Isoparaffin:olefin alkylation process and catalyst complex

FIELD OF THE INVENTION
The present invention relates to a novel alkylation catalyst complex and to an alkylation process employing such catalyst complex.
BACKGROUND OF THE INVENTION
Alkylation is a reaction in which an alkyl group is added to an organic molecule. Thus an isoparaffin can be reacted with an olefin to provide an isoparaffin of higher molecular weight. Industrially, the concept depends on the reaction of a C.sub.2 to C.sub.5 olefin with isobutane in the presence of an acidic catalyst producing a so-called alkylate. This alkylate is a valuable blending component in the manufacture of gasolines due not only to its high octane rating but also to its sensitivity to octane-enhancing additives.
Industrial alkylation processes have historically used hydrofluoric or sulfuric acid catalysts under relatively low temperature conditions. The sulfuric acid alkylation reaction is particularly sensitive to temperature, with low temperatures being favored to minimize the side reaction of olefin polymerization. Acid strength in these liquid acid catalyzed alkylation processes is preferably maintained at 88 to 94 weight percent by the continuous addition of fresh acid and the continuous withdrawal of spent acid. The hydrofluoric acid process is less temperature sensitive and the acid is easily recovered and purified.
U.S. Pat. No. 4,795,728 to Kocal, for example, teaches a hydrofluoric acid catalyzed alkylation process for producing motor fuel. The hydrofluoric acid catalyst complex includes from 0.5 to 5 weight percent of a cationic or anionic surfactant component enabling the process to be operated at an olefin:acid volumetric feed ratio of greater than 1.0 while maintaining acceptable alkylate quality.
Both hydrofluoric and sulfuric acid catalysts are gradually depleted in continuous alkylation processes and must be regenerated or replenished by mixture with fresh acid to maintain acid strength, reaction rate, and the resulting alkylate quality. Specifically, alkylate quality responds directly to increasing acid strength, and the acid makeup or regeneration rate is typically controlled together with other process variables such as temperature and space velocity, to meet a required alkylate quality specification.
Both sulfuric acid and hydrofluoric acid alkylation share inherent drawbacks including environmental and safety concerns, acid consumption, and sludge disposal. Research efforts have been directed to developing alkylation catalysts which are equally as effective as sulfuric or hydrofluoric acids but which avoid many of the problems associated with these two acids. For a general discussion of sulfuric acid alkylation, see the series of three articles by L. F. Albright et al., "Alkylation of Isobutane with C.sub.4 Olefins", 27 Ind. Eng. Chem. Res., 381-397, (1988). For a survey of hydrofluoric acid catalyzed alkylation, see 1 Handbook of Petroleum Refining Processes 23-28 (R. A. Meyers, ed., 1986).
Catalyst complexes comprising BF.sub.3 as well as BF.sub.3 :H.sub.3 PO.sub.4 adducts have been proposed, and are discussed in greater detail below. While these catalysts effectively overcome the safety and environmental drawbacks of sulfuric and hydrofluoric acid alkylation systems, the volume and quality of BF.sub.3 alkylates have not, in the past, proven comparable to that of sulfuric or hydrofluoric acid alkylates.
U.K. Patent 545,441, assigned to Standard Oil Development Company, teaches a BF.sub.3 :H.sub.3 PO.sub.4 catalyzed isoparaffin-olefin alkylation process.
U.S. Pat. No. 2,345,095 to Bruner teaches a paraffin-olefin alkylation process catalyzed by a boron trifluoride-water complex, represented by the formula BF.sub.3 :H.sub.2 O, where n is preferably from 1 to 1.5. The Bruner reference notes at page 2, left hand column, lines 13-23, that the BF.sub.3 :H.sub.2 O catalyst complex behaves similarly to sulfuric acid but is a superior alkylation catalyst because BF.sub.3 :H.sub.2 O does not promote oxidation to undesired byproducts.
U.S. Pat. Nos. 2,296,370 and 2,296,371 to Slotterbeck disclose a BF.sub.3 :H.sub.2 O:HF catalyst system and an isoparaffin-olefin alkylation process employing the same. The catalyst system is said to avoid yield loss due to oxidation of the resulting alkylate product. The Slotterbeck '370 and '371 patents also discuss loss of catalytic activity due to diminishing acid strength; see the Slotterbeck '370 patent at page 2, right hand column at line 75 through page 3, left hand column at line 55, and the Slotterbeck '371 patent at page 2, right hand column at line 66, through page 3, left hand column at line 41.
