A process for hydrodewaxing a petroleum or shale oil fraction, e.g., petroleum distillate, over a shape selective zeolite in the presence of hydrogen wherein the dewaxing is conducted in at least two stages, with some reheating of first effluent. Dewaxed feed, and a high octane gasoline byproduct are obtained as products.

FIELD OF THE INVENTION 
The invention relates to wax hydrocracking over shape selective zeolites. 
BACKGROUND OF THE INVENTION 
Hydrocarbn conversion processes utilizing crystalline zeolite catalysts 
have been the subject of extensive investigation during recent years as is 
clear from both the patent and scientific literature. Crystalline 
aluminosilicates have been found to be particularly effective for a wide 
variety of hydrocarbon conversion processes and have been described and 
claimed in many patents including U.S. Pat. Nos. 3,140,249; 3,140,252; 
3,140,251; 3,140,253; and 3,271,418. Aside from serving as general 
catalysts in hydrocarbon conversion processes, it is also known that the 
molecular sieve properties of zeolits can be utilized to preferentially 
convert one molecular species from a mixture of the same with other 
species. 
In a process of this type a zeolite molecular sieve is employed having 
catalytic activity within its internal pore structure and pore openings 
such that one component of a feed is capable of entering within the 
internal pore structure thereof and being converted to the substantial 
exclusion of another component which, because of its size, is incapable of 
entering within the pores of the zeolitic material. Shape selective 
catalytic conversion is also known in the art and is disclosed and claimed 
in U.S. Pat. Nos. 3,140,322; 3,379,640 and 3,395,094. 
Although a wide variety of zeolitic materials and particularly crystalline 
aluminosilicates have been successfully employed in various catalytic 
conversion processes, nevertheless, these prior art processes, in general, 
fell into one or two main categories. In one type of conversion process a 
zeolite was employed which had a pore size sufficiently large to admit the 
vast majority of components normally found in a charge, i.e., these 
materials are referred to as large pore size molecular sieves and they are 
generally stated to have a pore size of from 6 to 13 angstroms and are 
represented by zeolites X, Y and L. The other type of aluminosilicate was 
one which had a pore size of approximately 5 angstrom units and it was 
utilized to preferentially act upon normal paraffins to the substantial 
exclusion of other molecular species. Thus, by way of considerably 
over-simplification until recently, there were only two types of 
aluminosilicates which were available for hydrocarbon processing--those 
which would admit only normal paraffins and those which would admit all 
components normally present in a hydrocarbon feed charge. See U.S. Pat. 
No. 3,700,585 and Canadian Pat. No. 829,282. 
The cracking and/or hydrocracking of petroleum stocks is in general well 
known and widely practiced. It is known to use various zeolites to 
catalyze cracking and/or hydrocracking processes. 
Of particular recent interest has been the use of a novel class of 
catalysts to assist in the dewaxing of gas oils, lube base stocks, 
kerosenes and whole crudes, including syncrudes obtained from shale, tar 
sands and coal hydrogenation. U.S. Pat. No. 3,700,585 discloses the use of 
ZSM-5 zeolite to efficiently catalyze dewaxing of various petroleum 
feedstocks. 
U.S. Pat. No. 3,700,585 discloses and claims the cracking and hydrocracking 
of paraffinic materials from various hydrocarbon feedstocks by contacting 
such feedstock with a ZSM-5 zeolite at about 290.degree. to 712.degree. 
C., 0.5 to 200 LHSV and with a hydrogen atmosphere in some cases. This 
patent is based upon work on the dewaxing of gas oils, particularly virgin 
gas oils, and crudes although its disclosure and claims are applicable to 
the dewaxing of any mixture of straight chain, slightly branched chain and 
other configuration hydrocarbons. The catalyst may have a 
hydrogenation/dehydrogenation component incorporated therein. 
Other U.S. patents teaching dewaxing of various petroleum stocks are U.S. 
