Process for paraffin-olefin alkylation

Abstract

Olefin-paraffin alkylate is prepared by contacting C.sub.3 -C.sub.9 monoolefin with C.sub.4 -C.sub.6 iso-paraffin (which can, if desired, be prepared in situ from other paraffin isomers) in liquid phase with a substantially anhydrous acidic crystalline alumino-silicate zeolite, and stopping such contacting after substantial alkylation (which can include self-alkylation of the isoparaffin) has occurred but before the weight rate of production of unsaturated hydrocarbon becomes greater than the weight rate of production of saturated hydrocarbon. The degree of conversion of olefins and paraffins to saturate products can be increased by use of halide adjuvants containing bromine, chlorine or fluorine.

Parent Case Text

CROSS REFERENCE TO RELATED APPLICATION

This application is a continuation-in-part of Ser. No. 581,129, filed Aug. 25, 1966 by the present inventors, now abandoned and assigned to the Sun Oil Company to whom the present application is also assigned.

Description

BACKGROUND OF THE INVENTION

This invention relates to the production of normally liquid, saturated hydrocarbons, useful in gasoline blending, by reacting isoparaffins with olefins in liquid phase in the presence of a substantially anhydrous cystalline alumino-silicate zeolite. These zeolites, in hydrated form, are chemically characterized by the empirical formula, xM. X(AlO.sub.2).y(Si0.sub.2).zH.sub.2 0, wherein M is H.sup.+ and/or an equivalent valence of metal-containing cations and x, y and z are integers, the ratio x/y being usually (but not necessarily) from 1.0 to 0.2.

The invention also comprises the use of halide catalyst adjuvants to increase the degree of conversion of olefins and paraffins to alkylate. Novel catalysts and novel alkylate products are also within the scope of our invention and are described hereinafter.

The invention will be described more particularly in connection with a process for the preparation of an olefin-paraffin alkylate comprising:

A. contacting C.sub.3 -C.sub.9 monoolefin with C.sub.4 -C.sub.6 isoparaffin and with a substantially anhydrous acidic crystalline alumino-silicate zeolite, at a temperature below the critical temperature of the lowest boiling hydrocarbon reactant and at a pressure such that the reactants are at least partially in liquid phase, and

B. stopping such contacting after substantial alkylation has occurred but before the weight rate of production of unsaturated hydrocarbon becomes greater than the weight rate of production of saturated hydrocarbon.

It is usual in the laboratory to effect paraffin-olefin alkylation by means of strong acids, such as AlCl.sub.3, HF, and H.sub.2 S0.sub.4 ; however, in petroleum refining, aluminum trichloride catalysis is accompanied by equipment corrosion, cracking, sludge formation, and other side reactions and has not proven economical, except for ethylene-isoparaffin alkylation (for which there is no other satisfactory catalytic process). Commerical isoparaffin alkylation with C.sub.3 -C.sub.6 olefins to produce high octane components for gasoline utilizes either H.sub.2 S0.sub.4 or HF as the catalyst. These acids, although very effective as alkylation catalysts, are highly corrosive and are potentially hazardous to workmen; therefore, strict safety procedures must be adhered to in their use. In addition, as for example in U.S. 2,359,119, alkylation processes utilizing sulfuric acid normally require reaction temperatures from about 5.degree.-15.degree.C.; therefore, costly cooling is required.

The art has long sought to find a process for paraffin-olefin alkylation utilizing catalyst which does not have the above-mentioned disadvantages possessed by AlCl.sub.3, HF, or H.sub.2 S0.sub.4. In particular, a heterogeneous process utilizing a solid alkylation catalyst has been sought by the art since processes using a liquid catalyst require that the acid and feed hydrocarbons, which are mutually immiscible, be kept in homogeneous suspension. Such "homogenization" requires expensive agitation devices and consumes much power. In addition, emulsions can be formed which are different and costly to "break."

Heretofore, attempts to effect paraffin-olefin alkylation utilizing a solid catalyst have had but little success. In all such published attempts, the bromine numbers of the reported products have been high, indicating that olefin polymerization (or some other competing reaction) has occurred to a substantial extent rather than the hoped-for paraffin-olefin alkylation.

Aromatic-olefin alkylation utilizing a crystalline, alumino-silicate catalyst has been reported (e.g., U.S. Pat. No. 2,904,607). However, the art has long recognized that olefin-aromatic alkylation and olefin-paraffin alkylation are very different chemical reactions and that there is no equivalency between processes for these two dissimilar combinations.