U.K. Patent 550,711 teaches a process for increasing the activity of at least partially spent BF.sub.3 :H.sub.2 O catalyst systems for reuse in an organic condensation reaction. Briefly, the process volatilizes BF.sub.3 from the liquid catalyst mass to the extent required to promote separation of a distinct hydrocarbon phase from the catalyst mass. This hydrocarbon phase is then decanted off and fresh BF.sub.3 is added to restore catalytic activity.
Canadian Patent 424,000 teaches a process for producing gasoline boiling range hydrocarbons from isobutane and a normally gaseous olefin by absorbing the olefin in phosphoric acid of at least 75 weight percent concentration with an amount of isobutane equal to at least three moles of isobutane per mole of alkyl phosphate in the presence of a catalytic mixture of phosphoric acid and boron halide at temperature between 20.degree. C. and 60.degree. C.
U.S. Pat. No. 3,873,634 to Hoffman teaches a method of increasing the rate of ethylene alkylation by isobutane by carrying out the reaction simultaneously with the alkylation of a small amount of a higher weight olefin in the presence of a BF.sub.3 :H.sub.3 PO.sub.4 catalyst complex at low temperature and pressure.
U.S. Pat. No. 3,925,500 to Wentzheimer discloses a combined acid alkylation and thermal cracking process employing a BF.sub.3 :H.sub.3 PO.sub.4 acid catalyst in which unconverted propane and ethane from the alkylation process are converted, for example, to propylene and ethylene which are subsequently alkylated with isobutane to evolve a valuable liquid fuel.
Alkylate gasolines, such as those produced by the processes discussed above, are rich in isoparaffins and contain essentially no sulfur and aromatics. Generally exhibiting strong sensitivity to octane-enhancing additives, alkylate gasolines are prime candidates for blending into motor gasolines to meet the increasingly stringent environmental regulations restricting gasoline vapor pressure and aromatics content. However, the BF.sub.3 :H.sub.3 PO.sub.4 -catalyzed isoparaffin:olefin alkylation process has been found to require an induction, or equilibration period before the alkylate product meets industrial octane blending pool requirements. During the equilibration period, the alkylate product must be processed further before it can be blended into motor gasoline. For example, the low octane alkylate produced during the equilibration period can be charged to a catalytic reformer to upgrade the alkylate to reformate, a gasoline constituent enriched in substituted aromatics and isoparaffins. But regardless of how this low quality alkylate is processed, the fact remains that the refiner must incur incremental processing costs to upgrade the alkylate produced during the equilibration period before it is useful as a commercial gasoline blending stock. Thus it would be highly desirable to provide a liquid alkylation catalyst complex which avoids the safety and environmental concerns associated with hydrofluoric acid while continuously producing high quality alkylate without incurring an economic penalty for low octane alkylate production under process unit startup conditions.
SUMMARY OF THE INVENTION
In accordance with the present invention, it has been found that Lewis acids react with fluorophosphoric acids under certain conditions to form alkylation catalyst complexes which exhibit unusual properties. More particularly, the catalyst complexes of the present invention promote the reaction of an isoparaffin and an olefin which effects an unexpected improvement in product octane number over a particular range of excess Lewis acid concentrations. Further, the alkylate product resulting from isoparaffin:olefin alkylation in the presence of these catalyst complexes contains Lewis acid in sufficiently dilute concentrations that neutralization is a feasible, economically attractive alternative.
As used herein, the term "catalyst complex" refers to the reaction product of a specified liquid with Lewis acid, but is not presented to limit the scope of the invention by recitation or suggestion of a theoretical chemical structure for the catalytically active reaction product. The singular forms "catalyst complex" and "reaction product" are defined as including the plural forms "catalyst complexes" and "reaction products"; similarly, the plural forms are defined to include the singular.
The present invention includes both an isoparaffin:olefin alkylation catalyst complex and an isoparaffin:olefin alkylation process. In contrast to previously known alkylation catalysts and processes, it has surprisingly been found that a particular range of concentrations of excess Lewis acid in the presence of the reaction product of Lewis acid with a fluorophosphoric acid yield an unexpected improvement in alkylate quality. These results are of particular importance in the preferred embodiments of the process of the invention, which alkylates butene with isobutane to form isoparaffin-rich alkylate gasoline. The process not only improves alkylate octane and yield, but also decreases overall gasoline production costs in comparison to other Lewis acid-catalyzed alkylation processes by producing high octane alkylate at start-of-run conditions, thus eliminating the problem of upgrading low quality alkylate produced during the equilibration period. Still further, the catalyst complex and process of the present invention eliminate the need for costly separation and recycle of Lewis acid because the excess Lewis acid concentrations useful in the present invention are sufficiently dilute that the Lewis acid may economically be neutralized rather than recovered and recycled. Examples of suitable deacidification steps include neutralization, complexation, or sorption.
