The invention specifically relates to the alkylation of hydrocarbons
such as aromatics or paraffins to produce useful chemicals and motor fuel using
a fluidized, solid catalyst--for example, alkylation of isobutane to produce C8
isoparaffins useful as motor fuel blending components.
Large amounts of high octane gasoline are produced by the alkylation
of isobutane with butenes. Likewise, large amounts of valuable aromatic hydrocarbons
including cumene, ethylbenzene and C16-C21 linear alkylaromatics
are produced by the alkylation of benzene with olefins of the appropriate carbon
number. The variety of feed reactants and the passage of time has led to the development
of a number of effective alkylation technologies which are employed in large scale
One of the most widely used processes for the production of motor
fuel is HF alkylation as described in US-A-4139573. One of the advantages of the
use of liquid-phase HF as a catalyst is its resistance to deactivation, and the
relative ease with which a slipstream may be removed from an onstream reaction
zone for "regeneration". The HF itself is not chemically changed during use but
various organic reaction by-products such as "acid soluble oils" (ASO) accumulate
in the liquid-phase HF and are removed during this regeneration.
Regeneration is also necessary for all solid motor fuel alkylation
catalysts developed to date since they tend to suffer from a high deactivation
rate. Deactivation of solid catalysts is due to different, possibly multiple, causes
from those encountered with liquid HF as a catalyst and usually includes some accumulation
of hydrocarbonaceous deposits on the catalyst.
A common method of regenerating catalysts is by combustion of organic
deposits. This is often not desired for alkylation catalysts. US-A-3851004 describes
an alternative method for regenerating a solid bed alkylation catalyst comprising
a hydrogenation component on a zeolitic support which comprises contacting the
catalyst with a hydrogen-containing liquid-phase saturated hydrocarbon.
Any interruption in the operation of the reaction zone to regenerate
or replace catalyst is undesirable. Certain operating benefits are provided to
any process by an ability to operate in a continuous manner, which makes it desirable
to find a means to regenerate or replace the catalyst while the reaction zone is
kept in use. US-A-4973780 describes a moving bed benzene alkylation process in
which catalyst is continuously or periodically replaced with regenerated catalyst
to provide countercurrent catalyst-reactant flows. Cocurrent flow with catalyst
added to the bottom of the reactor is also disclosed.
It has also been proposed to provide continuous operation by simulating
the movement of the catalyst through the reaction and regeneration zones. US-A-4028430
describes the use of simulated countercurrent operations to perform a number of
alkylation reactions including the production of motor fuel. These references provide
separate reaction and catalyst reactivation zones, with an external regenerant
stream being employed for the reactivation. In both references the effluent of
the reaction zone is withdrawn from the alkylation zone immediately upon its exit
from the reaction zone. These references also teach the use of a "pump around"
stream to complete the simulation and provide a continuous liquid loop.
US-A-5157196 describes a moving bed paraffin alkylation process which
employs a plug flow in which the catalyst moves upward to a disengaging zone. Used
catalyst from the disengaging zone is passed into a wash zone.
The invention is a fluidized process for the alkylation of hydrocarbons.
The invention provides a continuous reaction zone which is not interrupted for
the periodic regeneration of catalyst. The invention also eliminates the need to
perform motor fuel or aromatic alkylation reactions using volatile and hazardous
liquid phase hydrofluoric acid. The invention is characterized by the use of a
fluidized riser-type reaction zone with the upper end of the reaction zone discharging
into a separate zone in which the reactants and products are separated from used
catalyst and the used catalyst is then stripped and recirculated to the riser.
A first portion of the used catalyst is mildly regenerated while a second portion
is drawn off for full regeneration in an external fluidized regeneration zone.
