Production of high octane gasoline

A moderate pressure hydrocracking process in which a highly aromatic, substantially dealkylated feedstock is processed directly to high octane gasoline by hydrocracking over a monofunctional acidic cracking catalyst, preferably comprising a large pore size, crystalline alumino-silicate zeolite hydrocracking catalyst such as zeolite Y. The feedstock which is preferably a light cycle oil obtained from catalytic cracking with an aromatic content of at least 50, usually at least 60 percent and an API gravity not more than 25. The hydrocracking typically operates at 600-1000 psig at moderate to high conversion levels to mazimize the production of monocyclic aromatics which provide the requiste octane value to the product gasoline.

FIELD OF THE INVENTION 
This invention relates to the production of high octane gasoline and more 
particularly to the production of high octane gasoline by hydrocracking 
highly aromatic fractions obtained from catalytic cracking and other 
petroleum refinery operations. 
BACKGROUND OF THE INVENTION 
Under present conditions, petroleum refineries are finding it necessary to 
convert increasingly greater proportions of crude to premium fuels such as 
gasoline and middle distillates such as diesel and jet fuel. Catalytic 
cracking processes, exemplified by the fluid catalytic cracking (FCC) 
process and Thermofor catalytic cracking (TCC) process together, account 
for a substantial fraction of heavy liquids convertion in modern 
refineries. Both are thermally severe processes which result in a 
rejection of carbon to coke and to residual fractions; during catalytic 
cracking high molecular weight liquids disproportionate into relatively 
hydrogen-rich light liquids and aromatic, hydrogen-deficient heavier 
distillates and residues. 
Catalytic cracking in the absence of hydrogen does not provide significant 
desulfurization nor is the nitrogen content of the feed selectively 
rejected with the coke. Both sulfur and nitrogen therefore concentrate 
appeciably in the heavier cracking products. Cracking therefore produces 
significant quantities of highly aromatic, hydrogen-deficient middle and 
heavy distillates that have high sulfur and nitrogen levels. Recycling 
these liquids to the catalytic cracker is often not an attractive option, 
because they are refractory and difficult to convert and often will impair 
conversion of the less refractory fresh feed. Generally, the level of 
heteroatom contaminants increases with the boiling point of the fraction, 
as shown in Table 1 below which gives the sulfur and nitrogen contents for 
two typical FCC product fractions, a light cycle oil and an FCC main 
column bottoms (proportions and percentages by weight, as in the remainder 
of this specification unless the contrary is stated). 
TABLE 1 
______________________________________ 
FCC Product Fractions 
Aromatics; 
S, wt. H, wt. 
wt. pct. 
pct. N, ppmw pct. 
______________________________________ 
Light Cycle Oil 
80 3.1 650 9.1 
Main Column Bottoms 
80+ 4.6 1500 6.8 
______________________________________ 
Present market requirements make refractory product streams such as these 
particularly difficult to dispose of as commercially valuable products. 
Formerly, the light and heavy cycle oil and FCC main column bottoms could 
be upgraded and sold as light or heavy fuel oil, such as No. 2 fuel oil or 
No. 6 fuel oil. Upgrading the light cycle oil was conventionally carried 
out by a relatively low severity, low pressure catalytic 
hydro-desulfurization (CHD) unit in which the cycle stock would be admixed 
with virgin mid-distillates from the same crude blend fed to the catalytic 
cracker. Further discussion of this technology is provided in the Oil and 
Gas Journal, May 31, 1982, pp. 87-94. 
Currently, however, the refiner is finding a diminished demand for fuel 
oil. At the same time, the impact of changes in supply and demand for 
petroleum has resulted in a lowering of the quality of the crudes 
available to the refiner; this has resulted in the formation of an even 
greater quantity of refractory cycle stocks. As a result, the refiner is 
left in the position of producing increased amounts of poor quality cycle 
streams from the catalytic cracker while having a diminishing market in 
which to dispose of these streams. 
At many petroleum refineries, the light cycle oil (LCO) from the FCC unit 
is a significant component of the feed to the catalytic 
hydrodesulfurization (CHD) unit which produces No. 2 fuel oil or diesel 
fuel. The remaining component is generally virgin kerosene taken directly 
from the crude distillation unit. The highly aromatic nature of LCO, 
particularly when the FCC unit is operated in the maximum gasoline mode, 
increases operational difficulties for the CHD and can result in a product 
having marginal properties for No. 2 fuel oil or diesel oil, as measured 
by cetane numbers and sulfur content. 
An alternative market for mid-distillate streams is automotive diesel fuel. 
However, diesel fuel has to meet a minimum cetaine number specification of 
about 45 in order to operate properly in typical automotive diesel 
engines. Because cetane number correlates closely and inversely with 
aromatic content, the highly aromatic cycle oils from the cracker 
typically with aromatic contents of 80% or even higher have cetane numbers 
as low as 4 or 5. In order to raise the cetane number of these cycle 
stocks to a satisfactory level by the conventional CHD technology 
described above, substantial and uneconomic quantities of hydrogen and 
high pressure processing would be required. 
Because of these problems associated with its use as a fuel, recycle of 
untreated light cycle oil to the FCCU has been proposed as a method for 
reducing the amount of LCO. Benefits expected from the recycle of LCO 
include conversion of LCO to gasoline, backout of kerosene from No. 2 fuel 
oil and diminished use of cetane improvers in diesel fuel. However, in 
most cases, these advantages are outweighed by disadvantages, which 
include increased coke make in the FCC unit, diminished quality of the 
resultant LCO and an increase in heavy cycle oil and gas. 
A typical LCO is such a refractory stock and of poor quality relative to a 
fresh FCC feed that most refineries do not practice recycle of the 
untreated LCO to any significant extent. One commonly practiced 
alternative method for upgrading the LCO is to hydrotreat severely prior 
to recycle to the catalytic cracker or, alternatively, to hydrotreat 
severely and feed to a high pressure fuels hydrocracker. In both such 
cases, the object of hydrotreating is to reduce the heteroatom content to 
low levels while saturating polyaromatics to increase crackability. 
Although this does enhance the convertibility of these aromatic streams 
considerably, the economic penalties derived from high hydrogen 
consumptions and high pressure processing are severe. In addition, in 
those instances where the production of gasoline is desired, the naphtha 
may require reforming to recover its aromatic character and meet octane 
specifications. 