The invention thus provides a catalyst complex for promoting alkylation of an isoparaffin with an olefin, said first catalyst complex comprising the reaction product of Lewis acid and a fluorophosphoric acid in molar ratios of from about 0.5:1 to about 1.5:1, wherein said fluorophosphoric acid is selected from the group consisting of H.sub.2 PO.sub.3 F and HPO.sub.2 F.sub.2, and wherein the catalyst complex additionally comprises excess Lewis acid in concentration of from about 10 ppm (weight) to about 10000 ppm (weight) based on the total isoparaffin:olefin feed. Suitable Lewis acids include, but are not limited to, BF.sub.3, BCl.sub.3, AlCl.sub.3, AlBr.sub.3, SbF.sub.5, TaF.sub.5, PF.sub.5, SiF.sub.4, SF.sub.4, FeCl.sub.3 and ZnCl.sub.2.
The invention further includes, in its process aspects a continuous process for the alkylation of an isoparaffin with an olefin comprising the steps of:
(a) providing a liquid alkylation catalyst complex comprising
(i) the reaction product of a fluorophosphoric acid selected from the group consisting of H.sub.2 PO.sub.3 F and HPO.sub.2 F.sub.2 with a Lewis acid in molar ratios of from about 0.5:1 to about 1.5:1, and, additionally,
(ii) excess Lewis acid in concentrations of from about 10 ppm (weight) to about 10000 ppm (weight) excess Lewis acid based on the total isoparaffin:olefin feed; and
(b) reacting an isoparaffin containing from 4 to 20 carbon atoms with an olefin containing from 2 to 12 carbon atoms in an alkylation reaction zone in the presence of said catalyst complex of step (a) under conversion conditions including temperature of from about -40.degree. C. to about 200.degree. C., pressure in the range of subatmospheric to about 5000 psig, and weight hourly space velocity from 0.01 to 100 hr.sup.-1 based on total olefin feed. The excess Lewis acid is preferably present in concentrations of from about 50 to about 8000 ppm, more preferably about 5000 ppm based on the total isoparaffin:olefin feed.
The process optionally further comprises withdrawing a product stream containing gasoline boiling range hydrocarbons and catalyst complex from the alkylation reaction zone of step (b) and recycling at least a portion of the catalyst complex to the alkylation reaction zone. In preferred process embodiments, the isoparaffin is isobutane and the olefins are butenes and propylene, the alkylated product is separated from the catalyst complex by decantation, and the hydrocarbon phase including the alkylated product, which is withdrawn from an upper decanted layer, is deacidified to remove residual Lewis acid. Suitable deacidification steps include neutralization, complexation, or sorption.

DESCRIPTION OF THE DRAWINGS
FIG. 1 schematically illustrates the major processing steps of a first embodiment of the process of the present invention in which an olefin is alkylated with an isoparaffin in the presence of a catalyst complex comprising the reaction product of Lewis acid and a fluorophosphoric acid, and, additionally, excess Lewis acid, in a riser reaction zone. FIG. 1 further shows an optional Lewis acid neutralization step.
FIG. 2 schematically illustrates the major processing steps of a second embodiment of the process of the present invention in which an olefin is alkylated with an isoparaffin in the presence of a catalyst complex comprising the reaction product of Lewis acid and a fluorophosphoric acid, and, additionally, excess Lewis acid, in a stirred tank reactor. FIG. 2 also shows the optional Lewis acid neutralization step.

DETAILED DESCRIPTION
The present invention includes both an isoparaffin:olefin alkylation catalyst comprising a Lewis acid with a fluorophosphoric acid, as well as an alkylation process employing the same which effects a surprising improvement in alkylate quality at the specified levels of excess Lewis acid.
FEEDSTOCKS
Feedstocks useful in the present alkylation process include at least one isoparaffin and at least one olefin. The isoparaffin reactant used in the present alkylation process has from about 4 to about 8 carbon atoms. Representative examples of such isoparaffins include isobutane, isopentane 3-methylhexane, 2-methylhexane, 2,3-dimethylbutane and 2,4-dimethylhexane.