One broad embodiment of the invention may be characterized as a process
for the alkylation of a feed hydrocarbon which comprises the steps of passing a
first catalyst stream, comprising fluidized regenerated catalyst, and a feed stream
comprising the feed hydrocarbon and an alkylating agent into the bottom of a vertical
riser-reaction zone maintained at reaction conditions and producing a reaction
zone effluent stream comprising used catalyst, the feed hydrocarbon and a product
hydrocarbon; discharging the reaction zone effluent stream into a separation zone
in which used catalyst is separated from liquid phase hydrocarbons and thereby
forming a liquid-phase separation zone effluent stream comprising the feed and
product hydrocarbons, with the thus separated used catalyst descending downward
within the separation zone; transferring a first portion of the used catalyst downward
through a liquid phase regeneration zone wherein the first portion of the used
catalyst is contacted with feed hydrocarbon containing dissolved hydrogen to form
regenerated, washed catalyst; transferring a smaller second portion of the used
catalyst or a portion of the regenerated, washed catalyst, together with feed hydrocarbon,
from the separation zone into a high temperature regeneration zone wherein the
catalyst is contacted with vapor phase hydrogen at vapor phase regeneration conditions
and withdrawing high temperature regenerated catalyst from the high temperature
regeneration zone as a second catalyst stream; charging at least a portion of the
second catalyst stream to the riser-reaction zone or to the liquid phase regeneration
zone or to both of these zones and employing at least a portion of the regenerated,
washed catalyst as the first catalyst stream which is then charged to the riser-reaction
zone; arid, recovering the product hydrocarbon from the separation zone effluent
BRIEF SUMMARY OF THE DRAWINGS
- Figure 1 is a simplified diagram illustrating an embodiment of the invention
in which the riser-reactor 2 discharges into a large separation vessel 4, which
retains a dense bed 34 of catalyst which is being subjected to mild liquid-phase
- Figure 2 illustrates a second embodiment of the invention in which the riser
reactor discharges into a hydrocyclone type separation zone 25 and the majority
of the used catalyst is then subjected to mild regeneration in a segregated washing
section 24 located below the hydrocyclone.
Hydrocarbon alkylation is widely used in the petroleum refining and
petrochemical industries to produce a variety of useful acyclic and cyclic hydrocarbon
products which are consumed in motor fuel, plastics, detergent precursors, and
petrochemical feedstocks. Much of the installed base of alkylation capacity uses
liquid phase hydrofluoric acid, generally referred to as HF, as the catalyst. The
use of HF in these applications has a long record of highly dependable and safe
operation. However, the potential damage from an unintentional release of any sizeable
quantity of HF and the need to safely dispose of some byproducts produced in the
process has led to an increasing demand for alkylation process technology which
does not employ liquid phase HF as the catalyst.
Numerous alkylation catalysts have been described in the open literature.
However, those that we have knowledge of all appear to suffer from unacceptably
high deactivation rates when employed at commercially feasible conditions. While
some catalysts have a sufficiently useful lifetime to allow the performance of
alkylation, the rapid change in activity results in a change in product composition
and also requires the periodic regeneration of the catalyst with the accompanying
removal of the reaction zone from operation. It is very desirable to provide a
continuous process for alkylation which is not subjected to periodic reaction zone
stoppages or variation in the product stream composition.
It is an objective of this invention to provide an alkylation process
which does not employ liquid phase HF as the catalyst. It is a further objective
of the subject invention to provide an alkylation process which utilizes a fluidized
solid catalyst. It is a specific objective of the invention to provide a solid
catalyst alkylation process for the production of motor fuel blending hydrocarbons.
A further objective of the subject process is to provide a continuous process which
delivers a uniform quality and quantity of product and which employs a fluidized
solid alkylation catalyst.
The subject process achieves these objectives by the use of unique
fluidized catalyst flow schemes in which a riser-type reactor delivers the product
hydrocarbons and used catalyst to a liquid phase separation zone from which catalyst
is removed for division between one of two regeneration zones.
The hydrocarbon feedstock to the subject process may be essentially
any hydrocarbon which is retained as an easily flowable liquid phase material at
the conditions employed in the reaction and mild regeneration zones and which may
be alkylated via solid catalyst at the conditions maintained in the riser reaction
zone. The feed hydrocarbon may therefore be an aromatic hydrocarbon such as benzene
or toluene. This feed hydrocarbon or substrate is often reacted with an alkylating
agent comprising an acyclic light olefin such as ethylene, propylene or butylene
to produce such chemicals as ethylbenzene and cumene. A large amount of benzene
is alkylated with higher carbon number olefins having from about 10 to about 15
carbon atoms per molecule to produce linear alkyl benzenes which are then sulfonated
to produce detergents. For motor fuel production the preferred feed hydrocarbons
are light paraffinic hydrocarbons such as the butanes. An especially preferred
paraffinic feed hydrocarbon is isobutane.
