Hydrocracking process for producing middle distillates

The selectivity of a midbarrel hydrocracking process for middle distillates is significantly increased by using a catalyst containing a Y zeolite that has been exchanged with rare earth cations. The rare earth-exchanged Y zeolite has a unit cell size below 24.45 angstroms and/or a water vapor sorptive capacity less than 10 weight percent at 25.degree. C. and a p/p.sub.o value of 0.1.

BACKGROUND OF THE INVENTION 
This invention relates to a catalytic hydrocracking process and a catalyst 
for use therein. The invention is particularly concerned with an improved 
process for producing middle distillate products using a catalyst highly 
selective for such products. 
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. Feedstocks 
most often subjected to hydrocracking are gas oils and heavy gas oils 
recovered from crude oil by distillation. A typical gas oil comprises a 
substantial proportion of hydrocarbon components boiling above about 
700.degree. F., usually at least about 80 percent by weight boiling above 
700.degree. F. A typical heavy gas oil has a boiling point range between 
about 600.degree. F. and 1050.degree. F. 
Hydrocracking is generally accomplished by contacting, in an appropriate 
reaction vessel, the gas oil or other feedstock to be treated with a 
suitable hydrocracking catalyst under conditions of elevated temperature 
and pressure in 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 the prime factor in determining such yields. 
At the present time, middle distillates are not in high demand relative to 
gasoline in the United States; however, marketing surveys indicate that 
there will be an increased demand for middle distillates as the year 2000 
approaches. For this reason, refiners have recently been focusing on 
midbarrel hydrocracking catalysts which selectively produce middle 
distillate fractions, such as turbine fuel and diesel fuel, that boil in 
the 300.degree. F. to 700.degree. F. range. 
The three main catalytic properties by which the performance of a midbarrel 
hydrocracking catalyst 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 about 60 percent, of products boiling below 
700.degree. F. The lower the activity temperature for a given catalyst, 
the more active such a catalyst is in relation to a catalyst of higher 
activity temperature. Selectivity of midbarrel or middle distillate 
hydrocracking catalysts may be determined during the foregoing described 
activity test and is measured as the percentage fraction of the 
700.degree. F.--product boiling in the desired midbarrel product range, 
e.g., 300.degree. F. to 700.degree. F. for diesel fuel and 300.degree. F. 
to 550.degree. F. for turbine fuel. 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 60 percent or other given conversion. 
U.S. Pat. Nos. 4,062,809 and 4,419,271, the disclosures of which are hereby 
incorporated by reference in their entireties, disclose two different 
types of very effective middle distillate hydrocracking catalysts. The 
catalyst of U.S. Pat. No. 4,062,809 contains molybdenum and/or tungsten 
plus nickel and/or cobalt on a support of silica-alumina dispersed in 
gamma alumina. U.S. Pat. No. 4,419,271 teaches that the catalyst of U.S. 
Pat. No. 4,062,809 can be improved by adding an aluminosilicate zeolite to 
the support, thereby producing a catalyst containing molybdenum and/or 
tungsten and nickel and/or cobalt supported on a mixture of an 
aluminosilicate zeolite, preferably an ultrahydrophobic zeolite known as 
LZ-10 zeolite, and a dispersion of silica-alumina in a gamma alumina 
matrix. The presence of the zeolite in this catalyst increases the 
activity of the catalyst without significantly affecting the selectivity. 
Although the catalysts of the above-discussed patents are highly effective 
middle distillate hydrocracking catalysts and have proven themselves in 
commercial environments, there is always a demand for new hydrocracking 
catalysts with superior overall activity, selectivity, and stability for 
middle distillate hydrocracking. 
SUMMARY OF THE INVENTION 
In accordance with the invention, it has now been surprisingly found that 
the activity and selectivity of middle distillate catalysts comprising one 
or more hydrogenation components and a Y zeolite having either a unit cell 
size below about 24.45 angstroms or a water vapor sorptive capacity less 
than about 10 weight percent at 25.degree. C. and p/p.sub.o value of 0.10 
can be substantially improved by exchanging the Y zeolite with rare 
earth-containing cations. 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, which product contains a relatively large 
proportion of components boiling in the midbarrel range of 300.degree. F. 
to 700.degree. F. The hydrocracking catalyst normally contains one or more 
hydrogenation components, such as one or more Group VIB or Group VIII 
metal components, in combination with the rare earth-exchanged Y zeolite 
and also usually contains an amorphous, inorganic refractory oxide, such 
as a dispersion of silica-alumina particles in an alumina matrix, and a 
porous, inorganic refractory oxide binder, such as alumina. As used herein 
"p/p.sub.o " represents the water vapor partial pressure to which the Y 
zeolite is exposed divided by the water vapor partial pressure at 
25.degree. C. 