The olefin component of the feedstock includes at least one olefin having from 2 to 12 carbon atoms. Representative examples of such olefins include butene-2, isobutylene, butene-1, propylene, ethylene, hexene, octene, and heptene, merely to name a few. The most preferred olefins include the C.sub.3 and C.sub.4 olefins, for example, propylene, butene-1, -, butene-2, isobutylene, or a mixture of one or more of these olefins. Suitable feedstocks for the process of the present invention are described in U.S. Pat. No. 3,862,258 to Huang et al. at column 3, lines 44-56, the disclosure of which is incorporated by reference as if set forth at length herein.
The overall molar ratio of isoparaffin to olefin in the reactor feed during normal alkylation operation is generally between 1:1 and 100:1, preferably between about 5:1 and about 20:1.
Process Conditions
The present alkylation process is suitably conducted at temperatures of from about -40.degree. C. to about 500.degree. C., preferably from below about 40.degree. C. to avoid undesirable side reactions, and most preferably from about 0.degree. C. to about 20.degree. C. Lower reaction temperatures are preferred to maximize alkylate octane.
Operating pressure is controlled to maintain the reactants in the liquid phase, and is suitably from about 50 to about 5000 psig, preferably from about 50 to about 1500 psig, and more preferably from about 80 to about 200 psig. Weight hourly space velocity ranges from about 0.01 to about 100 based on total olefin feed to the reaction zone. The most preferred weight hourly space velocity varies within this range as a function of reaction temperature and feedstock composition. Weight hourly space velocity is readily optimized to attain a desired octane specification with only minimum trial and error.
The particular operating conditions used in the present process will depend on the specific alkylation reaction being effected. Process conditions such as temperature, pressure, space velocity and molar ratio of the reactants will effect the characteristics of the resulting alkylate, and may be adjusted within the disclosed ranges by those skilled in the art with only minimal trial and error.
The process can be carried out in any suitable vessel, for example, in a continuous stirred tank reactor such as that typically employed in sulfuric acid alkylation or in a tubular riser type reactor such as is typically used for hydrofluoric acid alkylation.
Catalyst Complex
A first embodiment of the catalyst complex of the present invention comprises the reaction product of Lewis acid and at least one fluorophosphoric acid selected from the group consisting of H.sub.2 PO.sub.3 F and HPO.sub.2 F.sub.2 in molar ratios of BF.sub.3 : fluorophosphoric acid of from about 0.5:1 to about 1.5:1, and, additionally, excess Lewis acid from about 10 ppm (weight) to about 10000 ppm (weight) based on the total isoparaffin:olefin feed, preferably from about 25 ppm (weight) to about 5000 ppm (weight) excess Lewis acid, and more preferably about 1000 ppm (weight) excess Lewis acid. Effective overall concentrations of the BF.sub.3 : fluorophosphoric acid reaction product and excess Lewis acid constituents of the catalyst complex will therefore vary slightly as varying process conditions such as temperature and pressure affect the solubility of Lewis acid in the fluorophosphoric acid, but these concentrations will generally fall within the range of from about 5 to about 75 weight percent Lewis acid, from about 25 to about 95 weight percent fluorophosphoric acid.
Alkylation Process
The process of the present invention is most preferably operated in a continuous mode. In the preferred continuous process embodiment, Lewis acid addition rates to the alkylation reaction zone are controlled at rates of from about 10 ppm (weight) to about 10000 ppm (weight) Lewis acid, and typically fall within the range of from about 50 ppm (weight) to about 8000 ppm (weight) of the total isoparaffin:olefin feed to the alkylation reaction zone.
Product Deacidification
By removing residual acid in the alkylation reactor product stream before it reaches the product fractionation section, the present process provides the ancillary benefit of rendering unnecessary the costly corrosion control measures otherwise required for recovery of acid-containing product. The hydrocarbon product effluent from the deacidification step must be essentially free of both residual Lewis acid as well as alkaline wash solution to be a useful gasoline blending stock. More specifically, the hydrocarbon product effluent should contain no more than 50 ppm (weight) residual Lewis acid, and preferably contains less than 5 ppm (weight) of residual Lewis acid, based on the total hydrocarbon effluent from the reaction zone. Suitable deacidification process steps include one or more of sorption, neutralization, or complexation.
Sorbents suitable for use in the present process must preferentially sorb Lewis acids from hydrocarbon streams having such compounds dissolved therein, and retain the sorbed Lewis acid until the enriched sorbent is contacted with a lean stripping fluid. Examples of such sorbents include both solid and liquid Lewis bases. Suitable Lewis bases which are normally solid under sorption conditions include calcium fluoride (CaF.sub.2), as well as polyvinyl alcohol (PVA). Suitable Lewis bases which are normally liquid under sorption conditions include diethylether, dioxane and anisole.