Hydrocracking may be used to upgrade the higher-boilinhg more refractory 
products derived from catalytic cracking. The catalytic cracker is used to 
convert the more easily cracked paraffinic gas oils from the distillation 
unit while the hydrocracking accepts the dealkylated, aromatic cycle oils 
from the cracker and hydrogenates and converts them to lighter oils. See 
Petroleum Refining; Second Ed.; Gary, J. H and Handwerk, G. E.; Marcel 
Dekker, N.Y. 1984; pp. 138-151; Modern Petroleum Technology, Fourth Ed.; 
Hobson, G. D., Applied Science Publ. 1973; pp. 309-327. These 
hydrocracking processes using catalytically cracked feeds either on their 
own or mixed with virgin feeds have, however, generally been incapable of 
producing high octane gasoline directly. The reason for this is that they 
have conventionally been operated at high hydrogen pressures and at 
relatively high conversion levels so as to maximize the saturation of the 
aromatics (especially the refractory polynuclear aromatics), removal of 
heteroatoms in inorganic form and the subsequent conversion of the 
hydrogenated aromatics to paraffins. While this may produce acceptable 
diesel fuel (which benefits from the presence of n-paraffins) the octane 
quality of the gasoline has generally been poor as a consequence of the 
large quantities of low octane paraffin components. For present day use 
these gasolines will require extensive reforming with its consequent yield 
loss in order to conform to market produce specifications. To illustrate, 
U.S. Pat. No. 3,132,090 discloses the use of a two-stage hydrocracking 
scheme to produce gasoline. However, the octane number of the gasoline 
using a virgin distillate as charge is reported as 68 (RON+0). An octane 
of 80 (RON+3) is disclosed for a charge-stock of coker distillate and 
thermally cracked gas oils. The "high octane" gasolines described in this 
patent contain 3 ml/gallon of tetraethyl lead (TEL) and are in the range 
of 70-88 (RON+3). Because TEL adds about 4-6 octane numbers these 
gasolines have an octane rating on a clear basis (RON+0) in the range of 
65-83 (RON+0). 
Various low pressure hydrocracking processes have also been described. For 
example, U.S. Pat. Nos. 3,867,277 and 3,923,640 disclose low pressure 
hydrocracking processes using various high boiling feedstocks, generally 
of high (20-40) API gravity. The use of such feeds, coupled with the 
relatively high levels of conversion in those processes lead to naphthas 
of low octane rating since the alkyl groups present in the feeds come 
through into the naphtha together with the relatively straight chain 
paraffins produced by the ring opening and cracking of the aromatics. 
These processes have therefore been unsatisfactory for the direct 
production of high octane gasoline. 
Other low pressure hydrocracking processes producing aromatic products have 
been described in the past but their potential for producing high octane 
gasoline from low value, refractory cracking oils has not been 
appreciated. For example, U.S. Pat. No. 4,435,275 describes a method for 
producing aromatic middle distillates such as some heating oil from high 
gravity feeds under relatively low conversion conditions but with the 
objective of producing low-sulfur middle distillates, octane numbers of 
only about 78 (R+0P) are reported. 
A notable advance is described in U.S. Pat. No. 4,676,887 to which 
reference is made for details. It was found that highly aromatic, 
refractory feeds derived from catalytic cracking and other refinery 
operations could be converted directly to high octane gasoline by 
hydrocracking at relatively low pressures, typically 600-1000 psig (about 
4250-7000 kPa. abs.) and with low conversions, typically below 50 weight 
percent to 385.degree. F.-(195.degree. C.-) products. (All SI equivalents 
in this specification are rounded off to a convenient figure so as to 
permit convenient comparison; all pressures quoted in SI units are 
absolute pressures). By using a highly aromatic feed which has been 
substantially dealkylated, for example in a catalytic cracking operation, 
typically with an API gravity of 5-25, the hydrocracking proceeds with 
only a limited degree of aromatics saturation so that a large quantity of 
single-ring alkylaromatics (mainly benzene, toluene, xylenes and trimethyl 
benzenes) are obtained by ring opening of partial hydrogenation products 
of bicyclic aromatics. The single ring aromatics are not only in the 
gasoline boiling range but also possess high octane numbers so that a high 
octane gasoline is produced directly, suitable for blending into the 
refinery gasoline pool without prior reforming. 
A further development of this technique is described in U.S. application 
Ser. No. 940,382, in which high octane gasoline is produced from a 
relatively lower boiling fraction of the aromatic feed from the catalytic 
cracking operation. By the use of these light cut feeds, conversion may be 
raised to higher levels without adversely affecting gasoling octane or 
catalyst aging rate. 
In the process, the dealkylated feed is treated in two reactors either with 
or without interstage separation of inorganic nitrogen and sulfur and 
light ends. The first reactor in the sequence is a hydrotreating reactor 
containing a conventional hydrotreating catalyst comprising a metal with 
hydrogenation functionality on a non-acidic porous support, for example, 
nickel-molybdenum on alumina. The second reactor contains a hydrocracking 
catalyst with both acidic, cracking functionality and hydrogenation 
functionality. The hydrogenation functionality is provided by a metal 
usually of Groups VIA or VIIIA of the Periodic Table (IU Table) and the 
acidic functionality by a porous, inorganic, acidic solid, preferably a 
large pore size zeolite such as zeolite X or zeolite Y with a porous 
binder such as alumina, silica or silica-alumina. 
SUMMARY OF THE INVENTION 
We have now found that the hydrocracking may be carried out using a 
catalyst which is essentially free of the metal hydrogenation component. 
This is unexpected because hydrocracking is generally regarded as a 
process requiring a bifunctional catalyst: the hydrogenation (metal) 
functionality promotes the hydrogenation reactions and the acidic function 
promotes the cracking and ring-opening reactions. 
The hydrocracking process operates with a substantially dealkylated feed as 
described in U.S. Pat. No. 4,676,887 to produce a high octane gasoline 
product using relatively low hydrogen pressure and hydrogen consumptions. 
The process runs with acceptable cycle durations and may be operated in 
continuous flow process with cyclic catalyst regeneration. 
