates to catalyst compositions and their use in
hydrocarbon
conversion processes, particularly hydrocracking. The invention more
specifically relates to a
catalyst composition that comprises a Y zeolite and a beta zeolite as active
cracking
components. The invention specifically relates to a hydrocracking process that
produces
middle distillate.
BACKGROUND OF THE INVENTION
[0002] Petroleum refiners often produce desirable products such as turbine
fuel, diesel
fuel, and other hydrocarbon liquids known as middle distillates, as well as
lower boiling
liquids such as naphtha and gasoline, by hydrocracking a hydrocarbon feedstock
derived from
crude oil. Hydrocracking also has other beneficial results such as removing
sulfur and
nitrogen from the feedstock by hydrotreating. Feedstocks most often subject to
hydrocracking
are gas oils and heavy gas oils recovered from crude oil by distillation.
[0003] Hydrocracking is generally carried out by contacting, in an appropriate
reactor
vessel, the gas oil or other hydrocarbon feedstock with a suitable
hydrocracking catalyst under
appropriate conditions, including an elevated temperature and an elevated
pressure and the
presence of hydrogen so as to yield a lower overall average boiling point
product containing a
distribution of hydrocarbon products desired by the refiner. Although the
operating conditions
within a hydrocracking reactor have some influence on the yield of the
products, the
hydrocracking catalyst is a prime factor in determining such yields.
[0004] Hydrocracking catalysts are subject to initial classification on the
basis of the
nature of the predominant cracking component of the catalyst. This
classification divides
hydrocracking catalysts into those based upon an amorphous cracking component
such as
silica-alumina and those based upon a zeolitic cracking component such as beta
or Y zeolite.
Hydrocracking catalysts are also subject to classification on the basis of
their intended
predominant product of which the two main products are naphtha and
"distillate", a term
which in the hydrocracking refining art refers to distillable petroleum
derived fractions having
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a boiling point range that is above that of naphtha. Distillate typically
includes the products
recovered at a refinery as kerosene and diesel fuel. At the present time,
distillate is in high
demand. For this reason, refiners have been focusing on hydrocracking
catalysts which
selectively produce a distillate fraction.
[0005] The three main catalytic properties by which the performance of a
hydrocracking
catalyst for making distillate is evaluated are activity, selectivity, and
stability. Activity may
be determined by comparing the temperature at which various catalysts must be
utilized under
otherwise constant hydrocracking conditions with the same feedstock so as to
produce a given
percentage, normally 65 percent, of products boiling in the desired range,
e.g., below 371 C
(700 F) for distillate. The lower the temperature required for a given
catalyst, the more active
such a catalyst is in relation to a catalyst requiring a higher temperature.
Selectivity of
hydrocracking catalysts may be determined during the foregoing described
activity test and is
measured as a percentage of the fraction of the product boiling in the desired
distillate product
range, e.g., from 149 C (300 F) to 371 C (700 F). Stability is a measure of
how well a
catalyst maintains its activity over an extended time period when treating a
given hydrocarbon
feedstock under the conditions of the activity test. Stability is generally
measured in terms of
the change in temperature required per day to maintain a 65 percent or other
given
conversion.
[0006] Although cracking catalysts for producing distillate are known and used
in
commercial environments, there is always a demand for new hydrocracking
catalysts with
superior selectivity at a given activity and/or superior activity at a given
selectivity for producing
distillate.
BRIEF SUMMARY OF THE INVENTION
[0007] It has been found that hydrocracking catalysts containing a Y zeolite
having a unit
cell size or dimension ao of from 24.33 to 24.38 angstrom (hereinafter Y
Zeolite II) and
containing a beta zeolite preferably having an overall silica to alumina (Si02
to A1203) mole
ratio of less than 30 and a SF6 adsorption capacity of at least 28 weight-
percent (hereinafter
wt-%) have substantially improved selectivity at a given activity or
substantially improved
activity at a given selectivity compared to other hydrocracking catalysts now
commercially
available for use in hydrocracking processes for producing distillate. The
catalyst also
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contains a metal hydrogenation component such as nickel, cobalt, tungsten,
molybdenum, or
any combination thereof. The catalyst contains from 0.5 to 5 wt-% beta zeolite
based on the
combined weight of the beta zeolite, the Y Zeolite II, and the support on a
dried basis, and the
catalyst has a weight ratio of the Y Zeolite H to the beta zeolite of from 0.5
to 5 on a dried
basis. The Y Zeolite II has an overall silica to alumina mole ratio of from
5.0 to 11Ø In one
embodiment, the catalyst contains an additional Y zeolite and the Y Zeolite II
has a higher
unit cell size than that of the additional zeolite.
[0008] It is believed that a hydrocracking catalyst containing such a Y
zeolite and such a
beta zeolite is novel to the art.
[0009] Under typical hydrocracking conditions, including elevated temperature
and
pressure and the presence of hydrogen, such catalysts are highly effective for
converting gas
oil and other hydrocarbon feedstocks to a product of lower average boiling
point and lower
average molecular weight. In one embodiment, the product contains a relatively
large
proportion of components boiling in the distillate range, which as defined
herein is from
149 C (300 F) to 371 C (700 F).
BRIEF DESCRIPTION OF THE DRAWINGS
[0010] Fig. 1 is a graph of distillate selectivity versus relative catalyst
activity for several
hydrocracking catalysts.
[0011] Fig. 2 is a graph of the ratio of heavy distillate selectivity to light
distillate
selectivity versus relative catalyst activity for several hydrocracking
catalysts.
INFORMATION DISCLOSURE
[0012] Beta and Y zeolites have been proposed in combination as components of
several
different catalysts including catalysts for hydrocracking. For instance, US-A-
5,275,720; US-
A-5,279,726; and US-A-5,350,501 describe hydrocracking processes using a
catalyst
comprising a beta zeolite and a Y zeolite. US-A-5,350,501 describes a
hydrocracking process
using a catalyst comprising, among other components, zeolite beta and a Y
zeolite having a
unit cell size between 24.25 and 24.35 angstrom and a water vapor sorption
capacity at 4.6
mm water vapor partial pressure and 25 C less than 8.0 percent by weight of
the zeolite. US-
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A1-2004/0152587 describes a hydrocracking catalyst comprising a carrier
comprising a
zeolite of the faujasite structure having a unit cell size in the range of
from 24.10 to 24.40
angstrom, a bulk silica to alumina ratio above 12, and a surface area of at
least 850 m2/g, and
the catalyst may contain a second zeolite such as beta zeolite, ZSM-5 zeolite,
or a Y zeolite of
a different unit cell size. Two different Y zeolites have also been proposed
in combination as
components of several different catalysts including catalysts for
hydrocracking, as described
in US-A-4,661,239 and US-A-4,925,546.
DETAILED DESCRIPTION OF THE INVENTION
[0013] The process and composition disclosed herein may be used to convert a
feedstock
containing organic compounds into products, particularly by acid catalysis,
such as
hydrocracking organic compounds especially hydrocarbons into a product of
lower average
boiling point and lower average molecular weight. The composition, which may
be a catalyst
and/or a catalyst support, comprises a beta zeolite and Y Zeolite H. The
composition may also
comprise a refractory inorganic oxide. When used as a catalyst for
hydrocracking, the
composition contains a beta zeolite, Y Zeolite II, a refractory inorganic
oxide, and a
hydrogenation component.