Pat. No. Re. 28,398; U.S. Pat. Nos. 3,852,189; 3,891,540; 3,894,933; 
3,894,938; 3,894,939; 3,926,782; 3,956,102; 3,968,024; 3,980,550; 
4,067,797 and 4,192,734. 
Catalytic hydrodewaxing can be considered to be a relatively mild, shape 
selective cracking or hydrocracking process. It is shape selective because 
of the inherent constraints of the catalyst pore size upon the molecular 
configurations which are converted. It is mild because the conversions of 
gas oil feed to lower boiling range products is limited, e.g., usually 
below about 35 percent and more usually below about 25 percent. It is 
operative over a wide temperature range but is usually carried out at 
relatively low temperatures, e.g. start of run temperatures of about 
270.degree. C. are usual. 
An advance in hydrodewaxing was disclosed in U.S. Pat. No. 4,446,007 
(Smith), which is incorporated herein by reference. Smith recognized that 
the dewaxing process could be a source of high octane byproduct gasoline, 
provided that the temperature was raised relatively rapidly to at least 
360.degree. C. Rapid temperature increase after startup meant that there 
was some over dewaxing of the chargestock, but this was not harmful, and 
indeed even increases the blending value of the heavy fuel produced. More 
significantly, the byproduct gasoline was both high octane, and relatively 
low in aromatic content. 
Smith observed that during operation, the mild drop in temperature 
associated with fresh hydrodewaxing catalyst rapidly diminished, and that 
hydrogen consumption could be reduced, and reactor delta Ts would approach 
zero, within about a month after startup. 
Shape selective catalytic hydrodewaxing such as practiced in U.S. Pat. No. 
4,446,007, to produce heavy fuel oil product is not usually considered 
endothermic or exothermic. Usually reactor temperatures at the outlet 
roughly equal the inlet temperature. Although the process is a catalytic 
hydrocracking process, some catalytic hydrodewaxing units create hydrogen 
rather consume it. They can create H.sub.2 because a long chain paraffin 
in cracked into two or more olefinic fragments. This makes H.sub.2. The 
olefins may or may not be saturated before they leave the hydrocracking 
reaction zone, and this saturation consumes hydrogen. 
To summarize, shape selective catalytic hydrodewaxing to produce fuels is 
an unusual hydrocracking process in that there is not much temperature 
change through the reactor, there is not much hydrogen consumption, and it 
is usually conducted in a single stage. "Single stage" means that dewaxing 
is customarily conducted in one large reactor, or in several reactors in 
series, with no intermediate heating, cooling, removal of impurities, etc. 
between reactor beds. This is in contrast to conventional hydrocracking 
processes, which usually operate in several stages, with one or more 
quench stages to prevent temperature runaway. 
We realized that catalytic dewaxing unit as proposed in U.S. Pat. No. 
4,446,007 (Smith), would give an optimum startup, but not necessarily an 
optimum operation thereafter. The rapid start-up procedure of Smith solved 
the problem of making the dewaxing unit an efficient generator of high 
octane gasoline during startup, but did not solve the problem of working 
the catalyst to the maximum extent possible or extending the run length. 
We discovered a way to make the dewaxing unit produce even more high 
octane gasoline, and/or last for a longer period of time, than had 
heretofore been thought possible. 
BRIEF SUMMARY OF THE INVENTION 
Accordingly the present invention provides a process for catalytic 
hydrodewaxing of a wax containing feed in a reactor by contacting said 
feed with hydrogen in the presence of a catalyst comprising a shape 
selective crystalline zeolite having a silica to alumina mole ratio of at 
least 12 at a reactor inlet temperature above 300.degree. C., a liquid 
hourly spaced velocity of about 0.2 to 10, a reactor pressure of about 100 
psig to 3000 psig and a hydrogen to hydrocarbon mole ratio greater than 
zero to about 20, the improvement which comprises conducting the process 
in at least two stages, with an inlet temperature to the first stage in 
excess of 360.degree. C. to produce a first stage partially dewaxed 
effluent at a reduced temperature relative to said first stage inlet 
temperature, heating the effluent from the first stage by at least 
5.degree. C., and charging the heated first stage effluent to the second 
stage to produce a dewaxed hydrocarbon product. 
