Production of clean distillate fuels from heavy cycle oils

This invention discloses an enhanced process for the hydroprocessing of a feed, the feed comprising a highly aromatic refinery distillate stream boiling in the range between 300.degree. and 900.degree. F. The feed is separated into light and heavy streams such that the light stream contains from 0.1 to 5 wt. % dibenzothiophene, substituted dibenzothiophenes, and heavier polycyclic thiophenes. The lighter stream is hydrotreated at pressures from 300.degree. to 1000.degree. F. with a commercial catalyst having a hydrogenation component. The heavier stream is treated in the presence of hydrogen at higher pressure, from 600 to 2000 psig with a catalyst comprising active material having a Constraint Index of less than 2 in addition to a hydrogenation component in order to achieve over 35% conversion of material boiling above 630.degree. F. The active material of the catalyst is a highly siliceous zeolite or an acidic amorphous silica-alumina material.

FIELD OF THE INVENTION 
This invention relates to the hydroprocessing of highly aromatic refinery 
distillate streams, for manufacturing clean jet and diesel fuels as well 
as gasolines. More particularly, this invention relates to a process 
comprising segregation of distinct portions of such streams and a 
plurality of hydroprocessing zones operating at distinct operating 
conditions. 
BACKGROUND OF THE INVENTION 
In order to remain competitive, refiners have continuously sought to 
improve the quality of middle distillate products while simultaneously 
reducing processing costs. Refiners have recently sought to maximize 
existing equipment to achieve desired upgrades rather than build new 
equipment, in order to control costs. Such maximization is a continual 
challenge to refiners, since refining stocks have become heavier and 
poorer in quality. Upgrading capacity has been further strained by more 
stringent mandates on emissions. 
FCC cycle oil is a feed commonly used for the production of middle 
distillates and automotive diesel fuel. FCC cycle oil is a broad cut 
boiling between about 300.degree. F. and 900.degree. F. In addition to 
paraffins and cyclo paraffins, it contains both two and three ring 
aromatic structures and thiophenes. The thiophenes are generally multiple 
ring structures, such as benzothiophene, dibenzothiophene, substituted 
benzothiophenes and substituted dibenzothiophenes. 
Combined processing of both heavy and light portions of FCC cycle oil 
negatively affects hydroprocessing operations, such as catalytic 
hydrodesulfurization (CHD) and hydrocracking (HDC), including pressure 
requirements, flow rates, temperatures, product quality, and product 
yields. Operating conditions are generally dictated by the larger 
structures and are excessively severe for the lighter portion, which 
contains the smaller molecular structures. Catalysts tend to age 
relatively quickly when employed under excessively severe operating 
conditions, also. 
Zeolites have not been employed frequently as the support in commercial 
catalysts for mild hydroprocessing for heteratom removal and bond 
saturation (such as CHD), either on their own or combined with an 
amorphous matrix such as alumina because they tend to have a greater 
activity than alumina or other commonly used supports. With activity 
increase there is a concommitant increase in boiling range conversion and 
reduction in distillate yield. In addition, acidic zeolites are subject to 
coke formation and rapid aging under mild hydroprocessing conditions with 
feeds boiling above about 550.degree. F. 
Zeolites, especially zeolites X and Y, have long been used in more severe 
hydroprocessing operations such as hydrocracking, where their relatively 
greater activity is an asset. Under hydrocracking conditions they have 
excellent resistance to aging, particularly the more highly siliceous 
forms of zeolite Y, such as "ultra-stable" or USY. 
There are regulations throughout the world on the permissible quantity of 
sulfur in distillate products. The Environmental Protection Agency (EPA) 
and state environmental agencies, such as the California Air Resources 
Board (CARB) have established maximum standards of 0.05 wt % sulfur, for 
example. These standards went into effect in 1994. 
Various means have been proposed to upgrade feeds of high aromatic content. 
U.S. Pat. No. 4,789,457 discloses the recycling of full range cycle oils 
or cycle oil fractions to a catalytic cracking unit, where such feeds are 
subjected to low pressure hydrocracking in order to maximize the 
production of high octane gasoline. 