According to the present invention the process for producing a high octane 
gasoline boiling range product from a substantially dealkylated 
hydrocarbon feedstock employs a hydrocracking catalyst which is 
substantially free of hydrogenation functionality conferred by an added 
metal component. The hydrocracking catalyst therefore consists essentially 
of a porous, inorganic solid having acidic functionality, preferably a 
large pore size zeolite in a porous inorganic oxide matrix such as 
alumina, silica or silica-alumina. The hydrocracking step is preceded by a 
hydrotreating and an interstage separation of light ends and inorganic 
nitrogen and sulfur (ammonia and hydrogen sulfide) may optionally be 
interposed.

DETAILED DESCRIPTION 
Feedstock 
The feeds used in the present process are hydrocarbon fractions which are 
highly aromatic and hydrogen deficient. They are fractions which have been 
substantially dealkylated, as by a catalytic cracking operation, for 
example, in an FCC or TCC unit. Dealkylated feeds may also be obtained 
from other refinery units such as delayed, contact or fluid bed coking 
units. However, catalytic cracking units will generally provide the bulk 
of the feed and the invention will be described with reference to feeds 
from such units. It is a characteristic of catalytic cracking 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 
cracking. It is these detached alkyl groups which lead to the bulk of the 
gasoline product from the cracker. The aromatic moieties such as benzene, 
naphthalene, benzothiophenes, dibenzothiophenenes and polynucler aromatics 
(PNAs) such as anthracene and phenanthrene from 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. 
Feedstocks of this type have an aromatic content in excess of 50 wt. 
percent; for example, 70 wt. percent or 80 wt. percent or more, aromatics. 
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. below 
10 wt. percent or 9 wt. percent. The API gravity is also 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 with corresponding hydrogen contents from 8.5-12.5 wt. 
percent. Sulfur contents are typically from 0.5-5 wt. percent and nitrogen 
from 50-1000 ppmw. 
The feeds for the present process may therefore be full range cycle oils 
with end points up to about 750.degree. F. (400.degree. C.) or a light out 
cycle oil with an end point up to about 650.degree. C. (345.degree. C.). 
The use of light cut cycle oils with end points below 650.degree. F. 
(345.degree. C.) preferably below 600.degree. (315.degree. C.) is 
desirable because their high conversion levels may be tolerated during the 
hydrocracking without excessive catalyst aging. If full range cycle oils 
are used with end points above about 650.degree. F. (345.degree. C.), it 
is desirable to limit conversion to the levels specified in U.S. Pat. No. 
4,676,887, i.e. to below 50% per pass to gasoline boiling range products 
(385.degree. F.-, 196.degree. C.-) and preferably to a value related to 
the hydrogen pressure employed (not more than 0.05 times the hydrogen 
pressure expressed in psig). Reference is made to U.S. Pat. No. 4,676,887 
for a description of the applicable process conditions. 
Suitable feeds for the present process therefore include substantially 
dealkylated cracking product fractions with an end point below 650.degree. 
F. (345.degree. C.), preferably below 600.degree. F. (315.degree. C.). 
Initial boiling point will usually be 300.degree. F. (150.degree. C.) or 
higher, e.g. 330.degree. F. (165.degree.) or 385.degree. F. (195.degree. 
C.). Light cut light cycle oils (LCOs) within these boiling ranges ae 
highly suitable. 
Full range light cycle oils (FRLCO) generally have a boiling point range 
between 385.degree. and 750.degree. F. (195.degree.-400.degree. C.) and 
may also be used. Light cycle oils generally contain from about 60 to 80% 
aromatics and, as a result of the catalytic cracking process, are 
substantially dealkylated. Other examples of suitable feedstocks include 
the dealkylated liquid products from delayed or fluid bed coking 
processes. 
The use of the dealkylated feeds is a significant feature of the process. 
It will not produce high octane gasoline from predominantly virgin or 
straight run oils and which have not been previously dealkylated by 
processes such as catalytic cracking or coking. If the feed used in the 
present process has not been previously dealkylated, the large alkyl 
groups found in the feed will be cracked off during the hydrocracking and 
will be found in the resulting naphtha fraction. Because these groups are 
relatively straight chain, a low octane gasoline product will result. 
Smaller, i.e., C.sub.1 -C.sub.3, alkyl side groups, if present do not 
appear in the naphtha boiling range products from the hydrocracker (even 
if conditions are severe enough to remove them) and so they have no effect 
on product octane. If a mixture of dealkylated and non-dealkylated 
feedstock is used, the octane number will be intermediate between the 
octane numbers of the feeds used separately. A mixture of alkylated and 
dealkylated feedstocks can be used in commercial operation but if so, it 
is likely that the gasoline will have to be subjected to a reforming 
process in order to achieve the desired octane. 
Hydrocracking Catalyst 
The catalyst used for the hydrocracking step is a monofunctional, 
heterogeneous, porous solid catalyst possessing acidic (cracking) and 
functionality but, at least in its initial stage prior to use, is 
essentially free of hydrogenation functionlity. Because the highly 
aromatic feed contains relatively bulky bicyclic and polycyclic components 
the catalyst should have a pore size which is sufficiently large to admit 
these materials to the interior structure of the catalyst where cracking 
can take place. A pore size of at least about 7.4 A (corresponding to the 
pore size of the large pore size zeolites X and Y) is sufficient for this 
purpose but because the end point of the feed is limited, the proportion 
of bulky, polynuclear aromatics is quite low and for this reason, very 
large pore sizes greatly exceeding those previously mentioned are not 
required. Crystalline zeolite catalysts which have relatively limited pore 
size range, as compared to the so-called amorphous materials such as 
alumina or silica-alumina, may therefore be used to advantage is view of 
their activity and resistance to poisoning. Catalysts having aromatic 
selectivity, i.e. which will crack aromatics in preference to paraffins 
are preferred because of the highly aromatic character of the feed. 
The preferred hydrocracking catalysts are the crystalline catalysts, 
generally the zeolites, and, in particular, the large pore size zeolites 
having a Constraint Index less than 2. For purposes of this invention, the 
term "zeolite" is meant to represent the class of porotectosilicates, 
i.e., porous crystalline silicates, that contain silicon and oxygen atoms 
as the major components. Other components are also present, including 
aluminum, gallium, oron, boron and the like, with aluminum being preferred 
in order to obtain the requisite acidity. Minor components may be present 
separately, in mixtures in the catalyst or intrinsically in the structure 
of the catalyst. 
