[0001] This invention relates to a process for catalytically hydrocracking hydrocarbon chargestocks
to produce low pour point distillate products of reduced viscosity.
[0002] Hydrocracking is a process which has achieved widespread use in petroleum refining
for converting various petroleum fractions into lighter and more valuable products,
especially gasoline and distillates such as jet fuels, diesel oils and heating oils.
Hydrocracking is generally carried out in conjunction with an initial hydrotreating
step in which the heteroatom-containing impurities in the feed are hydrogenated without
a substantial bulk conversion of tne feed. During this step, the heteroatoms, principally
nitrogen and sulfur, are converted into ammonia and hydrogen sulfide and these gases
may be removed prior to the subsequent hydrocracking step although the two stages
may be combined in cascade without interstage separation, for example, as in the Unicracking-JHC
Process and as described in U.S. Patent 4,435,275. However, the presence of large
quantities of ammonia in the hydrotreating effluent may result in a significant suppression
of cracking in the subsequent hydrocracking step although this may be compensated
by an increase in severity.
[0003] In the subsequent hydrocracking step, the hydrotreated feedstock is contacted with
a catalyst which has both an acidic function and a hydrogenation function. In the
first step of the reaction, the polycyclic aromatics in the feedstock are hydrogenated,
after which cracking takes place together with further hydrogenation. Depending upon
the severity of the reaction conditions, the polycyclic aromatic in the feedstock
will be hydrocracked down to paraffinic materials or, under less severe conditions,
to monocyclic aromatics as well as paraffins.
[0004] The acidic function in the catalyst is provided by a carrier such as alumina, silica-alumina
or a crystalline zeolite such as faujasite, zeolite X, zeolite Y or mordenite. Large
pore zeolites have proved to be highly useful catalysts for this purpose because they
possess a high degree of intrinsic cracking activity and, for this reason, are capable
of producing a good yield of gasoline. They also possess a better resistance to nitrogen
and sulfur compounds than the amorphous materials such as alumina and silica-alumina.
[0005] The hydrogenation function is provided by a metal or combination of metals. Noble
metals of Group VIIIA of the Periodic Taole (the Periodic Table being that approved
by IUPAC), especially platinum or palladium may be used, as may base metals of Groups
IVA, VIA and VIIIA, especially chromium, molybdenum, tungsten, cobalt and nic
Kel. Combinations of metals such as nickel-molybdenum, cobalt-molybdenum, cobalt-nickel,
nickel-tungsten, cobalt-nickel-molybdenum and nickel-tungsten-titanium have been shown
to be very effective and useful.
[0006] In EP-A-94827, there is described a hydrocracking process which, besides achieving
a bulk conversion of the feedstock, also dewaxes it. The zeolite catalyst component
used in tnat process, zeolite beta, has a number of highly useful and characteristic
properties. First, it shows a significant distillate selectivity; that is, it tends
to produce hydrocracked products boiling in the distillate range (about 165°-345°C)
as opposed to conventional hydrocracking catalysts which are naphtha-directing and
which tend to
' produce a gasoline boiling range (about C
5 to 165°C) product. Although this behavior is shared by other highly siliceous zeolites
such as high-silica Y and high silica ZSM-20, zeolite beta also has the unique ability
to hydroisomerize and hydrocrack the paraffinic components of the feed. This is in
marked contrast to the behavior of other zeolites such as zeolite Y: if a waxy feedstock
is hydrocracked with a conventional large pore catalyst such as zeolite Y, the viscosity
of the oil is reduced by cracking most of the 343°C+ material into lower boiling products.
Tne remainder of the 345°C+ material that is not converted, however, contains the
majority of the paraffinic components in the feedstock because with these conventional
catalysts the aromatics are converted preferentially as compared to the paraffins.
The unconverted 345°C+ material therefore retains a high pour point so that the final,
hydrocracked product containing the unconverted paraffins will also have a relatively
high pour point. Thus, although the viscosity is reduced, the pour point mignt still
be unacceptable. Even if the conditions are adjusted to give complete or nearly complete
conversion, the higher molecular weight hydrocarbons which are present in the product
will contain a substantial proportion of straight chain components (n-paraffins).
If these are of sufficiently high molecular weight themselves (as they often are)
they will constitute a waxy component in the product. The final product may therefore
be proportionately more waxy than the feedstock (because the non- paraffinic components
have been selectively removed by cracking) and, consequently, may have a pour point
whicn is equally unsatisfactory or even more so. Attempts to reduce the molecular
weight of tnese straight chain paraffinic products will only serve to produce very
light fractions, for example propane, butanes and light naphtha, so decreasing the
desired liquid yield.
[0007] Zeolite beta, by contrast, removes the paraffinic components so that a dewaxing effect
is achieved simultaneously with the bulk conversion. So, if a gas oil containing paraffins,
naphthenes and aromatics is treated under hydrocracking conditions with a zeolite
beta catalyst, all three types of hydrocarbon will be converted whereas other zeolites
would selectively hydrocrack the naphthenes and aromatics and concentrate the paraffins.