In order to avoid aging, conversion should be limited when operating with 
full range light cycle oil, and lower boiling fractions are preferred. 
U.S. Pat. No. 5,011,593 discloses the treating of full range cycle oils 
(boiling in the range of 385.degree.-750.degree. F.) or fractions thereof 
by catalytic hydrodesulfurization employing zeolite beta and a 
hydrogenation component. 
U.S. Pat. No. 3,957,625 discloses that sulfur impurities tend to 
concentrate in the heavier portion of a product fraction. It proposes a 
method of removing the sulfur from catalytically cracked gasoline by 
hydrodesulfurization of the heavy portion of the gasoline. The octane 
contribution of the olefins found in the lighter fraction is therefore 
retained. The light and heavy gasoline fractions are then recombined 
following separate treating. 
U.S. Pat. No. 4,990,242 is concerned with enhanced removal of sulfur from 
fuels. This patent discloses the fractionation of a feedstock and the 
separate removal of sulfur from the lighter fraction. On splitting the 
full range cycle oil at 570.degree.-575.degree. F. and separately 
hydrotreating the two fractions, the light portion attains a sulfur level 
below the 0.05 wt. % S standard. When combined with the separately 
hydrotreated heavy fraction, however, the standard is not met. 
SUMMARY OF THE INVENTION 
In conventional refinery operation, broad boiling streams with an aromatic 
content greater than 40%, such as FCC cycle oil, are hydroprocessed, 
usually in either a hydrodesulfurization (CHD) or hydrocracking (HDC) 
unit. In the instant invention the heavier portion of such a stream, which 
boils between about 600.degree. and about 900.degree. F. is separately 
processed over a catalyst or catalyst mixture comprising at least one 
highly siliceous zeolite or acidic amorphous silica-alumina having at 
least one hydrogenation component. The specific combination of zeolites 
and hydrogenation components is determined by the sulfur, nitrogen, 
aromatics and n-paraffin content of the feed and by the desired product 
slate. 
Separate processing of the heavy stream results in significant benefits in 
desulfurization effectiveness (thereby enabling governmental 
specifications to be met), in kerosene, diesel and gasoline product yield, 
in refinery operating cost and in some instances capital investment. 
DETAILED DESCRIPTION OF THE INVENTION 
Feedstock 
The feeds used in the present process are hydrocarbon fractions which are 
highly aromatic and hydrogen deficient. They are fractions which have an 
aromatic content in excess of at least 40 wt. percent and often 60 wt. 
percent or 80 wt. percent or more. Highly aromatic feeds of this type 
typically have hydrogen contents below 14 wt. percent, usually below 12.5 
wt. percent or even lower, e.g. 8-10 wt. percent or 8-9 wt. percent. The 
API gravity is often a measure of the aromaticity of the feed, usually 
being below 30 and in most cases below 25 or even lower, e.g. below 20. In 
most cases the API gravity will be in the range 5 to 25 e.g. 5-15, with 
corresponding hydrogen contents from 8.5-12.5 wt. percent. Sulfur contents 
are typically from 0.5-5 wt. percent and nitrogen from 50-3000 ppmw, more 
usually 100-1000 ppmw. 
The feeds of this type which are especially useful in the present process 
are the dealkylated cycle oil fractions produced by catalytic cracking 
operations, for example, in an FCC or TCC unit. A characteristic of 
catalytic cracking is that the alkyl groups, generally bulky, relatively 
large alkyl groups (typically but not exclusively C.sub.5 -C.sub.9 
alkyls), which are attached to aromatic moieties in the feed become 
removed during the course of the cracking. It is these detached alkyl 
groups which contribute to the gasoline fraction produced from the 
cracker. Aromatic moieties such as naphthalenes, benzothiophenes, 
dibenzothiophenes and polynuclear aromatics (PNAs) such as anthracene and 
phenanthrene are among the high boiling products from the cracker. The 
mechanisms of acid-catalyzed cracking and similar reactions remove side 
chains of greater than 5 carbons while leaving behind short chain alkyl 
groups, primarily methyl, but also ethyl groups on the aromatic moieties. 