[0014] The hydrocracking process and composition disclosed herein centers on
using a
catalyst containing a particular beta zeolite and a particular Y zeolite. The
composition may
optionally contain an additional Y zeolite. The beta zeolite preferably has a
relatively low
silica to alumina mole ratio and a relatively high SF6 adsorption capacity. Y
Zeolite II has a
unit cell size of from 24.33 to 24.38 angstrom. When the additional zeolite is
present, Y
Zeolite II has a higher unit cell size than that of Y Zeolite I. It has been
found that differing
performance results when such a beta zeolite and such Y zeolites are
incorporated in a
hydrocracking catalysts in this way. Compared to catalysts containing one or
two Y zeolites,
the selectivity of product boiling in the distillate range is higher at a
given activity or the
activity is higher at a given selectivity of product boiling in the distillate
range.
[0015] Beta zeolite is well known in the art as a component of hydrocracking
catalysts.
Beta zeolite is described in US-A-3,308,069 and US Reissue No. 28341, which
are hereby
incorporated by reference herein in their entireties. The beta zeolite that is
used in the process
and composition disclosed herein has a silica to alumina mole ratio of less
than 30 in one
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embodiment, less than 25 in another embodiment, more than 9 and less than 30
in yet another
embodiment, more than 9 and less than 25 in a further embodiment, more than 20
and less
than 30 in another embodiment, or more than 15 and less than 25 in still
another embodiment.
As used herein, unless otherwise indicated, the silica to alumina (Si02 to
A1203) mole ratio
of a zeolite is the mole ratio as determined on the basis of the total or
overall amount of
aluminum and silicon (framework and non-framework) present in the zeolite, and
is
sometimes referred to herein as the overall silica to alumina (Si02 to A1203)
mole ratio.
[0016] Beta zeolite is usually synthesized from a reaction mixture containing
a templating
agent. The use of templating agents for synthesizing beta zeolite is well
known in the art. For
example, US-A-3,308,069 and US Reissue No. 28341 describe using
tetraethylammonium
hydroxide and US-A-5,139,759, which is hereby incorporated herein by reference
in its
entirety, describes using the tetraethylammonium ion derived from the
corresponding
tetraethylammonium halide. Another standard method of preparing beta zeolite
is described
in the book titled Verified Synthesis of Zeolitic Materials, by H. Robson
(editor) and K. P.
Lillerud (XRD Patterns), second revised edition, ISBN 0-444-50703-5, Elsevier,
2001. It is
believed that the choice of a particular templating agent is not critical to
the success of the
process disclosed herein. In one embodiment the beta zeolite is calcined in
air at a
temperature of from 500 to 700 C (932 to 1292 F) for a time sufficient to
remove to remove
the templating agent from the beta zeolite. Calcination to remove the
templating agent can be
done before or after the beta zeolite is combined with the support and/or the
hydrogenation
component. Although it is believed that the templating agent could be removed
at calcination
temperatures above 700 C (1292 F), very high calcination temperatures could
significantly
decrease the SF6 adsorption capacity of beta zeolite. For this reason it is
believed that
calcination temperatures above 750 C (1382 F) for removing the templating
agent should be
avoided when preparing the beta zeolite for use in the process disclosed
herein. It is critical to
the process disclosed herein that the SF6 adsorption capacity of the beta
zeolite is at least 28
wt-%.
[00171 While it is known that steaming a zeolite such as beta results in
changes to the
actual crystalline structure of the zeolite, the abilities of present day
analytical technology
have not made it possible to accurately monitor and/or characterize these
changes in terms of
important structural details of the zeolite. Instead, measurements of various
physical
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properties of the zeolite such as surface area are used as indicators of
changes that have
occurred and the extent of the changes. For instance, it is believed that a
reduction in the
zeolite's capacity to adsorb sulfur hexafluoride (SF6) after being steamed is
caused by a
reduction in the crystallinity of the zeolite or in the size or accessibility
of the zeolite's
micropores. It is, however, an indirect correlation of the changes in the
zeolite that may be
undesirable, since the SF6 adsorption capacity in the catalyst used in the
process and
composition disclosed herein is relatively high. In embodiments of the process
and
composition disclosed herein, the SF6 adsorption capacity of the beta zeolite,
whether steam
treated or not, should be at least 28 wt-%.
[0018] Accordingly, the beta zeolite of the process and composition disclosed
herein may
be characterized in terms of SF6 adsorption. This is a recognized technique
for the
characterization of microporous materials such as zeolites. It is similar to
other adsorption
capacity measurements, such as water capacity, in that it uses weight
differences to measure
the amount of SF6 which is adsorbed by a sample which has been pretreated to
be
substantially free of the adsorbate. SF6 is used in this test since because
its size and shape
hinders its entrance into pores having a diameter of less than 6 angstrom. It
thus can be used
as one measurement of available pore mouth and pore diameter shrinkage. This
in turn is a
measurement of the effect of steaming on the zeolite. In a simplistic
description of this
measurement method, the sample is preferably first predried in a vacuum at 300
C (572 F)
for one hour, then heated at atmospheric pressure in air at 650 C (1202 F) for
two hours, and
finally weighed. It is then exposed to the SF6 for one hour while the sample
is maintained at a
temperature of 20 C (68 F). The vapor pressure of the SF6 is maintained at
that provided by
liquid SF6 at 400 torr (53.3 kPa (7.7 psi)). The sample is again weighed to
measure the
amount of adsorbed SF6. The sample may be suspended on a scale during these
steps to
facilitate these steps.
[0019] In any mass production procedure involving techniques such as steaming
and
heating there is a possibility for individual particles to be subjected to
differing levels of
treatment. For instance, particles on the bottom of a pile moving along a
rotary kiln may not
be subjected to the same atmosphere or temperature as the particles which
cover the top of the
pile. This factor must be considered during manufacturing and also during
analysis and
testing of the finished product. It is, therefore, recommended that any test
measure done on
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the material is performed on a representative composite sample of the entire
quantity of
finished product to avoid being misled by measurements performed on individual
particles or
on a non-representative sample. For instance, an adsorption capacity
measurement is made on
a representative composite sample.
[0020] Although the process and the composition disclosed herein can use a
beta zeolite
that has not been subjected to a steaming treatment, the process and the
composition
disclosed herein can also use beta zeolite that is subjected to steaming,
provided that the
steaming is relatively mild in comparison to steaming of beta zeolite in the
literature. Under
the proper conditions and for the proper time, steaming beta zeolite can yield
a catalyst that
can be used in the process and composition disclosed herein.