[0008] As mentioned in EP-A-94827, a preliminary hydrotreating step is desirable in order
to remove nitrogen and sulfur and to saturate aromatics to naphthenes without substantial
bulk conversion. The zeolite beta-catalyzed hydrocracking process is therefore adaptable
to otherwise conventional hydrocracking operations.
[0009] It has now been found that zeolite beta hydrocracking catalysts offer a number of
advantages in hydrocracking processes carried out at unconventionally low pressures,
typically below 10,000 kPa hydrogen partial pressure. In particular, compared to a
similar process using a conventional catalyst system, product quality is improved
while hydrogen consumption is decreased. Specific product characteristics in which
improvements may be noted include low sulfur content, Smoke Point (which is higher,
that is to say better) and Diesel Index (which is higher). The products of the process
are especially useful as diesel and jet fuels because of their highly paraffinic compositions
coupled witn their low pour points; however, they may also be used as highly satisfactory
heating oils and kerosenes because of their low pour point and sulfur content. The
catalyst is, moreover, capable of maintaining hydrocracking activity over long cycles,
typically about one year or longer. The zeolite beta catalyst is, in particular, noted
for its ability to maintain its activity for extended periods of time even when operating
at relatively high conversions under low pressures. Conventional catalysts tend, by
contrast, to undergo rapid aging when operated at high conversions under low pressures,
typically below 10,000 kPa.
[0010] According to the invention, there is provided a hydrocracking process for making
a highly paraffinic, low pour point distillate product from a heavy hydrocarbon oil
feed boiling substantially above 345°C, which comprises:
(i) passing the heavy oil feed over a hydrotreating catalyst in the presence of hydrogen
at elevated temperature and at a hydrogen partial pressure of not more than 10,000
kPa to hydrotreat the oil;
(ii) passing the hydrotreated oil from step (i) over a hydrocracking catalyst comprising
zeolite beta and a hydrogenation-dehydrogenation component in the presence of hydrogen
at elevated temperature and at a hydrogen partial pressure of not more than 10,000
kPa to hydrocrack the oil at a bulk conversion of at least 40 weight percent and;
(iii) recovering from the hydrocracked product from step (ii) a fraction boiling below
345°C.
[0011] The process is operated with an initial hydrotreating step and this is followed by
the hydrocracking step either with or without interstage separation. Both steps are
carried out at the low pressures which characterize the process.
[0012] The feedstock for the process is a heavy oil fraction having an initial boiling point
of at least 200°C and normally of at least 345°C or higher. Suitable feedstocks of
this type include gas oils such as vacuum gas oil, coker gas oil, visbreaker oil,
deasphalted oil and catalytic cracker cycle oil. Normally, the feedstock will have
an extended boiling range, for example 345 to 590°C but may be of more limited ranges
with certain feedstocks. For reasons which are explained below, the nitrogen content
is not critical; generally it will be in the range 200 to 1000 ppmw. Likewise, the
sulfur content is not critical and typically may range as high as 5 percent by weight.
Sulfur contents of 2.0 to 3.0 percent by weight are common.
[0013] These heavy oil feeds will comprise high molecular weight long chain paraffins and
high molecular weight aromatics with a large proportion of fused ring aromatics. During
the processing, the fused ring aromatics and naphthenes are cracked by the acidic
catalyst and the paraffinic cracking products, together with paraffinic components
of the initial feedstock undergo isomerization to iso-paraffins with some cracking
to lower molecular weight materials. Hydrogenation of unsaturated side chains on the
monocyclic cracking residues of the original polycyclics is catalyzed by the hydrogenation-dehydrogenation
component of the hydrocracking catalyst to form substituted monocyclic aromatics which
are highly desirable end products. The heavy hydrocarbon oil feedstock will normally
contain a substantial amount boiling above 230°C and will normally have an initial
boiling point of at least about 290°C, more usually about 345°C. Typical boiling ranges
will be about 345 to 565°C or about 345 to 510°C but oils with a narrower boiling
range may, of course, be processed, for example those with a boiling range of about
345 to 455°C. Heavy gas oils are often of this kind as are cycle oils and other non-residual
materials. It is possible to co-process materials boiling below 260°C but the degree
of conversion will be lower for such components. Feedstocks containing lighter ends
of this kind will normally have an initial boiling point above 150°C.
[0014] The process is of particular utility with highly paraffinic feeds because, with feeds
of this kind, the greatest improvement in pour point may be obtained. However, most
feeds will contain a certain content of polycyclic aromatics; it is a notable feature
of zeolite beta that it retains its ability to remove the waxy components of the feed
even in the presence of substantial amounts of aromatics, for example 10 percent or
more aromatics. However, the aromatic content of the feed will normally not exceed
50 percent by weight of the feedstock. Typically the aromatic content will be 20-50,
more usually about 30, weight percent of the feed.
[0015] The feedstock may contain relatively large proportions of w
3xy hydrocarbons in the 345°C+ fraction; these waxy hydrocarbons be characterized chemically
as straight chain and slightly branched cnain paraffins, that is n-paraffins and iso-paraffins
having short chain branches. The higher molecular weignt paraffins will generally
be slightly branched chain materials as opposed to being wholly straight chain but
because the branchings are relatively short, the linear chain characteristics will
predominate, so that the material will be of a waxy nature, contributing to a high
pour point.
