Note: Descriptions are shown in the official language in which they were submitted.
21~'~212
FIELD OF THE INVENTION
This invention relates to hydroisomerization processes
including hydrocracking wherein the reaction is conducted in the
presence of carbon dioxide, and terminal cracking, i.e., methane
formation, is substantially minimized without substantial effect on
C2-C4 yields.
BACKGROUND OF THE INVENTION
Hydroisomerization processes and catalysts therefor are well
known, catalysts including noble metals, Pt, Pd, Rh supported on
flourided alumina, and Group VIII non-noble metals with or without one
or more Group VI metals supported on silica, alumina or silica-
alumina. These catalysts are usually bifunctional; they contain a
metal hydrogenation catalyst and an acidic cracking function.
Carbon oxides, carbon dioxide and particularly carbon
monoxide have have been disclosed in United States Patent No.
3,711,399 as inhibitors of hydrocracking in isomerization processes
using highly acidic fluorine containing catalyst, the carbon oxide
being added in relatively small amounts. Hydrocracking is virtually
completely suppressed and C4- yields are virtually negligible.
Nevertheless, LPG and light liquids yields in these processes
are desirable to insure appropriate pour points for diesel and jet
fuels. Methane, by itself, is the particularly undesirable product
since, for example, isomerized products can be made from Fischer-
Tropsch waxes which, in turn, ultimately come from methane via synthe-
sis gas production. Consequently, a desire exists for isomerization
processes that suppress or substantially eliminate methane formation
without substantial affect on LPG and light liquid yields.
SUMMARY OF THE INVENTION
In accordance with this invention terminal cracking of C5+
hydrocarbons is substantially suppressed, e.g., to less than about 1.0
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weight percent based on feed, by conducting a hydroisomerization
process incorporating carbon dioxide in the reaction mixture in the
presence of a catalyst comprising one or more Group VIII non-noble
metals or one or more Group VI metals, or both, the metals being
supported on an acidic support comprising alumina or silica alumina.
As a consequence of this invention, LPG and light liquid
yields are substantially unaffected while methane yield -- resulting
from terminal cracking is substantially suppressed. The known litera-
ture reports suppression of cracking not only of terminal bonds but
also of disubstituted bonds which have a higher activation rate for
cracking than terminal bonds. Thus, this invention provides a very
selective mechanism for suppressing hydrocracking than the gross
operations of the known literature. Further, the literature uses
exceedingly low amounts of carbon oxides with highly acidic catalysts.
The process of this invention uses greater amounts of carbon dioxide
(carbon monoxide not being useful) with a much less acidic catalyst --
a decidedly counter-intuitive approach. As a consequence of this
invention LPG yields, e.g., C2-C4, and light liquid yields, e.g.,
C5-320°F (160°C), 320-500°F (160-260°C) are
unaffected while C1 yields
are suppressed to less than about 1 wt%, preferably less than about
0.5 wt%.
DESCRIPTION OF THE DRAWINGS
Figures 1 a-1 f shows plots of various product yields when carbon
monoxide is added to the feed. Time is always on the abscissa.
Specifically Figures la-if are plots of 700°F+ (371°C+) wax
conversion,
methane yield, C2-C4 yield, C5-320°F (160°C) yield, 320-
500°F (160-
260°C) yield, and 500-700°F (260-371°C) yield, all v.
time. The first
vertical dotted line shows CO in at 320 hours and the second dotted
line shows CO out at about 650 hours.
Figures 2a-2f shows plots of various product yields when carbon
dioxide is added to the feed. Time is always on the abscissa.
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Specifically Figures 2a-2f show yields for the same products as in
Figures la-1~ The dotted line shows COz in at about 320 hours.
The amount of carbon dioxide used in conjunction with feed is
at least about 0.2 mole % based in feed, preferably at least about 0.3
mole %, preferably about 0.3 mole % to about 1.0 mole %. Of interest
ie the fact that while carbon oxides are often lumped together as
catalyst poisons, only carbon dioxide suppressed terminal cracking
with the non-noble metal, functional catalyst of this invention;
carbon monoxide had virtually no effect on the process.
Total conversion of feed during the process is 20-90%,
preferably 30-70%, and more preferably 40-60%.
The active hydroisomerization metals are non-noble metals
selected from Group VIII of the Periodic chart of the Elements.
Preferred metals are nickel and cobalt or mixtures thereof and mix-
tures thereof with molybdenum, a Group VI metal. The Group VIII
metals may be present on the catalyst in amounts sufficient to be
catalytically active for hydroisomerization. Specifically, metal
concentrations ranging from about 0.05 to about 20 wt%, preferably
about 0.1 to 10 wt%, still more preferably 2.0 to 5.0 wt% may be used.
For example, in a preferred catalyst the cobalt loading may be 1-4
wt%, and the nickel loading may be 0.1-1.5 wt%. A Group VI metal such
as molybdenum also can be employed in amounts more or less than or
equal to the non-noble Group VIII metal, e.g., 1.0 to 20 wt%, prefer-
ably 8-15 wt%, in all cases by total weight of catalyst.