Preliminary tests indicate that the catalyst of the invention, when used in 
hydrocracking to produce middle distillate products such as diesel fuel 
and turbine fuel, has a somewhat higher activity and a surprisingly 
greater selectivity than other middle distillate catalysts now 
commercially available for use in midbarrel hydrocracking processes. These 
tests surprisingly indicate that, at 60 percent conversion, greater than 
about 86 volume percent of the 700.degree. F.--product, typically greater 
than about 88 volume percent, boils in the range between 300.degree. F. 
and 700.degree. F. while greater than about 80 volume percent, frequently 
greater than about 83 volume percent, boils in the range between 
300.degree. F. and 550.degree. F. 
DETAILED DESCRIPTION OF THE INVENTION 
The hydrocracking process of the invention is directed to the production of 
high yields of middle distillates such as diesel fuel, which, as defined 
herein, boils in the 300.degree. F. to 700.degree. F. range, and turbine 
fuel, which, as defined herein, boils in the 300.degree. F. to 550.degree. 
F. range. These high yields are obtained utilizing a catalyst containing 
one or more hydrogenation components in combination with a Y zeolite 
having a unit cell size below about 24.45 angstroms and/or a water vapor 
sorptive capacity less than about 10 weight percent at 25.degree. C. and 
p/p.sub.o value of 0.10, which zeolite has been exchanged with rare 
earth-containing cations. The Y zeolite preferably has an overall 
silica-to-alumina mole ratio less than 6.0, usually between about 4.5 and 
5.6. Preferably, the catalyst also comprises a porous, inorganic 
refractory oxide binder and an amorphous, inorganic refractory oxide 
diluent, which may or may not have catalytic activity. 
The Y zeolite which comprises the midbarrel hydrocracking catalyst of the 
invention has either (1) a unit cell size less than about 24.45 angstroms 
or (2) a sorptive capacity for water vapor at 25.degree. C. and a 
p/p.sub.o value of 0.10 of less than 10 weight percent, preferably less 
than 5 weight percent. Preferred Y zeolites meet both of the foregoing 
requirements. The term "Y zeolite" as employed herein is meant to 
encompass all crystalline zeolites having either the essential X-ray 
powder diffraction pattern set forth in U.S. Pat. No. 3,130,007 or a 
modified Y zeolite having an X-ray powder diffraction pattern similar to 
that of U.S. Pat. No. 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 zeolite 
into a catalytically active and stable form. The present invention 
requires a Y zeolite having either or both of the two properties mentioned 
above, such Y zeolites being modified Y zeolites in comparison to the Y 
zeolite taught in U.S. Pat. No. 3,130,007. 
The Y zeolites used in the catalyst of the invention are large pore 
zeolites having an effective pore size greater than 7.0 angstroms. Since 
some of the pores of the Y zeolites are relatively large, the zeolites 
allow molecules relatively free access to their internal structure. Thus, 
the Y zeolites useful in the composition of the invention generally have a 
low Constraint Index, typically below 1.0, preferably below 0.75, and 
usually below about 0.5. 
The preferred Y zeolite for use in the hydrocracking catalyst of the 
invention is a UHP-Y zeolite, an ultrahydrophobic Y zeolite, that has been 
exchanged with rare earth-containing cations. The composition and 
properties of UHP-Y zeolites are disclosed in U.S. Pat. No. 4,401,556 
herein incorporated by reference in its entirety. See also Great Britain 
Patent 2 014 970 B which is also herein incorporated by reference in its 
entirety. UHP-Y zeolites and similar zeolites are, in essence, produced by 
a four step procedure in which a Y zeolite in the alkali metal form 
(usually sodium) and typically having a unit cell size of about 24.65 
angstroms is cation exchanged with ammonium ions, calcined in the presence 
of water vapor (preferably in the presence of at least 0.2 psia water 
vapor, even more preferably at least 1.0 psia water vapor, and more 
preferably still, at least 10 psia water vapor, and most preferably of 
all, an atmosphere consisting essentially of or consisting of steam) so as 
to produce a unit cell size in the range of 24.40 to 24.64 angstroms, 
preferably 24.42 to 24.62 angstroms, then ammonium exchanged once again, 
and then calcined again in the presence of sufficient water vapor 
(preferably in an atmosphere consisting essentially of steam, and most 
preferably consisting of steam) so as to yield a unit cell size below 
24.40, and most preferably no more than 24.35 angstroms. The first 
ammonium exchange step typically reduces the sodium content of the 
starting sodium Y zeolite from a value usually greater than about 8.0 
weight percent, usually between 10 and 13 weight percent, calculated as 
Na.sub.2 O, to a value in the range between about 0.6 and 5 weight 
percent, while the second ammonium exchange further reduces the sodium 
content to less than about 0.5 weight percent, usually less than 0.3 
weight percent. It will be seen from the above-discussed manufacturing 
procedure that UHP-Y zeolites differ from the Y zeolite taught in U.S. 
Pat. No. 3,929,672 by the addition of the final steam calcination step, 
some of the zeolites of said patent being known under the designations 
Y-82 or LZY-82 and Y-84 or LZY-84. 