[0016] In the process, the feedstock is heated to an elevated temperature and is then passed
over the hydrotreating and hydrccracking catalysts in the presence of hydrogen. Because
the thermodynamics of hydrocracking become unfavorable at temperatures above about
450°C, temperatures above this value will not normally be used. In addition, because
the hydrotreating and hydrocracking reactions are net exothermic, the feedstock need
not be heated to the temperature desired in the catalyst bed which is normally in
the range 360 to 440°C. At the beginning of the process cycle, the temperature employed
will be at the lower end of this range but as the catalyst ages, the temperature may
be increased in order to maintain the desired degree of activity.
[0017] The heavy oil feedstock is passed over the catalysts in the presence of hydrogen.
The space velocity of the oil is usually in the range 0.1 to 10 LHSV, preferably 0.2
to 2.0 LHSV and the hydrogen circulation rate from 250 to 1000 n.1.1 (liters of hydrogen
- measured at standard temperature and pressure - per liter of oil), and more usually
from 300 to 800 n.l.l
-1. Hydrogen partial pressure is usually at least 75 percent of the total system pressure
with reactor inlet pressures normally being in the range of 3000 to 10000 kPa, more
commonly from 5000 to 7000 kPa. When operating at low conversions, for example, less
than 50 volume percent conversion to 345°C- products, the pressure may be considerably
lower than normal, conventional practices. It has been found tnat pressures of 5000
to 7000 kPa are satisfactory, as compared to the pressures of at least 10,500 kPa
normally used in commercial hydrocracking processes. However, if desired, low conversion
may be obtained by suitable selection of other reaction parameters, for example temperature,
space velocity, choice of catalyst, even lower pressures may be used. Low pressures
are desirable from the point of view of equipment design since less massive and consequently
cheaper equipment will be adequate. Similarly, lower pressures usually influence less
aromatic saturation and thereby permit economy in the total amount of hydrogen consumed
in the process. However, certain catalysts may not be sufficiently active at very
low pressures, for example 3000 kPa and higher pressures may then be necessary at
the space velocities desired in order to maintain a satisfactory throughput.
[0018] In the first stage of the process, the feed is passed over a hydrotreating catalyst
to convert nitrogen and sulfur containing compounds into gaseous ammonia and hydrogen
sulfide. At this stage, hydrocracking is minimized but partial hydrogenation of polycyclic
aromatics proceeds, together with a limited degree of conversion to lower boiling
(345
0C-) products. The catalyst useo in this stage may be a conventional denitrogenation
(denitrification) catalyst. Catalysts of this type are relatively immune to poisoning
by the nitrogenous and sulfurous impurities in the feedstock and, generally comprise
a non-noble metal component supported on an amorphous, porous carrier such as silica,
alumina, silica-alumina or silica-magnesia. Because extensive cracking is not desired
in this stage of the process, the acidic functionality of tne carrier may be relatively
low compared to that of the subsequent hydrocracking catalyst. The metal component
may be a single metal from Groups VIA and VIIIA of the Periodic Table such as nickel,
cobalt, chromium, vanadium, molybdenum, tungsten, or a combination of metals such
as nickel-molybdenum, cobalt-nickel-molybdenum, cobalt-molybdenum, nickel-tungsten
or nickel-tungsten-titanium. Generally, the metal component will be selected for good
hydrogen transfer activity; the catalyst as a whole will have good hydrogen transfer
and minimal cracking characteristics. The catalyst should be pre-sulfided in the normal
way in order to convert the metal component (usually impregnated into the carrier
and converted to oxide) into the corresponding sulfide.
[0019] In the hydrotreating (denitrogenation) stage, the nitrogen and sulfur impurities
are converted into ammonia and hydrogen sulfide. At the same time, the polycyclic
aromatics are partially hydrogenated to form substituted aromatics which are more
readily cracked in the second stage to form alkyl aromatics. Because the process may
be operated with only a limited degree of overall conversion, the effluent from the
first stage may be passed directly to the second or hydrocracking stage without the
conventional interstage separation of ammonia or hydrogen sulfide, although hydrogen
quenching may be carried out in order to control the effluent temperature and to control
the catalyst temperature in the second stage. However, interstage separation of ammonia
and hydrogen sulfide and light fractions may be carried out, especially with the noble
metal hydrocracKing catalysts which are more sensitive to the impurities.
[0020] The effluent from the denitrogenation/desulfurization stage is passed to the hydrocracking
step to crack partially hydrogenated aromatics and carry out the other characteristic
reactions which take place over the hydrocracking catalyst.
[0021] The hydrocracking catalyst comprises zeolite beta at least partly in the hydrogen
form as an acidic component, together with a hydrogenation-dehydrogenation component.
The hydrogenation-dehydrogenation component is provided by a metal or combination
of metals. Noble metals of Group VIIIA, especially platinum, or base metals of Groups
IVA, VIA and VIIIA, especially chromium, molybdenum, tungsten, cobalt and nickel,
may be used. Base metal combinations such as nickel-molybdenum, cobalt-nickel, nickel-tungsten,
cobalt-nickel-molybdenum and nickel-tungsten-titanium are useful, although for certain
applications platinum is preferred.