Thus, the "substantially dealkylated" cracking products include those 
aromatics with small alkyl groups, such as methyl, and ethyl, and the like 
still remaining as side chains, but with relatively few large alkyl 
groups, i.e., the C.sub.5 -C.sub.9 groups, remaining. More than one of 
these short chain alkyl groups may be present, for example, one, two or 
more methyl groups. 
Cycle oil feeds include full range cycle oils which typically have a 
boiling range within the range of about 300.degree.-900.degree. F. and 
preferably in the range of about 350.degree.-800.degree. F. Fractionation 
of a full range cycle oil or adjustment of the cut points on the cracker 
fractionation column may be used to obtain two portions of cycle oil. The 
lower end temperature of the lighter fraction may be as low as 300.degree. 
F., preferably between about 320.degree. and 350.degree. F., and possibly 
as high as 400.degree. F., and the top end temperature of the lighter 
fraction may range from about 500.degree. F. to about 700.degree. F., 
preferably from about 550.degree. to 675.degree. F. and most preferably 
from 600.degree. to 650.degree. F. The heavier fraction will boil 
generally in the range above the top temperature of the lower fraction, 
but below about 900.degree. F. and preferably below between about 
750.degree. and 850.degree. F. It will be understood that some portion of 
the lighter fraction will carry over into the heavier fraction in any 
commercial distillation process. The precise temperature of the split will 
depend on the total sulfur content, the relative amounts of 
benzothiophenes, substituted benzothiophenes, dibenzothiophene, 
substituted dibenzothiophenes, and heavier polycyclic thiophenes present 
in the cycle oil, the desired product slate from the refinery, and the 
operating capabilities of the available hydroprocessing equipment. In 
general, the greater the sulfur content of a full range cycle oil and the 
higher the percentage of dibenzothiophenes, substituted dibenzothiophenes, 
and heavier polycyclic thiophenes, the lower the preferred temperature 
will be for the split. When a 0.3% S light product is desired, for 
example, the temperature will be chosen such that the light fraction 
contains no more than 5 wt. % and preferably less than 3 wt. % of 
dibenzothiophenes, substituted dibenzothiophenes and heavier polycyclic 
thiophenes. When a 0.05% S light product is desired, the temperature for 
the split will be chosen such that the light fraction contains less than 1 
wt. % and preferably less than 0.5 wt. % of dibenzothiophenes, substituted 
dibenzothiophenes and heavier polycyclic thiophenes. 
It will thus be understood that an optimum content of dibenzothiophene, 
substituted dibenzothiophenes, and heavier polycyclic thiophenes in the 
light fraction will exist, in relation to the desired light product sulfur 
level. For products currently envisioned, the optimum content will be 
between about 0.1 and 5 wt. %, preferably from 0.1 to 2 wt. %. 
Catalyst 
The catalysts used in the processing of the lighter portion of the cycle 
oil are of a conventional nature. Without being limited to any particular 
catalyst, typical catalysts are in the form of extrudates and include 
molybdenum on alumina, cobalt molybdate on alumina, nickel molybdate on 
alumina, nickel tungstate or combinations thereof. Catalyst choice may 
depend on the particular application. Cobalt molybdate catalyst is 
generally used when sulfur removal is the primary interest. The nickel 
catalysts find application in the treating of cracked stocks for olefin or 
aromatic saturation. The preparation of these catalysts is now well known 
in the art. 
The catalysts used for hydroprocessing the heavier portion of the cycle oil 
comprise highly siliceous zeolites or acidic amorphous silica-alumina 
materials as active components. They are bifunctional, heterogenous, 
porous solid catalysts which possess both acidic and hydrogenation 
functionality. Because the aromatic feeds contain relatively bulky 
bicyclic and tricyclic aromatic components the catalyst is required to 
have a pore size which is sufficiently large to admit these materials to 
the interior structure of the catalyst where the acid-catalyzed ring 
opening reactions can take place in order to effect removal of the 
heteroatoms under deep desulfurization conditions. Zeolite beta possesses 
a pore size of the requisite magnitude provided by the twelve-membered 
ring system. Zeolite beta is a known zeolite and is described in U.S. Pat. 
