Note: Descriptions are shown in the official language in which they were submitted.
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HYDRODEWAXING METHOD
This invention relates to a method for
hydrodewaxing a hydrocarbon feed, such as a middle
distillate or a lubricant hydrocarbon fraction.
The dewaxing of hydrocarbons to produce liquid
products of lower pour point is a process of great
commercial significance. Although alternatives exist,
the use of shape-selective catalysts, such as the
intermediate pore size æeolite catalysts, and in
particular, ZSM-5, to selectively convert those
paraffins that contribute the most to high pour point
has many advantages over other methods.
Catalytic dewaxing over intermediate pore size
zeolites and in particular, over ZSM-5, is a known
process and is described in, for example, U.S. Patent
No. Reissue 28,398. Such a dewaxing process produces
significant quantities of olefins, including C3 and C4
olefins as well as C5+ olefins in the gasoline boiling
range, as a result of which a relatively high octane
olefinic naphtha is one of the by-products of the
process.
Catalytic dewaxing is typically effected in the
presence of hydrogen to retard catalyst aging.
However, the conditions employed are not conducive to
olefin saturation so that significant quantities of
gasoline range and lighter olefins are produced in the
process. Moreover, in the dewaxing of lube boiling
range hydrocarbons, the effluent from the dewaxing
reactor may be cascaded directly into a hydrotreating
reactor in order to saturate and stabilize lube range
olefins in the dewaxing products. Similarly, in
distillate dewaxing, a hydrotreater may be provided in
order to remove unsaturation unless the dewaxed
distillate is combined with virgin distillate and the
combined distillate treated in the refinery CHD
(catalytic hydrodesulfurization) unit. In both of
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these cases, the hydrogen consumption of the
hydrotreating step is needlessly increased by the
saturation of olefinic components outside the boiling
range of the desired products, principally of C5- and
gasoline range olefins. To reduce hydrogen consumption
it would be possible to arrange for separation between
the dewaxing reactor and the hydrotreater but this may
still leave lower olefins to be carried out over into
the hydrotreater.
It has now been found that improved separation of
the lower olefinic materials may be provided by
stripping the dewaxed products prior to hydrotreating,
preferably using an oil solvent such as naphtha which
is fed into the top of a stripper/separator so that the
recycle gas is essentially free of wet gas and heavier
fractions. Operation in this manner confers several
benefits. One is a significant reduction in the
hydrogen consumption in the hydrotreating reactor since
the majority of the light olefins are no longer
hydrotreated in the hydrotreating reactor. In
addition, the hydrogen circulation rate for the
hydrotreating reactor can be better controlled so as to
reduce the pressure drop across the hydrotreating
reactor.
Alternatively, the hydrotreating reactor can be
operated at a higher pressure level than the
hydrodewaxing reactor; this affords the opportunity to
modify the pour point of dewaxed lube products.
Accordingly, the invention resides in a process
for dewaxing a hydrocarbon feed comprising:
1. dewaxing the feed by contacting the feed in
the presence of hydrogen with a dewaxing catalyst;
2. obtaining a dewaxed effluent;
3. stripping the dewaxed effluent with a gaseous
stripping medium to form a dewaxed fraction, a
light olefinic gas fraction and an olefinic
gasoline fraction;
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4. separating the dewaxed fraction from the olefinic
gasoline fraction and the light olefinic gas
fraction, and;
5. hydrotreating the dewaxed fraction in the presence
of hydrogen with a hydrotreating catalyst.
The separation of the light hydrodewaxing products
may be enhanced by stripping with effluent vapor from
the hydrotreating reactor, hydrogen make-up and/or
hydrogen recycle. Overall hydrogen make-up
requirements are reduced in that the light
hydrodewaxing products including the light C5- olefins
and olefinic gasoline, are not subject to hydrotreating
in the hydrotreating reactor. Recoveries up to about
90% of the C5+ gasoline, as well as up to 20% C3 and up
to 45% C4 upstream of the hydrotreating reactor can be
attained by the present invention.