Solid sorbents are preferably packed into fixed beds equipped with flow distributors to promote uniform contact with the solid sorbent. Lewis acid sorption by a liquid sorbent, on the other hand, may be carried out using various gas-liquid contacting means well known to those skilled in the art, for example, a tower containing trays or packing to improve liquid-vapor contact. The enriched sorbent is then stripped of Lewis acid by gas-liquid stripping of the sorbent liquid with an isoparaffin-containing alkylation feedstream. For a general survey of suitable vapor-liquid contacting techniques and process equipment, see A. L. Kohl, "Absorption and Stripping", HANDBOOK OF SEPARATION TECHNOLOGY, 340 (1987).
The residual acid may also be removed from the hydrocarbon product effluent stream via neutralization via contact with a base such as KOH or NaOH. A suitable apparatus for conducting this process step is a packed bed defluorinator containing substantially inert Al.sub.2 O.sub.3 beads, as is commonly used to treat the depropanizer overhead stream in many commercial HF alkylation process units.
Process Flow
The reactor configuration for the process of the present invention is not critical; the process may be carried out in any suitable vessel, for example, a riser reactor or a stirred tank reactor. The liquid mixture from the stirred tank reactor is piped to a liquid-liquid separator, for example, a decanter, where the more dense liquid alkylation catalyst rapidly migrates to the bottom of the settler, while the gasoline-containing product floats to the top of the settler. The liquid catalyst is withdrawn from the bottom of the settler, and may be disposed and/or combined with fresh catalyst and sent back to the reactor. The hydrocarbon product is collected from the top of the settler and piped to a deacidification step, for example, a caustic scrubber where the hydrocarbon product contacts an aqueous caustic solution (NaOH) to remove residual acid.
Referring now to FIG. 1, a first embodiment of the process of the present invention is schematically illustrated. Feedstock containing isoparaffin and olefin flows through line 10 and is mixed with recycled isoparaffin from line 11 to provide a weight ratio of about 10:1 isoparaffin:olefin in the total feed to alkylation reactor 20. The isoparaffin is most preferably isobutane and the olefin is most preferably 2-butene. In the first embodiment of the invention as described herein, alkylation reactor 20 is a riser reactor as is commonly used in hydrofluoric acid alkylation processes. See, for example, 1 Handbook of Petroleum Refining Processes 23-28 (R. A. Meyers, ed., 1986), cited above, which is incorporated by reference as if set forth at length herein, for a general description of alkylation in a riser reactor.
A product stream containing partially spent catalyst complex, alkylated conversion products, and unreacted isoparaffin is withdrawn from alkylation reactor 20 through line 22 and charged to settler 30, preferably below the level of the heavier catalyst complex layer 34. The mixture readily separates due to the relative densities of the hydrocarbons and the BF.sub.3 : fluorophosphoric acid alkylation catalyst complex, which are approximately 0.6 g/cc and 1.7 g/cc, respectively.
The heavier catalyst complex accumulated in the lower section of 34 of settler 30, is withdrawn through line 36 and flows to mixing vessel 40 where the catalyst is admixed with fresh Lewis acid, or a mixture of Lewis acid and a fluorophosphoric acid, which is added via line 42. The Lewis acid charge rate to mixing vessel 40 is preferably controlled to restore the catalytic activity of the BF.sub.3 : fluorophosphoric acid complex with little or no excess Lewis acid. Thus typical excess Lewis acid levels fall within the range of about 25 ppm (weight) to about 5000 ppm (weight) Lewis acid based on the total isoparaffin:olefin feed to the reactor.
The catalyst then flows to catalyst cooler 50 before being returned to reactor 20. The alkylation reaction is typically exothermic, and for this reason, the catalyst complex is precooled before it is returned to the reaction zone to maintain the alkylation reaction temperature within the preferred range. As noted above, temperature control in the alkylation reaction zone effects alkylate quality and is therefore an essential process control consideration. A suitable cooling medium withdraws heat from the catalyst complex, preferably by indirectly contacting the catalyst complex as the cooling medium flows through tubes of a shell-and-tube heat exchanger as illustrated. Any suitable cooling medium known to those skilled in the art may be employed, for example, chilled water or a solution of ethylene glycol in water.
Because the activity of the catalyst complex gradually diminishes over time, a portion of the catalyst complex is continuously or intermittently withdrawn from catalyst cooler 50 through line 52 for treatment in a regeneration facility or disposal in suitable facilities (not shown).