"UHP-Y" zeolites are defined herein as zeolite aluminosilicates having an 
overall silica-to-alumina mole ratio greater than 4.5, the essential X-ray 
powder diffraction pattern of zeolite Y, a unit cell size or dimension 
a.sub.o of less than 24.45 angstroms, a surface area of at least 300 
m.sup.2 /g (BET), a sorptive capacity for water vapor at 25.degree. C. and 
a p/p.sub.o value of 0.10 of less than 10 weight percent, and a Residual 
Butanol Test value of not more than 0.40 weight percent. Preferred UHP-Y 
zeolites have one or more of the following properties: an overall 
silica-to-alumina mole ratio from 4.5 to 35; a surface area of at least 
350 m.sup.2 /g, and a sorptive capacity for water vapor at 25.degree. C. 
and a p/p.sub.o value of 0.10 of less than 5 weight percent. Especially 
preferred are UHP-Y zeolites having an overall silica-to-alumina mole 
ratio of 4.5 to 9 and/or a sorptive capacity for water vapor at 
25.degree. C. and a p/p.sub.o value of 0.10 of less than 4 weight percent. 
Although UHP-Y zeolites having silica-to-alumina mole ratios below 6.0 may 
be most preferred, UHP-Y zeolites that have been treated with a mineral 
acid to remove aluminum and thereby increase their overall 
silica-to-alumina molar ratio may also be used. Such acid treated UHP-Y 
zeolites are discussed in detail in U.S. Pat. No. 5,047,139, the 
disclosure of which is herein incorporated by reference in its entirety. 
The more preferred UHP-Y zeolites for use in the present invention have a 
unit cell size or dimension less than about 24.40 angstroms, and even more 
preferably no more than 24.35 angstroms. LZ-10 zeolite is the most 
preferred UHP-Y zeolite, LZ-10 zeolite being available from UOP. LZ-10 
zeolite usually has a unit cell size or dimension at or above 24.20 
angstroms, preferably between 24.20 and 24.40, and most preferably between 
about 24.25 and 24.35 angstroms, and has a water vapor sorptive capacity 
at 4.6 mm water vapor partial pressure and 25.degree. C. less than 8.0 
percent by weight of the zeolite. See U.S. Pat. No. 4,419,271 which 
previously has been incorporated by reference in its entirety. 
As discussed above, the Y zeolites used in the catalyst of the invention 
are typically made by a process which involves two ammonium exchange steps 
to reduce the sodium or other alkali metal content of the starting Y 
zeolite to a value less than 0.5 weight percent sodium, usually less than 
about 0.3 weight percent, calculated as Na.sub.2 O. These zeolites of 
reduced sodium content possess catalytic cracking activity and can be used 
as components of hydrocracking catalysts. In accordance with this 
invention, it has now been found that the selectivity of catalysts 
containing these Y zeolites for middle distillate production can be 
substantially increased by ion exchanging the Y zeolites of reduced sodium 
content with rare earth-containing cations. 
The rare earth metals selected for exchange into the zeolite may be any one 
or any combination of the lanthanide elements having atomic numbers 
according to the Periodic Table of Elements between 57 and 71. Thus, the 
metals suitable for ion exchange herein include lanthanum, cerium, 
praseodymium, neodymium, samarium, europium, gadolinium, terbium, 
dysprosium, holmium, erbium, thulium, ytterbium, and lutetium. In the most 
preferred embodiment of the invention, a mixture of rare earth cations is 
introduced into the zeolite, with the mixture often containing rare earth 
metals in a distribution similar to that of the rare earth ore (e.g., 
bastnasite, monazite, xenotime, and the like) from which the metals were 
derived. 
There are many known methods by which one can exchange rare earth cations 
for sodium and other cations, particularly hydrogen ions, in a crystalline 
aluminosilicate Y zeolite. The most usual way is to contact the zeolite 
with an aqueous solution containing multivalent cations of the rare earth 
element or elements to be exchanged into the zeolite. Most often, the 
solution contains more than about 20 grams per liter of rare earth metal 
cations (calculated as RE.sub.2 O.sub.3 where RE is the sum of all rare 
earth metals under consideration, regardless of whether any one or more of 
such metals actually forms a trioxide of equivalent formula), and the 
contacting is usually accomplished by immersing the zeolite into the 
ion-exchange solution and stirring at ambient temperature or above but 
usually at no more than about 100.degree. C. If desired, the solution may 
also contain ammonium ions, and the solution may further contain any of a 
number of anions that will not interfere with the cation exchange, e.g. 
chloride, nitrate, sulfate, etc. 
For best results, the ion exchange is performed in a manner such that the 
rare earth-exchanged zeolite contains at least about 1 percent, preferably 
at least 2 percent, and usually between about 4 and 6 percent, by weight 
of rare earth metals, calculated as RE.sub.2 O.sub.3. Although a small 
proportion of the rare earth metals exchanged into the zeolite will 
replace some of the residual sodium ions at exchange sites in the zeolite, 
the largest proportion will exchange with hydrogen ions and/or ammonium 
ions because of their relatively high concentration versus the low 
concentration, usually below 0.3 weight percent, calculated as Na.sub.2 O, 
of sodium cations present. Sometimes, only a single immersion of the 
zeolite into the ion exchange solution will be sufficient for the 
necessary exchange. However, in some cases it may be necessary to carry 
out the ion exchange by several immersions into a solution containing rare 
earth metal cations, or by immersion serially into several solutions of 
differing rare earth element content, or by other known methods for 
introducing rare earth metal cations into a zeolite. 