[0021] Hydrothermally treating zeolites for use in hydrocracking catalysts is
a relatively
blunt tool. For any given zeolite, steaming decreases the acidity of the
zeolite. When the
steamed zeolite is used as a hydrocracking catalyst, the apparent result is
that the overall
distillate yield increases but the catalyst's activity decreases. This
apparent tradeoff between
yield and activity has meant that to achieve high activity means not to steam
the beta zeolite,
but at the expense of lower product yields. This apparent tradeoff between
yield and activity
must be considered and is a limit to the improvement that appears to be
obtainable by
steaming the beta zeolite. When the steamed beta zeolite is used in the
catalysts disclosed
herein, the improvement in activity over catalysts containing only Y zeolite
would appear
limited while the improvement in yield over such catalysts would appear more
enhanced.
[0022] If the beta zeolite is to be steamed, such steaming can be performed
successfully
in different ways, with the method which is actually employed commercially
often being
greatly influenced and perhaps dictated by the type and capability of the
available equipment.
Steaming can be performed with the beta zeolite retained as a fixed mass or
with the beta
zeolite being confined in a vessel or being tumbled while confined in a
rotating kiln. The
important factors are uniform treatment of all beta zeolite particles under
appropriate
conditions of time, temperature and steam concentration. For instance, the
beta zeolite should
not be placed such that there is a significant difference in the amount of
steam contacting the
surface and the interior of the beta zeolite mass. The beta zeolite may be
steam treated in an
atmosphere having live steam passing through the equipment providing low steam
concentration. This may be described as being at a steam concentration of a
positive amount
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less than 50 mol-%. Steam concentrations may range from 1 to 20 mol-% or from
5 to 10
mol-%, with small-scale laboratory operations extending toward higher
concentrations. The
steaming may be performed for a positive time period of less than or equal to
1 or 2 hours or
for 1 to 2 hours at a temperature of less than or equal to 600 C (1112 F) at
atmospheric
pressure and a positive content of steam of less than or equal to 5 mol-%. The
steaming may
be performed for a positive time period of less than or equal to 2 hours at a
temperature of
less than or equal to 650 C (1202 F) at atmospheric pressure and a positive
content of steam
of less than or equal to 10 mol-%. The steam contents are based on the weight
of vapors
contacting the beta zeolite. Steaming at temperatures above 650 C (1202 F)
appears to result
in beta zeolite that is not useful in the process disclosed herein since the
SF6 adsorption
capacity of the resulting beta zeolite is too low. Temperatures below 650 C
(1202 F) can be
used, and the steaming temperature can be from 600 C (1112 F) to 650 C (1202
F), or less
than 600 C (1112 F). It is taught in the art that there is normally an
interplay between time
and temperature of steaming, with an increase in temperature reducing the
required time.
Nevertheless, if steaming is done, for good results it appears a time period
of 1/2 to 2 hours or
1 to 1 1/2 hours can be used. The method of performing steaming on a
commercial scale may
be by means of a rotary kiln having steam injected at a rate which maintains
an atmosphere of
10 mol-% steam.
[0023] An exemplary lab scale steaming procedure is performed with the zeolite
held in a
6.4 cm (2-1/2 inch) quartz tube in a clam shell furnace. The temperature of
the furnace is
slowly ramped up by a controller. After the temperature of the zeolite reaches
150 C (302 F)
steam generated from deionized water held in a flask is allowed to enter the
bottom of the
quartz tube and pass upward. Other gas can be passed into the tube to achieve
the desired
steam content. The flask is refilled as needed. In the exemplary procedure the
time between
cutting in the steam and the zeolite reaching 600 C (1112 F) is one hour. At
the end of the set
steam period the temperature in the furnace is reduced by resetting the
controller to 20 C
(68 F). The furnace is allowed to cool to 400 C (752 F) (about 2 hours) and
the flow of
steam into the quartz tube is stopped. The sample is removed at 100 C (212 F)
and placed in
a lab oven held overnight at 110 C (230 F) with an air purge.
[0024] The beta zeolite of the process and composition disclosed herein is not
treated
with an acid solution to effect dealumination. In this regard it is noted that
essentially all raw
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(as synthesized) beta zeolite is exposed to an acid to reduce the
concentration of alkali metal
(e.g., sodium) which remains from synthesis. This step in the beta zeolite
manufacture
procedure is not considered part of the treatment of manufactured beta zeolite
as described
herein. In one embodiment, during the treatment and catalyst manufacturing
procedures, the
beta zeolite is exposed to an acid only during incidental manufacturing
activities such as
peptization during forming or during metals impregnation. In another
embodiment, the beta
zeolite is not acid washed after the steaming procedure as to remove aluminum
"debris" from
the pores.
[0025] Also included in the process and composition disclosed herein is a Y
zeolite
having a unit cell size of from 24.33 to 24.38 angstrom. This Y zeolite is
sometimes referred
to herein as Y Zeolite II in order to distinguish this Y zeolite from an
optional additional Y
zeolite having a different unit cell size and described hereinafter. Y Zeolite
11 preferably has a
unit cell size of from 24.34 to 24.36 angstrom. Y Zeolite 11 can have an
overall silica to
alumina mole ratio of from 5.0 to 12.0 in one embodiment, from 5.0 to 11.0 in
another
embodiment, and from 5.0 to 10.0 in yet another embodiment. The process and
composition
disclosed herein require a Y Zeolite 11.
[0026] Optionally and in addition to Y Zeolite 11, the process and composition
disclosed
may include an additional Y zeolite, which is sometimes referred to herein as
Y Zeolite I. Y
Zeolite I has a different unit cell size from the unit cell size of Y Zeolite
II. The unit cell size
of the Y Zeolite I is preferably at least 0.04 angstrom smaller than the unit
cell size of Y
Zeolite 11. The unit cell size of Y Zeolite I is more preferably from 24.25 to
24.32 angstrom,
and even more preferably from 24.26 to 24.30 angstrom. Y Zeolite I can have an
overall silica
to alumina mole ratio of from 5.0 to 12.0 in one embodiment, from 5.0 to 11.0
in another
embodiment, and from 5.0 to 10.0 in yet another embodiment.
[0027] The option of adding Y Zeolite I during the manufacturing process gives
catalyst
producers flexibility to make products that meet the individual requirements
of hydrocracking
unit operators. The presence of Y Zeolite I in the catalyst changes the
properties of the
catalyst without the need to change how the Y Zeolite 11 itself is prepared or
the amount of Y
Zeolite H used in the catalyst. In some instances, however, adding Y Zeolite I
decreases the
requirement for Y Zeolite 11, which is an additional advantage when sufficient
quantities of Y
Zeolite 11 are not available. Hydrocracking unit operators, especially those
producing
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distillate, can use catalysts containing both Y Zeolite I and Y Zeolite H as a
tool to satisfy
their particular and sometimes unique requirements for hydrocracking catalyst
activity and
selectivity.
[0028] The term "Y zeolite" as used herein is meant to encompass all
crystalline zeolites
having either the essential X-ray powder diffraction pattern set forth in US-A-
3,130,007 or a
modified Y zeolite having an X-ray powder diffraction pattern similar to that
of US-A-
3,130,007 but with the d-spacings shifted somewhat due, as those skilled in
the art will
realize, to cation exchanges, calcinations, etc., which are generally
necessary to convert the Y
zeolite into a catalytically active and stable form. Y Zeolite I and Y Zeolite
II are modified Y
zeolites in comparison to the Y zeolite taught in US-A-3,130,007. As used
herein, unit cell
size means the unit cell size as determined by X-ray powder diffraction.