The entering feed hydrocarbon is typically alkylated with a linear
olefin having from 2 to 15 carbon atoms per molecule. The feed hydrocarbon may
also be reacted with an alkylating agent chosen from a variety of compounds other
than olefins including monohydric alcohols. Examples of the alcohols which may
be employed as the alkylating agent include ethanol and methanol. For instance,
methanol is widely described in the literature as being useful in the paraselective
methylation of benzene and toluene.
The operation of the subject invention may be best discerned by reference
to the Drawings. In both figures, isobutane is the feed hydrocarbon and it is reacted
with C4 olefins to produce C8 hydrocarbons which may be recovered
by normal product recovery methods such as fractional distillation. Although there
are differences between the two embodiments, the same numbering system has been
employed on each one to the extent that the same or analogous equipment is employed.
As used herein the term "substantially free" means a molar concentration
less than 1.5 mole percent. The term "rich" is intended to indicate a concentration
of the specified compound or class of compounds greater than 50 mole percent.
Referring to Figure 1, a liquid phase feed stream comprising an admixture
of isobutane and C4 olefins enters the process through line 1 at the
bottom of the riser-reactor 2. The injection of this liquid results in the upward
flow of the contents of the riser-reactor 2 including solid catalyst which travels
downward through the transfer line 20 at a rate controlled by the slide valve 21.
Liquid phase isobutane flows into the transfer line 20 through line 22 at a rate
sufficient to cause a continuous minor net upward liquid flow through the conduit
20. This upward liquid flow is intended to strip hydrogen from the catalyst and
surrounding liquid to prevent the entrance of hydrogen into the riser and to wash
heavy hydrocarbon contaminants from the surface of the catalyst. Conduit 20 will
therefore deliver a stream of hydrogen-free freshly regenerated, washed catalyst
to the bottom of the riser reactor. This catalyst is admixed with the entering
reactant feed stream and catalyzes the reaction of olefins with the entering C4
isobutane or recirculating isobutane to form C8 product hydrocarbons.
The reaction products, the residual isobutane and the now used catalyst exit from
the top 3 of the riser reactor 2 and enter into a large volume cylindrical separation
chamber 4. The flush stream of line 22 may be passed into the conduit above the
slide valve 21.
The low liquid velocities present within the separation vessel 4 allow
the solid particulate catalyst to settle downward and form a dense fluid bed 34
located in a lower portion of the vessel. This bed is preferably maintained in
a dense fluidized state by the passage of a liquid phase stream comprising isobutane
and dissolved hydrogen into the vessel 4 through line 19. A distribution grid 18
is employed near the bottom of vessel 4 to achieve a uniform distribution of the
entering liquid-phase isobutane containing dissolved hydrogen throughout the dense
fluid bed. The catalyst retained within the fluid bed is therefore subjected to
a mild regeneration procedure by contact with hydrogen saturated isobutane. This
added isobutane of line 19 together with that from line 22 and the olefin-free
liquid phase material exiting from the riser reactor 2 will gradually travel upward
through the vessel 4 and enter into a cyclone type ("hydroclone") liquid-solids
separator 5. The cyclone 5 effects the further separation of any entrained catalyst
or catalyst fragments from the liquid phase hydrocarbons before the liquid is discharged
into a plenum 37 at the top of the vessel 4. The plenum may be utilized to facilitate
the installation of two or more separate cyclones into the separation vessel 4.
The thus collected liquid phase hydrocarbons are removed from the process through
line 6 as a product stream and transferred to the appropriate product recovery
facilities not shown on the drawing.
The major portion of the used catalyst retained in the dense bed is
withdrawn as a continuous stream through line 20 at the rate set by the slide valve
21. This first catalyst stream flows downward countercurrent to some of the hydrogen-free
isobutane charged to the process via line 22. The remainder of this isobutane flows
from line 22 into the riser 2. The purpose of this procedure is to prevent the
entrance of hydrogen into the riser where it could saturate olefins added by line
1. If the catalyst employed in the process does not promote the hydrogenation of
the olefins, then this washing procedure may be eliminated. Slide valve 21 will
need to always be slightly open, or other means provided, to allow the flow of
liquid upward through line 20 and ensure hydrogen does not enter the reactor 2.