In another embodiment the present invention provides a process for the 
selective cracking of wax in a heavy feed in a hydrogen containing 
atmosphere over a shape selective zeolite wax cracking catalyst at a 
temperature in excess of about 360.degree. C. to produce a dewaxed heavy 
feed and a gasoline boiling range product having a research clear octane 
number of at least 90, the improvement comprising hydrocracking the wax in 
at least a first stage reaction zone and at least a second stage reaction 
zone, and the first produces a first stage effluent which is heated and 
charged to the second stage reaction zone. 
In a more limited embodiment, the present invention provides in a process 
for catalytic hydrodewaxing of a wax containing feed in a reactor by 
contacting said feed with hydrogen in the presence of a catalyst 
comprising a shape selective crystalline zeolite having a silica to 
alumina mole ratio of at least 12 at a reactor inlet temperature above 
300.degree. C., a liquid hourly space velocity of 0.2 to 10, a reactor 
pressure of 100 psig to 3000 psig at least 1500 SCFB of hydrogen, the 
improvement comprising conducting the wax hydrocracking in at least a 
first and at least a second stage reaction zone, and wherein endothermic 
wax cracking reactions predominate in the first stage reacton zone which 
endothermic reactions cause a reduction in temperature across the first 
stage reaction zone of at least 10.degree. C. and wherein the temperature 
in the second stage zone is increased by the addition of 400-1500 SCFB 
H.sub.2 of hot hydrogen to the first zone effluent thereby increasing the 
temperature in said second stage reaction relative to the temperature of 
the first stage effluent.

DETAILED DESCRIPTION 
The process of out invention involves may aspects which are conventional 
(such as feedstock, dewaxing catalyst, etc.) and some aspects which are 
new to shape selective catalytic dewaxing (multistage operation, with heat 
added intermediate the stages). The conventional aspects will be briefly 
discussed, followed by a more detailed discussion of the multistage, 
reheating aspects of our invention. 
FEEDSTOCKS 
Any waxy material which has heretofore been processed in shape selective 
catalytic dewaxing processes can be used. This includes gas oils, lube 
stocks, kerosenes, whole crudes, synthetic crudes, tar sand oils, shale 
oils, etc. These heavy feeds may be subjected to one or more conventional 
pretreatment steps, such as hydrotreating, to remove excessive amounts of 
nitrogen impurities, metals, etc. 
The preferred chargestocks are gas oils and vacuum gas oils derived from 
paraffinic crudes. Gas oils contemplated for use herein will have boiling 
ranges of 350.degree.-850.degree. F., while vacuum gas oils typically have 
boiling ranges of 500.degree.-900.degree. F. 
Pour points are typically 75.degree.-100.degree. F., or more, frequently, 
85.degree.-90.degree. F., with cloud points perhaps 5.degree. F. above the 
pour point. 
The feed preferably is slightly heavier, re end point, than the 
specification end point of the desired product. This is somewhat heavier 
than the conventional feed (usually an atmospheric gas oil) to shape 
selective catalytic dewaxing units making fuel oil products. Some light 
vacuum gas oil, or material boiling in this range, is preferably present 
in the feed. 
The dewaxing process can convert some feeds boiling beyond the diesel or 
No. 2 fuel oil boiling range into materials boiling within the desired 
range. The dewaxing process used herein is not an efficient converter of 
heavy feeds to lighter feeds, and will leave some fractions of the feed 
(primarily the aromatic and naphthenic fractions) relatively untouched, so 
although these non-paraffinic materials can be tolerated in the feed, they 
are not efficiently converted by the shape selective zeolite catalyst. A 
relatively heavy feed, with product end point specs satisfied by 
downstream fractionation, maximizes production of more valuable light 
products from less valuable heavy feed. 