[0029] The Y zeolites used in the process and composition disclosed herein are
large pore
zeolites having an effective pore size greater than 7.0 angstrom. Since some
of the pores of
the Y zeolites are relatively large, the Y zeolites allow molecules relatively
free access to
their internal structure. The pores of the Y zeolites permit the passage
thereinto of benzene
molecules and larger molecules and the passage therefrom of reaction products.
[0030] One group of Y zeolites that may be used in the process and composition
disclosed herein as Y Zeolite I, Y Zeolite II, or both includes zeolites that
are sometimes
referred to as ultrastable or ultrahydrophobic Y zeolites. The composition and
properties of
this group of Y zeolites are, in essence, prepared by a four step procedure.
First, a Y zeolite in
the alkali metal form (usually sodium) and typically having a unit cell size
of 24.65 angstrom
is cation exchanged with ammonium ions. The ammonium exchange step typically
reduces
the sodium content of the starting sodium Y zeolite from a value usually
greater than 8 wt-%,
usually from 10 to 13 wt-%, calculated as Na20, to a value in the range from
0.6 to 5 wt-%,
calculated as Na20. Methods of carrying out the ion exchange are well known in
the art.
[0031] Second, the Y zeolite from the first step is calcined in the presence
of water vapor.
For example, the Y zeolite is calcined in the presence of at least 1.4
kPa(absolute) (hereinafter
kPa(a)) (0.2 psi(absolute) (hereinafter psi(a))), at least 6.9 kPa(a) (1.0
psi(a)), or at least 69
kPa(a) (10 psi(a)) water vapor, in three embodiments. In two other
embodiments, the Y
zeolite is calcined in an atmosphere consisting essentially of or consisting
of steam. The Y
zeolite is calcined so as to produce a unit cell size in the range of 24.40 to
24.64 angstrom.
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[0032] Third, the Y zeolite from the second step is ammonium exchanged once
again.
The second ammonium exchange further reduces the sodium content to less than
0.5 wt-%,
usually less than 0.3 wt-%, calculated as Na20.
[0033] Fourth, the Y zeolite from the third step is treated further so as to
yield Y zeolite
having a unit cell size from 24.25 to 24.32 angstrom or preferably from 24.26
to 24.30
angstrom in the case of Y Zeolite I. In the case of Y Zeolite II, the
treatment yields a Y zeolite
having a unit cell size from 24.33 to 24.38 angstrom or preferably from 24.34
to 24.36
angstrom. The zeolite Y resulting from the fourth step has an overall silica
to alumina mole
ratio from 5.0 to 12.0 in one embodiment, from 5.0 to 11.0 in another
embodiment, and from
5.0 to 10.0 in yet another embodiment. The treatment of the fourth step can
comprise any of
the well known techniques for dealuminating zeolites in general and
ultrastable Y zeolite in
particular so as to yield the desired unit cell size and overall silica to
alumina mole ratio. The
fourth treatment step may change the unit cell size and/or the framework
silica to alumina
mole ratio, with or without changing the overall silica to alumina mole ratio.
Generally,
zeolite dealumination is accomplished by chemical methods such as treatments
with acids,
e.g., HCI, with volatile halides, e.g., SiC14, or with chelating agents such
as
ethylenediaminetetraacetic acid (EDTA). Another common technique is a
hydrothermal
treatment of the zeolite in either pure steam or in air/steam mixtures,
preferably such as
calcining in the presence of sufficient water vapor (for example, in an
atmosphere consisting
essentially of steam, and most preferably consisting of steam) so as to yield
the desired unit
cell size and overall silica to alumina mole ratio.
[0034] The above-discussed preparation procedure for Y zeolites used in the
process and
composition disclosed herein differs from the procedure for the Y zeolites
taught in US-A-
3,929,672 by the addition of the fourth treatment step. US-A-3,929,672, which
is hereby
incorporated herein by reference in its entirety, discloses a method for
dealuminating an
ultrastable Y zeolite. US-A-3,929,672 teaches a preparation procedure wherein
a sodium Y
zeolite is partially exchanged with ammonium ions, followed by steam
calcination under
controlled temperature and steam partial pressure, followed by yet another
ammonia
exchange and then by an optional calcination step in a dry atmosphere. The
exchange and
steam calcination steps can be repeated to achieve the desired degree of
dealumination and
unit cell size reduction. The zeolites of US-A-3,929,672 are known under the
designation Y-
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84 or LZY-84 commercially available from UOP LLC, Des Plaines, Illinois,
U.S.A. Y-84 or
LZY-84 zeolites may be produced by the first three steps just mentioned, but
optionally one
may include a further calcination step in a dry atmosphere, e.g., a
calcination in water- and
steam-free air, at 482 C (900 F) or higher.
[0035] The above-discussed preparation procedure for Y zeolites used in the
process and
composition disclosed herein is similar to the procedure for the Y zeolites
taught in US-A-
5,350,501. However, particular conditions in the above-discussed fourth
treatment step are
selected in order to produce critical ranges of unit cell size for Y Zeolite
II and optional Y
Zeolite I. US-A-5,350,501, which is hereby incorporated herein by reference in
its entirety,
discloses a fourth step that involves calcining the resulting zeolite from the
third treatment
step in the presence of sufficient water vapor (in an atmosphere consisting
essentially of
steam or consisting of steam) so as to yield a unit cell size below 24.40, and
most preferably
no more than 24.35 angstrom, and with a relatively low sorptive capacity for
water vapor. The
Y zeolite produced by the four-step procedure in US-A-5,350,501 is a UHP-Y
zeolite, an
ultrahydrophobic Y zeolite as defined in US-A-5,350,501. US-A-5,350,501
defines a "UHP-
Y" zeolites as zeolite aluminosilicates having among other properties, a unit
cell size or
dimension as of less than 24.45 angstrom and a sorptive capacity for water
vapor at 25 C and
a p/po value of 0.10 of less than 10.00 weight percent. The most preferred UHP-
Y zeolite in
US-A-5,350,501 is LZ-10.
[0036] Another group of Y zeolites which may be used in the process and
composition
disclosed herein as Y Zeolite I, Y Zeolite II, or both may be prepared by
dealuminating a Y
zeolite having an overall silica to alumina mole ratio below 5 and are
described in detail in
US-A-4,503,023; US-A-4,597,956 and US-A-4,735,928, which are hereby
incorporated
herein by reference in their entireties. US-A-4,503,023 discloses another
procedure for
dealuminating a Y zeolite involving contacting the Y zeolite with an aqueous
solution of a
fluorosilicate salt using controlled proportions, temperatures, and pH
conditions which avoid
aluminum extraction without silicon substitution. US-A-4,503,023 sets out that
the
fluorosilicate salt is used as the aluminum extractant and also as the source
of extraneous
silicon which is inserted into the Y zeolite structure in place of the
extracted aluminum. The
salts have the general formula:
(A)2/b SiF6
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wherein A is a metallic or nonmetallic cation other than H+ having the valence
"b." Cations
represented by "A" are alkylammonium, NH4+, Mg++, Li+, Na+, K+, Ba++ Cd++,
Cu++,
H+, Ca++, Cs+, Fe++, Co++, Pb++, Mn++, Rb+, Ag+, Sr++, Ti+, and Zn++.