After the ion exchange described above, the rare earth-containing Y zeolite 
is dried and then embodied into support particles which serve to carry one 
or more hydrogenation components. In the preferred method, this is 
accomplished by combining the rare earth-exchanged zeolite with (1) a 
material such as an alumina hydrogel or peptized, Catapal alumina, which, 
upon calcination, will yield a porous, inorganic refractory oxide binder 
or (2) with a material which itself is a porous, inorganic refractory 
oxide binder, for example, alumina, silica-alumina, silica magnesia, a 
clay, such as kaolin, as well as combinations of such materials. A 
sufficient amount of the Y zeolite is normally used such that the support 
comprises between about 1 and about 50 weight percent, preferably between 
about 2 and 30 weight percent, more preferably between about 5 and 20 
weight percent, of the rare earth-exchanged Y zeolite. Perhaps the most 
convenient method for physically integrating the zeolite and the binder is 
to comull the porous, inorganic refractory oxide binder or precursor 
material with the zeolite, and subsequently extrude the comulled material 
through a die having small openings therein of desired cross-sectional 
size and shape, e.g., circle, trilobal clover-leaf, quadralobal clover 
leafs, etc., breaking or cutting the extruded matter into appropriate 
lengths, e.g., 1/16 to 3/4 inch, drying the extrudates, and then calcining 
at a temperature between 800.degree. F. and 1200.degree. F. to produce a 
material suitable for use in high temperature hydrocracking reactions. At 
present it is preferred that the support be produced in cylindrical form; 
however, as stated above, other cross-sectional shapes are possible, such 
as cloverleafs of polylobal design, for example, trilobal or quadralobal 
shapes, as shown, for example, in FIGS. 8 and 10, respectively, in U.S. 
Pat. No. 4,028,227 herein incorporated by reference in its entirety. 
It will be understood, of course, in the foregoing description that the 
porous, inorganic refractory oxide is used as a binder material to hold 
the rare earth-exchanged zeolite particles together in the support, and 
accordingly, if desired, other materials can also be incorporated into the 
comulled mixture, including for example, amorphous, inorganic refractory 
oxide diluents, which may or may not possess some type of catalytic 
activity. Examples of preferred diluents are those having cracking 
activity such as alumina, silica-alumina and a heterogeneous dispersion of 
finely divided silica-alumina particles in an alumina matrix, which 
dispersion is described in detail in U.S. Pat. Nos. 4,062,809 and 
4,419,271. Additionally and alternatively, hydrogenation component 
precursors can also be comulled into the mixture, as will be discussed in 
more detail hereinafter. 
The heterogeneous dispersion of finely divided silica-alumina in an alumina 
matrix mentioned above is a preferred diluent for use in the catalyst and 
may be prepared by comulling an alumina hydrogel with a silica-alumina 
cogel in hydrous or dry form. Alternatively, the alumina hydrogel may be 
comulled with a "graft copolymer" of silica and alumina that has been 
prepared, for example, by first impregnating a silica hydrogel with an 
aluminum salt and then precipitating alumina gel in the pores of the 
silica hydrogel by contact with ammonium hydroxide. In the usual case, the 
cogel or copolymer is mixed or mulled with the alumina hydrogel such that 
the cogel or copolymer comprises between about 5 and 75 weight percent, 
preferably 20 to 65 weight percent, of the mixture. The overall silica 
content of the resulting dispersion on a dry basis is normally between 
about 5 and about 75 weight percent, preferably between about 15 and about 
45 weight percent. Typically, the silica-alumina is dispersed in a gamma 
alumina matrix. 
It will be further understood that producing the catalyst support in 
extrudate form, while certainly the most highly preferred method, is still 
but one option available to those skilled in the art. The catalyst support 
may also be produced in tablet, granules, spheres, and pellets as desired, 
by any known method for combining zeolites with porous, inorganic 
refractory oxide components. Regardless of how the support particles are 
produced, they typically contain between 10 and 40, preferably from 15 to 
30, weight percent binder and between 25 and 90, usually between 40 and 
80, weight percent amorphous diluent. 
After the catalyst support particles are produced, they are converted to 
catalyst particles by compounding, as by impregnation of the particles, 
with one or more precursors of at least one catalytically active 
hydrogenation metal component. The impregnation may be accomplished by any 
method known in the art, including spray impregnation wherein a solution 
containing the hydrogenation metal precursors in dissolved form is sprayed 
onto the support particles. Another method involves soaking the support 
particles in a large volume of the impregnation solution. Yet another 
method is the pore volume or pore saturation technique wherein the support 
particles are introduced into an impregnation solution of volume just 
sufficient to fill the pores of the support. On occasion, the pore 
saturation technique may be modified so as to utilize an impregnation 
solution having a volume between 10 percent less and 10 percent more than 
that which would just fill the pores. If the active metal precursors are 
incorporated by impregnation, a subsequent or second calcination, as for 
example at temperatures between 700.degree. F. and 1200.degree. F., will 
convert the metals to their respective oxide forms. In some cases, 
calcinations may follow each impregnation of individual active metals. 