CATALYST 
Any conventional shape selective zeolite which can be used to selectively 
crack normal paraffins in a heavy hydrocarbon stream can be used herein. 
More details on suitable zeolites, and their properties are disclosed in 
U.S. Pat. No. 4,446,007. 
As disclosed in U.S. Pat. No. 4,446,007, the preferred zeolites have a 
Constraint Index of 1-12. 
Of the zeolite materials useful in the present process, zeolites ZSM-5, 
ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38 and ZSM-48 are noted. Zeolite ZSM-5 
is preferred. ZSM-5 is described in U.S. Pat. No. No. 3,702,886 and U.S. 
Pat. No. Re 29,948, each being incorporated by reference. ZSM-11 is 
described in U.S. Pat. No. 3,709,979, which is incorporated by reference. 
ZSM-12 is described in U.S. Pat. No. 3,832,449, which is incorporated by 
reference. ZSM-23 is described in U.S. Pat. No. 4,076,842, which is 
incorporated herein by reference. U.S. Pat. Nos. 4,016,245 and 4,046,859, 
describing ZSM-35 and ZSM-38, respectively, are incorporated herein by 
reference. 
CATALYTIC DEWAXING CONDITIONS 
The conditions in each stage of the catalytic dewaxing reaction zone are 
broadly within those conditions heretofore found suitable for shape 
selective catalytic hydrodewaxing. More details of preferred conditions 
are recited in U.S. Pat. No. 4,446,007 (Smith). 
Briefly, the shape selective catalytic dewaxing occurs at temperature from 
316.degree.-454.degree. C. (600.degree.-850.degree. F.), at LHSVs ranging 
from 0.1-10. Preferred conditions include temperature of at least 
360.degree. C. 
Pressures are usually mild, typically on the order of prior art 
hydrotreating processes ranging around 100-1000 psig. Operation with 400 
pounds of hydrogen partial pressure gives good results. 
OVERALL DEWAXING SEVERITY 
In the process of our invention, higher conversions of feed to lighter 
products are obtained, relative to what was the norm for prior art shape 
selective catalytic hydrodewaxing processes. Operation at high 
temperatures suggested in U.S. Pat. No. 4,446,007 (Smith) converts 
catalytic dewaxing (CDW) from a process which merely dewaxes and produces 
some low octane gasoline boiling range byproduct, to a process which not 
only dewaxes but also efficiently generates high octane gasoline which can 
be added directly to the refinery blending pool. Even more desirably, the 
high octane gasoline by-product produced by catalytic dewaxing is 
relatively low in aromatics. 
Accordingly, we prefer to operate with somewhat higher temperatures, and 
conversions than was heretofore considered the norm. Conversions of at 
least 15 wt. % of the feed to products boiling in the gasoline and lighter 
range should be achieved, and most preferably conversions of 20-45% of the 
feed to lighter products should be achieved. Depending on paraffin content 
of the chargestock, conversions of 30-50 wt. % of the feed, or more, to 
lighter products are contemplated herein. Gasoline yields of 20-25 wt. % 
of the feed may be achieved. 
Expressed as wt. % conversions of waxy paraffins in the feed, it is 
preferred to crack most of the paraffins, more preferably, 75 wt. % of the 
normal and slightly branched paraffins, with 90.sup.+ % conversions being 
possible in some cases. 
Expressed as gasoline octane, the overall severity should be enough to 
produce a gasolie boiling range fraction having an octane number (Research 
Clear) of 90 or higher, preferably above 91, and most preferably above 92. 
The average reactor temperature (weight average bed temperature) will be 
somewhat higher in our process as compared to the prior art, although the 
average inlet temperature to the first reactor will not change so much. 