The lower boiling components of the dewaxer
effluent can be further treated to upgrade the naphtha
fraction to high octane olefinic naphtha having an
octane number of 90RON + 0. The resulting olefinic
gasoline can be directly blended into the existing
gasoline pool. The recovered C3/C4 olefins can be
utilized as alkylation unit feed, or as a feedstock to
an oligomerization unit such as that described in U.S.
Patent No. 4,695,364.
The present olefin separation/stripping technique
may be employed either with lube boiling range feeds or
to produce middle distillate fuels, e.g., kerosene and
jet fuel. Thus, the feed may be a neutral or residual
lube feed, e.g., light neutral, heavy neutral or bright
stock, or an atmospheric or vacuum gas oil feed for
distillate fuel production.
The invention will now be more particularly
described with reference to the accompanying drawings,
in which:
Figure 1 is a schematic representation of a
conventional hydrodewaxing process.
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Figure 2 is a schematic representation of a
hydrodewaxing process according to a first example of
the invention.
Figure 3 is a schematic representation of a
hydrodewaxing process according to a second example of
the invention.
Figure 4 is a schematic representation of a
hydrodewaxing process according to a third example.
Figure 5 is a schematic representation of a
fractionator for achieving a higher olefinic C3/C4
recovery.
Referring to the drawings, Figure 1 shows a
simplified schematic of a conventional lube
hydrodewaxing process in which a hydrocarbon fraction
such as a lube range raffinate feed is fed via conduit
lO through heat exchanger 12 to a hydrodewaxing reactor
14 containing a ZSM-5 catalyst. A source of hydrogen
(not shown) is fed through conduit 16, to the
hydrodewaxing reactor 14 with make-up hydrogen supplied
through line 18. The dewaxed reactor effluent exits
the reactor 14 through conduit 20 and passes through
heat exchanger 22 and conduit 26 into hydrotreating
reactor 24. The hydrotreater reactor effluent exits
through conduit 28 and then passes through heat
exchanger 30 and conduit 32 to separator 34. The
bottoms from separator 34 are fed by conduit 36 to
naphtha stripper 38 which operates in a conventional
manner with steam introduced through conduit 40 into
the lower portion of the stripper to separate naphtha
and light gases from the lube oil and kerosene which
leave the stripper 38 as the bottoms fraction through
conduit 56. The overhead from the stripper 38 is fed
through conduit 44 to a heat exchanger 46 and then
through conduit 49 to a settling tank 50. Unstabilized
naphtha is withdrawn from the settling tank 50 through
conduit 52 while lighter gases are withdrawn through
conduit 54 for use as fuel or as a feed to an
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oligomerization unit as noted above. The bottoms
effluent from stripper 38 comprising lube oil and
kerosene, is conveyed in conduit 56, to downstream
processing units, such as a vacuum stripper (not
shown).
The lighter effluent leaving separator 34 through
conduit 56 passes through heat exchanger 60 and conduit
62 to separator 64. The bottoms from separator 64 are
fed through conduit 66 to an intermediate level in the
naphtha stripper 38. Contaminants such as hydrogen
sulfide plus nitrogen are removed from the effluent
leaving separator 64 through conduit 68 in scrubber 70.
The hydrogen leaving scrubber 70 via line 72 is partly
recycled through conduit 74 and recycle compressor 75
to hydrodewaxing reactor 14 and partly removed through
conduit 76.