Although the chemistry of the alkylation reactions is complex and not completely understood, it is very probable that one major reason for the non-equivalency of olefin-aromatic alkylation and olefin-paraffin alkylations is that aromatic-olefin alkylation is dependent upon proton ejection from the intermediate carbonium ion whereas proton ejection in paraffin olefin reactions yields unsaturated products. In contrast, the production of saturated paraffin-olefin "alkylate" requires hydride transfer to the intermediate carbonium ion and, thus, a catalyst and process condition which favor hydride transfer over proton ejection.

In U.S. Pat. No. 3,251,902, claims are directed to the alkylation of C.sub.4 and C.sub.5 isoparaffins with C.sub.2 -C.sub.5 olefins. The examples, however, show only ethylene or propylene as feed olefins. Ethylene-paraffin reactions do not teach how to alkylate paraffins with C.sub.3 -C.sub.6 olefins. With propylene and isobutane, the reported characteriszations of the products in the examples of U.S. Pat. No. 3,251,902 are what would be expected in view of the above-discussed prior art, particularly the art dealing with olefin-paraffin alkylations. That is, with propylene and isobutane as the feed hydrocarbons, the reported products of the alkylation process of U.S. Pat. No. 3.251,902 are highly unsaturated, and, in fact, the inventors admit that such unsaturation indicates that "polymerization of the olefin is more pronounced than alkylation." They attribute such polymer production in their process to the thermal stability of the feed olefin and the high concentration of acid sites in certain of their zeolite catalysts. In no case do they recognize the above-discussed importance of hydride transfer in the production of saturated rather than unsaturated "alkylate."

SUMMARY OF THE INVENTION

As is further disclosed herein, we have discovered a process for the production of highly saturated alkylate from C.sub.3 -C.sub.9 monoolefins which requires not only a catalyst with a large number of acid sites of sufficient strength for hydride transfer but which also utilizes conditions which favor hydride transfer, such as introducing the olefin to the reactor in the liquid phase and in intimate admixture with C.sub.4 -C.sub.6 isoparaffin and, preferably, controlling the addition of the feed olefin such that the unreacted olefin in the hydrocarbon-catalyst reaction mixture is maintained at less than 12 mole percent (and most preferably less than 7 mole percent) based on the total paraffin content of the reaction mixture. In contrast to the process U.S. Pat. No. 3,251,902, we have also discovered that better yields of superior products can be obtained in our process with the more highly acid catalysts than with the less acid catalysts.

The use of dilute olefin feed streams has been suggested in conjunction with processes for olefin-aromatic alkylation (e.g., U.S. Pat. No. 3,251,897); however, there is no prior art suggestion or teaching of our liquid phase, acid zeolite-catalyzed paraffin-olefin alkylation process wherein C.sub.3 -C.sub.9 feed monoolefins are intimately premixed with feed paraffin, nor of our control of the concentration of unreacted olefin in the reaction mixture with acidic zeolite catalysts in order to obtain paraffin-olefin alkylation rather than polymerization.

We have further discovered that the production of saturated alkylation (and self-alkylation) products rather than unsaturated hydrocarbons is effected when the mean residence time (or retention or holding time) of the reaction mixture with the catalyst is in the range of 0.05 to 0.5 hours per (gram of hydrocarbon per gram of catalyst). More preferably, the mean residence time is 0.1-0.4 hours.

Preferably, there is present in the reactor mixture a halide adjuvant containing fluorine, chlorine or bromine. Particularly favorable alkylation conditions involve a temperature in the range of 25.degree. to 120.degree.C. (more preferably 50.degree.-100.degree.C.) and sufficient pressure to maintain a substantial part of each reactant in liquid phase (since mixed gas-liquid phase conditions are more likely to result in poor mixture of the feed paraffin and feed olefin, thus promoting olefin homopolymerization).

In general, unsaturated reaction products are indicative of olefin homopolymerization rather than paraffin-olefin alkylation. True alkylation, including isoparaffin self-alkylation, produces saturated hydrocarbons, However, apart from these primary reactions, i.e., polymerization and alkylation, the acidic zeolites catalyze many secondary reactions, such as cracking, disproportionation, and aromatization. These reactions can transform unsaturated "polymer" to saturated hydrocarbons and can cause unsaturated hydrocarbons to be formed.

In our process we desire to maximize the production of saturated hydrocarbons, and particularly the trimethylpentanes, since, as is shown by C. R. Cupit, et al., Petrol. Chem. Eng., Dec., 1961 at pages 204- 5, these have the more desirable antiknock characteristics, such as low sensitivity. With C.sub.4 olefins and isobutane, therefore, we desire to maximize the percentage yield, based on the weight of olefin charged, of C.sub.5 .sup.+ saturates and the yield of trimethylpentanes.