This is because the primary effect of out invention is higher temperatures 
in the second stage, rather than higher temperatures in the first stage. 
MULTISTAGE-SEVERITY CONTROL 
Our invention requires that the dewaxing process be operated in at least 
two stages, with some control of severity in each stage. We can shift 
H.sub.2 addition or adjust severity, or preferably, do both. We will first 
review one stage operation (prior art, U.S. Pat. No. 4,446,007), then 
address our two stage process. 
Some prior art catalytic dewaxing processes, when operated as efficient 
generators of high octane gasoline, had a significant endothermic reaction 
in the first 25% or so of the catalyst bed, which was largely offset by 
exothermic reactions occurring in the latter portions of the bed. After 
the unit lined out (within 30 days) an observer looking at the process 
would see little or no temperature change between feed and products. Very 
little hyrogen consumption was shown in, FIG. 3 of U.S. Pat. No. 4,446,007 
(Smith). A casual observer who was familiar with shape selective catalytic 
dewaxing would not expect any significant temperature change, because 
although the cracking of wax is endothermic, the hydrogenation of the 
resulting olefinic fragments would be an offsetting, exothermic reaction. 
We then observed the operation of a commercial dewaxing unit which happened 
to be conducted in two reactors. Although the reaction was conducted in 
two separate reactor vessels, they functioned as a single stage because 
there was no intermediate heating or cooling, nor addition or removal of 
anything in between reactor 1 and reactor 2. Two reactors held the 
catalyst, but it was a one stage process. By analyzing the temperature of 
the feed, and the reactor 1 and 2 outlets we realized that there was a 
significant drop in temperature in the first reactor, and temperature 
would increase slightly, if at all, in the second reactor. Sometimes there 
would be a slight further decrease in temperature (3.degree. to 7.degree. 
F.) in the second reactor as well. 
We wanted to have a measure of control over the reaction severity in 
reactor 1 and 2. Our belief was that the existing process was not being 
operated as efficiently as possible, because the catalyst temperature was 
reduced excessively through much of the second reactor, and probably in 
some of the first reactor too. 
We realized that by adding more hydrogen to the reactor 1 effluent, and 
adding the H.sub.2 hot, it would be possible to heat the material going to 
the second reactor and make the second reactor work harder. 
Based on our commercial operating experience, we realized that the inlet 
temperature of the second stage should be increased preferably by at least 
5.degree. F. and, if possible, increased to within 15.degree. C. of the 
reactor 1 inlet, and preferably within 10.degree. C., and most preferably 
have a temperature approaching that of the first reactor inlet. 
It would be possible to conduct the process in 3, 4, or even more reactors, 
but in most instances the benefits to be gained by such an elaborate 
reactor design will not be offset by the cost associated with the 
elaborate design. 
When a two reactor system is used, the first reactor should contain 25-70% 
of the total inventory of shape selective zeolite catalysts, while the 
second reactor should contain 30-75%. 
Based on our experimental work with one commercial unit (with about a 
1/3-2/3 split between catalyst in the first and second reactors) we 
believe that in future installations the optimum catalyst loading will be 
about the same. 
For a three reactor system, the first reactor could contain 10-40% of the 
total catalyst inventory while the second reactor could contain 20-40%, 
with the remainder being in the third reactor. 
Heat can be added in many ways to the second reactor. The easiest method 
for a retrofit is addition of a hot hydrogen stream. Any other 
conventional means of getting heat into the second stage can be used, 
e.g., indirect heat exchange, addition of some hot material which is not 
harmful to the process, or passing the first reactor effluent through a 
fired heater. 
DETAILED DESCRIPTION OF FIG. 1 
The present invention can be better understood with reference to the 
attached figures. FIG. 1 shows a considerably simplified process flow 
diagram of one embodiment of the invention, while FIG. 2 shows average 
reactor temperatures versus days on stream during several commercial tests 
of a dewaxing unit. 