The hydrogen-free isobutane of line 21 can alternatively be passed into line 20
at a point above the slide valve 21.
A second and smaller stream of the catalyst present in the dense fluid
bed at the bottom of the separation vessel 4 is withdrawn through line 8 at a flow
rate controlled by the slide valve 7. This smaller stream comprises both solid
catalyst and liquid-phase hydrocarbons and is passed into an external regeneration
vessel 9, with the catalyst being retained in the regeneration vessel 9 for some
average time set by the transfer rate in line 8. It is currently preferred that
the second catalyst stream has a uniform flow rate but a variable rate could be
used to facilitate batch regeneration. While in the regeneration vessel the catalyst
is agitated and fluidized by the addition of a high temperature vapor phase stream
comprising hydrogen and isobutane through line 10. This stream has been heated
by means not shown to a sufficient temperature to cause the vaporization of at
least a major portion of the liquid phase hydrocarbons which enter the regeneration
vessel through line 8 in the second catalyst stream. There is thereby formed a
vapor phase regeneration zone effluent stream comprising hydrogen, isobutane and
any other hydrocarbons which enter the regeneration zone through line 8 or result
from the regeneration process. This higher temperature hydrogen-rich stripping
is a much more intense regeneration procedure and is preferably performed at conditions
yielding a catalyst residence time of at least 30 minutes within the regeneration
A stream of fully regenerated catalyst is removed from the regeneration
zone via line 12 at a rate controlled by slide valve 13. This rate is preferably
approximately equal to the rate at which catalyst is fed into the regeneration
zone but may fluctuate over short periods. The highly regenerated catalyst first
flows through a catalyst cooler 14 which receives low temperature isobutane through
line 15 and is then passed into the bottom of a second riser 17. The catalyst is
preferably cooled to a temperature below 38°C. A stream of liquid-phase isobutane
from line 16 then fluidizes the highly regenerated catalyst and causes it to flow
upward via line 17 into the separation vessel 4 where it is commingled with the
catalyst which has been subjected to the mild regeneration. It has been noted that
highly regenerated catalyst tends to produce lower octane number product and the
use of a blend of fresh and mildly regenerated used catalyst should reduce this
tendency. As shown on the drawing, the catalyst from the transfer riser 17 preferably
enters the separation vessel 4 in the upper half of the vessel 4 to avoid the pressure
head of the catalyst retained in the vessel 4. Alternatively, the transfer riser
17 may direct the highly regenerated catalyst into the dense bed of catalyst present
in the lower portion of the separation vessel. This may be done to increase the
degree of fluidization at the bottom of the separation vessel. Alternatively, a
portion or all of the highly regenerated catalyst could also be fed directly into
the riser 2. The higher volume of mildly regenerated catalyst and its greater rate
of circulation overwhelms the addition rate from line 17.
The circulation of the catalyst through the high temperature regeneration
zone requires the catalyst to be heated and cooled. The utility requirements of
the process also require that the heat of reaction of the alkylation reaction be
removed. These activities can be integrated with the operation of the products
recovery section of the process. For instance, the heat available in the vapor
phase stream discharged via line 11 from the high temperature regeneration vessel
9 can be used to aid in reboiling a fractionation column. Heat can also be supplied
to the product recovery section from the cooler 14 used to cool catalyst being
returned to the riser-reactor. Alternatively, the heated coolant may be passed
into the regeneration zone 9.
Figure 2 illustrates a different embodiment of the subject process.