No. 3,308,069 (Wadlinger) to which reference is made for a description of 
this catalyst, its properties and preparation. Its use in catalytic 
dewaxing processes is described in U.S. Pat. No. 4,419,220 to which 
reference is also made for a further description of this catalyst and its 
use in dewaxing processes. 
Acidity in a potential zeolite or amorphous silica-alumina suitable for use 
in this invention can be conveniently measured by the alpha test. The 
alpha value is an approximate indication of the catalytic cracking 
activity of the catalyst compared to a standard catalyst. The alpha test 
gives the relative rate constant (rate of normal hexane conversion per 
volume of catalyst per unit time) of the test catalyst relative to the 
standard catalyst which is taken as an alpha of 1 (Rate Constant =0.016 
sec -1). The alpha test is described in U.S. Pat. No. 3,354,078 and in J. 
Catalysis, 4, 527 (1965); 6, 278 (1966); and 61, 395 (1980), to which 
reference is made for a description of the test. The experimental 
conditions of the test used to determine the alpha values referred to in 
this specification include a constant temperature of 538.degree. C. and a 
variable flow rate as described in detail in J. Catalysis, 6.1, 395 
(1980). In general, acidic materials useful in this invention will have an 
alpha of at least 1, preferably at least 5, and most preferably 10 or 
above. 
As indicated above, the preferred catalysts of this invention comprise 
either highly siliceous zeolites or an amorphous silica-alumina material 
having an acidic functionality. The latter materials are well-known in the 
hydroprocessing art. If the zeolite desired may be produced in the desired 
highly siliceous form by direct synthesis, this is often the most 
convenient method for obtaining it. Zeolite beta, for example, is known to 
be capable of being synthesized directly in forms having silica:alumina 
ratios up to 100:1, as described in U.S. Pat. Nos. 3,308,069 and Re 28,341 
which describe zeolite beta, its preparation and properties in detail. 
Even higher silica:alumina ratios are possible, as would be recognized by 
those skilled in the art. Zeolite Y, on the other hand, can be synthesized 
readily only in forms which have silica:alumina ratios up to about 5:1. In 
order to achieve higher ratios, various techniques may be employed to 
remove structural aluminum so as to obtain a more highly siliceous 
zeolite. The same is true of mordenite which, in its natural or directly 
synthesized form has a silica:alumina ratio of about 10:1. Zeolite ZSM-20 
may be directly synthesized with silica:alumina ratios of 7:1 or higher, 
typically in the range of 7:1 to 10:1, as described in U.S. Pat. Nos. 
3,972,983 and 4,021,331. Zeolite ZSM-20 also may be treated by various 
methods to increase its silica:alumina ratio. In general, any zeolite or 
amorphous silica-alumina material having an acidic functionality which 
exhibits a Constraint Index below 2.0 can be considered for this 
invention. The method by which Constraint Index is determined is described 
fully in U.S. Pat. No. 4,016,218, incorporated herein by reference for 
details of the method. Constraint Index (CI) values for typical zeolites 
which are suitable as catalysts in the process of this invention are as 
follows: 
______________________________________ 
CI (at test temperature) 
______________________________________ 
ZSM-4 0.5 (316.degree. C.) 
MCM-22 0.6-1.5 (399.degree. F.-454.degree. C.) 
TEA Mordenite 0.4 (316.degree. C.) 
REY 0.4 (316.degree. C.) 
Amorphous Silica-alumina 
0.6 (538.degree. C.) 
Dealuminized Y 0.5 (510.degree. C.) 
Zeolite Beta 0.6-2.0 (316.degree. C.-399.degree. C.) 
ZSM-20 0.5 (371.degree. C.) 
Mordenite 0.5 (316.degree. C.) 