Zeolites with a silica-to-alumina mole ratio of at least 10:1 are useful, 
it is preferred to use zeolites having much higher silica-to-alumina mole 
ratios, i.e., ratios of at leat 50:1. The silica-to-alumina mole ratio 
referred to may be determined by conventional analysis. This ratio is 
meant to represent, as closely as possible, the ratio in the rigid anionic 
framework of the zeolite crystal and to exclude aluminum in the binder or 
in cationic or other forms within the channels. 
A convenient measure of the extent to which a zeolite provides control to 
molecules of varying sizes to its internal structure is the Constraint 
Index of the zeolite. Zeolites which provide a highly restricted access to 
and egress from its internal structure have a high value for the 
Constraint Index, and zeolites of this kind usually have pores of small 
size, e.g., less than 5 Angstroms. On the other hand, zeolites which 
provide relatively free access to the internal zeolite structure have a 
low value for the Constraint Index and usually pores of lrge size, e.g., 
greater than 8 Angstroms. The method by which Constraint Index is 
determined is described fully in U.S. Pat. No. 4,016,218, to which 
reference is made for details of the method. A Constraint Index of less 
than 2 and preferably less than 1 is a characteristic of the hydrocracking 
catalysts used in the present process. 
Constraint Index (CI) values for some typical large pore materials are 
shown in Table 2 below: 
TABLE 2 
______________________________________ 
Constraint Index 
CI (Test Temperature) 
______________________________________ 
ZSM-4 0.5 (316.degree. C.) 
ZSM-20 0.5 (371.degree. C.) 
TEA Mordenite 0.4 (316.degree. C.) 
Mordenite 0.5 (316.degree. C.) 
REY 0.4 (316.degree. C.) 
Amorphous Silica-Alumina 
0.6 (538.degree. C.) 
Dealuminized Y (Deal Y) 
0.5 (510.degree. C.) 
Zeolite Beta 0.6-2 (316.degree.-399.degree. C.) 
______________________________________ 
The nature of the CI parameter and the technique by which it is determined 
admit of the possibility that a given zeolite can be tested under somewhat 
different conditions and thereby exhibit different Constraint Indices. 
Constraint Index may vary with severity of operation (conversion) and the 
presence or absence of binders. Other variables, such as crystal size of 
the zeolite, the presence of occluded contaminants, etc., may also affect 
the Constraint Index. It may be possible to select test conditions, e.g., 
temperature, as to establish more than one value for the Constraint Index 
of a particular zeolite, as with zeolite beta. A zeolite is considered to 
have a Constraint Index within the specific range if it can be brought 
into the range under varying conditions. 
The large pore zeolites, i.e., those zeolites having a Constraint Index 
less than 2 have a pore size sufficiently large to admit the vast majority 
of components normally found in the feeds. These zeolites are generally 
stated to have a pore size in excess of 7 Angstroms and are represented by 
zeolites having the structure of, e.g., Zeolite Beta, Zeolite X, Zeolite 
Y, faujasite, Ultrastable Y (USY), Dealuminized Y (Deal Y), Mordenite, 
ZSM-3, ZSM-4, ZSM-18 and ZSM-20. Zeolite ZSM-20 resembles faujasite in 
certain aspects of structure, but has a notably higher silica/alumina 
ratio than faujasite, as do the various forms of zeolite Y, especially USY 
and De-Aly. Zeolite Y is the preferred catalyst, and it is preferably used 
in one of its more stable forms, especially USY or De-AlY. 
Although Zeolite Beta has a Constraint Index less than 2, it does not 
behave exactly like a typical large pore zeolite. Zeolite Beta satisfies 
the pore size requirements for a hydrocracking catalyst for use in the 
present process but it is not preferred because of its paraffin-selective 
behavior. 
Because they are aromatic selective and have a large pore size, the 
amorphous hydrocracking catalysts such as alumina and silica-alumina may 
be used although they are not preferred. 
Zeolite ZSM-4 is described in U.S. Pat. No. 3,923,639; Zeolitte ZSM-20 in 
U.S. Pat. No. 3,972,983; Zeolite Beta in U.S. Pat. Nos. 3,308,069 and Re 
28,341; Low sodium Ultrastable Y molecular sieve (USY) is described in 
U.S. Pat. Nos. 3,293,192 and 3,449,070; Dealuminized Y zeolite (Deal Y) 
may be prepared by the method found in U.S. Pat. No. 3,442,795; and 
Zeolite UHP-Y is described in U.S. Pat. No. 4,401,556. Reference is made 
to these patents for details of these zeolite catalysts. 
The catalyst should have some acidity, i.e., an alpha value greater than 1 
for the cracking function. The alpha value, a measure of zeolite acidic 
functionality, is described together with details of its measurement in 
U.S. Pat. No. 4,016,218 and in J. Catalysis, Vol. VI, pages 278-287 (1966) 
and reference is made to these for such details. However, because the 
catalyst is being used in a fixed bed operation with a highly aromatic 
feed at low hydrogen pressure, it must have a low coking tending in order 
to reduce aging and for this reason, a low alpha value is preferred. Alpha 
values between 1 and 200, preferably not more than 100 are preferred, with 
values not more than 75 e.g. 50 being useful. 
Catalyst stability during the extended cycle life is essential and this may 
be conferred by suitable choice of catalyst structure and composition, 
especially silica:alumina ratio. This ratio may be varied by initial 
zeolite synthesis conditions, or by subsequent dealuminization as by 
steaming or by substitution of frame work aluminum with other trivalent 
species such as boron, iron or gallium. Because of its convenience, 
steaming is a preferred treatment. In order to secure satisfactory 
catalyst stability, high silica:alumina ratios, e.g. over 50:1 are 
preferred, e.g. about 200:1 and these may be attained by steaming. The 
alkali metal content should be held at a low value, preferably below 1% 
and lower, e.g. below 0.5% Na. This can be achieved by successive 
sequential ammonium exchange followed by calcination. 
Improved selectivity and other beneficial properties may be obtained by 
subjecting the zeolite to treatment with steam at elevated temperatures 
ranging from 500.degree. to 1200.degree. F. (399.degree.-538.degree. C.), 
and preferably 750.degree. to 1000.degree. F. (260.degree.-694.degree. 
C.). The treatment may be accomplished in an atmosphere of 100% steam or 
an atmosphere consisting of steam and a gas which is substantially inert 
to the zeolites. A similar treatment can be accomplished by lower 
temperatures and elevated pressure, e.g. 350.degree. to 700.degree. F. 
(177.degree.-371.degree. C.) at 10 to about 200 atmospheres. 