[0037] A preferred member of this group is known as LZ-210, a zeolitic
aluminosilicate
molecular sieve commercially available from UOP LLC, Des Plaines, Illinois,
U.S.A. LZ-210
zeolites and the other zeolites of this group are conveniently prepared from a
Y zeolite
starting material. The LZ-210 zeolite has an overall silica to alumina mole
ratio from 5.0 to
12.0 in one embodiment, from 5.0 to 11.0 in another embodiment, and from 5.0
to 10.0 in yet
another embodiment. The unit cell size can be preferably from 24.25 to 24.32
angstrom or
more preferably from 24.26 to 24.30 angstrom in the case of Y Zeolite I. In
the case of Y
Zeolite Il, the unit cell size can be from 24.33 to 24.38 angstrom or
preferably from 24.34 to
24.36 angstrom. The LZ-210 class of zeolites used in the process and
composition disclosed
herein have a composition expressed in terms of mole ratios of oxides as in
the following
formula:
(0.85-1.1)M2/nO : A1203 : xSiO2
wherein "M" is a cation having the valence "n" and "x" has a value from 5.0 to
12Ø
[0038] In general, LZ-210 zeolites may be prepared by dealuminating Y-type
zeolites
using an aqueous solution of a fluorosilicate salt, preferably a solution of
ammonium
hexafluorosilicate. The dealumination can be accomplished by placing a Y
zeolite, normally
but not necessarily an ammonium exchanged Y zeolite, into an aqueous reaction
medium
such as an aqueous solution of ammonium acetate, and slowly adding an aqueous
solution of
ammonium fluorosilicate. After the reaction is allowed to proceed, a zeolite
having an
increased overall silica to alumina mole ratio is produced. The magnitude of
the increase is
dependent at least in part on the amount of fluorosilicate solution contacted
with the zeolite
and on the reaction time allowed. Normally, a reaction time of between 10 and
24 hours is
sufficient for equilibrium to be achieved. The resulting solid product, which
can be separated
from the aqueous reaction medium by conventional filtration techniques, is a
form of LZ-210
zeolite. In some cases this product may be subjected to a steam calcination by
methods well
known in the art. For instance, the product may be contacted with water vapor
at a partial
pressure of at least 1.4 kPa(a) (0.2 psi(a)) for a period of between 1/4 to 3
hours at a
temperature between 482 C (900 F) and 816 C (1500 F) in order to provide
greater
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crystalline stability. In some cases the product of the steam calcination may
be subjected to an
ammonium-exchange by methods well known in the art. For instance, the product
may be
slurried with water after which an arnmonium salt is added to the slurry. The
resulting
mixture is typically heated for a period of hours, filtered, and washed with
water. Methods of
steaming and ammonium-exchanging LZ-210 zeolite are described in US-A-
4,503,023; US-
A-4,735,928 and US-A-5,275,720.
[0039] Optional Y Zeolite I prepared by the above-discussed preparation
procedures and
used in the process and composition disclosed herein have the essential X-ray
powder
diffraction pattern of zeolite Y and a unit cell size or dimension ao of
preferably from 24.25
to 24.32 angstrom, more preferably from 24.26 to 24.30 angstrom. The Y Zeolite
11 prepared
by the above-discussed preparation procedures and used in the process and
composition
disclosed herein have the essential X-ray powder diffraction pattern of
zeolite Y and a unit
cell size or dimension ao of from 24.33 to 24.38 angstrom, preferably from
24.34 to 24.36
angstrom. Y Zeolite I, Y Zeolite II, or both can have an overall silica to
alumina mole ratio of
from 5.0 to 12.0 in one embodiment, from 5.0 to 11.0 in another embodiment,
and from 5.0 to
10.0 in yet another embodiment. Y Zeolite I and/or Y Zeolite II may have a
surface area
(BET) of at least 500 m2/g, less than 800 m2/g, often less than 700 m2/g and
typically from
500 to 650 m2/g.
[0040] Another method of increasing the stability and/or acidity of the Y
zeolites is by
exchanging the Y zeolite with polyvalent metal cations, such as rare earth-
containing cations,
magnesium cations or calcium cations, or a combination of ammonium ions and
polyvalent
metal cations, thereby lowering the sodium content until it is as low as the
values described
above after the first or second ammonium exchange steps. Methods of carrying
out the ion
exchange are well known in the art.
[0041] The catalyst used in the process disclosed herein is intended primarily
for use as a
replacement catalyst in existing commercial hydrocracking units. Its size and
shape is,
therefore, preferably similar to those of conventional commercial catalysts.
It is preferably
manufactured in the form of a cylindrical extrudate having a diameter of from
0.8 -3.2 mm
(1/32 - 1/8 in). The catalyst can however be made in any other desired form
such as a sphere
or pellet. The extrudate may be in forms other than a cylinder such as the
form of a well-
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known trilobal or other shape which has advantages in terms or reduced
diffusional distance
or pressure drop.
[0042] Commercial hydrocracking catalysts contain a number of non-zeolitic
materials.
This is for several reasons such as particle strength, cost, porosity, and
performance. The
other catalyst components, therefore, make positive contributions to the
overall catalyst even
if not as active cracking components. These other components are referred to
herein as the
support. Some traditional components of the support such as silica-alumina
normally make
some contribution to the cracking capability of the catalyst. In embodiments
of the process
and composition disclosed herein, the catalyst contains a relatively small
content of beta
zeolite. The catalyst contains from 0.5 to 5 wt-%, preferably from 0.7 to 2.6
wt-%, of beta
zeolite based on the combined weight of the beta zeolite, Y Zeolite I (if
any), Y Zeolite II, and
the support all on a dried basis. As used herein, the weight on a dried basis
is considered to be
the weight after heating in dry air at 500 C (932 F) for 6 hours. The catalyst
has a weight
ratio of the Y Zeolite I to the beta zeolite of from 0.5 to 5, preferably from
0.5 to 2.0, on a
dried basis. When the optional Y Zeolite I is present, the catalyst has a
weight ratio of the Y
Zeolite I to the Y Zeolite II of from 1.5 to 8, preferably from 2 to 6.5, on a
dried basis. When
the optional Y Zeolite I is present, the catalyst contains from more than 5 wt-
% to at most 15
wt-%, of the Y Zeolite I and the Y Zeolite II based on the combined weight of
the beta
zeolite, the Y Zeolite I, the Y Zeolite II, and the support, all on a dried
basis.
[0043] The remainder of the catalyst particle besides the zeolitic material
may be taken up
primarily by conventional hydrocracking materials such as alumina and/or
silica-alumina. The
presence of silica-alumina helps achieve the desired performance
characteristics of the
catalyst. In one embodiment the catalyst contains at least 25 wt-% alumina and
at least
wt-% silica-alumina, both based on the combined weight of the zeolites and the
support,
25 all on a dried basis. In another embodiment, the silica-alumina content of
the catalyst is above
40 wt-% and the alumina content of the catalyst is above 20 wt-%, both based
on the
combined weight of the zeolites and the support, all on a dried basis.