Turning now to processing the alkylated hydrocarbon products, the lighter hydrocarbon phase contained in the alkylation reactor effluent is charged to separator 30 and readily forms a discrete hydrocarbon phase in the upper section 32 of the separator. This hydrocarbon phase including not only alkylated product but also unreacted isobutane as well as small amounts of Lewis acid, is withdrawn from separator 30 and charged to a deacidification vessel 60, which typically comprises a caustic (aqueous NaOH) wash. Fresh caustic enters the deacidification vessel 60 through line 64 and contacts the BF.sub.3 -containing hydrocarbon mixture. Because the hydrocarbon phase and the aqueous caustic phase are substantially immiscible, the two liquids are readily separated, with spent caustic and treated hydrocarbon being withdrawn from deacidification vessel 60 via lines 66 and 68, respectively. The minor amount of Lewis acid contained in the hydrocarbon phase fed to the deacidification zone typically forms precipitate which is withdrawn with the spent caustic phase through line 66.
The deacidified hydrocarbon phase containing unreacted isobutane and lighter C.sub.4 - aliphatics together with alkylated product flows to deisobutanizer separator 70 which separates the hydrocarbon stream into an overhead stream containing isobutane and lighter hydrocarbons and a bottom stream containing normal butane and gasoline boiling range alkylated product. The bottom stream is withdrawn from deisobutanizer 70 through line 76 for further fractionation (not shown) to reduce the alkylate gasoline vapor pressure or may be sent directly to product storage facilities (not shown). The overhead stream flows through line 72 to depropanizer 80 which splits off propane as an overhead stream through line 84 for resale or for use as fuel gas and recycles isobutane to line 10 through line 11 as described above. The deisobutanizer and depropanizer separators may comprise any suitable separation means known to those skilled in the art, for example, conventional multiple tray fractionation towers with reboilers and overhead condensers.
Referring now to FIG. 2, another embodiment of the process of the present invention is schematically illustrated in which the isoparaffin:olefin feedstock contacts the catalyst complex in a well stirred tank. A mixture of isoparaffin and olefin flows through line 110 and is mixed with recycled isoparaffin from line 111 to provide a weight ratio of about 10:1 isoparaffin:olefin in the total feed. The isoparaffin is most preferably isobutane and the olefin is most preferably 2-butene.
The mixed feedstream flowing through line 110 enters charge cooler 120 where it is cooled to a temperature slightly below the most preferred alkylation temperature as described above. The exothermic alkylation reaction then raises the temperature of the reactants to the preferred range as the conversion reaction progresses. As described above in the description of the first embodiment, the cooling medium preferably flows through the tubes 121 of a shell-and-tube heat exchanger. The cooling medium may comprise water with optionally added freezing point depressants such as ethylene glycol if alkylation process conditions require precooling the mixed feed to temperatures near the lower extremes of the preferred alkylation temperature ranges, e.g. 0.degree. C.
The cooled mixed feed is then withdrawn from charge cooler 120 through line 122, and enters stirred tank reactor 130. Catalyst loading in stirred tank reactor 130 ranges from about 0.1 to about 75 volume percent catalyst complex, preferably from about 25 to about 75 volume percent catalyst complex.
If the exothermic heat load in reactor 130 so requires, a reactor cooling coil 132 may be positioned within the reactor to maintain temperature within the preferred range to assure consistent alkylate quality. Because the coolant temperatures required in the tubes 121 of cooler 12 and the reactor cooling coil 132 are approximately equivalent, the two heat exchangers may be supplied in parallel from a central cooler or chiller. As used herein, the term "cooler" refers to any suitable source of process cooling water, for example an underground well, a natural draft cooling tower, a forced air cooling tower, or a refrigeration unit. The term "chiller" as used herein refers more specifically to refrigeration units, particularly those refrigeration units which supply a chilled solution of water and a freezing point depressant such as ethylene glycol.
Generally, cooling towers are suitable for use in conjunction with the present process to supply water at temperatures as low as about 24.degree. C. (75.degree. F.) under cool to moderate atmospheric conditions. Reaction temperatures as low as about 10.degree. C. (50.degree. F.) may be achieved via autorefrigeration, i.e., flashing i-C.sub.4 for cooling. For process temperatures below about 10.degree. C. (50.degree. F.), refrigeration units are typically employed to circulate a chilled aqueous mixture containing a freezing point depressant, for example, ethylene glycol.