The zeolites are preferably composited with a matrix comprising another 
material resistant to the temperature and other conditions employed in the 
process. The matrix material is useful as a binder and imparts greater 
resistance to the catalyst for the severe temperature, pressure and 
reactant feed stream velocity conditions encountered in the process. 
Useful matrix materials include both synthetic and naturally occurring 
substances, such as clay, silica and/or metal oxides. The latter may be 
either naturally occurring or in the form of synthetic gelatinous 
precipitates or gels including mixtures of silica and metal oxides such as 
alumina and silica-alumina. The matrix may be in the form of a cogel. 
Naturally occurring clays which can be composited with the zeolite include 
those of the montmorillonite and kaolin families. Such clays can be used 
in the raw state as originally mined or initially subjected to 
calcination, acid treatment or chemical modification. The relative 
proportions of zeolite component and the matrix, on an anhydrous basis, 
may vary widely with the zeolite content ranging from between about 1 to 
about 99 wt %, and more usually in the range of about 5 to about 80 wt % 
of the dry composite. If the feed contains greater than 20% 650.degree. 
F.+material, that the binding matrix itself be an acidic material having a 
substantial volume of large pore size material, not less than 100 
A.degree.. The binder is preferably composited with the zeolite prior to 
treatments such as steaming, impregnation, exchange, etc., in order to 
preserve mechanical integrity and to assist impregnation with 
non-exchangeable metal cations. 
The original cations associated with each of the crystalline silicate 
zeolites utilized herein may be replaced by a wide variety of other 
cations, according to conventional techniques. Typical replacing cations 
including hydrogen, ammonium and metal cations, including mixtures of 
these cations. Useful cations include metals such as rare earth metals, 
e.g., manganese, as well as metals of Group IIA and B of the Periodic 
Table, e.g., zinc, and Group VIII of the Periodic Table, e.g., platinum 
and palladium, to promote stability (as with the rare earth cations) or a 
desired functionality (as with the Group VI or VIII metals). Typical 
ion-exchange techniques are to contact the particular zeolite with a salt 
of the desired replacing cation. Although a wide variety of salts can be 
employed, particular preference is given to chlorides, nitrates and 
sulfates. Representative ion-exchange techniques are disclosed in a wide 
variety of patents, including U.S. Pat. Nos. 3,140,249; 3,140,251; and 
3,140,253. 
Following contact with a solution of the desired replacing cation, the 
zeolite is then preferably washed with water and dried at a temperature 
ranging from 150.degree. to about 600.degree. F. (65.degree.-315.degree. 
C.), and thereafter calcined in air, or other inert gas, at temperatures 
ranging from about 500.degree. to 1500.degree. F. (260.degree.-815.degree. 
C.) for periods of time ranging from 1 to 48 hours or more. 
The hydrocracking catalyst is essentially free prior to use of metal 
components which provide hydrogenation-dehydrogenation functionality. It 
is therefore essentially free of metals of Groups VIA and VIIIA of the 
Periodic Table (IU Table) such as tungsten, vanadium, zinc, molybdenum, 
rhenium, nickel, cobalt, chromium, manganese, or a noble metal such as 
platinum or palladium, although these metals may be present in amounts 
which do not confer significant hydrogenation functionality on the 
catalyst at least prior to use. During operation some metals especially 
nickel and vanadium may be picked up from the feed and this may confer 
some hydrogenation activity but this is not essential to the operation of 
the process. 
Process Configuration 
A process using a full range LCO feed is illustrated schematically in FIG. 
1. A gas oil or resid feed to an FCC unit 10 is cracked in the FCC unit 
and the cracking products are fractionated in the cracker fractionator 11 
to produce the various hydrocarbon fractions which leave the fractionator 
in the conventional manner. A full range of a lighter cut cycle light 
cycle oil (FRLCO or LCLCO)) is withdrawn from the fractionator 11 through 
draw-off conduit 12 and is passed to hydrotreater 15 which forms the first 
stage of the hydrocracking unit. The feed is hydrotreated in unit 15 to 
effect some aromatics saturation and to hydrogenate residual heteroatoms, 
especially nitrogen and sulfur, which are removed in interstage separator 
16 as ammonia hydrogen sulfide together with excess hydrogen which is 
returned, after purificaytion, in the hydrogen circuit line 17. The 
interstage separation and gas purification may not be necessary but is 
shown here as an optional feature. The hydrotreated cycle oil then passes 
to hydrocracker 18 which forms the second stage of the unit in which ring 
opening and cracking take place to form a hydrocracked product which is 
rich in monocyclic aromatics in the gasoline boiling range. After hydrogen 
separation in separator 19, the hydrocracker effluent is fractionated in 
the conventional manner in distillation tower 20 to form the products 
including dry gas, gasoline, middle distillate and a bottoms fraction 
which may be withdrawn and blended into low sulfur fuel oil, or optionally 
recycled to FCCU 10 through recycle conduit 21. The gasoline range product 
from tower 20 is of high octane rating and is suitable for being blended 
directly into the refinery gasoline product pool without reforming or 
other treatment to improve octane number. 
While the monofunctional catalysts used in the present hydrocracking 
process achieve unexpectedly long operational cycles, it may be desirable 
to provide some method of periodic or continuous catalyst regeneration 
because the cycles may not be as long as the cycles encountered in 
conventional hydrocracking operation, typically six months to one year. 
Catalyst regeneration may be provided by means of a fixed-bed, swing 
reactor configuration. In this type of unit two reactors are loaded with 
the hydrocracking catalyst; liquid is charged through the hydrofining 
reactor and then over one of the two hydrocracking reactors. When the 
second reactor attains end-of-cycle, for example, at the end of thirty 
days, the flow is switched to other hydrocracking reactor and the reactor 
containing the aged catalyst is regenerated. This procedure is continued 
over the entire operation with one of the hydrocracking reactors in use 
while the other is regenerated. 
An alternative regenerative technique is to provide for continuous catalyst 
regeneratin by operating with a moving-bed or fluidized bed hydrocracking 
reactor and continuously withdrawing a portion of the catalyst from the 
reactor and regenerating it in a separate vessel. The hydrocracker will 
then have a configuration similar to that shown in U.S. Pat. No. 4,797,478 
which describes a fluidized bed reactor for the conversion of methanol to 
hydrocarbons, using continuous catalyst regeneration in a separate vessel. 