[0022] The content of the metal component will vary according to its catalytic activity.
Thus, the highly active noble metals may be used in smaller amounts than the less
active base metals. For example, about 1 weight percent or less platinum will be effective
and in a preferred base metal combination, about 7 weight percent nickel and about
2.1 to about 21 weight percent tungsten, expressed as metal. The hydrogenation component
can be exchanged onto the zeolite, impregnated into it or physically admixed with
it. If the metal is to be impregnated into or exchanged onto the zeolite, it may be
done, for example, by treating the zeolite with a platinum metal-containing ion. Suitable
platinum compounds include chloroplatinic acid, platinous chloride and various compounds
containing the platinum ammine complex. The metal compounds may be either compounds
in which the metal is present in the cation of the compound and compounds in wnich
it is present in the anion of the compound. Both types of compounds can be used. Platinum
compounds in which the metal is in the form of a cation or cationic complex, for example
Pt(NH
3)
4C1
2 are particularly useful, as are anionic complexes such as the vanadate and metatungstate
ions. Cationic forms of other metals are also very useful since they may be exchanged
onto the zeolite or impregnated into it.
[0023] The acidic component of the hydrocracking catalyst is zeolite beta. Zeolite beta
is a crystalline zeolite having a pore size greater than 5 Angstrom units (5 x 10*
m). Its composition and X-ray structure are described in U.S. Patents 3,308,069 and
Re 28,341, to which reference may be made for a description of this zeolite, its preparation
and properties. Hydrocracking catalysts based on zeolite beta are described in EP-A-94827,
to which reference may be made for a description of them.
[0024] When it is used in the catalysts for the process of the invention, the zeolite is
at least partly in the hydrogen form in order to provide the desired acidic functionality
for the cracking reactions which are to take place. It is normally preferred to use
the zeolite in a form which has sufficient acidic functionality to give it an alpha
value of 1 or more. The alpha value, a measure of zeolite acidic functionality, is
described, together with details of its measurement in U.S. Patent 4,016,218 and in
J. Catalysis, Vol. VI, pages 278-287 (1966) and reference may be made to these for
such details. The acidic functionality may be controlled by base exchange of the zeolite,
especially with alkali metal cations such as sodium, by steaming or by control of
the silica:alumina ratio of the zeolite.
[0025] It has been found that steamed zeolite beta catalysts having an alpha value of from
100 to 200, preferably about 150, are preferred for the process, as compared to unsteamed
catalysts having alpha equal to about 600 to 800, particularly with base metal hydrogenation-dehydrogenation
components, especially nickel-tungsten, the steamed catalysts having been found to
be more stable for conversion.
[0026] Because the nydrogenation functionality may also be varied by choice of metal and
its relative quantity, the balance between the hydrogenation and cracking functions
may be adjusted as circumstances require. The ammonia produced in the first stage
will, to some degree, tend to reduce the acidic functionality of the hydrocrcking
catalyst but in the present process only a limited degree of conversion is desired
and so the reduced cracking consequent upon the diminution of acidic functionality
is not only acceptable but also useful.
[0027] The preferred forms of zeolite beta for use in the present process are the high silica
forms, having silica:alumina ratios of at least 30:1. It has been found, in fact,
that zeolite beta may be prepared with silica:alumina ratios above the 200:1 maximum
specified in U.S. Patents 3,308,069 and Re. 28,341. Ratios of at least 50:1 and preferably
at least 100:1 or even higher, for example 250:1, 500:1 may be used.
[0028] The silica:alumina ratios referred to are the structural or framework ratios, related
to the ratio of the Si0
4 to the A10
4 tetrahedra which together constitute the structure of which the zeolite is composed.
This 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 the 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 method
described below which result in the presence of ionic aluminum free of the zeolite
structure, are employed. Due care should therefore be taken to ensure that the framework
silica:alumina ratio is correctly determined.
[0029] The preparation of the highly siliceous forms of zeolite beta is described in U.S.
Patent 4,419,220, to which reference may be made for a description of these forms
and their preparation.
[0030] Prior to use the zeolite should be dehydrated at least partially. This can be done
by heating to a temperature in the range of 200 to 600°C in air or an inert atmosphere
such as nitrogen for 1 to 48 hours. Dehydration can also be performed at lower temperatures
merely by using a vacuum, but a longer time is required to obtain a sufficient amount
of dehydration.
[0031] It may be desirable to incorporate the catalyst in another material resistant to
the temperature and other conditions employed in the process. Such matrix materials
include synthetic and naturally occurring substances such as inorganic materials,
for example clay, silica and metal oxides, as described in U.S. Patent 4,419,220.
Matrix materials may themselves possess catalytic properties, generally of an acidic
nature.
[0032] The catalyst may be treated oy conventional pre-sulfiding treatments, for example
by heating in the presence of hydrogen sulfide, to convert oxide forms of the metals
such as Co0 and NiO into their corresponding sulfides.