Like the embodiment of Figure 1, the feed hydrocarbon and feed olefin reactants
enter the bottom of a riser reactor 2 through line 1. Regenerated catalyst flowing
downward through the transfer conduit 20 at a rate controlled by the slide valve
21 is admixed with the entering feed hydrocarbons from line 1 and fluidized upward
through the riser 2. At the upper terminal end of the riser 2, the olefin-free
reactant-catalyst admixture is directed horizontally into a hydrocyclone 25 which
functions as the separation zone. The hydrocyclone is the sole solids-liquid separation
device employed in this embodiment. The residual feed hydrocarbon which has not
been converted in the riser-reactor and the product hydrocarbons exit from the
upper end of the hydrocyclone through line 6 for transfer to the product recovery
zone. The used catalyst separated in this manner from the reactants passes downward
through the lower portion of the hydrocyclone. A first small portion of this used
catalyst is diverted from the bottom of the hydrocyclone through conduit 8 at a
rate controlled by slide valve 7. This small portion of the used catalyst is passed
directly into the high intensity external regeneration vessel 9. In comparison,
the embodiment of Figure 1 diverts catalyst which has been subjected to mild regeneration.
Regeneration vessel 9 is operated in a manner similar to the external
regeneration zone of the embodiment of Figure 1. The catalyst is preferably confined
within this regeneration zone for an average residence time of at least 30 minutes
while being contacted with a heated stream of hydrogen and isobutane fed to the
bottom of the regeneration zone 9 through line 10. This hot hydrogen-hydrocarbon
stripping removes liquid phase hydrocarbons and deposits from the catalyst and
produces a vapor phase regeneration zone effluent stream removed from the regeneration
zone 9 through line 11. This regeneration zone effluent stream is preferably cooled
sufficiently to condense substantially all of the hydrocarbons contained within
this stream and then subjected to vapor-liquid separation. The recovered liquids
are passed into the products recovery zone and the hydrogen may be recycled to
the bottom of the regeneration zone.
Catalyst which has been subjected to the high temperature stripping
is withdrawn from vessel 9 through line 12 at a rate controlled by the slide valve
13. This hot catalyst is admixed with liquid phase isobutane from line 32 and then
passed into the heat exchanger 14. Isobutane coolant supplied by line 15 is used
to cool the catalyst to less than about 38° C and the catalyst then flows into
the stripping section 26. Catalyst cooling may be used to heat and/or vaporize
the isobutane. Vaporization has some advantages.
The majority of the catalyst collected in the bottom of the hydrocyclone
25 passes downward through a liquid-filled wash section 24 which functions as the
mild regeneration zone in this embodiment. The descending catalyst preferably passes
through a series of funnel-shaped baffles 23 intended to admix and stir the catalyst
and promote uniform contacting of the descending catalyst with a rising stream
of hydrogen saturated isobutane injected into the bottom of the wash section through
line 29. The countercurrent contacting within the wash zone 24 imparts a mild regeneration
to the used catalyst descending from the hydrocyclone and washes off some of the
deactivating deposits. The hydrogen saturated isobutane rises into the hydrocyclone
and is removed with the product stream of line 6.
The regenerated, washed catalyst descending through the wash section
enters the stripping section 26 where it is contacted with an additional quantity
of upward flowing isobutane from line 31. The catalyst descends through the stripping
zone 26 countercurrent to the rising hydrogen-free isobutane. A second series of
inclined conical or funnel-like baffles 30 is provided in the stripping zone at
various locations to ensure admixing of the rising isobutane with the descending
catalyst and a thorough removal of hydrogen and deactivating deposits from the
catalyst. Although the baffles 30 are shown only above the junction with the transfer
conduit 12 they may be located below this point also. At the midpoint of the stripping
section, the descending catalyst stream from the wash section is joined and admixed
with the stream of highly regenerated catalyst removed from the external regeneration
zone 9. Additional cooling is provided to the bottom of the stripping section 26
by means of indirect heat exchanger 35 located at the bottom of the stripping section
which receives coolant through lines 36. This cooling results in the catalyst being
brought to the desired reaction zone inlet temperature before passage into the
riser 2. The upward flow of hydrogen-free isobutane from line 31 is relied upon
to flush hydrogen from the catalyst stream of line 12.
The steps of the subject invention include the regeneration of catalyst
located in one regeneration zone by contact with a liquid-phase hydrocarbon, which
is preferably the feed hydrocarbon such as isobutane. Hydrogen is preferably dissolved
in this liquid-phase stream up to the point of the stream being saturated with
hydrogen. The average residence of catalyst particles in the liquid-phase hydrocarbon
regeneration zone is preferably from about 0.5 to 15 minutes. The liquid-phase
or "mild" regeneration is performed in a vessel or conduit in relatively open communication
with the reaction zone. The temperature and pressure conditions employed in this
regeneration zone will therefore be very similar to those in the reaction zone.