However, the alumina is
believed to function only as a binder and to not be an active cracking
component. The catalyst
support may contain over 50 wt-% silica-alumina or over 50 wt-% alumina based
on the
weight of the support on a dried basis. Approximately equal amounts of silica-
alumina and
alumina are used in an embodiment. Other inorganic refractory materials which
may be used
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as a support in addition to silica-alumina and alumina include for example
silica, zirconia,
titania, boria, and zirconia-alumina. These aforementioned support materials
may be used
alone or in any combination.
[0044] Besides the beta zeolite, the Y zeolite, and other support materials,
the subject
catalyst contains a metallic hydrogenation component. The hydrogenation
component is
preferably provided as one or more base metals uniformly distributed in the
catalyst particle.
The hydrogenation component is one or more element components from Groups 6,
9, and 10
of the periodic table. Noble metals such as platinum and palladium could be
applied but best
results have been obtained with a combination of two base metals.
Specifically, either nickel
or cobalt is paired with tungsten or molybdenum, respectively. The preferred
composition of
the metal hydrogenation component is both nickel and molybdenum or both nickel
and
tungsten. The amount of nickel or cobalt is preferably between 2 and 8 wt-% of
the finished
catalyst. The amount of tungsten or molybdenum is preferably between 8 and 22
wt-% of the
finished catalyst. The total amount of a base metal hydrogenation component is
from 10 to 30
wt-% of the finished catalyst.
[0045] The catalyst of the subject process can be formulated using industry
standard
techniques. This can, with great generalization, be summarized as admixing the
beta zeolite
and the Y zeolite with the other inorganic oxide components and a liquid such
as water or a
mild acid to form an extrudable dough followed by extrusion through a
multihole die plate.
The extrudate is collected and preferably calcined at high temperature to
harden the extrudate.
The extruded particles are then screened for size and the hydrogenation
components are added
as by dip impregnation or the well known incipient wetness technique. If the
catalyst contains
two metals in the hydrogenation component these may be added sequentially or
simultaneously. The catalyst particles may be calcined between metal addition
steps and again
after the metals are added.
[0046] In another embodiment, it may be convenient or preferred to combine the
porous
inorganic refractory oxide, the beta zeolite the Y zeolite, and compound(s)
containing the
metal(s), then to comull the combined materials, subsequently to extrude the
comulled
material, and finally to calcine the extruded material. The comulling is
effected with a source
of metal, such as ammonium heptamolybdate or ammonium metatungstate and
another source
of another metal, such as nickel nitrate or cobalt nitrate, with both source
compounds
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generally being introduced into the combined materials in the form of an
aqueous solution or
as a salt. Other metals can be similarly introduced in dissolved aqueous form
or as a salt.
Likewise, non-metallic elements, e.g., phosphorus, may be introduced by
incorporating a
soluble component such as phosphoric acids into the aqueous solution when
used.
[0047] Yet other methods of preparation are described US-A-5,279,726 and
US-A-5,350,501, which are hereby incorporated herein by reference in their
entireties.
[0048] Catalysts prepared by the above-discussed procedures contain the
hydrogenation
metals in the oxide form. The oxide form is generally converted to the sulfide
form for
hydrocracking. This can be accomplished by any of the well known techniques
for sulfiding,
including ex situ presulfiding prior to loading the catalyst in the
hydrocracking reactor,
presulfiding after loading the catalyst in the hydrocracking reactor and prior
to use at an
elevated temperature, and in situ sulfiding, i.e., by using the catalyst in
the oxide form to
hydrocrack a hydrocarbon feedstock containing sulfur compounds under
hydrocracking
conditions, including elevated temperature and pressure and the presence of
hydrogen.
[0049] The hydrocracking process disclosed herein will be operated within the
general
range of conditions now employed commercially in hydrocracking processes. The
operating
conditions in many instances are refinery or processing unit specific. That
is, they are dictated
in large part by the construction and limitations of the existing
hydrocracking unit, which
normally cannot be changed without significant expense, the composition of the
feed and the
desired products. The inlet temperature of the catalyst bed should be from 232
C (450 F) to
454 C (850 F), and the inlet pressure should be from 5171 kPa(g) (750 psi(g))
to 24132
kPa(g) (3500 psi(g)), and typically from 6895 kPa(g) (1000 psi(g)) to 24132
kPa(g) (3500
psi(g)). The feed stream is admixed with sufficient hydrogen to provide a
volumetric
hydrogen circulation rate per unit volume of feed of 168 to 1684 normal
ltr/ltr measured at
0 C (32 F) and 101.3 kPa(a) (14.7 psi(a)) (1000 to 10000 standard it3/barrel
(SCFB)
measured at 15.6 C (60 F) and 101.3 kPa(a) (14.7 psi(a))) and passed into one
or more
reactors containing fixed beds of the catalyst. The hydrogen will be primarily
derived from a
recycle gas stream which may pass through purification facilities for the
removal of acid
gases although this is not necessary. The hydrogen rich gas admixed with the
feed and in one
embodiment any recycle hydrocarbons will usually contain at least 75 mol
percent hydrogen.
For hydrocracking to produce distillate the feed rate in terms of LHSV will
normally be
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within the broad range of 0.3 to 3.0 hr 1. As used herein, LHSV means liquid
hourly space
velocity, which is defined as the volumetric flow rate of liquid per hour
divided by the
catalyst volume, where the liquid volume and the catalyst volume are in the
same volumetric
units.
[0050] The typical feed to the process disclosed herein is a mixture of many
different
hydrocarbons and coboiling compounds recovered by fractional distillation from
a crude
petroleum. It will normally have components that boil higher than the upper
end of the range
of the 149 C (300 F) to 371 C (700 F) boiling range in order to produce
distillate. Often it
will have a boiling point range starting above 340 C (644 F) and ending in one
embodiment
below 482 C (900 F), in another embodiment below 540 C (1004 F), and in a
third
embodiment below 565 C (1049 F). Such a petroleum derived feed may be a blend
of
streams produced in a refinery such as atmospheric gas oil, coker gas oil,
straight run gas oil,
deasphalted gas oil, vacuum gas oil, and FCC cycle oil. A typical gas oil
comprises
components that boil in the range of from 166 C (330 F) to 566 C (1050 F).
Alternatively,
the feed to the process disclosed herein can be a single fraction such as a
heavy vacuum gas
oil. A typical heavy gas oil fraction has a substantial proportion of the
hydrocarbon
components, usually at least 80 percent by weight, boiling from 371 C (700 F)
to 566 C
(1050 F). Synthetic hydrocarbon mixtures such as recovered from shale oil or
coal can also
be processed in the subject process. The feed may be subjected to
hydrotreating or treated as
by solvent extraction prior to being passed into the subject process to remove
gross amounts
of sulfur, nitrogen or other contaminants such as asphaltenes.