Alternative methods of introducing the active metal components into the 
catalyst support include (1) mixing an appropriate solid or liquid 
containing the metal components with the materials to be extruded through 
the die and (2) impregnating the materials to be extruded with the desired 
metal components prior to carrying out the extrusion. Such methods may 
prove less expensive and more convenient than the impregnation methods 
discussed above and will also result in the active hydrogenation 
components being intimately mixed with the components of the support. 
Hydrogenation components suitable for incorporation into the extruded 
catalyst support particles comprise metals selected from Group VIII and/or 
Group VIB of the Periodic Table of Elements. As used herein "Periodic 
Table of Elements" refers to the version found in the inside front cover 
of the Handbook of Chemistry and Physics, 65th Edition, published in 1984 
by the Chemical Rubber Company, Cleveland, Ohio. Preferred hydrogenation 
components comprise metals selected from the group consisting of platinum, 
palladium, cobalt, nickel, tungsten, chromium, and molybdenum. Preferably, 
the catalyst contains at least one Group VIII metal component and at least 
one Group VIB metal component, with cobalt or nickel and molybdenum or 
tungsten being preferred combinations of active components and nickel and 
tungsten being most preferred. The catalyst typically contains up to about 
15, usually between about 1 and 10 weight percent, preferably between 2 
and 8 weight percent, of a non-noble Group VIII metal, calculated as the 
monoxide, and up to 30, usually from about 2 to 28 weight percent, and 
preferably between about 10 and 25 weight percent, of the Group VIB metal, 
calculated as the trioxide. If the hydrogenation component comprises a 
noble metal such as platinum or palladium, it is generally desired that 
the catalyst contain between about 0.2 and about 10 weight percent, 
preferably between about 0.30 and 2.0 weight percent, calculated as the 
metal. 
By the foregoing procedures or their equivalents, catalysts with the 
hydrogenation metals present in the oxide form are prepared as 
particulates. The finished hydrocracking catalysts typically have a BET 
surface area ranging between about 100 and 300 m.sup.2 /g. When used to 
selectively produce middle distillates, these catalysts usually comprise 
(1) between about 1 and 40 weight percent, preferably between about 3 and 
20 weight percent, rare earth-containing Y zeolite having a unit cell size 
below about 24.45 angstroms and/or a water vapor sorptive capacity less 
than about 10 weight percent at 25.degree. C. and p/p.sub.o value of 0.10, 
(2) between about 5 and 40 weight percent porous, inorganic refractory 
oxide binder, preferably between about 10 and 25 weight percent, (3) 
between about 20 and about 90 weight percent amorphous, inorganic 
refractory oxide diluent, such as a dispersion of silica-alumina particles 
in a gamma alumina matrix, preferably between about 30 and 80 weight 
percent, (4) between about 2 and 28 weight percent Group VIB metal 
hydrogenation component, preferably between about 10 and 25 weight 
percent, and (5) between about 0.2 and 15 weight percent Group VIII 
hydrogenation metal component, preferably between 2 and 8 weight percent. 
Catalysts prepared in the oxide form as described above are generally 
converted to the sulfide form for hydrocracking purposes. This can be 
accomplished by presulfiding the catalyst prior to use at an elevated 
temperature, e.g., 300.degree. to 700.degree. F., with, for example, a 
mixture consisting of 10 volume percent H.sub.2 S and 90 volume percent 
H.sub.2. The catalyst can be presulfided ex situ by various sulfiding 
processes; as an illustration, see "Sulficat.RTM.: Off-Site Presulfiding 
of Hydroprocessing Catalysts from Eurecat" by J. H. Wilson and G. Berrebi, 
Catalysis 87, Studies in Surface Science and Catalysts Vol. 38, Elsevier 
Science Publishers B. V., 1988, pages 393-398. Alternatively, the 
sulfiding is accomplished in situ, 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. 
The catalysts described above are useful in the conversion of a wide 
variety of hydrocarbon feedstocks via hydrocracking to more valuable 
hydrocarbon products of lower average boiling point and lower average 
molecular weight, which products typically boil in the range between about 
300.degree. F. and about 700.degree. F. The feedstocks that may be 
subjected to hydrocracking by the process of the invention include mineral 
oils and synthetic oils such as shale oil, oil derived from tar sands, 
coal liquids, and the like. Examples of appropriate feedstocks for 
hydrocracking include atmospheric gas oils, vacuum gas oils, and catalytic 
cracker cycle oils. Preferred hydrocracking feedstocks include gas oils 
and other hydrocarbon fractions having at least 50 weight percent of their 
components boiling above 700.degree. F. Such feedstocks typically contain 
individual concentrations of nickel, vanadium, iron and copper less than 
about 8.0 ppmw, preferably less than about 5.0 ppmw, and most preferably 
less than about 1.0 ppmw. Normally, heavy hydrocarbon oils such as a heavy 
crude oil, a reduced crude oil, vacuum distillation residues and similar 
heavy materials are not suitable feedstocks for the process of the 
invention. 