In either case, whether with swing reactor or continuous regeneration, it 
is preferred to operate under conditions which provide a cycle life of at 
least five and preferably at least ten days. In fluid bed operation it is 
desirable that only a small portion of the catalyst inventory require 
regeneration at one time. The regenerator generally requires a separate 
vessel as well as an air compressor; a substantial economic penalty would 
arise if the catalyst deactivated at a fast rate requiring a very large 
regeneration facility. The detrimental economic impact is compounded by 
the need to carry out the regeneration in a manner that avoids exposure of 
the catalyst to excessive heat or water vapor which would result in 
hydrothermal damage to the zeolite. Thus it is desirable that the catalyst 
does not have a fast deactivation rate. Similarly, in the case of swing 
reactor operation, the downtime taken up during regeneration reduces unit 
capacity so that again, extended cycle life is desirable. 
The fact that unexpectedly long cycle durations have been achieved 
notwithstanding, the absence of the metal component may be attributed to 
the nature of the process. The preferred light cycle oil feeds contain 
limited amounts of high molecular weight PNA coke precursors so that coke 
laydown is minimized even at the low hydrogen pressures which prevail in 
the unit. In addition, the initial hydrotreating step over the 
hydrogenation catalyst effects a sufficient degree of hydrogenation to 
obviate the necessity of a metal function in the second stage. In order to 
maintain the longest cycle durations the hydrocracking catalyst should be 
selected for minimum coking propensity and for this purpose the Y 
zeolites, especially USY, are preferred. In addition, low sodium content 
is desirable and this may suitably be achieved by successive ammonium 
exchange followed by calcination. 
A preliminary hydrotreating step before the hydrocracking is preferred in 
order to saturate polynuclear aromatics (PNAS) and to remove nitrogen and 
sulfur containing impurities. The bulk of the PNAS are found in the higher 
boiling portion of the cycle together with the bulk of the heteroatoms and 
accordingly the hydrotreating step may be carried out at the lower 
hydrotreating severity with light cut LCO feeds. 
Hydrocracking Conditions 
During the hydrocracking process the objective is to create monocyclic 
aromatics of high octane value from the aromatics in the feed. Because the 
feed contains principally bicyclic aromatics such as naphthalene, 
benzothiophene, etc., the degree of saturation during the hydrocracking 
step must be limited so as to avoid complete hydrogenation of these 
components. For this reason, relatively low to moderate hydrogen pressures 
are used, usually not more than 1000 psig (7000 kPa), with minimum 
pressures usually being about 400 psig (about 2860 kPa) and typical 
pressures in the range of 600-1000 psig (about 4250-7000 kPa). The exact 
pressure is selected according to feed characteristics (aromatic and 
heteroatom content), extent of preliminary hydrotreatment, catalyst 
stability and aging resistance and the desired product characteristics. 
Similarly, because ring opening is also to be limited in order to preserve 
the aromatic character of the gasoline product, severity (temperature, 
residence time, conversion) is also limited. Conversion to 385.degree. 
F.-(195.degree. C.-) gasoline should be below 80 volume percent and 
preferably below 65 volume percent. Although conversion may exceed 75 
volume percent, conversion levels between 55 and 70 volume percent 
represent a typical maximum. Optimum conversion levels are from about 20 
to 50 e.g. 20 to 30, volume percent for highest product octane. Because 
the absence of heteroatoms and PNAS from the feed reduces catalyst 
deactivation from heteroatom and PNA induced inhibition and coking, there 
is a reduced degree of necessity to relate conversion to hydrogen pressure 
with the LCO feeds (as described in application Ser. No. 940,328). With 
full range feeds, however, conversion should preferably be limited to no 
more than one-twentieth the hydrogen partial pressure expressed in psig, 
as described in U.S. Pat. No. 4,676,887. Pressures between 400 and 1000 
psig (2860-7000 kPa), usually in the range 600-1000 psig (4250-7000 kPa) 
with conversions up to 70 volume percent are preferred. Hydrocracking 
temperatures are typically up to 850.degree. F. (450.degree. C.) although 
higher temperatures up to about 900.degree. F. (480.degree. C.) may be 
employed, commonly with temperature minima of about 600.degree. F. 
(315.degree. C.) or higher, e.g. 700.degree. F. (370.degree. C.) being a 
recommended minimum. Space velocity will vary with temperature and the 
desired level of conversion but will typically be 0.25-2.5 hr..sup.-1, 
more usually 0.5-1.5 hr..sup.-1 (LHSV, 20.degree. C.). Hydrogen 
circulation rates of 500-5000 SCF/Bbl (90-900 n. 1.1..sup.-1) are 
suitable. Heat for the endothermic cracking reactions is readily supplied 
by the feed from the exothermic hydrotreating step and this factor, 
together with the fact that low coke make is achieved, makes fixed bed 
operation feasible, eliminating the need for continuous regeneration in 
favorable cases. 
A notable feature of the present process is the low hydrogen consumption 
during the hydrocracking step. Compared to processes carried out under 
similar conditions employing a conventional bifunctional hydrocracking 
catalyst, the hydrogen consumption is markedly lower, especially at higher 
conversions. In fact, contrary to conventional hydrocracking techniques, 
hydrogen consumption may decline with increasing conversion. Hydrogen 
consumptions of less than 1300 SCF/Bbl may be readily obtained at 
conversions (420.degree. F.+, 215.degree. C.) over 20 vol. percent and 
consumptions below 1200 SCF/Bbl at conversions (420.degree. F.+, 
215.degree. C.) of 20 to 40 vol. percent. 
Hydrotreating 
Although, as stated above, the use of two-stage hydrocracking, i.e. 
hydrotreating followed by hydrocracking is not preferred with light cut 
feeds, it is preferred with all feeds and is, in practical terms, a 
necessity with the higher boiling feeds, e.g. full range cycle oil. 
Hydrotreating will also be useful if the feed has a relatively high 
heteroatom content since hydrotreating with interstage separation of 
inorganic nitrogen and sulfur will enable extended cycle life to be 
obtained in the hydrocracking unit. Preliminary hydrotreating may be 
carried out with or without interstage separation before the hydrocracking 
step. If interstage separation is omitted, and cascaded operation in fixed 
bed reactors is employed, the hydrotreating catalyst may simply be loaded 
on top of the hydrocracking catalyst in the reactor. 