In FIG. 1, a combined heavy feed, comprising a Heavy Atmospheric Gas Oil 
(HAGO), Light Vacuum Gas Oil (LVGO) and FCC Intermediate Cycle Oil (ICO) 
are added via line 4, mixed with makeup H.sub.2 rich gas in line 2, 
recycle H.sub.2 rich gas in line 22 and passed through heat exchanger 20 
and line 5 into heater 6. The heated feed is charged via line 8 into the 
first stage reactor 10. The first stage effluent is removed via line 12. 
The first reactor effluent, within a month after startup, usually is at 
least 10.degree. C. cooler than the feed in line 8. There is a drop in 
temperature because of the endothermic wax cracking reactions occurring in 
first stage reactor 10. First stage effluent is heated, by adding hot 
hydrogen from line 13. The resulting mixture is passed into second stage 
reactor 15. The dewaxed heavy feed, cracked products and H.sub.2 are 
removed via line 19, passed through heat exchanger 20 and discharged via 
line 21 into high pressure separator 25. 
Actually, the figure is quite a simplification of the actual refinery 
operation. All of the feed, and much of the recycle and make-up H.sub.2 
was charged to a heater and then passed to reactor 1. The heater 6 could 
not handle the load, so we used a separate heater 16, which happened to be 
available. This second heater was used to heat a portion of the recycle 
and make-up H.sub.2. The material from both heaters entered the first 
reactor 10. The first reactor functioned just as it would have with a 
conventional single heater for the combined stream of oil and recycle and 
make-up gas. Our unusual heater system (with two heaters instead of the 
more normal single heater) was then further modified, by the addition of a 
line equivalent to line 13 shown in FIG. 1, to allow us to add hot H.sub.2 
directly to the reactor 1 effluent, and thereby heat the feed to the 
second reactor. 
Our actual point of make-up H.sub.2 addition was also somewhat different 
than that shown in FIG. 1, but the difference is not important to the 
process. The make-up H.sub.2 purity from a Pt reformer varies from day to 
day because of normal fluctuations of the reformer. The shape selective 
catalytic dewaxing unit usually cleans up the hydrogen which is a little 
unusual for a hydrogen consuming unit. Thus there is very little 
difference in hydrogen purity of the make-up H.sub.2 to the dewaxing unit 
and the H.sub.2 purity of the dewaxing unit recycle gas, so the point of 
addition of make-up H.sub.2 is not critical. In our plant, the make-up 
H.sub.2 was added to the recycle H.sub.2. Some of the combined H.sub.2 
stream (recycle+make-up H.sub.2) was mixed with the oil feed, and the 
remainder of the combined H.sub.2 was sent to the second heater 16, i.e., 
to the heater with no oil feed. 
High pressure separator 25 operates at a temperature of 
60.degree.-130.degree. F. and pressure of about 525 psig. A hydrogen rich 
gas stream is withdrawn via line 24 and removed as a fuel gas by-product 
in line 71, recycled to mix with fresh feed via line 22 or sent via line 
23 to heater 16 to produce the hot hydrogen rich gas in line 13. 
Liquid is removed from high pressure separator 25 via line 28 and 
discharged into low pressure separator 30, operating at a temperature of 
60.degree.-130.degree. F. and a pressure of 175-180 psig. A fuel gas 
stream is removed via line 29. Flashed liquid is removed via line 31 and 
charged to stabilizer or debutanizer 35. C.sub.4 and lighter hydrocarbons 
are removed overhead via line 39, cooled in cooling means not shown, and 
charged to overhead accumulator 40. The figure is also somewhat simplified 
re this and other distiallation columns, i.e., reflux lines, coolers 
associated with column overhead vapor lines, pumps, etc. have been omitted 
for clarity. A fuel gas stream is removed via line 41 while a C.sub.3 
/C.sub.4 rich liquid is discharged via line 72, for further processing in 
the FCC depropanizer. 