Usually, the feedstocks described above are hydrotreated before being 
subjected to the hydrocracking process of the invention. The hydrotreating 
is performed in conjunction with hydrocracking, usually by a method 
referred to as "integral operation." In this process, the hydrocarbon 
feedstock is introduced into a catalytic hydrotreating zone wherein, in 
the presence of a suitable catalyst and under suitable conditions, 
including an elevated temperature (e.g., 400.degree. to 1000.degree. F.) 
and an elevated pressure (e.g., 100 to 5000 p.s.i.g.) and with hydrogen as 
a reactant, the organonitrogen components and the organosulfur components 
contained in the feedstock are converted to ammonia and hydrogen sulfide, 
respectively. Suitable hydrotreating catalysts include zeolite- or 
sieve-free, particulate catalysts comprising a Group VIII metal component 
and a Group VIB metal component on a porous, inorganic, refractory oxide 
support most often composed of alumina. The entire effluent removed from 
the hydrotreating zone is subsequently treated in the hydrocracking zone 
maintained under suitable conditions of elevated temperature, pressure, 
and hydrogen partial pressure, and containing the hydrocracking catalyst 
of the invention. Usually, the hydrotreating and hydrocracking zones in 
integral operation are maintained in separate reactor vessels, but, on 
occasion, it may be advantageous to employ a single, downflow reactor 
vessel containing one or more upper beds of the hydrotreating catalyst 
particles and one or more lower beds of the hydrocracking catalyst 
particles. Examples of integral operation may be found in U.S. Pat. Nos. 
3,159,564, 3,655,551, 4,040,944, and 4,584,287 all of which are herein 
incorporated by reference in their entireties. In some cases, the effluent 
from the hydrocracking zone is subjected to hydrotreating in a manner 
similar to that described above in order to remove trace mercaptans from 
the product. 
The catalyst of the invention is usually employed as a fixed bed of 
catalytic extrudates in a hydrocracking reactor into which hydrogen and 
the feedstock are introduced and passed in a downwardly direction. The 
reactor vessel is maintained at conditions so as to convert the feedstock 
into the desired product, which is normally a hydrocarbon product 
containing a substantial portion of turbine fuel and diesel fuel 
components boiling in the range between 300.degree. F. and 700.degree. F. 
In general, the temperature of the reaction vessel is maintained between 
about 450.degree. F. and about 850.degree. F., preferably between about 
500.degree. F. and 800.degree. F. The pressure normally ranges between 
about 750 p.s.i.g. and about 3500 p.s.i.g., preferably between about 1000 
and about 3000 p.s.i.g. The liquid hourly space velocity (LHSV) is 
typically between about 0.3 and 5.0, preferably between about 0.5 and 3.0, 
reciprocal hours. The ratio of hydrogen gas to feedstock utilized usually 
ranges between 1,000 and about 10,000 standard cubic feet per barrel, 
preferably between about 2,000 and 8,000 standard cubic feet per barrel, 
as measured at 60.degree. C. and 1 atmosphere. 
The typical gas oil feedstock contains no more than about 35 volume 
percent, usually less than 15 volume percent, constituents boiling in the 
300.degree. F. to 700.degree. F. range. When middle distillates are 
desired, the hydrocracking operation conditions are chosen so that at 
least 80 volume percent, preferably at least 86 volume percent, and most 
preferably at least 88 volume percent, of the 700.degree. F.--product 
boils in the range between 300.degree. F. and 700.degree. F. Usually, the 
700.degree. F.--product contains greater than 75 volume percent, 
preferably at least 80 volume percent, and more preferably greater than 83 
volume percent, hydrocarbons boiling in the range between 300.degree. F. 
and 550.degree. F. 
Based on presently available data, the catalyst of the present invention as 
compared to a commercial middle distillate hydrocracking catalyst 
containing a non-rare earth-exchanged LZ-10 zeolite having a unit cell 
dimension of 24.30 angstroms provides for enhanced results when used to 
selectively produce turbine and diesel fuel. In particular, the catalyst 
of the invention provides for a higher activity and significant increases 
in the yield of hydrocarbon distillates boiling in the 300.degree. F. to 
550.degree. F. range and the 300.degree. F. to 700.degree. F. range. These 
achievements, and others, are proven in the following example which is 
provided for illustrative purposes and not to limit the invention as 
defined by the claims.