[0033] The relative proportions of the hydrocracking and the hydrotreating catalysts may
be varied according to the feedstock in order to convert the nitrogen in the feedstock
into ammonia before the charge passes to tne hydrocracking step; the object is to
reduce the nitrogen level of the charge to a point where the desired degree of conversion
by the hydrocracking catalyst is attained with the optimum combination of space velocity
and reaction temperature. The greater the amount of nitrogen in the feed, the greater
then will be the proportion of hydrotreating (denitrogenation) catalyst relative to
the hydrocracking catalyst. If the amount of nitrogen in the feed is low, the catalyst
ratio may be as low as 10:90 (by volume, denitrogenation:hydrocracking). In general,
however, ratios between 25:75 to 75:25 will be used. with many stocks an approximately
equal volume ratio will be suitable, for example 40:60, 50:50 or 60:40.
[0034] In addition to the denitrogenation function of the hydrotreating catalyst another
and at least as important function is desulfurization since the sulfur content of
the distillate product is one of the most important product specifications which have
to be observed. The low sulfur products are more valuable and are often required by
environmental regulation; the degree of desulfurization achieved is therefore of considerable
significance. The degree of desulfurization obtained will be dependent in part upon
the ratio of the hydrotreating catalyst to the hydrocracking catalyst and appropriate
choice of the ratio will be an important factor in the selection of process conditions
for a given feedstock and product specification. The degree of desulfurization will
increase as the proportion of the hydrotreating catalyst increases and the lowest
sulfur contents consistent with the required conversion may be obtained with an appropriate
selection of the catalyst ratio. Another function of the hydrotreating catalyst is
to aid in the saturation of polycyclic coke precurors and this, in turn, helps in
extending the life of the hydrocracking catalyst.
[0035] The degree of desulfurization is, of course, dependent upon factors other than the
choice of catalyst ratio. It has been found that the sulfur content of the distillate
product is dependent in part upon the conversion and regulation of the conversion
will therefore enable the sulfur content of the distillate to be further controlled:
greater desulfurization is obtained at higher conversions and therefore the lowest
sulfur content distillates will be obtained near the desired maximum conversion. Alternatively,
it may be possible to increase the degree of desulfurization at a given conversion
by raising the temperature of the hydrotreating bed while holding the temperature
of the hydrocracking bed constant. This may be accomplished by appropriate use of
hydrogen quenching.
[0036] The overall conversion may be maintained at a low level, less than 50 volume percent
to lower boiling products, usually 340°C- products from the heavy oil feedstocKs used
while still maintaining satisfactory product quality. The conversion may, of course,
be maintained at even lower levels, for example 30 or 40 percent by volume. The degree
of cracking to gas (C
4-) which occurs at these low conversion figures is correspondingly low and so is the
conversion to naphtha (200°C-); the distillate selectivity of the process is accordingly
high and overcracking to lighter and less desired products is minimized. It is believed
that in cascade operation this effect is procured, in part, by the effect of the ammonia
carried over from the first stage. Control of conversion may be effected by conventional
expedients such as control of temperature, pressure, space velocity and other reaction
parameters.
[0037] Surprisingly, it has been found that the presence of nitrogen and sulfur compounds
in the second stage feed does not adversely affect catalyst aging in the absence of
interstage separation provided that sufficient denitrogenation catalyst is employed.
Catalyst life before regeneration in this process may typically be one year or even
longer. Tne extended operational life of the hydrocracking catalyst in the presence
of nitrogen and sulfur, present as ammonia and hydrogen sulfide, respectively, in
the second stage feed is a surprising aspect of operation in the cascade mode. Further,
the stability of tne catalyst is even more remarkable at the relatively low hydrogen
partial pressures utilized in low conversion operation. Generally, the activity of
cracking catalysts is adversely and severely affected by nitrogen poisoning and carbon
(coke) deposition to such an extent that with an FCC catalyst, for example, the coke
deposition is so rapid that regeneration must be carried out continuously in order
to maintain sufficient activity. In hydrocracking, the experience is that low hydrogen
partial pressures are conducive to more rapid coke accummulation as the polycyclic
coke precursors undergo polymerization; higher hydrogen pressure, on the other hand,
tends to inhibit coke formation by saturating these precursors before polymerization
takes place. For these reasons, the excellent stability of the hydrocracking catalyst
in this process is quite unexpected. The zeolite beta hydrocracking catalysts are
notable for their ability to operate at relatively high conversions, for example at
least 50% conversion to 345
0C- products under low hydrogen pressures; conventional catalysts, by contrast, age
rapidly under such conditions. When regeneration is, however, necessary, for example
after one year, it may be carried out oxidatively in a conventional manner.