The invention also includes a second regeneration operation in which catalyst is
contacted with a vapor-phase gas stream at an elevated temperature in the range
of 80 to 500°C and more preferably from 100 to 250°C. The zone in which this "hydrogen
stripping" or severe regeneration step is performed is operated in a manner which
provides a longer average residence time for the catalyst particles than the liquid-phase
regeneration step. The average residence time of a catalyst particle should be
at least 30 minutes and can reach 12 to 24 hours. This regeneration step is performed
using a vapor-phase hydrogen rich gas stream. The presence of some isobutane in
this gas stream may be desirable to increase the heat capacity of the gas and therefore
increase catalyst heat up rates. The longer residence time required for this regeneration
step allows the high temperature gas charged to the regeneration zone to vaporize
liquid which flows into the severe regeneration zone.
All of the catalyst passing from the separation zone to the bottom
of the riser is preferably subject to one of the two forms of regeneration. A much
smaller quantity of catalyst flows through the hydrogen stripping regeneration
zone compared to the flow through the liquid-phase regeneration. The flow through
the high temperature regeneration zone will be only about 0.2 to about 20 weight
percent, and preferably from about 0.4 to about 5 weight percent of the total catalyst
flow through the riser.
The catalyst flow into the bottom of the riser is preferably as close
to a continuous steady state flow as the equipment and catalyst system allow. The
flow can, however, be in the form of numerous small quantities of catalyst transferred
in rapid sequence.
The Drawing and above description are presented in terms of controlling
catalyst flow through the use of slide valves. Alternative means can be used for
this purpose including, for example, other types of valves, lockhoppers, fluid
flow control (reverse flow of liquid), screw conveyors, etc. One particular alternative
is the use of an "L valve", which would reduce the amount of isobutane flush needed
in the process.
The embodiment shown in the Drawing may also be varied by the use
of other types of heat exchangers and by the use of other coolants. While the use
of isobutane as coolant and integration with the product fractionation zone is
preferred, other coolants including water, air or other hydrocarbons can be employed.
A further variation encompasses the use of countercurrent fluid flow to simultaneously
cool newly regenerated catalyst and to flush hydrogen from the catalyst and liquid
surrounding the catalyst.
The subject process can be performed using any solid, that is, heterogeneous,
catalyst which is stable and has the required activity and selectivity for the
desired reaction at the conditions needed to maintain liquid-phase reactants in
the reaction zone. A large number of catalysts have been proposed for the production
of motor fuel by alkylation including various zeolites and superacid catalysts.
For instance, US-A-4384161 describes the use of a large pore zeolite and a Lewis
acid. The zeolites referred to include ZSM-4, ZSM-3, the faujasites including zeolite
Y and mordenite. The Lewis acids mentioned in this reference include boron trifluoride
and aluminum chloride. The alkylation of isoparaffins using a somewhat similar
catalyst system comprising a large pore crystalline molecular sieve such as a pillared
silicate or an aluminophosphate or silicoaluminophosphate together with a gaseous
Lewis acid is disclosed in US-A-4935577. The use of these Lewis acids is not preferred
in the subject process as they provide their own waste handling and safety problems.
They also will probably require provisions for the circulation of the Lewis acid,
which complicates the process. US-A-5157200 describes an isoparaffin alkylation
process using a catalyst comprising a crystalline transition alumina, preferably
eta or gamma alumina, which has been treated with a Lewis acid under anhydrous
conditions. Previously referred to US-A-5157196 describes an isoparaffin alkylation
process using a slurried solid catalyst, with the preferred catalyst being an acid
washed silica which has been treated with antimony pentafluoride. Both of these
last two references describe a number of prior art solid alkylation catalysts.
A preferred alkylation solid catalyst comprises a refractory inorganic
oxide impregnated with a monovalent cation, especially an alkali metal cation or
an alkaline earth metal cation, and whose bound surface hydroxyl groups have been
at least partially reacted with a Friedel-Crafts metal halide. Analogs of these
catalysts without the metal cations are described in US-A-2999074 and US-A-3318820
which describe preparation techniques which can be applied to the preferred catalysts.