The hydrotreating catalyst may be any suitable hydrotreating catalyst, many 
of which are commercially available. These are generally constituted by a 
metal or combination of metals having hydrogenation/dehydrogenation 
activity and a relatively inert, i.e. non-acidic refractory carrier having 
large pores (20.degree. A or more). Suitable carriers are alumina, 
silica-alumina or silica and other amorphous, large pore size amorphous 
solids such as those mentioned above in connection with the hydrocracking 
catalyst binder materials. Suitable metal components are nickel, tungsten, 
cobalt, molybdenum, vanadium, chromium, often in such combinations as 
cobalt-molybdenum, nickel-molybdenum or nickel-cobalt-molybdenum. Other 
metals of Groups VI and VIII of the Periodic Table may also be employed. 
About 0.1-20 wt percent metal, usually 0.1-10 wt. percent, is typical. 
Because the catalyst is relatively non-acidic (although some acidity is 
necessary in order to open heterocyclic rings to effect hetero atom 
removal) and because temperature is relatively low, conversion during the 
hydrotreating step will be quite low, typically below 10 volume percent 
and in most cases below 5 volume percent. Temperatures will usually be 
from 600.degree. to 800.degree. F. (315.degree.-425.degree. C.), mostly 
from 625.degree. to 750.degree. F. (330.degree. to 400.degree. C.). Space 
velocity (LHSV at 20.degree. C.) will usually be from 0.25 to 4.0 
hr..sup.-1, preferably 0.4 to 2.5 hr..sup.-1, the exact space velocity 
selected being dependent on the extent of hydrotreating desired and the 
selected optional temperature. Hydrogen pressures of 200-1000 psig 
(1500-7000 kPa), preferably 400-800 psig (2860-5620 kPa) are typical with 
hydrogen circulation rates of 500-5000 SCF/Bbl (90-9000 n.1.1..sup.-1) 
being appropriate. If cascade operation is employed, the hydrotreating 
pressure will be slightly higher than that desired in the hydrocracking 
step to allow for bed pressure drop. 
The hydrotreating catalyst, like the hydrocracking catalyst, may be 
disposed as a fixed, fluidized, or moving bed of catalyst, although a 
downflow, fixed bed operation is preferred because of its simplicity. 
When a preliminary hydrotreatment is employed, conditions in the 
hydrocracking step may be adjusted suitably to maintain the desired 
overall process objective, i.e. incomplete saturation of aromatics with 
limited ring opening of hydroaromatic components to form high octane 
gasoline boiling range products. Thus, if some saturation of bicyclic 
aromatics such as naphthalene, methyl naphthalenes and benzothiophenes is 
taken in the hydrotreating step, hydrogen consumption in the hydrocracking 
step will be reduced so that a lower temperature will result if space 
velocity is kept constant (since the extent of the exothermic hydrogenatio 
reactions will be less for the same throughput in the second stage). In 
order to maintain the desired level of conversion (which is dependent on 
temperature, it may be necessary to decrease space velocity 
commensurately. 
Hydrocracker Products 
As described above, the objective of the present process is to produce a 
high octane gasoline directly. The boiling range of the gasoline will 
typically be C.sub.5 -385.degree. F. (C.sub.5 -196.degree. C.) (end point) 
but gasolines of higher or lower end points may be encountered, depending 
on applicable product specifications, e.g. C.sub.5 -330.degree. F. 
(C.sub.5 -165.degree. C.) (end point) of C.sub.5 -450.degree. F. (C.sub.5 
-232.degree. C.). Minimum target octane number is 85 clear or higher, e.g. 
87 (RON+0). In most cases, higher octane ratings are attainable, for 
example, clear ratings of at least 90 or higher, e.g. 95. In favorable 
cases, clear octane ratings of 100 or higher may be attained. In all 
cases, the gasoline boiling range product may be blended directly into the 
refinery gasoline pool without reforming or other treatment to improve 
octane. As mentioned above, the hydrocracker bottoms fraction may be 
recycled to the catalytic cracking unit where its enhanced crackability as 
a consequence of its increased hydrogen content will further improve the 
total gasoline yield, this time by increasing the yield from the cracker. 
The hydrocracker bottoms may also be combined with the high boiling cut of 
the cycle oil (from fractionator 13) after it has been hydrotreated, e.g. 
in a conventional CHD unit form a fuel oil or diesel fuel or, 
alternatively, the combined stream can be recycled to the FCCU, as 
previously described. 
The present process is notable for the production of high octane gasoline 
directly from the highly aromatic product from the catalytic cracking 
unit. The use of lower hydrogen pressures and moderate processing 
conditions in the hydrocracker enables this result to be achieved with low 
hydrogen consumption and low utility requirements. 
The invention is illustrated in the following Examples. 
EXAMPLE 1 
A light cycle oil from a fluid catalytic cracking operation was used as 
feed. It had the following properties as set out in Table 3. 
TABLE 3 
______________________________________ 
LCO Hydrocracking Feed 
______________________________________ 
Nominal Boiling Range, .degree.F. (.degree.C.) 
400-725 (204-385) 
Gravity, API 12.1 
Sulfur, wt. % 2.8 
Nitrogen, ppmw 660 
Hydrogen, wt. % 9.3 
______________________________________ 
The feed was subjected to two stage cascade hydrotreating/hydrocracking (no 
interstage separation) using a commercial Ni--Mo/Al.sub.2 O.sub.3 
hydrotreating catalyst at 675.degree. F. (355.degree. C.), 600 psig (4240 
kPa) H.sub.2 (inlet), 1 LSHV for the hydrotreating catalyst and a 
silica-bound, steamed ultrastable Y zeolite as the hydrocracking catalyst 
with no added metal component for hydrogenation functionality. Properties 
of the hydrocracking catalyst were shown in Table 4 below. 
TABLE 4 
______________________________________ 
Monofunctional HC Catalyst 
Type USY/SiO.sub.2 
______________________________________ 
USY, wt. % 65 
SiO.sub.2, wt. % 
35 
Alpha 55 
______________________________________ 
The hydrocracking was carried out for 38 days at 600 psig (4240 kPa) 
H.sub.2 inlet pressure, 1 hr.sup.-1 LHSV and temperatures ranging from 
690.degree. F. (365.degree. C.) to 810.degree. F. (430.degree. C.) to 
maintain a constant conversion of 20 vol. % to 385.degree. F.- 
(196.degree. C.-) gasoline. The aging of the catalyst was followed by 
plotting the temperature required to maintain the specified conversion and 
the results are shown in FIG. 2. An end-of-cycle (EOC) temperature of 
about 800.degree. F. (425.degree. C.) was reached only after 30 days. 