Figure 2 illustrates a dewaxing process similar to
that shown in Figure 1 but with intermediate separation
and stripping, according to a first example of the
present invention. The same references as those in
Figure 1 are used to indicate like elements in Figure 2
and the remaining drawings. The intermediate
separation and stripping section of the first example
is indicated by the dotted line DL of Figure 2 and
comprises an olefin fractionator 78 to which is fed the
effluent from hydrodewaxing reactor 14. A source of
make-up hydrogen is introduced through inlet 80 to
fractionator 78 to strip the lighter olefinic dewaxing
products from the hydrodewaxer effluent. These olefins
leave the fractionator 78 as overhead and are conveyed
through outlet 82, heat exchanger 84, and conduit 88 to
a gas/liquid phase separator 86. The bottoms from the
separator 86, comprising Unstabilized olefinic naphtha,
are removed through conduit 90 and partially withdrawn
through conduit 92. Fractionator reflux is provided by
pump 93 and conduit 94 which return part of the
olefinic naphtha to olefinic fractionator 78. Light
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gases from separator 86 are removed via conduit 96 and
are partially fed through conduit 98 and heat exchanger
99 to provide a source of feed hydrotreating reactor
24. The remainder of the light gases pass through
conduit 100, flow control valve 102, and conduit 104 to
a scrubber 70 so as to provide a source of high purity
recycle hydrogen to the hydrodewaxing reactor 14,
conveyed through conduits 74 and 18.
Variable flow control valve 106 is used to control
the flow of dewaxed effluent through heat exchanger 22
to fractionator 78. For increased fractionation
efficiency, heat exchanger 22 may be by-passed by
controller 106 so that the flow may be sent directly to
the fractionator 78, preferably to a tray below the
mainfeed tray where stream 26 from heat exchanger 22 is
fed to the fractionator. Variable flow control valve
106 is actuated in response to a signal from
temperature sensor 108 on line 118 via communication
line 110. An additional flow control valve 112 is
connected by communication line 116 to sensor 114 at
the bottom of fractionator 78 to control the flow of
the stripped, dewaxed effluent through conduit 118 form
fractionator 78 to hydrotreater 24.
As can be appreciated from the process
schematically illustrated in Figure 2 the effluent of
hydrodewaxing reactor 14 is stripped in fractionator 78
to remove olefins and the light product leaving
fractionator 78 through outlet 82 primarily comprised
C3/C4 olefins and C5+ gasoline. The light products
recovered from the fractionator overhead can be sent to
the unsaturated gas plant of an FCC unit for further
separation onto C3, C4 and C5+ components. The removal
of the olefinic fraction also permits better control of
the hydrogen circulation rate for the hydrotreating
reactor by feeding hydrogen-containing overhead streams
through conduit 98 to hydrotreater 24 thereby
permitting a reduction in the pressure in the
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hydrotreating reactor and its associated separation
system pressure drop.
Unrecovered olefinic C3/C4 components plus ethene
leaving the fractionator section through conduit 104
are recycled with the hydrogen, through conduit 18, to
hydrodewaxing reactor 14 and may undergo conversion to
olefinic gasoline over the ZSM-5 catalyst used in
hydrodewaxing reactor 14 by olefin-to-gasoline
oligomerization. This is not possible in the
conventional process illustrated in Figure 1 since the
entire dewaxer effluent is processed in the
hydrotreating reactor and this converts all the olefins
in the effluent to paraffins.
Additional advantages may be achieved by the
alternative embodiment in Figure 3, in which the dotted
line DL2 illustrates separation and stripping section
of this second example. In Figure 3, the hydrotreating
reactor 24 can be operated at a pressure higher than
that of the hydrodewaxing reactor 14 by providing a
pump 120 to increase the pressure of the stripped
hydrodewaxing effluent leaving fractionator 78 through
variable flow control valve 112 and conduit 118. High
purity hydrogen for the hydrotreater 24 is provided at
an appropriate pressure from recycle compressor via
conduit 128. Moreover, effluent vapor from the
hydrotreater 24 after passage through the separators 34
and 64 is recycled to the hydrodewaxing reactor 14
via conduit 68.