______________________________________ 
Control of the silica:alumina ratio to the zeolite in its as-synthesized 
form may be exercised by an appropriate selection of the relative 
proportions of the starting materials, especially the silica and alumina 
precursors, a relatively smaller quantity of the alumina precursor 
resulting in a higher silica:alumina ratio in the product zeolite, up to 
the limit of the synthetic procedure. If higher ratios are desired and 
alternative synthesis affording the desired high silica:alumina ratios are 
not available, other techniques such as those described below may be used 
in order to prepare the desired highly siliceous zeolites. 
The silica:alumina ratios referred to in this specification are the 
structural or framework ratios. This is the ratio for the SiO.sub.4 to the 
AlO.sub.4 tetrahedra which together constitute the structure of which the 
zeolite is composed. This ratio may vary from the silica:alumina ratio 
determined by various physical and chemical methods. For example, a gross 
chemical analysis may include aluminum which is present in the form of 
cations associated with the acidic sites on the zeolite, thereby giving a 
low silica:alumina ratio. Similarly, if the ratio is determined by 
thermogravimetric analysis (TGA) of ammonia desorption, a low ammonia 
titration may be obtained if cationic aluminum prevents exchange of the 
ammonium ions onto the acidic sites. These disparities are particularly 
troublesome when certain treatments such as the dealuminization methods 
described below which result in the presence of ionic aluminum free of the 
zeolite structure are employed. Care should therefore be taken to ensure 
that the framework silica:alumina ratio is correctly determined. 
A number of different methods are known for increasing the structural 
silica:alumina ratio of various zeolites. Many of these methods rely upon 
the removal of aluminum from the structural framework of the zeolite by 
chemical agents appropriate to this end. A considerable amount of work on 
the preparation of aluminum deficient faujasites has been performed and is 
reviewed in Advances in Chemistry Ser. No. 121, Molecular Sieves, G. T. 
Kerr, American Chemical Society, 1973. Specific methods for preparing 
dealuminized zeolites are described in the following, and reference is 
made to them for details of the method: Catalysis by Zeolites 
(International Symposium on Zeolites, Lyon, Sept. 9-11, 1980), Elsevier 
Scientific Publishing Co., Amsterdam, 1980 (dealuminization of zeolite Y 
with silicon tetrachloride); U.S. Pat. No. 3,442,795 and G. B. Pat. No. 
1,058,188 (hydrolysis and removal of aluminum by chelation); G. B. Pat. 
No. 1,061,847 (acid extraction of aluminum); U.S. Pat. No. 3,493,519 
(aluminum removal by steaming and chelation): U.S. Pat. No. 3,591,488 
(aluminum removal by steaming); U.S. Pat. No. 4,273,753 (dealuminization 
by silicon halides and oxyhlides); U.S. Pat. No. 3,691,099 (aluminum 
extraction with acids); U.S. Pat. No. 4,093,560 (dealumination by 
treatment with salts); U.S. Pat. No. 3,937,791 (aluminum removal with 
Cr(III) solutions); U.S. Pat. No. 3,506,400 (steaming followed by 
chelation); U.S. Pat. No. 3,640,681 (extraction of aluminum with 
acetyl-acetonate followed by dehydroxylation): U.S. Pat. No. 3,836,561 
(removal of aluminum with acid); DE-OS Pat. No. 2,510,740 (treatment of 
zeolite with chlorine or chlorine-contrary gases at high temperatures), 
N.L. Pat. No. 7,604,264 (acid extraction), JA Pat. No. 53,101,003 
(treatment with EDTA or other materials to remove aluminum) and J. 
Catalysis 54 295 (1978) (hydrothermal treatment followed by acid 
extraction). 
Because of their convenience and practicality, the preferred 
dealuminization methods for preparing the present highly siliceous 
zeolites are those which rely upon acid extraction of the aluminum from 
the zeolite. Zeolite beta may be dealuminized by acid extraction using 
mineral acids such as hydrochloric acid. Highly siliceous forms of zeolite 
Y may be prepared by steaming or by acid extraction of structural aluminum 
(or both). Because zeolite Y in its normal, as-synthesized conditions, is 
unstable to acid, it must first be converted to an acid-stable form. 