[0051] The subject process is expected to convert a large portion of the feed
to more
volatile hydrocarbons such as distillate boiling range hydrocarbons. Typical
conversion rates
vary from 50 to 100 volume-percent (hereinafter vol-%) depending greatly on
the feed
composition. The conversion rate is between from 60 to 90 vol-% in an
embodiment of the
process disclosed herein, from 70 to 90 vol-% in another embodiment, from 80
and to 90 vol-
% in yet another embodiment, and from 65 to 75 vol-% in still another
embodiment. The
effluent of the process will actually contain a broad variety of hydrocarbons
ranging from
methane to essentially unchanged feed hydrocarbons boiling above the boiling
range of any
desired product. The effluent of the process typically passes from a reactor
containing a
catalyst and is usually separated by methods known to a person of ordinary
skill in the art,
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including phase separation or distillation, to produce a product having any
desired final
boiling point. The hydrocarbons boiling above the final boiling point of any
desired product
are referred to as unconverted products even if their boiling point has been
reduced to some
extent in the process. Most unconverted hydrocarbons are recycled to the
reaction zone with a
small percentage, e.g. 5 wt-% being removed as a drag stream. For producing
distillate, at
least 30 wt-%, and preferably at least 50 wt-%, of the effluent boils below
371 C (700 F).
[0052] The process and composition disclosed herein can be employed in what
are
referred to in the art as single stage and two stage process flows, with or
without prior
hydrotreating. These terms are used as defined and illustrated in the book
titled
Hydrocracking Science and Technology, by J. Scherzer and A.J. Gruia, ISBN 0-
8247-9760-4,
Marcel Dekker Inc., New York, 1996. In a two-stage process the subject
catalyst can be
employed in either the first or second stage. The catalyst may be preceded by
a hydrotreating
catalyst in a separate reactor or may be loaded into the same reactor as a
hydrotreating catalyst
or a different hydrocracking catalyst. An upstream hydrotreating catalyst can
be employed as
feed pretreatment step or to hydrotreat recycled unconverted materials. The
hydrotreating
catalyst can be employed for the specific purpose of hydrotreating polynuclear
aromatic
(PNA) compounds to promote their conversion in subsequent hydrocracking
catalyst bed(s).
The subject catalyst can also be employed in combination with a second,
different catalyst,
such as a catalyst based upon Y zeolite or having primarily amorphous cracking
components.
[0053] In some embodiments of the process disclosed herein, the catalyst is
employed
with a feed or in a configuration that the feed passing through the catalyst
is a raw feed or
resembles a raw feed. The sulfur content of crude oil, and hence the feed to
this process,
varies greatly depending on its source. As used herein, a raw feed is intended
to refer to a feed
which has not been hydrotreated or which still contains organic sulfur
compounds which
result in a sulfur level above 1000 wt-ppm or which still contains organic
nitrogen
compounds that result in a nitrogen level above 100 wt-ppm (0.01 wt-%).
[0054] In other embodiments of the process disclosed herein, the catalyst is
used with a
feed that has been hydrotreated. Persons of ordinary skill in the art of
hydrocarbon processing
know and can practice hydrotreating of a raw feed to produce a hydrotreated
feed to be
charged to the process disclosed herein. Although the sulfur level of the
hydrotreated feed
may be between 500 and 1000 wt-ppm, the sulfur level of the hydrotreated feed
is less than
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500 wt-ppm in one embodiment of the process disclosed herein and from 5 to 500
wt-ppm in
another embodiment. The nitrogen level of the hydrotreated feed is less than
100 wt-ppm in
one embodiment and from I to 100 wt-ppm in another embodiment.
[00551 All references herein to the groups of elements of the periodic table
are to the
IUPAC "New Notation" on the Periodic Table of the Elements in the inside front
cover of the
book titled CRC Handbook of Chemistry and Physics, ISBN 0-8493-0480-6, CRC
Press,
Boca Raton, Florida, U.S.A., 80th Edition, 1999-2000. All references herein to
surface area
are to single-point surface areas at a nitrogen partial pressure of p/po of
0.03 as determined by
the BET (Brunauer-Emmett-Teller) method using nitrogen adsorption technique as
described
in ASTM D4365-95, Standard Test Method for Determining Micropore Volume and
Zeolite
Area of a Catalyst, and in the article by S. Brunauer et al., J. Am. Chem.
Soc., 60(2), 309-319
(1938). All references herein to boiling points are to boiling points as
determined by ASTM
D2887, Standard Test Method for Boiling Range Distribution of Petroleum
Fractions by Gas
Chromatography. ASTM methods are available from ASTM International, 100 Barr
Harbor
Drive, P.O. Box C700, West Conshohocken, Pennsylvania, U.S.A.
[00561 The following examples are provided for illustrative purposes and not
to limit the
process and composition as defined in the claims.
EXAMPLE 1
Sample 1
[00571 A modified Y zeolite was prepared by steaming an ammonium exchanged Y
zeolite sold by UOP LLC (Des Plaines, Illinois, USA) and referred to in the
literature as Y-84
having a sodium content of less than 0.2 wt-% calculated as Na20. The
resulting modified Y
zeolite is referred to herein as Sample 1 and had an overall silica to alumina
(Si02 to A1203)
mole ratio of 5.0 to 5.5, a unit cell size of 24.28 angstrom, and a surface
area of 540 to 640
m2/g. Sample 1, which is an example of Y Zeolite I, is referred to in the
Table as Y1.
Sample AW 1
[0058] A sample of Sample 1 was acid washed. The resulting acid-washed
modified Y
zeolite is referred to herein as Sample AW 1 and had an overall silica to
alumina (Si02 to
A1203) mole ratio of 11.0, a unit cell size of 24.28 angstrom, and a surface
area of 570 to 750
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m2/g. Sample AW 1, which is an example of Y Zeolite I, is referred to in the
Table as AW
Y l.
Sample 2
[0059] A modified Y zeolite was prepared in a manner similar to that described
for
Sample 1, except the steaming conditions were different. The resulting
modified Y zeolite is
referred to herein as Sample 2 and had an overall silica to alumina (Si02 to
A1203) mole
ratio of from 5.0 to 5.5, a unit cell size of 24.35 angstrom, and a surface
area of 630 to 730
m2/g. Sample 2, which is an example of Y Zeolite II, is referred to in the
Table as Y2.
EXAMPLE 2
[0060] Eight catalysts (A-H) were prepared by mixing Sample 1 if present,
Sample AW 1
if present, Sample 2 if present, a beta zeolite having an overall silica to
alumina (Si02 to
A1203) mole ratio of 23.8 and an SF6 adsorption capacity of 29 wt-% if
present, amorphous
silica-alumina, and HNO3-peptized CatapalTM C boehmite alumina in a muller.