EXAMPLE 
Catalyst 1 
Catalyst 1, a catalyst of the invention, was prepared by mixing 10 weight 
percent of a rare earth-exchanged LZ-10 zeolite, 70 weight percent of a 
dispersion of silica-alumina particles in a gamma alumina matrix (Aero 
5545 obtained from Criterion Catalyst Company L. P.), which dispersion 
contained about 55 weight percent alumina and about 45 weight percent 
silica, and 20 weight percent peptized Catapal alumina binder. The mixture 
was mulled and then extruded through a 1/16-inch cylindrical die to form 
cylindrical extrudates that were cut into 1/8 to 1/2 inch lengths. The 
extrudates were dried at 100.degree. C. and then calcined at 900.degree. 
F. The dried and calcined extrudates were then impregnated by the pore 
saturation method with an aqueous solution containing nickel nitrate and 
ammonium metatungstate in sufficient quantities such that, after the 
impregnated extrudates were dried at 100.degree. C. and calcined at 
900.degree. F., the resultant catalyst particles contained about 5 weight 
percent nickel, calculated as NiO, and about 22 weight percent tungsten, 
calculated as WO.sub.3, on a support consisting of 10 weight percent rare 
earth-exchanged LZ-10 zeolite, 70 weight percent dispersion, and 20 weight 
percent alumina binder. 
The LZ-10 zeolite used to make Catalyst 1 had a unit cell size of 24.30 
angstroms, an effective pore size above about 7.0 angstroms and an overall 
silica-to-alumina mole ratio of about 5.2. The LZ-10 zeolite, which was 
obtained from UOP, was exchanged with rare earth cations by slurrying the 
zeolite in a solution of mixed rare earth chlorides. One thousand (1,000) 
grams of the LZ-10 zeolite were stirred at room temperature for one hour 
with a solution containing 100 grams of rare earth chlorides in 2,500 ml 
of water. The slurry was filtered and the resultant solids washed with 
about 5 liters of water. The process was then repeated once and the solid 
product was dried at 110.degree. C. overnight. The rare earth chloride 
solution contained a mixture of rare earth elements in the following 
proportions calculated as the oxides: 50 weight percent CeO.sub.2 ; 33 
weight percent La.sub.2 O.sub.3 ; 12 weight percent Nd.sub.2 O.sub.3 ; 4 
weight percent Pr.sub.6 O.sub.11 ; and 1 weight percent other rare earth 
elements, calculated as RE.sub.2 O.sub.3. After the rare earth exchange 
was completed, the LZ-10 zeolite contained about 5.1 weight percent rare 
earths calculated as RE.sub.2 O.sub.3. 
Catalyst 2 
Catalyst 2, a comparative catalyst, was prepared similarly to Catalyst 1 
except that, prior to mixing the LZ-10 zeolite with the other catalyst 
components, the zeolite was exchanged with aluminum cations instead of 
rare earth cations. The aluminum ion exchange was carried out by slurrying 
1,000 grams of the LZ-10 zeolite with a solution prepared by dissolving 
100 g of aluminum ammonium sulfate in 2,000 ml of water. The slurry was 
stirred at room temperature for one hour after which time it was filtered 
and washed and the procedure repeated. The solid product was then dried at 
110.degree. C. overnight. The finished catalyst contained the nickel and 
tungsten in the proportions above specified for Catalyst 1 on a support 
consisting of 10 weight percent aluminum-exchanged LZ-10 zeolite, 70 
weight percent dispersion of silica-alumina particles in a gamma alumina 
matrix, and 20 weight percent alumina binder. After the aluminum cation 
exchange, the LZ-10 zeolite contained about 3.1 weight percent more 
aluminum, calculated as Al.sub.2 O.sub.3, than it contained prior to the 
aluminum cation exchange. 
Catalyst 3 
Catalyst 3 was a sample of a commercial middle distillate hydrocracking 
catalyst obtained from UOP. It was prepared similarly to Catalysts 1 and 2 
except the LZ-10 zeolite was not rare earth- or aluminum-exchanged and 
contained no rare earth or aluminum cations at exchange sites. The 
catalyst contained about 5 weight percent nickel, calculated as NiO, and 
about 22 weight percent tungsten, calculated as WO.sub.3 on a support 
consisting of 10 weight percent LZ-10 zeolite, 70 weight percent 
dispersion of silica-alumina particles in a gamma alumina matrix, and 20 
weight percent alumina binder. 
Each of the above-described three catalysts was presulfided by passing a 
gas stream consisting of 10 volume percent hydrogen sulfide and the 
balance hydrogen through a bed of the catalyst at a temperature initially 
of about 300.degree. F. and slowly increased to 700.degree. F. and held at 
that temperature for about one hour. 
The three catalysts were then tested for activity and selectivity in middle 
distillate hydrocracking using a nonhydrotreated light Arabian vacuum gas 
oil having an API gravity of 23.degree., an initial boiling point of 
494.degree. F., a final boiling point of 1048.degree. F. and a 50 percent 
boiling point of 844.degree. F., with about 20 volume percent boiling 
below about 786.degree. F. and 5 volume percent boiling below 741.degree. 
F., as determined by a modified ASTM D1160 distillation. The gas oil, 
which contained about 0.085 weight percent nitrogen, calculated as the 
element, and 2.1 weight percent sulfur, calculated as the element, was 
passed on a once-through basis through an isothermal reactor containing 
about 140 ml of the catalyst mixed with 95 ml of six to eight mesh quartz. 