[0038] The conversion of the organic nitrogen compounds in the feedstock over the hydrotreating
catalyst to inorganic nitrogen (as ammonia) enables the desire degree of conversion
to be maintained under relatively moderate and acceptable conditions, even with relatively
nitrogenous feedstocks. Severe problems would be encountered with nitrogenous feedstocks
if the hydrotreating catalyst were not used: in order to maintain the desired conversion
it would oe necessary to raise the temperature but if the feedstock is highly nitrogenous,
it might be necessary to go to temperatures at which the hydrocracking reactions become
thermodynamically unfavored. Furthermore, the volume of catalyst is fixed because
of the design of the plant and this imposes limits on the extent to which the space
velocity can be varied, thereby imposing additional processing restrictions. The hydrotreating
catalyst, on the other hand, shifts the nitrogen content of the feedstock into inorganic
form in which it does not inhibit the activity of the catalyst as much as it would
if it were in its original organic form, even though some reduction in activity is
observed, as mentioned above. Thus, higher conversion may be more readily achieved
at reduced temperatures, higher space velocities or both. Product distribution will,
however, remain essentially unaffected at constant conversion.
[0039] The present process has the advantage that it may be operated in existing low pressure
equipment. For example, if a desulfurizer is available, it may be used with relatively
few modifications since the present process may be operated at low pressures comparable
to the low severity conditions used in desulfurization. This may enable substantial
savings in capital costs to be made since existing refinery units may be adapted to
increase the pool of distillate products. If new units are to be built there is still
an economic advantage because the equipment does not have to be designed for such
high pressures as are commonly used in conventional hydrocracking processes. However,
minor modifications may be necessary to existing equipment in order to maintain operation
within the nominal limits selected. For example, a hydrodesulfurizer may require quench
installation in order to keep the temperature in the hydrocracking bed to the desired
value; alternatively, an additional reactor may be provided with appropriate quenching.
The precise reactor configuration used will, of course, depend upon individual requirements;
the skilled person will be able to appreciate and design the plant appropriately.
An exemplary hydrocracking unit without interstage separation is shown in simplified
form in U.S. Patent 4,435,275, to wnich reference may tie made for a description of
it.
[0040] A particularly surprising attribute of the present process is that it is possible
to use Pt/beta as the hydrocracking catalyst without interstage separation of heteroatoms.
As is well known, platinum-containing catalysts are particularly prone to poisoning
and it is therefore surprising that the Pt/beta functions as it does; in fact, compared
to single stage hydrocracking without hydrotreating, less Pt/beta is required, other
conditions being equal. Furthermore, the Pt/beta catalyst has been shown to be superior
in hydroisomerization activity, as compared to NiW/beta, although its hydrocracking
activity is lower than that of NiW/beta either in the cascade (no interstage separation)
or two stage (interstage separation) modes. Because of this, the pour point obtained
with the Pt/beta hydrocracking catalyst may be lower, for example by about 10°C than
with the NiW/beta. The distillate selectivity of the Pt/beta catalyst may also be
higher than that of the NiW/
beta version under comparable conditions although the NiW/beta catalysts are also notable
for giving higher rates of desulfurization and lower hydrogen consumption at equivalent
conversions, depending upon the hydrotreating conditions. The Ni-W containing catalysts
are also more active for conversion than the Pt-containing catalysts.
[0041] Compared to single stage hydrocracking, that is hydrocracking with a single catalyst,
two-stage hydrotreating/hydrocracking (with interstage separation) achieves greater
activity, permitting lower temperatures to be used, while using less zeolite, typically
about 30 percent less. An improvement in distillate selectivity may also result. An
additional benefit is that desulfurization and paraffin isomerization selectivity
are also improved. Cascade processing without interstage separation provides product
yields and qualities similar to those of separate two-stage processing without any
major loss of conversion activity, although some loss of isomerization selectivity-may
ensue. However, any such loss may be compensated for by using a Pt/beta catalyst instead
of a base metal catalyst, for example NiW/beta.
[0042] The hydrocracked products are low sulfur, low pour point distillates, generally containing
less than 0.3 weight percent sulfur. Compared to the distillate products obtained
by hydrocracking over conventional catalysts at low pressures (less than 7,000 kPa)
using limited conversions, the product has a lower aromatic content. Zeolite beta
hydrocracking catalysts gives a more paraffinic distillate product than conventional
catalysts because it hydrocracks paraffins in preference to aromatics; conventional,
amorphous hydrocracking catalysts by contrast, tend to act on the aromatics and at
low pressures produce an even more aromatic product which is generally unsuitable
for use as a jet fuel (aromatics conversion requires hydrogen for ring saturation
and at low hydrogen pressures, the aromatics conversion will be limited, so that the
product will oe more aromatic). The products from zeolite beta are therefore highly
satisfactory for use as jet and diesel fuels by reason of their highly paraffinic
nature and low pour point; besides being relatively rich is iso-paraffins. The paraffinic
character of the distillate also tends, to a certain extent, to dissolve any waxy
paraffins which are present, so contributing further to good low temperature mobility
characteristics. Thus, with the zeolite oeta hydrocracking catalysts, aromatics are
rejected in the bottoms (unconverted) fraction and a more highly paraffinic product
is obtained, either in cascade or two-stage operation.