The preferred refractory oxide is alumina having a surface area greater than 50
m2/g, but the use of other oxides including titania, zirconia, silica,
boria and aluminum phosphate is contemplated. The preferred catalyst also contains
a metal component active for olefin hydrogenation deposited on the inorganic oxide
prior to reaction of the bound surface hydroxyl groups with the metal halides.
This metal may be chosen from the group consisting of nickel, platinum, palladium,
and ruthenium with the first three of these metals being preferred. The catalyst
contains one or more monovalent metal or alkaline earth metal cations selected
from the group consisting of lithium, sodium, potassium, cesium, silver, copper,
beryllium, magnesium, calcium and barium. Subsequent to the deposition of these
metals and the controlled calcination of the composite, the composite is reacted
with a Friedel-Crafts metal halide. The metal may be aluminum, zirconium, tin,
tantalum, gallium, antimony or boron. Suitable halides are the fluorides, chlorides
The presence of a highly active metal hydrogenation component on the
catalyst will promote hydrogenation of the feed olefin if both the olefin and hydrogen
simultaneously contact the catalyst. This potential waste of the olefin and hydrogen
can be avoided by careful design and operation of the process to avoid having both
the olefin and hydrogen in simultaneous contact with the catalyst. This can be
done by flushing the hydrogen or olefin from the catalyst before inserting it into
a zone containing the other compound as described above.
Silicalites have been described as useful alkylation catalysts for
the production of monoalkylbenzenes in US-A-4489214 and as useful in methylating
toluene to produce paraxylene in US-A-4444989. The use of ZSM-5 zeolites in aromatic
alkylation is described in US-A-3751506. ZSM-5 zeolites that have been treated
with one or more compounds or elements to improve their selectivity for para-selective
alkylation of aromatic hydrocarbons are described in US-A-4420418. The use of zeolite
L, zeolite Omega and zeolite beta as alkylation catalysts for the selective alkylation
of benzene is described in US-A-4301316. The use of a number of natural and synthetic
zeolites including clinoptilolite and zeolite Y as alkylation catalysts is described
The catalyst may be in the form of any suitable shape and size which
results in a solid catalyst which flows readily in both dry and wet states and
which is readily fluidized at the moderate liquid flow rates employed in the riser.
The catalyst can therefore be present as small irregular particles or as uniformly
shaped particles. It is preferred that the catalyst is present as "microspheres"
having an average diameter less than 0.16 cm and more preferably less than 0.08
Suitable operating conditions for the reaction zone include a temperature
of about -50 to 100°C, preferably 20 to 50°C, and a pressure as required to maintain
the hydrocarbons present as a liquid. A moderate pressure in the general range
of about 120 to about 3500 kPa is preferred with 2000-3000 kPa being highly preferred.
The weight hourly space velocity for the olefin may range from about 0.1 to 5.0
hr-1. The riser reaction zone is preferably designed and operated in
a manner intended to promote plug flow (no backmixing) of the reactants, products
and catalyst within the riser. However, the liquid must flow upward faster than
the catalyst in order to transport it.
It is generally preferred that the reaction zone is operated with
an excess of the feed hydrocarbon compared to the alkylating agent. That is, it
is preferred to operate with a ratio of the feed paraffinic or aromatic hydrocarbon
to a feed olefin at the reactor entrance greater than 1:1, and preferably from
2:1 to 5:1 or higher as measured by the flow rates into the reaction zone. It is
highly preferred to operate with an abundance of isoparaffin compared to alkylating
agent in a motor fuel alkylation process. Specifically, it is preferred that the
molar ratio of isoparaffin to olefin being charged to the reaction zone is greater
than 2:1 and more preferably greater than 3:1. Ratios from 10:1 to about 100:1
or higher can be employed for motor fuel alkylation. The known technique of feeding
the olefin at a number of points along the flow path of the feed hydrocarbon may
be employed to maintain a higher average paraffin to olefin ratio.
Provisions may be made for removing used catalyst from the reaction
zone and to replace the used catalyst with fresh catalyst. Conventional valved
lockhopper systems may be used for this purpose.