The reactor was operated at a liquid hourly space velocity (LHSV) of 1.0 
reciprocal hour, a total pressure of 2,000 psig and a once-through 
hydrogen flow rate of 10,000 standard cubic feet per barrel. The 
temperature of the reactor was adjusted to provide a 60 volume percent 
conversion to materials boiling below 700.degree. F. The results of these 
tests are set forth in Table I below: 
TABLE I 
__________________________________________________________________________ 
Activity (.degree.F.) 
Reactor Temp. 
Selectivity 
to Provide 
(Vol. % of 700.degree. F.-Product) 
Catalyst 
Composition of 
60% Conversion 
Trubine 
Diesel 
Designation 
Support (Wt. %) 
to 700.degree. F.- 
300-500.degree. F. 
300-700.degree. F. 
__________________________________________________________________________ 
1 70% silica-alumina 
748 85.1 89.8 
in alumina 
20% binder 
10% rare earth- 
exchanged 
LZ-10 zeolite 
2 70% silica-alumina 
742 76.7 84.1 
in alumina 
20% binder 
10% aluminum-exchanged 
LZ-10 zeolite 
3 70% silica-alumina 
754 79.5 85.6 
in alumina 
20% binder 
10% LZ-10 zeolite 
__________________________________________________________________________ 
As can be seen from the data in Table I, the catalyst of the invention, 
i.e., Catalyst 1, is 6.degree. F. (748.degree. F.-754.degree. F.) more 
active than the commercial catalyst, i.e., Catalyst 3. From kinetic 
considerations, it is known that an activity differential of about 
30.degree. F. of one catalyst over another roughly translates into a 
doubling in catalytic activity. Thus, a 6.degree. F. differential between 
Catalyst 1 and Catalyst 3 represents about a 20 percent improvement in 
activity for Catalyst 1 and means that this catalyst can be used for 
hydrocracking a given feedstock under the same operating conditions as the 
commercial catalyst but at a feed rate 20 percent higher. Alternatively, 
the catalyst of the invention could be used to produce the same conversion 
as the commercial catalyst at the same feed rate but initially at a 
temperature 6.degree. F. lower. 
In addition to possessing an increased activity over the commercial 
catalyst, Catalyst 1 is also 5.6 volume percent (85.1-79.5) more selective 
for turbine fuel boiling in the range between 300.degree. F. and 
550.degree. F. and 4.2 volume percent (89.8-85.6) more selective for 
diesel fuel boiling in the range of 300.degree. F. to 700.degree. F. Such 
increases in selectively are extremely significant in view of the fact a 
4.0 percent increase is considered commercially to be very significant 
because substantially more of the feed is converted to the desired product 
and less to lower valued products. For example, in a once-through 
hydrocracking process which yields 14,000 barrels per day of diesel fuel, 
a 4.2 percent increase in selectivity associated with Catalyst 1 will 
yield 588 more barrels per day and approximately 185,000 more barrels per 
year of the desired product. Assuming diesel fuel is priced at $40 per 
barrel, this increase in selectively is worth almost 7.5 million dollars 
per year to the refiner--a significant amount of money. 
A comparison of Catalyst 1 with Catalyst 2, which differs from Catalyst 1 
in that the LZ-10 zeolite is exchanged with aluminum cations instead of 
rare earth cations, indicates that Catalyst 2 is 6.degree. F. more active 
(742.degree. F.-748.degree. F.) then Catalyst 1. However, Catalyst 1 is 
surprisingly and unexpectedly much more selective. According to the data, 
Catalyst 1 yields 8.4 volume percent (85.1-76.7) more turbine fuel boiling 
in the range between 300.degree. F. and 550.degree. F. and 5.7 volume 
percent (89.8-84.1) more diesel fuel boiling in the range between 
300.degree. F. and 700.degree. F. These selectivity differences are much 
greater than those observed between Catalyst 1 and the commercial catalyst 
and clearly indicate that the selectivity differences between Catalysts 1 
and 2 are commercially very significant. If an analysis similar to that 
set forth above for the selectivity differences between Catalyst 1 and the 
commercial catalyst, i.e., Catalyst 3, is followed, the use of Catalyst 1 
in lieu of Catalyst 2 results in an income of about 10 million dollars per 
year more to the refiner. The fact that the presence of the multivalent 
rare earth cations in the LZ-10 zeolite of Catalyst 1 would yield such 
significant and surprising increases in selectivity of both turbine and 
diesel fuel over a similar catalyst in which the LZ-10 zeolite is 
exchanged with multivalent aluminum cations is unexpected. 
Although the invention has been primarily described in conjunction with an 
example and by references to embodiments thereof, it is evident that many 
alternatives, modifications, and variations will be apparent to those 
skilled in the art in light of the foregoing description. Accordingly, it 
is intended to embrace within the invention all such alternatives, 
modifications, and variations that fall within the spirit and scope of the 
appended claims.