[0043] The process may be operated in different modes so as to alter.the properties of the
products. In particular, the process may be operated in a jet fuel mode to produce
a highly paraffinic product with a notably low pour point; alternatively, it may be
operated in a diesel fuel mode to produce an excellent diesel fuel with a high Diesel
Index (Diesel Index is the product of the Aniline Point in degrees Fanrenheit (°F
= [°C.9/5] + 32) and API Gravity/ 100), typically about 50-65, as compared to about
35 for known hydrocracking processes; although small quantities of gas and naphtha
will be produced, the proportion of distillate range material will be enhanced relative
to conventional processes operating at higner pressures and with higher conversions
in multi-stage operations with interstage separation to remove ammonia.
[0044] The removal of sulfur in the higher boiling distillate oils is usually at least 90
percent complete so that these products will readily meet specifications for non-polluting
fuel oils. The naphtha which is produced is characterized, like the other products,
by a low heteroatom (sulfur and nitrogen) content and is an excellent feed for subsequent
naphtha processing units, especially reforming units because of its high cycloparaffin
content; the low heteroatom content enables it to be used in platinum reformers without
difficulty. Because the zeolite beta hydrocracking catalysts can be operated at higher
conversions under lower hydrogen pressures than conventional catalysts (without excessive
aging) they are capable of achieving the desired conversion levels at lower severities
than conventional catalysts (lower temperature, nigher space velocity) and this means
that the degree of desulfurization may be lower for a given conversion, resulting
in a reduced hydrogen consumption, so that a smaller gas plant is required. The products
can be made to meet applicable specification, for example 0.3 percent sulfur maximum,
but the excessive desulfurization which accompanied high conversions with conventional
zeolites under low pressures is avoided.
[0045] A particularly notable feature of the present process is that the bottoms product
(345°C+) fraction is dewaxed during the hydrocracking so that it, too, has a lower
pour point. Because of this, it has become possible to extend the distillate end point
so as to increase the distillate pool yield. In the past, distillates such as jet
fuel, diesel and heating oil have usually been held to a
' 345°C end point because the presence of waxy components in the higher boiling fractions
precluded their inclusion in the distillate pool as pour point specifications would
not have been met. With the present zeolite beta catalysts, however, the relatively
low pour point of the higher boiling fraction permits some of it to be included in
the distillate pool without exceeding pour point limitations. Generally, the distillate
end point may be extended to about 400 or 415°C. Although, as mentioned above, the
zeolite beta catalyst tends to reject aromatics to the high boiling fraction, the
presence of significantly higher amounts of paraffins in the 345
0C-(about 650°F-) fraction will reduce the final aromatics concentration in the product
if an extended end point is taken. Either the noble metal-containing catalysts, for
example Pt/beta, or base metal containing catalysts, for example NiW/beta, will permit
a higher product end point to be used for any given conversion than a conventional
amorphous catalyst which may be conversion limited because of aging (aging is rapid
with conventional catalysts at conversions over about 40 percent but zeolite beta
based catalysts are capable of running at 70% conversion even at pressures below 7,000
kPa). Thus, tne distillate pool is increased not only by the better yield and the
possibility of using higher conversions without excessive catalyst aging but also
by the prospect of raising the distillate end point by a significant extent, all with
a lower hydrogen consumption and with improved product quality'at any given level
of conversion.
[0046] As mentioned above, a number of improved product qualities may be noted for the distillate
fractions, including Pour Point (as measured by ASTM D-97), Diesel Index and Smoke
Point (ASTMO-1322). These improvements are obtained, moreover, with reduced hydrogen
consumption and higher conversion and distillate yield. Although the Aniline Points
(ASTM D-611) of the products may be only slightly higher than those obtained with
conventional catalysts, their API gravities are notably higher so that their Oiesel
Indices are markedly higher. Tne distillate products are therefore highly satisfactory
diesel fuels.
[0047] The ability to improve the Diesel Index is notable since conventional catalysts operating
at low to moderate pressures below
7,000 kPa tend to give a relatively poor quality product, typically of about 35 Diesel
Index, as compared to a specification minimum of 53. Increasing the endpoint of the
conventional product from about 345°C to a nominal 370°C would result in an improvement
to about 01=40. Further increases in end point would lead to additional improvements
in Diesel Index, typically to DI=42 but typical product pour point specifications
(-8°C) would preclude extension of the end point in this manner. Upgrading of the
product to the required specification would then require blending with straight-run
kerosene. Use of the zeolite beta based catalyst would, by contrast, result in a distillate
product of_significantly improved quality even at conversions of 35 to 40%. For a
full range distillate (165
0-345
0C) a Diesel Index of about 45 can be obtained and this can be increased to about 50
by raising the end point to a nominal 400°C, while the dewaxing function of the catalyst
reduces the Cloud Point (ASTM D-2500) to about -7°C.. In addition, the pour point
of the 400
oC+ fraction is reduced to about 4°C. These improvements are attained, moreover, with
no increases in net hydrogen consumption (about 80 n.l.l.
-1) relative to the process with the conventional catalyst. These improvements in product
quality and operating economics may be attributed to the preferential conversion of
paraffins over aromatics obtained with the zeolite beta catalysts.