Methods for doing this are known and one of the most common forms of 
acid-resistant zeolite Y is known as "Ultrastable Y" (USY). USY is 
described in U.S. Pat. Nos. 3,293,192 and 3,402,996 and the publication, 
Society of Chemical Engineering (London) Monograph Molecular Sieves, page 
186 (1968) by C. V. McDaniel and P. K. Maher. Reference is made to these 
for details of the zeolite and its preparation. In general, "ultrastable" 
refers to Y-type zeolite which is highly resistant to degradation of 
crystallinity by high temperature and steam treatment and is characterized 
by a R.sub.2 O content (wherein R is Na, K or any other alkali metal ion) 
of less than 4 weight percent, preferably less than 1 weight percent, and 
a unit cell size less than 24.5 Angstroms and a framework silica:alumina 
ratio above about 5, e.g., ratios of 15, 50 or 200 or more. The 
ultrastable form of Y-type zeolite is obtained primarily by a substantial 
reduction of the alkali metal ions and the unit cell size reduction of the 
alkali metal ions and the unit cell size reduction. The ultrastable 
zeolite is identified both by the smaller unit cell and the low alkali 
metal content in the crystal structure. 
The ultrastable form of the Y-type zeolite can be prepared by successively 
base exchanging a Y-type zeolite with an aqueous solution of an ammonium 
salt, such as ammonium nitrate, until the alkali metal content of the 
Y-type zeolite is reduced to less than 4 weight percent. The base 
exchanged zeolite is then calcined at a temperature of 540.degree. C. to 
800.degree. C. for up to several hours, cooled and successively base 
exchanged with an aqueous solution of an ammonium salt until the alkali 
metal content is reduced to less than 1 weight percent, followed by 
washing and calcination again at a temperature of 540.degree. C. to 
800.degree. C. to produce an ultrastable zeolite Y. The sequence of ion 
exchange and heat treatment results in the substantial reduction of the 
alkali metal content of the original zeolite and results in a unit cell 
shrinkage which is believed to lead to the ultra high stability of the 
resulting Y-type zeolite. 
The ultrastable zeolite Y may then be extracted with acid to produce a 
highly siliceous form of the zeolite. 
Other methods for increasing the silica:alumina ratio of zeolite Y by acid 
extraction are described in U.S. Pat. Nos. 4,218,307, 3,591,488 and 
3,691,099, to which reference is made for details of these methods. 
In addition to the highly siliceous zeolite or acidic amorphous 
silica-alumina having a Constraint Index below 2, the catalyst or 
catalysts used in this mixture may also contain a binder. The binder is 
typically an amorphous inorganic oxide material such as alumina, 
silica-alumina or silica and this binder may comprise from about 20 to 80 
percent, and preferably 40 to 60 percent of the catalyst (excluding metal 
hydrogenation component). Because the zeolite or acidic amorphous 
silica-alumina provides the desired acidic functionality to the catalyst, 
the matrix, if present, may be essentially non-acidic. Non-selective 
active material conversion during the process is thus maintained at a 
desirably low level. A further description of suitable matrix materials 
and of compositing methods may be found in U.S. Pat. No. 4,789,457 
(Fischer) to which reference is made for such a description. 
The catalysts of this invention which comprise an active material also have 
a metal component to provide the necessary hydrogenation functionality. 
Suitable hydrogenation components include the metals of Groups VIA and 
VIIIA of the Periodic Table (IU Table) specifically tungsten, vanadium, 
zinc, molybdenum, rhenium, nickel, cobalt, chromium or manganese. The 
hydrogenation component is generally present in an amount between 0.1 and 
about 25 wt %, normally 0.1 to 5 wt %, especially for noble metals, and 
preferably 0.3 to 3 wt %. This component can be exchanged or impregnated 
into the composition, using a suitable compound of the metal. The 
compounds used for incorporating the metal component into the catalyst can 
usually be divided into compounds in which the metal is present in the 
cation of the compound and compounds in which it is present in the anion 
of the compound. Compounds which contain the metal as a neutral complex 
may also be employed. The compounds which contain the metal in the ionic 
state are generally used, although cationic forms of the metal have the 
advantage that they will exchange onto the active material. Anionic 
complex ions such as vanadate or metatungstate which are commonly employed 
can however be impregnated onto the zeolite/-binder composite without 
difficulty in the conventional manner since the binder is able to absorb 
the anions physically on its porous structure. Higher proportions of 
binder will enable higher amounts of these complex ions to be impregnated. 