The beta
zeolite, which had an overall silica to alumina (Si02 to A1203) mole ratio of
23.8 and an SF6
adsorption capacity of 29 wt-%, either contained the template used during its
synthesis or had
been subsequently calcined at mild conditions to remove the aforementioned
template. In the
Table, the beta zeolite containing the template is referred to as Beta 1, and
the calcined beta
zeolite is referred to as Beta 2. The amorphous silica-alumina was either CCIC
silica-alumina
which had a nominal composition of 75 wt-% silica and 25 wt-% alumina, or
Sira140 silica-
alumina, which had a nominal composition of 40 wt-% silica and 60 wt-%
alumina. CCIC
silica-alumina is available from Catalysts & Chemicals Industries Co. Ltd.
(CCIC), and
Catapal C alumina and Sira140 silica-alumina are available from Sasol Germany
GmbH. The
amounts of these components on a dried basis in each final catalyst are listed
in the Table.
The resulting mixture was extruded into 1.6 mm (1/16 in) diameter cylindrical
particles of
between 3.2 mm (1/8 in) and 12.7 mm (1/2 in) in length. The wet extrudates
were dried at
104 C (220 F) for a minimum of 4 hr and then calcined at temperatures in
excess of 550 C
(1022 F) for a minimum of 90 minutes. For catalysts A-F and H, sufficient
nickel nitrate to
provide 4 wt-% nickel (calculated as Ni) in the final catalyst and sufficient
ammonium
metatungstate to provide 14 wt-% tungsten (calculated as W) in the final
catalyst were then
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added to the calcined extrudates to incipient wetness, while for catalysts G
and H the
corresponding amounts were 5 wt-% nickel and 17.5 wt-% tungsten. The
extrudates were
then dried to be free-flowing, and then oxidized by calcining at 500 C (932 F)
for a minimum
of 90 minutes. Catalyst I is a standard hydrocracking catalyst containing on
average 5.5 wt-%
nickel and 17.5 wt-% tungsten. It is believed that the differences in nickel
and tungsten
contents do not have a significant effect on the hydrocracking activity and
selectivity results
described in these examples.
EXAMPLE 3
[0061] Each of the above-described nine catalysts was pre-sulfided by passing
a gas
stream consisting of 10 vol-% H2S and the balance H2 through a bed of the
catalyst at a
temperature initially of 149 C (300 F) and slowly raised to 413 C (775 F) and
held at the
temperature for 6 hours.
[0062] The nine catalysts were compared for hydrocracking activity and
selectivity (i.e.,
product yields) in simulated first stage testing. Specifically, the nine
catalysts were separately
tested for hydrocracking a hydrotreated light Arabian vacuum gas oil (VGO)
feed having a
specific gravity of 0.877 at 15.6 C (60 F) (API gravity of 30.05 ), an initial
boiling point of
107 C (224 F), a 5 wt-% boiling point of 195 C (382 F), a final boiling point
of 550 C
(1021 F), and a 50 wt-% boiling point of 24 C (795 F), with 13 wt-% boiling
below 288 C
(550 F) and 26 wt-% boiling below 371 C (700 F).
[0063] Each catalyst was tested for simulated first stage operation by passing
the
feedstock through a laboratory size reactor at a LHSV of 1.5 hr-1, a total
pressure of 13786
kPa(g) (2000 psi(g)), and a volumetric hydrogen feed rate per unit volume of
feed of 1684
normal ltr/ltr measured at 0 C (32 F) and 101.3 kPa(a) (14.7 psi(a)) (10000
SCFB measured
at 15.6 C (60 F) and 101.3 kPa(a) (14.7 psi(a))). Sufficient di-tert-butyl
disulfide was added
to the feed to provide 2.1 wt-% sulfur and thereby to simulate a hydrogen
sulfide-containing
atmosphere as it exists in commercial first stage hydrocracking reactors. In
addition,
sufficient cyclohexylamine was added to the feed to provide 780 wt-ppm
nitrogen and thereby
to simulate an ammonia-containing atmosphere as it exists in commercial first
stage
hydrocracking reactors.
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CA 02657311 2009-01-08
WO 2008/011289 PCT/US2007/073043
[0064] For hydrocracking tests to produce distillate, the temperature
conditions were
adjusted as necessary to maintain a 65 wt-% net conversion to materials
boiling below 371 C
(700 F), over the course of 100 hours. Net conversion is the effluent boiling
below 371 C
(700 F) as a percentage of the feed minus the percentage of the feed boiling
below 371 C
(700 F). At the end of the 100 hours, the temperature required to maintain the
65 wt-% net
conversion was recorded, and the relative activities and selectivities of each
catalyst were
calculated. These data are summarized in the Table. The selectivity values for
each catalyst
were total distillate (i.e., 149 C (300 F) to 371 C (700 F)), light distillate
(i.e., 149 C
(300 F) to 288 C (550 F)), and heavy distillate (i.e., 288 C (550 F) to 371 C
(700 F)). The
relative activity value for each catalyst is entered as the difference between
the required
temperature of the catalyst to maintain the 65 wt-% net conversion and a
reference
temperature that was the same for all nine catalysts. The lower the value for
relative activity,
the more active is the catalyst.
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CA 02657311 2009-01-08
WO 2008/011289 PCT/US2007/073043
O O O O v v ¾¾¾ N N v t~ ~
- O O O ~ z z z + v M O
O O O O O O O ¾¾ r - O-' t~ M~
N Q +
M O
O O O O O O Q¾¾ ? v O ^' oo
oo O v
C7 O O O O OM z z z ++~ M O
O v O v v~ n~ p ~ O M M ~ M O
oo ~O tn m Ln M t\ o
Ln O O ~ Ln N N
(n ~ O p Mrn M ~ oo Ln
++ v M O
Ln v~ Ln v~ O~ O O O t~ t~ M v v -O
W O Ln o~ Lr'
+ + n M o
~ N O V't V't N N ~ M O O ~~O N M O~ N
U oo
+ v't M
Ln v~ O O O O O MN v~ O O O
Q ~} r- p Lq
+L/=t M O
Ln V't ¾ O O V't N M --~
¾ M M N N N z + + iPt M
U
0
U o U o M V
4-+ O '" '" N N o pp ^ 0 a
00
bA ~o =_ . U ~ .~ M ~ N M U 00 ~ oo
+ 00
73 0 w ~v
~ 3 N N ~ ~ 3 3 w )
0
C ~ N O U ~ U ++ M U= d M ULn U O
.~ = ~ ~ ~ -
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u U1tV v~ ~~ .. a g O, Ln oo gõ: Q
U ~ ~"'
.~ -- N ~~ z
CA 02657311 2009-01-08
WO 2008/011289 PCT/US2007/073043
[0065] Fig. 1 is a chart of the 149 C (300 F) to 371 C (700 F) cut distillate
selectivity of
Catalysts A-I plotted versus the relative catalyst activity expressed in terms
of reactor
temperature above the reference temperature required to achieve 65 wt-% net
conversion of
the VGO to the total distillate cut. Catalysts A-F (squares) show more total
distillate
selectivity at a given relative activity than Catalysts G-I (diamonds).
[0066] Fig. 2 is a chart of the weight ratio of the heavy distillate cut
selectivity to the light
distillate cut selectivity versus the relative activity. Catalysts A-F
(squares) show a
significantly higher selectivity of heavy distillate relative to light
distillate compared to
Catalysts G-I (diamonds).
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