The metals are impregnated onto or added to the support as
suitable metal salts or acids, e.g., nickel or cobalt nitrate, etc.
The catalyst is then dried and calcined in well known fashions.
The base silica and alumina materials used in this invention
may be, for example, soluble silicon containing compounds such as
alkali metal silicates (preferably where Na20:Si02=1:2 to 1:4),
tetraalkoxysilane, orthosilicic acid ester, etc.; sulfates, nitrates,
or chlorides of aluminum alkali metal aluminates, or inorganic or
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organic salts of alkoxides or the like. When precipitating the
hydrates of silica or alumina from a solution of such starting
materials, a suitable acid or base is added and the pH is set within
the range of about 6.0 to 11Ø Precipitation and aging are carried
out, with heating, by adding an acid or base under reflux to prevent
evaporation of the treating liquid and change of pH. The remainder of
the support producing process is the same as those commonly employed,
including filtering, drying and calcination of the support material.
The support may also contain small amounts, e.g., 1-30 wt% of
materials such as magnesia, titania, zirconia, hafnia, or the like.
A preferred support is an amorphous silica-alumina carrier,
containing less than about 35 wt% silica, preferably about 2-35 wt%
silica, more preferably 5 to 30 wt% silica, and having the following
pore-structural characteristics:
Pore Radius
(A1 Pore Volume
0-300 >0.03 ml/g
100-75,000 <0.35 ml/g
0-30 <25% of the volume of the pores with 0-300 A radius
100-300 <40% of the volume of the pores with 0-300 A radius
Such materials and their preparation are described more fully
in U.S. Patent No. 3,843,509 incorporated herein by reference. The
materials have a surface area ranging from about 180-400 m2g, prefer-
ably 230-375 m2/g, a pore volume of 0.3 to 1.0 ml/g, preferably 0.5 to
0.95 ml/g, bulk density of about 0.5-1.0 g/ml, and a side crushing
strength of about 0.8 to 3.5 kg/mm.
The feed materials that are isomerized with the catalyst of
this invention are waxy feeds, i.e., C5+, preferably boiling above
about 350°F (177°C) preferably above about 550°F
(288°C) and may be
obtained either from a Fischer-Tropsch process which produces substan-
tially normal paraffins or from slack waxes. Slack waxes are the
by-products of dewaxing operations where a diluent such as propane or
a ketone (e. g., methylethyl ketone, methyl isobutyl ketone) or other
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diluent is employed to promote wax crystal growth, the wax being
removed from the lubricating oil base stock by filtration or other
suitable means. The slack waxes are generally paraffinic in nature,
boil above about 600°F (316°C), preferably in the range of
600°F
(316°C) to about 1050°F (566°C), and may contain from 1
to 35 wt% oil.
Waxes with low oil contents, e.g., 5-20 wt% are preferred; however,
waxy distillates or raffinates containing 5-45% wax may also be used
as feeds. Slack waxes are usually freed of polynuclear aromatics and
heteroatom compounds by techniques known in the art, e.g., mild
hydrotreating as described in U.S. Patent No. 4,900,707, which also
reduces sulfur and nitrogen levels preferably to less than 5 ppm and
less than 2 ppm, respectively. Fischer-Tropsch waxes are preferred
feed materials, having negligible amounts of aromatics, sulfur and
nitrogen compounds.
Isomerization conditions include temperatures of 300-400°C,
500-3000 psig hydrogen, 1000-10,000 SCF/bbl hydrogen treat and space
velocity of 0.1-10.0 LHSV. Preferred conditions include 320-385°C,
750-1500 psig hydrogen, 0.5-2 v/v/hr.
The catalyst is generally employed in a particulate form,
e.g., cylindrical extrudates, trilobes, quadrilobes, and ranging in
size from about 1-5 mm. The hydroisomerization can be carried out in
a fixed bed reactor and the products may be recovered by distillation.
The following examples will illustrate this invention but are
not meant to be limiting in any way.
All of the hydroisomerization studies were carried out in a
small upflow pilot plant. The catalyst was evaluated at 750 psig,
0.50 LHSV, 690-700°F (366-377°C), and with a nominal H2 treat
rate of
2500 SCF/B. A 10 cc charge of catalyst crushed and sized to 14/35
mesh was employed in each case. The catalyst comprised 15.2 wt% Mo03
and 3.2 wt% Co0 on a silica-alumina Co gel with 20-30 wt% bulk silica.
Balances were typically collected at 24-72 hour intervals. The
reaction temperature was set to meet a target of 50% 700°F+ wax
conversion and was not adjusted during the run. The Fischer-Tropsch
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wax employed in these studies had a nominal composition of 0.70%
IBP-500°F (260°C), 20.48% 500-700°F (260-
371°C), 78.82% 700°F+
(371°C). Typical run lengths were 800-1000 hours. Boiling range
distributions for gas, naphtha, distillate range products, and lubes
were obtained by a combination of simulated gas chromatography distil-
lation and gas chromatography-mass spectroscopy.
The effect of carbon monoxide was evaluated and the catalyst
was activated using the following procedure:
1. Pressure test at 100°F (38°C) at ca. 750 psi hydrogen
pressure.