The properties of the product were as follows at 12 days TOS. 
TABLE 5 
______________________________________ 
Hydrocracked Products 
______________________________________ 
Feed LCO 
TOS, days 12 
Reactor temp., .degree.F., (.degree.C.) 
727 (386) 
H.sub.2 Pressure, psig (kPa) 
600 (4240) 
LHSV, overall 0.5 
Product Distribution, wt. %: 
C.sub.5 -- 4.2 
C.sub.5 --385.degree. C. (C.sub.5 --195.degree. C.) 
17.8 
385.degree.-420.degree. F.(195-215.degree. C.) 
4.3 
420.degree. F. + Distillate 
73.7 
385.degree. F. Gasoline RON: 
93.9 
______________________________________ 
The results above and in FIG. 2 show that it is possible to employ a 
monofunctional catalyst for cracking dealkylated feeds to high octane 
gasoline products. 
EXAMPLES 2-4 
To compare performance of the dual catalyst system and one catalyst system, 
three additional experiments were conducted. The dual catalyst system 
consists of a hydrotreating catalyst and an unpromoted USY catalyst so as 
to separate hydrotreating function and cracking function in two individual 
catalysts. In the one catalyst system, a NiMo promoted USY was used to 
provide both hydrotreating and cracking functions. The experiments used a 
light cut LCO as the feedstock as in Table 6. Catalyst properties are 
given in Table 7. Through this summary, HDT refers to a commercial 
NiMo/Al.sub.2 O.sub.3 hydrotreating catalyst; USY, a unpromoted USY 
catalyst; and NiMO USY is the NiMo version of the USY catalyst. Prior to 
metal additions, extrudates of the USY and NiMo USY catalysts were steamed 
to 55 alpha. 
TABLE 6 
______________________________________ 
Feedstock Properties - Light Cut LCO 
______________________________________ 
Nominal Boiling Range, .degree.F. (.degree.C.) 
360-600 (182-315) 
Gravity, API 16.8 
Sulfur, wt. % 2.7 
Nitrogen, ppmw 190 
Hydrogen, wt. % 9.5 
Composition, vol. % 
Aromatics 87 
Olefins 3 
Saturates 10 
______________________________________ 
TABLE 7 
______________________________________ 
Catalyst Properties 
Catalyst HDT USY NiMo/USY 
______________________________________ 
Type Hydro- Monofunctional 
Bifunctional 
treating Hydrocracking 
Hydrocracking 
Composition, wt % 
USY zeolite.sup.(2) 
non 65.sup.(1) 65.sup.(1) 
Nickel/oxide 
4.5 -- 4.8 
Molybdenum 20.0 -- 9.3 
Aluminum balance balance balance 
______________________________________ 
Notes: 
.sup.(1) prior to metal additions 
.sup.(2) steamed to 65 alpha activity measured prior to metal additions 
Three experiments were conducted: Examples 2 and 3 used the dual catalyst 
system, and Example 4 used the one catalyst system. For Examples 2 and 3, 
the hydrotreating catalyst and USY catalyst were loaded in separate 
reactors in a 50/50 volumetric ratio and operated in cascade mode without 
an interstage separation of H.sub.2 S and NH.sub.3. During the 
experiments, the USY catalyst temperature was varied to cover a wide range 
conversion, while hydrotreating catalyst temperature was kept constant at 
610.degree. and 685.degree. F. (320.degree. and 365.degree. C.), for 
Example 2 and Example 3, respectively. Detailed experimental conditions 
are summarized in Table 8 below. 
TABLE 8 
______________________________________ 
Experimental Conditions 
Example 2 
Example 3 Example 4 
______________________________________ 
Pressure, psig (kPa abs) 
600(4240) 600(4240) 600(4240) 
Catalyst: HDT/USY HDT/USY NiMo/USY 
Catalyst Ratio 
1:1 1:1 -- 
H.sub.2 circ., SCF/B(n.1.1..sup.-1 
5000(890) 5000(890) 5000(890) 
Reactor Temp, .degree.F.(.degree.C.) 
HDT 610(320 685(363) -- 
USY 600-750 700-772 600-730 
(315-400) (370-110) (315-390) 
Overall LHSV, hr.sup.-1 
0.8 0.8 0.5-0.8 
420.degree. F..sup.+ conversion, pct. 
15-40 25-35 20-60 
______________________________________ 
The results are given in Table 9 below and in FIGS. 3 and 4. Table 9 gives 
the C.sub.4 hydrocarbon distribution at fixed conversion levels while 
FIGS. 3 and 4 show the gasoline product octane (R+0) and hdrogen 
consumption, respectively, at the different conversion levels employed. 
TABLE 9 
______________________________________ 
C.sub.4 Hydrocarbon Distribution 
Example 2 
Example 3 Example 4 
______________________________________ 
420.degree. F..sup.+ conv, pct 
40 29 38 
Catalyst HDT/USY HDT/USY NiMo/USY 
C.sub.4 Distribution, pct. 
i-C.sub.4 75 65 35 
C.sub.4.sup.= 
5 23 10 
N--C.sub.4 20 12 65 
______________________________________ 
Results showed unexpectedly that the HDT/USY catalyst system (Examples 2 
and 3) produces naphthas having an octane (R+O) higher than the NiMo USY 
catalyst (Example 4) over the conversion range studied (FIG. 3). (Note: 
conversion is defined as percentage of 420.degree. F..sup.+ (215.degree. 
C..sup.+) materials in the feed converted to the 420.degree. F..sup.- 
(215.degree. C..sup.-) materials. At equivalent conversion, the HDT/USY 
catalyst system makes slightly more C1-C4 light gases. However, the C4s 
produced from the HDT/USY catalyst system (Examples 2 and 3) are rich in 
iso-butane and butenes, and can be directly blended into the alkylation 
feeds to further increase gasoline yield. Table 9 above compares C4 stream 
compositions for the three examples. Total iso-butane and butenes 
concentrations for Examples 2 and 3 are greater than 80%, compared to 45% 
for Example 4. 
Additionally, the HDT/USY catalyst system consumes less hydrogen than the 
NiMo USY catalyst, particularly at high conversion as shown in FIG. 4.