The emulsion of unreacted isobutane, other light hydrocarbons, alkylation reaction products, and catalyst complex is withdrawn from stirred tank reactor 130 via line 134 and enters separator 140, preferably below the level of accumulated catalyst complex 144. The hydrocarbon phase, including not only alkylated product but also unreacted isobutane as well as small amounts of Lewis acid readily accumulates in the upper section 142 of the separator and is withdrawn from separator 140 via overhead line 146. The hydrocarbon phase is charged to a deacidification vessel 150, where the hydrocarbon phase is treated to remove residual acid, typically with a caustic (aqueous NaOH) wash entering vessel 150 through line 152. This deacidification step proceeds generally as described above with reference to the first embodiment. The hydrocarbon phase and the aqueous caustic phase are readily separated, with spent caustic and treated hydrocarbon being withdrawn from neutralization vessel 150 via lines 156 and 158, respectively.
The heavier acid catalyst complex phase 144 is withdrawn from separator 140 through line 148 and flows to mixing vessel 180 through line 149 where the catalyst is admixed with fresh Lewis acid, or a mixture of Lewis acid and a fluorophosphoric acid, which is added via line 182. The Lewis acid charge rate to mixing vessel 180 is preferably controlled to restore the catalytic activity of the BF.sub.3 :a fluorophosphoric acid complex with low levels of excess Lewis acid. Typical excess Lewis acid levels range from about 10 ppm (weight) to about 1000 ppm (weight) Lewis acid based on the total isoparaffin:olefin feed to the reactor.
The treated hydrocarbon phase then enters separator 160 where it is fractionated into a bottom alkylate-rich stream, a normal-butane rich side draw, and an isobutane-rich overhead stream, flowing through lines 162, 164, and 168, respectively. The bottom alkylate-rich stream contains isoparaffinic gasoline-boiling range hydrocarbons and is routed to storage or to gasoline blending facilities (not shown). The normal butane side draw may be resold or further upgraded elsewhere in the refinery complex (not shown). The overhead stream containing isobutane and lighter hydrocarbons flows through line 168 to depropanizer 170 where propane and lighter hydrocarbons are split off overhead and routed to fuel gas or for further upgrading. The depropanized isobutane stream from the bottom of depropanizer 170 is then recycled to line 110 via line 111 where it enters the charge cooler 120 and is recycled through the alkylation process.
EXAMPLES
The following examples illustrate the range of excess Lewis acid levels which provide the unexpected octane improvement of the present invention.
EXAMPLE 1
40 grams of monofluorophosphoric acid (H.sub.2 PO.sub.3 F, purchased from Alpha Chemical Company) were loaded into a 300 cc autoclave reactor. The reactor was filled with isobutane, cooled to 50.degree. F., pressurized to 150 psig while stirring at 1000 rpm and the reactor contents were then saturated with BF.sub.3 by passing BF.sub.3 through the isobutane for 24 hours at a rate of 25 cc Lewis acid per minute. A pre-mixed 10/1 wt/wt isobutane/2-butene feed was introduced at a rate of 170 grams per hour, and the BF.sub.3 co-feed rate was reduced to 0.17 grams per hour.
Throughout the run, liquid product was withdrawn from an upper section of the autoclave reactor and collected in a receiver chilled to 0.degree. C. (32.degree. F.). The receiver off-gas was passed through calcium hydroxide scrubbers to remove BF.sub.3, and finally through a wet test meter. The total product composition was determined from gas chromatographic analyses of the liquid and off-gas samples.
EXAMPLE 2
40 grams of orthophosphoric acid (85 weight % H.sub.3 PO.sub.4 ; 15 weight % H.sub.2 O; purchased from Baker Chemical Company) were loaded into a 300 cc autoclave reactor. The reactor was filled with isobutane, cooled to 50.degree. F., pressurized to 150 psig while stirring at 1000 rpm and the reactor contents were then saturated with BF.sub.3 overnight at a rate of 25 cc BF.sub.3 per minute. A pre-mixed 10/1 wt/wt isobutane/2-butene feed was introduced at a rate of 150 grams per hour and the BF.sub.3 co-feed rate was reduced to 0.75 grams per hour.
Throughout the run, liquid product was withdrawn from an upper section of the autoclave reactor and collected in a receiver chilled to 0.degree. C. (32.degree. F.). The hydrocarbon product phase and the catalyst complex phase were found to separate readily within the reactor by decantation. The hydrocarbon liquid product withdrawn from the autoclave reactor was found to be relatively free from catalyst complex, and loss of catalyst complex to the hydrocarbon product stream during the run was only minor.