In the third example shown in Figure 4, the
hydrotreating reactor effluent vapor is used as a
stripping medium in the fractionator 78 and is fed from
separator 34 through conduit 122 to the bottom of
fractionator 78. This is especially beneficial if
severe stripping of the hydrodewaxing reactor effluent
is desired in olefinic fractionator 78 since the
quantity of effluent vapor from separator 34 will
usually exceed the quantity of make-up hydrogen.
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As noted, with regard to the embodiments described
in Figures 2-4, the light overhead fractions (C3, C4)
from the fractionator are mostly carried into the
gaseous phase. Higher C3/C4 recoveries can be obtained
by employing the fractionator illustrated in Figure 5.
Thus, referring to Figure 5, the fractionator 278 has
input conduits 222 for introducing the effluent from a
hydrodewaxing reactor (not shown) as well as a
stripping medium conduit 280 for introducing a
stripping gas, such as hydrogen or hydrotreating
reactant effluent vapor. A lean oil is also introduced
through conduit 282 to the top of fractionator 278 to
dissolve the C3/C4 hydrocarbons and the unstabilized
olefinic naphtha plus rich oil is withdrawn from
fractionator 278 through conduit 284. A portion of the
unstabilized olefinic naphtha plus rich oil may be
returned as reflux through pump 286, cooler 290 and
conduit 292, while the remainder is withdrawn via
conduit 294 to further downstream treatment sections
(not shown). The bottoms from fractionator 278
comprise an effluent including lube oil and kerosene
which leave through conduit 296 while the top of olefin
fractionator 278 produces an overhead comprising treat
gas withdrawn through conduit 298. The hydrogen purity
of the treat gas is significantly improved as compared
with the conventional fractionators.
It is possible to upgrade about 90% (based on
Figure 2 configuration) of the hydrodewaxing reactor
effluent naphtha to high octane olefinic naphtha with
an octane number of 90 (R+O). This is an improvement
of approximately 15 to 20 octane numbers over that
obtainable by hydrotreating the naphtha. The olefinic
gasoline so produced can be directly blended into the
gasoline pool after treating which reduces the load on
the hydrotreater, which permits processing more
hydrocarbons per catalyst (weight basis).
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The present invention permits the recovery of
approximately 20% of olefinic C3 and up to about 45~ of
the C4 produced in the hydrodewaxing reactor (based on
Figure 2 configuration). By removing these olefins
from the dewaxing effluent before introduction into the
hydrotreating reactor, hydrogen consumption is
significantly reduced because hydrogen is not consumed
in saturating the olefins in the hydrotreater 24.
Significant advantage also results from the fact
that by-products of the present invention include
recycle gas of higher hydrogen purity than previously
obtainable. In addition, hydrogen circulation rate and
overall operating pressures may be selected so that
hydrotreating may be accomplished at a higher pressure
than the hydrodewaxing step. Moreover, stripping steam
requirements and stripper size for the naphtha stripper
can be significantly reduced as a result of the removal
of the olefinic fraction at an earlier stage in the
unit.
By effecting olefin separation between the dewaxer
and the hydrotreater the naphtha stripper diameter can
be reduced although the treat gas compressor size will
need to be increased to handle the increased gas volume
at this stage. The olefins recovery does not affect
the lube recovery or lube quality. All the lube
components leave the olefin fractionator as bottoms and
are mixed with the required volume of hydrogen before
entering the hydrotreating reactor. The naphtha end
point can be controlled by use of reflux as previously
described.
In a similar way to that described above for lube
protection, the interstage separation of the olefinic
cracking products produced by the dewaxing reactions
may also be effected in distillate dewaxing.
Application of the present olefin separation-stripping
technique to distillate dewaxing will follow the same
lines as described in detail above, except that a
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relatively taller fractionator tower will be required
in order to achieve the desired degree of product
separation in view of the relative closeness of boiling
points with distillate products and feeds. A higher
reflux ratio, with a consequent effect on tower size
may also be needed to separate products adequately
although the light olefins will be removed in the C4
olefin fractionator in the manner described above.