2. Reactor temperature increased to 700°F (371°C) while
maintaining
hydrogen pressure at 750 psi and a flow rate of 2500 SCF/bbl.
3. Reactor temperature held at 700°F (371°C) for approximately
18
hours.
4. Hydrogen feed and pressure were adjusted to standard operating
conditions and feed was cut in feed to start operation.
After feed was cut in, balances were collected on a regular
basis for 16 days to ensure that the catalyst had lined out, i.e.,
reached steady state conditions. At this point, the gas was switched
from pure hydrogen to a mixture containing 0.405 mole % carbon
monoxide in a balance of hydrogen. After 13 days on the CO/H2 mix-
ture, the gas was switched back to pure hydrogen for the remainder of
the run.
In this case there was an increase in C1 (methane), likely
due to~the hydrogenation of CO under reaction conditions.
The effect of carbon dioxide was studied also. In this case
a somewhat different but similar catalyst activation procedure was
used and is outlined below. This activation procedure leads to high
methane yields.
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1. Reactor temperature increased to 250°F (121°C) at atmospheric
pressure under nitrogen.
2. Pressure test with nitrogen, followed by pressure test with
hydrogen, while maintaining reactor temperature at 250°F
(121°C).
3. Reactor pressure increased to 1250 psia with hydrogen flow rate
at 3400 SCF/B. Reactor temperature increased at a rate of 30°F
(16°C) per hour to 700°F (371°C). Once reactor
temperature
reaches 700°F (371°C), hold for approximately 4-5 hours.
4. Pressure test hot with hydrogen; reduce hydrogen feed and
pressure to standard operating conditions; cut in feed to start
operation.
As in the previous run, once feed was cut-in balances were
collected on a regular basis for 13.5 days to ensure that the catalyst
had lined out. At this point the gas was switched from pure hydrogen
to a mixture of 0.604 mole % carbon dioxide, the balance being
hydrogen for the remainder of the run.
The effect of carbon monoxide on the catalyst performance is
displayed graphically in Figure 1. In general, the CO seemed to have
very little impact on the catalyst performance. The most significant
effect was a decrease in the 700°F+ (371°C+) wax conversion
which was
observed almost immediately after the CO was introduced. This
decrease continued until the conversion leveled off at about 55
percent and stayed at this level even when pure hydrogen was reintro-
duced to the system.
A small change in the methane yield was also detected. The
methane yield actually increased when the CO was introduced. This
occurred even though the conversion level was decreasing. Generally,
methane yield tracks reasonably well with conversion (i.e., an
increase in conversion usually leads to an increase in methane).
However, in this case, the slight increase in methane yield may be due
to CO hydrogenation, particularly since the methane level drops
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significantly when pure hydrogen is reintroduced and corresponds
almost exactly with the amount of methane which would be produced if
the CO was quantitatively converted to methane.
Examination of the remaining products (e.g., C2-C4, C5-320°F
(160°C), 320-500°F (160-260°C), 500-700°F (260-
371°C)) reveal little
or no effect from the CO other than differences attributed to the
change in conversion.
The effect of carbon dioxide on the catalyst performance is
displayed graphically in Figure 2. Although the product selectivities
for this run are significantly different than those obtained in the CO
experiment (primarily due to the different activation procedures), it
is the relative effect of the C02 that is of primary importance.
Shortly after the introduction of the Co2/H2 mixture, the
700°F+ (371°C+) wax conversion decreased by about 17%. The
conversion
slowly started to increase thereafter but did not reached the original
level.
The methane yield shows the most dramatic change as a result
of the C02. The activation procedure used in this run caused an
extremely high methane yield of about 2 wt%. Introduction of C02
caused this level to drop to less than 0.30 wt% where it remained for
the duration of the run. A small reduction in the methane yield would
be expected due to the decrease in the conversion; however, the effect
is too great to account for the total reduction.
Examination of the remaining products (e.g., C2-C4, C5-320°F
(160°C), 320-500°F (160-260°C), 500-700°F (260-
371°C) reveal little or
no effect from the C02 other than differences attributed to the
changes in conversion.
The following table illustrates the actual product yields for
the C02 experiment.
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_ g _
Total Time On Run
(Hours) 297.0 320.5 434.0 650.0 770.5
Delta Time ----- t=0 t=113.5 t=329.5 t=450
700F+ Conversion46.88 52.12 41.64 40.67 45.71
CH4 2.005 2.022 0.265 0.256 0.237
C2H6 0.196 0.190 0.033 0.037 0.043
C3H8 0.416 0.414 0.358 0.408 0.472
C4H10 1.140 1.182 1.054 1.289 1.555
t = 0 is the pointat whichcarbon dioxidewas added
These data were taken after line-out had been achieved.
From the table it is clear that C1 (methane) was substantial-
ly suppressed; C2 was suppressed somewhat, C3 and C4 were virtually
unaffected and as a result C2-C4 Was substantially unaffected.
Additionally, total conversion was suppressed at the outset of C02
addition, and recovered somewhat as the reaction proceeded. Thus,
C2-C4 cracked products can range from about 1 wt% to about 3 wt%.