Base metal components, especially cobalt either alone or with molybdenum, 
or nickel either alone or mixed with tungsten or molybdenum are 
particularly preferred in the present process. 
As indicated previously, hydroprocessing catalysts of the instant invention 
comprise preferably large pore, highly siliceous zeolites such as zeolite 
beta and USY. Base metal components, especially cobalt either alone or 
with molybdenum, or nickel either alone or mixed with tungsten or 
molybdenum are particularly preferred in the present process. They may be 
used however in conjunction with amorphous catalysts such Co/Mo on 
alumina. 
PROCESS CONDITIONS 
In this invention, mild and conventional hydrotreating conditions, suitable 
for the removal of heteroatoms such as S, N and O, are appropriate for 
processing the lighter portion of the cycle oil. Thus total pressures will 
normally be in the range of 300-1000 psig, preferably 400-600 psig, 
hydrogen circulation rates will be 500-6000 SCF/B, preferably 1000-2000 
SCF/B, temperatures will be 400.degree.-800.degree. F., preferably 
500.degree.-700.degree. F., and WHSV will be from about 0.5 to 6 hr.sup.-1 
preferably from 1 to 4. 
Conditions in the hydroprocessing of the heavier portion of the cycle oil 
(involving the use of the catalyst comprising the active material) are 
more severe. Total pressure will be between about 600 and 2000 psig, 
preferably 900 to 1500 psig, hydrogen circulation rates will be 1000-8000 
SCF/B, preferably 2000-5000, temperatures will be 500.degree.-800.degree. 
F., preferably 600-750, and WHSV will be from about 0 5 to 6 hr.sup.-1, 
preferably from 1 to 4. 
Precise operating conditions will be selected on the basis of desired 
product slate, sulfur and aromatics specifications, if any, and available 
refinery hydroprocessing equipment. Products containing less than 0.05 wt. 
% S can be made under the conditions of this invention, if desired.

EXAMPLES 
Example 1 
This example demonstrates the disadvantage of hydroprocessing full range 
cycle oil and the particular care which must be taken in choosing the 
cutpoint temperature in segregating a full range cycle oil into light and 
heavy portions. 
The feed is a light portion of cycle oil which boils between about 
300.degree. to 650.degree. F. and contains 2.1% S and 74% aromatics. The 
cutpoint for its separation from full range cycle oil is 575.degree. F. It 
contains 0.6% dibenzothiophene, substituted dibenzothiophenes, and heavier 
polycyclic thiophenes. 
When this feed is processed over a commercial NiMo/alumina catalyst at 
568.degree. F., 2.6 WHSV, 900 psig, and 6000 SCF/B hydrogen, a liquid 
product is obtained which contains 0.10% S and does not meet a 0.05% S 
automotive diesel fuel specification. 
When the experiment is repeated with a feed which boils between 300.degree. 
and 630.degree. F. and contains approximately 0.1% dibenzothiophene and 
substituted dibenzothiophenes, the liquid product meets the 0.05% S 
specification. The cutpoint for separating this light portion of cycle oil 
is 550.degree. F. 
Cutpoint in the separation of light and heavy portions of the cycle oil may 
thus be tailored, based on the dibenzothiophene and substituted 
dibenzothiophene content of the light portion, to meet a desired light 
product sulfur specification and to accommodate available hydroprocessing 
equipment. It will also be appreciated by those skilled in the art, that 
an optimum cutpoint will exist. 
Example 2 
This example shows that only a minor amount of the heavy portion of a cycle 
oil can be converted to distillate boiling below 600.degree. F. and 
meeting a 0.05% S specification by a catalyst which contains only a 
hydrogenation component. 
The feed is a heavy portion of a cycle oil and boils between about 
560.degree. and 800.degree. F. It contains 2.8% S, virtually all of it in 
the form of dibenzothiophene, substituted dibenzothiophenes, and heavier 
polycyclic thiophenes. It contains 74% aromatics. 