Stabilizer 35 is reboiled using conventional reboiler 36. The net bottoms 
products is removed via line 37, passed through heater 46 and discharged 
via line 44 into splitter column 45. Gasoline boiling range hydrocarbons 
are removed overhead via line 47 and discharged into overhead accumulator 
55. Gasoline boiling range hydrocarbons are removed via line 73 as a 
product. 
An intermediate boiling range stream is removed from column 45 via line 49 
and charged to steam side stripper 50. Light materials are discharged 
overhead via line 52 and sent back to the main column 45, while a diesel 
fraction is removed via line 74 as a product. 
A bottoms product is withdrawn via line 59 from column 45 and charged to 
vacuum flash 60. An overhead vapor stream is removed via line 63 and 
charged to overhead accumulator 55 for recovery of gasoline boiling range 
components. An intermediate boiling range stream is withdrawn via line 62 
and charged to steam side stripper 50, while a vacuum gas oil fraction is 
withdrawn via line 75. 
EXAMPLE 
The invention was tested in a commercial dewaxing unit. As is common in all 
operating commercial units, the unit was being run to make a product, not 
to generate data. There are always changes in operation, and problems so 
there is quite a scatter in the data generated by a commercial plant. The 
commercial test occurred at a refinery which runs heavy paraffinic crudes, 
with attendant distillate fluidity problems. 
The refinery chose shape selective catalytic dewaxing as the most cost 
effective way of eliminating distillate cold flow problems and improving 
plant profitability. The refinery had an idle high pressure hydrotreating 
unit which was built in 1972 to pretreat 17,000 BPSD of heavy FCC naphtha 
prior to reforming. For a number of reasons, the unit was mothballed. Thus 
unit contained most of the equipment required by the CDW process except 
for the addition of one major vessel. Changes were made to the piping and 
reactor internals and the unit pressure was dropped to 525 psig. FIG. 1 is 
a schematic of the revamped unit. 
A typical feedstock is reported below. 
TABLE 1 
______________________________________ 
TYPICAL CDW FEEDSTOCK 
______________________________________ 
API Gravity 29.4 
Pour Point, .degree.F. 
85 
Aniline Point, .degree.F. 
180 
K Factor 11.8 
Distillation, .degree.F. (D-1160) 
10% 631 
30% 672 
50% 699 
70% 733 
90% 800 
______________________________________ 
Although the operation of the CDW reactor section is similar to a 
hydrodesulfurizer (HDS), that is, oil and hydrogen are passed over a fixed 
bed of catalyst, the disposition of the products and by-products is 
different. The unsaturated light liquid hydrocarbons from the stabilizer 
are sent to the FCC gas plant for further recovery. The butenes become 
alkylation feed. Propenes are polymerized. The CDW naphtha is sent 
directly to gasoline blending. The distillate product is blended directly 
to diesel fuel, and the bottoms are recycled to the FCC unit. 
Direct blending of CDW naphtha into the gasoline pool is possible because 
of its high octane number (typically 92 RONC) and low mercaptan level. 
Table 2 lists the properties of this stream. Caustic and water washing 
equipment were added to the unit to handle high mercaptan levels, but the 
low sulfur crudes run to date have made it unnecessary to use the 
facilities. If high sulfur crudes are processed, this equipment will have 
to be activated. 