[0048] The present process is particularly useful in the production of jet fuels which are
characterized by a low Pour Point and good Smoke points. This is particularly surprising
given the low hydrogen pressures at which the process operates and which, with conventional
catalysts, produces distillates of a pronounced aromatic character which cannot be
used as jet fuels. The present jet fuels have an aromatic content below 40 volume
percent, and usually below 30 volume percent, typically 20-30 volume percent. Smoke
Points (ASTMD-1322) are usually in the range 15 to 30. Particularly surprising is
that fuels of this kind are obtained at high conversions, for example over 40 or over
50 percent and at low specific hydrogen consumptions, for example not more than 100
or even 7
5 n.l.l.
-1.
[0049] This desirable result may be attributed to the aromatics rejection (to higher boiling
fractions, specifically above 345°C), resulting from the use of zeolite beta.
[0050] The advantages of the present process may be perceived, moreover, not only in terms
of product quality and yield and economy and stability of operation but also in overcoming
certain limitations which may inhere in existing hydrocracking units. In general terms,
these limits include:
1. Hydrogen consumption (the size of the gas plant-may preclude greater output),
2. Pressure limitation (equipment limits),
3. End Point limitation (limited by the Pour Point of the product which has to conform
to specifications)
4. Limits on distillation capacity
5. Conversion limits (set by catalyst aging and hydrogen pressure)
[0051] Because the hydrogen consumption of the process is lower than that with the use of
conventional catalysts, it may be possible to overcome the first limitation by using
the present process - without a loss of product quality. Also, because the zeolite
beta catalysts can operate at higher conversions under lower hydrogen pressures, e.g.
below 10,000 kPa, even below 7,000 kPa, without excessive aging, limitations (2) and
(5) can be circumvented. The End Point limitation can also be raised because, as mentioned
above, the higher boiling fraction is effectively dewaxed as well as the distillate
fraction. Thus, a number of the limitations inherent in the operation of existing
hydrocracking units may be overcome by the use of the zeolite beta-containing catalysts.
[0052] The following Examples illustrate the invention.
Examples 1 & 2
[0053] These Examples compare single stage hydrocracking (Ex. 1) with two stage, separated
hydrotreating/hydrocracking, that is with interstage separation (Ex. 2).
[0054] An Arab Light heavy vacuum gas oil (HVGO) (345°-580°C) was treated by the two processes,
using NiW/beta catalysts (NiW/beta with Al
2O
3 binder, 50:50 zeolite:binder). The composition of the HVGO is given in Table 1 oelow.
The conditions used and the results obtained are shown in Table 2 below.
Examples 3 - 5
[0056] Hydrocracking was carried out using the same HVGO feed as in Examples 1 and 2, but
with a Pt/beta catalyst. The operation modes were single stage (Ex. 3), and two stage,
separated (Ex. 4). For comparison, a cascade operation (no interstage separation)
is shown in Example 5, using a slightly different feed containing about 5 percent
345°C- products, but with the results normalized to a 345
0C+ basis. The conditions used and results obtained are shown in Table 3 below.
Example 6
[0057] An Arab Light HVGO having the properties shown in Table 4 below was evaluated in
a cascade mode hydrotreating/hydrocracking process using a commercial Ni-Mo/Al
20
3 hydrotreating catalyst in the first stage. A number of different catalysts were used
in the second half of the reactor, including Pt/beta, NiW/beta and two conventional,
commercially available hydrocracking catalysts (identified as A and B). The process
conditions used were 5960 kPa, 0.50 LHSV and 535 n.1.1 H2 (inlet) with the reactor
temperature adjusted between 400 and 430°C to obtain various conversions.
[0058]
[0059] The results are shown in Figure 1 which relates the pour point of the 345°C+ fraction
to the conversion attained, using the different hydrocracking catalysts. For comparison,
the Figure also shows the results obtained using interstage separation.
[0060] The figure shows that the hydrocracking catalysts containing zeolite beta as the
acidic component dewax unconverted gas oil. The Pt/beta catalyst has a greater selectivity
for paraffin isomerization resulting in low pour point products at moderate boiling
range conversions. Hydrogen consumption with the Pt/beta catalyst at 50% conversion
was about 85 n.l.l.
-1.
Example 7
[0061] The Arab Light HVGO II of Table 4 was subjected to cascade hydrotreating - hydrocracking
(no interstage separation) using a conventional hydrotreating catalyst and a Pt/beta
hydrocracking catalyst (0.6% Pt). The process was operated at about 5960 kPa pressure,
with conditions (temperature, space velocity) adjusted to give 46.5 weight percent
345°C+ conversion. The product qualities are set out below in Table 5.
Example 8
[0062] A hydrocracking operation was carried out using a NiW/beta hydrocracking catalyst
(4% Ni, 10% W, zeolite: alumina 50:50; steamed 24 hours in H-Na form 540°C, 90% steam
in air, atmospheric pressure). The feed was the Arab Light HVGO II of Table 4. The
stability of the operation together with selected product properties can be seen from
Table 6 below.
Example
[0063] Using the same feed and process conditions as in Example 8, the hydrocracking operation
was repeated, but using a Pt/beta hydrocracking catalyst (0.6% Pt; zeolite:alumina
50:50; steamed 72 hours in H-Na form, 540°C, 90% steam in air, atmospheric pressure).
The results are shown in Table 7 below.