When this feed is processed over the same commercial NiMo/alumina catalyst 
as in Example 1, but at 900 psig, 6000 SCF/B hydrogen, 1 to 3 WHSV, and at 
temperatures ranging from 550.degree. to 800.degree. F., a limit of about 
35% is reached in the conversion of material boiling above about 
630.degree. F. to lighter hydrocarbons. Selectivity to 
420.degree.-630.degree. F. distillate is 86%. 
Example 3 
This example shows that the limitation on conversion of material boiling 
above 630.degree. F. is lifted and that selectivity to 
420.degree.-630.degree. F. distillate in some cases increases when highly 
siliceous zeolite is included in the catalyst as the active material. 
The feed is the heavy portion of a cycle oil and boils between 590.degree. 
and 815.degree. F. It contains 3.7% S, 1500 ppm N, and 78% aromatics. 
When this feed is processed under the conditions of Example 2 at 1.1 WHSV 
over a commercial NiMo/USY catalyst (the catalyst containing approximately 
40% of USY zeolite and the USY having a unit cell parameter of 24.3.ANG.), 
conversion of material boiling above 630.degree. F. increases steadily 
with increasing temperature, namely, from 39% at 701.degree. F. to 45% at 
722.degree. F. and to 60% at 753.degree. F. Selectivity for 
420.degree.-620.degree. F. distillate is 91% at 701.degree. F., 83% at 
722.degree. F., and 64% at 753.degree. F. The enhanced selectivity is 
attributed to a zeolite-induced partial breakdown of high boiling 
three-ring aromatics, to yield distillate range two-ring structures. Much 
of the material boiling below 420.degree. F. is low-sulfur gasoline. It 
will be understood that material boiling above 630.degree. F. may be 
recycled to extinction if so desired. 
Example 4 
This example supports the assertion in Example 3 that high boiling 
three-ring structures are yielding 420.degree.-630.degree. F. distillate 
in the presence of zeolite. 
The experiment of Example 3 is repeated at 753.degree. F., but with a feed 
containing 7.5% phenanthrene, a three-ring aromatic which boils at 
640.degree. F. Under these conditions, conversion of material boiling 
above 630.degree. F. is 61%, vs 60% for feed at the same conditions 
without the phenanthrene present (see Example 3). Selectivity to 
420.degree.-630.degree. F. distillate is 72%, vs 64% in Example 3. 
Example 5 
This example illustrates even higher conversion than in Example 3 when a 
heavy portion is hydrotreated to reduce S and N levels before contact with 
zeolite. It will be appreciated by those skilled in the art of 
distillation that commercial hydrotreating of the heavy portion will 
require conditions not unlike those preferred for hydroprocessing of this 
material. 
The feed is the same as in Example 3 except that it has been hydrotreated 
with a commercial hydrotreating catalyst which contains 3% Ni and 13% Mo 
to achieve a sulfur content of 0.32% and a nitrogen content of 760 ppm. 
When this feed is processed at 755.degree. F. under the same conditions as 
in Example 3, 71% of the material boiling above 630.degree. F. is 
converted and the selectivity to 420.degree.-630.degree. F. distillate is 
68%. The product liquid contains 0.002% S. Here, too, it is recognized 
that the material boiling above 630.degree. F. may be recycled to 
extinction over the hydrotreating and hydroprocessing sequence of 
catalysts. 
Example 6 
This example demonstrates that USY may be replaced by another large pore 
siliceous zeolite. The catalyst composition is similar to that of Example 
3, but zeolite Beta is used in place of USY. The catalyst contains 
approximately 40 wt. % zeolite. 
When the heavy feed used in Example 3 is processed under the same 
conditions over this NiMo/Beta catalyst at 700.degree. F., 37% of the 
material boiling above 630.degree. F. is converted to lighter 
hydrocarbons, and selectivity to 420.degree.-630.degree. F. distillate is 
85%. At 750.degree. F., conversion is 51% and selectivity is 76%.