TABLE 2 
______________________________________ 
TYPICAL CDW GASOLINE PRODUCT PROPERTIES.sup.(1) 
______________________________________ 
Properties 
API Gravity 70 
Sulfur, wt. % 0.01 
Mercaptans, ppmw 35 
Bromine Number 140 
Octanes 
RON-Clear 92 
MON-Clear 79 
Composition,.sup.(2) Vol. % 
Paraffins 32 
Olefins 60 
Naphthenes 5 
Aromatics 3 
Distillation, .degree.F. (D-86) 
IBP 95 
5 Vol % 115 
10 Vol % 130 
30 Vol % 155 
50 Vol % 180 
70 Vol % 220 
90 Vol % 275 
95 Vol % 300 
EP 325 
______________________________________ 
.sup.(1) Derived from waxy crudes 
.sup.(2) Debutanized sample 
The CDW diesel oil is a blend of slide draws from the splitter and the 
vacuum flash unit. The target pour point is typically minus 10.degree. F., 
but it is adjusted to meet pool fluidity requirements. The low pour point 
CDW product is blended with FCC light cycle oil and virgin distillates to 
meet No. 2 and diesel fuel specifications. Properties of these three 
blending stocks are shown on Table 3. 
TABLE 3 
______________________________________ 
PROPERTIES OF DIESEL BLENDING COMPONENTS.sup.(1) 
CDW FCC 
Distillate 
LAGO LOO 
______________________________________ 
Properties 
API Gravity 28.0 35.7 28.7 
Sulfur, wt % 0.12 0.08 0.10 
Nitrogen, ppmw 250 68 140 
Basic Nitrogen, ppmw 
102 47 72 
Kinematic Vis. @ 40.degree. C., cs 
6.92 3.16 2.79 
Kinematic Vis. @ 100.degree. C., cs 
2.04 1.26 1.13 
Flash Point (COC), .degree.F. 
150 194 150 
Carbon Residue, wt % 
0.01 0.03 0.01 
Aniline Point, .degree.F. 
150 156 125 
Bromine Number 9.0 0.9 12 
Cetane Index 48 54 40 
Fluidity 
Pour Point, .degree.F. 
-10 10 10 
Cloud Point, .degree.F. 
0 14 18 
Distillation, .degree.F. (D-86) 
IBP 348 359 396 
5 Vol % Distilled 
435 425 410 
10 Vol % Distilled 
479 449 472 
30 Vol % Distilled 
587 503 472 
50 Vol % Distilled 
618 535 512 
70 Vol % Distilled 
641 564 558 
90 Vol % Distilled 
671 602 614 
95 Vol % Distilled 
685 625 635 
EP 701 650 659 
______________________________________ 
.sup.(1) Derived from waxy crudes 
CDW Catalyst Aging 
As shown in FIG. 2, the catalyst has an initial high aging rate, but then 
it lines out to provide a long cycle. The temperature variations on FIG. 2 
are due to the many shifts in crude quality that the refinery experiences. 
Variations due to throughput (space velocity) and product pour point have 
been accounted for by normalizing the data to a pour point of minus 
10.degree. F. and a design throughput of 17,000 BPDS. FCC variations as 
well as crude shifts have not been accounted for in normalizing the data. 
When the CDW unit was streamed some downstream fractionation equipment was 
not ready for operation. These were commissioned later in the first cycle. 
The run length for this cycle exceeded six months. The second cycle was 
cut short due to a scheduled crude unit turn around. 
A hot hydrogen reheat line was added before the start of the third cycle. 
There was also an improvement in virgin feed quality, because of the crude 
unit modifications. With hot hydrogen reheat, and better feed, the third 
cycle length was increased to 264 days on stream. 
FIG. 2 thus shows the reduced catalyst aging rates achieved through the 
process of the present invention. Cycle 1 and cycle 2 represent prior art 
dewaxing processes, i.e., with no reheating of the first stage effluent 
from reactor 10. Cycle 3 represents the present invention, namely adding 
about 1100-1200 SCFB of hot, H.sub.2 rich gas to increase the inlet 
temperature about 7.degree. to 35.degree. F., depending on charge rate, to 
the second stage dewaxing reactor 15. 
Although the data show quite a lot of scatter, as is to be expected in 
commercial units, the data nonethless show the much longer cycle lengths 
(reduced catalyst aging rates) can be achieved by conducting the dewaxing 
process in at least two stages, with hot hydrogen addition to the second 
stage.