COMPARISON OF EXAMPLES 1 AND 2
At 48 hours on-stream, BF.sub.3 /H.sub.2 PO.sub.3 F (Example 1) afforded 86 wt % C.sub.8 paraffins compared to only 79 wt % C.sub.8 paraffins with a BF.sub.3 /H.sub.3 PO.sub.4 /H.sub.2 O catalyst (Example 2). Alkylates at 48 HOS with BF.sub.3 /H.sub.2 PO.sub.3 F (Example 1) and BF.sub.3 /H.sub.3 PO.sub.4 /H.sub.2 O (Example 2) catalysts had octanes of 96.0 and 93.4 R+M/2, respectively. These results clearly show the advantage of BF.sub.3 /H.sub.2 PO.sub.3 F compared to BF.sub.3 /H.sub.3 PO.sub.4 H.sub.2 O at early time on-stream. The alkylate product with BF.sub.3 /H.sub.2 PO.sub.3 F at 414 HOS was comparable to that produced at 48 HOS. However, the alkylate with BF.sub.3 /H.sub.3 PO.sub.4 /H.sub.2 O dramatically improved with time on-stream. At 576 HOS, the alkylate with BF.sub.3 /H.sub.3 PO.sub.4 /H.sub.2 O had an octane of 95.5 R+M/2. Thus, after an "equilibration" period, the performance of BF.sub.3 /H.sub.3 PO.sub.4 /H.sub.2 O is comparable to BF.sub.3 /H.sub.2 PO.sub.3 F. These results are summarized in Table 1.
TABLE 1__________________________________________________________________________Results With Liquid BF.sub.3 -Promoted Catalysts__________________________________________________________________________Catalyst BF.sub.3 /H.sub.2 PO.sub.3 F BF.sub.3 /H.sub.2 PO.sub.3 F BF.sub.3 /H.sub.3 PO.sub.4 /H.sub.2 BF.sub.3 /H.sub.3 PO.sub.4 H.sub.2 OHours On-Stream 48 414 48 576Alkylate Compositions, wt %C.sub.5 -C.sub.7 8.1 6.4 11.8 7.5C.sub.8 86.1 88.5 78.9 86.9C.sub.9 + 5.8 5.1 9.3 5.6TMP/DMH 6.2 6.4 3.5 4.9Olefin Conversion, % 99.4 99.3 100.0 99.2Yield (Gms C.sub.5 +/Gm C.sub.4 .dbd. Conv.) 1.9 1.9 2.1 2.0C.sub.5 + Octane (R + M/2) 96.0 96.0 93.4 95.5__________________________________________________________________________
EXAMPLE 3
Excess BF.sub.3 is required to maintain stable performance with BF.sub.3 /H.sub.2 PO.sub.3 F. To demonstrate the effect of BF.sub.3 concentration, continuous alkylation data were obtained with 0, 1000, 5000 and 12,500 ppmw BF.sub.3 co-feed. The results are shown in Table 2. Alkylate quality decreased from 96 to 94 R+M/2 upon raising BF.sub.3 co-feed from 1000 to 12,500 ppmw, but increased from 94.6 to 96.0 when excess BF.sub.3 was raised from 0 to 1000 ppm.
TABLE 2__________________________________________________________________________Effect of BF.sub.3 Concentration On BF.sub.3 /H.sub.2 PO.sub.3 FPerformance__________________________________________________________________________BF.sub.3 Co-feed (ppmw on total HC) 0 1000 5000 5000 12500 12Hours On-Stream 400 414 438 462 486 50Alkylate Composition, wt %C.sub.5 -C.sub.7 13.3 6.4 7.3 7.1 8.5 6.3C.sub.8 72.2 88.5 85.4 85.6 83.3 85.4C.sub.9 + 14.5 5.1 7.3 7.3 8.2 8.3TMP/DMH 5.2 6.4 5.1 5.2 4.5 4.5Olefin Conversion, % 99.6 99.3 100 100 100 100Yield (Gms C.sub.5 +/Gm C.sub.4 .dbd. Conv.) 1.9 1.9 1.9 1.9 1.8 1.9C.sub.5 + Octane (R + M/2) 94.6 96.0 95.0 95.4 93.7 94.5__________________________________________________________________________
Changes and modifications in the specifically described embodiments can be carried out without departing from the scope of the invention which is intended to be limited only by the scope of the appended claims.

Number	Name	Date	Kind
2469335	Johnson et al.	May 1949
2842605	Appell	Jul 1958

Isoparaffin:olefin alkylation process and catalyst complex

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