One means of maximizing this percent yield of C.sub.5 .sup.+ saturates is to prepare suitably active acid catalysts which favor hydride transfer and to conduct the process under conditions (as described herein) such that, as primary reactions, paraffin-olefin alkylation and isoparaffin self-alkylation are favored over polymerization. The reaction should also be controlled in a manner which reduces the occurrence of undesirable reactions.

Another method of increasing C.sub.5 .sup.+ saturate yield is to use reaction conditions which favor the secondary reaction of octanes with isobutane to form C.sub.7, C.sub.6 and C.sub.5 paraffins.

We have found that with isobutane and butene-2 feeds, even when such secondary reactions have occurred to some extent, the molecular ratio of trimethylpentanes/dimethylhexanes (TMP/DMH.sub.x) in the reaction mixture indicates the relative degree to which the primary reaction was alkylation or polymerization. That is, dimethylhexanes arise from olefin dimerization followed by hydride transfer; whereas, trimethylpentane formation is largely dependent upon paraffin-olefin combination to form a carbonium ion species followed by hydride abstraction from the isobutane. Therefore, the higher the ratio TMP/DMH.sub.x, the greater the effect of alkylation reactions, in contrast to olefin homopolymerization.

In general, the greater the tendency for the catalyst to initiate hydride transfer, the greater the ratio TMP/DMH.sub.x. For example, with isobutane-butene-2 feeds, AlCl.sub.3 catalyst (at 30.degree.C.) produces reaction products where TMP/DMH.sub.x is about 2/1. With HF or with H.sub.2 S0.sub.4, alkylates can be obtained with TMP/DMH.sub.x ratios about 8/1. commercial H.sub.2 S0.sub.4 alkylates have TMP/DMH.sub.x ratios between 3/1 and 6/1.

Prior art publications, such as those previously referred to, fail to teach how to use solid catalyst to obtain a C.sub.4 -C.sub.6 isoparaffin C.sub.3 -C.sub.9 olefin reaction product in which saturated products predominate rather than unsaturate. They also do not teach how to obtain products having high TMP/DMH.sub.x ratios. In contrast, we have discovered, and disclose herein, a paraffin-olefin alkylation process which utilizes acidic cystalline zeolite catalyst to obtain a predominantly saturated product in which the TMP/DMH.sub.x ratio can be greater than 7. We further teach how our process can be used to produce a paraffin-olefin alkylate which contains only negligible amounts of unsaturated reaction products. We have also discovered that maximum conversion of olefin to saturated products can be obtained with this process under conditions where the reaction mixture contains but a minor amount of unsaturated reaction products.

We have further discovered, unexpectedly (in view of prior art), that the conversion of olefin to saturated products can also be increased by the use of halide catalyst adjuvants. That is, the incorporation in the reaction mixture of small amounts of certain halides containing bromine, chlorine or fluorine allows the production of as much as 100 percent more saturated hydrocarbons from the same quantity of olefin reactant than can be produced under the same reaction conditions in the absence of the halide adjuvant.

BRIEF DESCRIPTION OF THE DRAWINGS

In the attached drawings, FIG. 1 illustrates the variation in the yield of C.sub.5 .sup.+ paraffins based on the olefin reactant as the reaction time is increased in our process.

FIG. 2 illustrates (by the solid curve) the weight percent of C.sub.5 .sup.+ unsaturates produced, based on the olefin reactant (here, butene-2), as the reaction time is increased. Also illustrated in FIG. 2, by the broken curve, is the moles of n-butane produced per mole of butene-2 converted, as a function of time.

FIGS. 3, 4 and 5 illustrate an apparatus which is particularly useful for effecting the continuous production of alkylate from a paraffin-olefin feed, utilizing the process of the present invention. This apparatus comprises three sections, a feed-mixing section, FIG. 3, a stirred, slurry reactor section, FIG. 4, and a product recovery section, FIG. 5.

In the feed section (FIG. 3), C.sub.3 -C.sub.9 monoolefin is admixed with C.sub.4 -C.sub.6 isoparaffin and transported, as by pumping, to the reactor section. The reactor section comprises a pressure vessel with means for maintaining the catalyst in suspension, such as a turbine mixer and baffles, means for introducing feed hydrocarbon and adjuvants such as a halide promoter, and means for separating a catalyst-free portion of the reaction mixture and transporting it from the reactor to the product recovery section.

The reactor section (FIG. 4), also includes means for maintaining sufficient pressure in the reaction vessel to insure that the reactants and the reaction mixture are in liquid phase, means (such as a differential pressure cell) for maintaining the liquid inside the reactor at a desired level and means for maintaining the reactor at the desired temperature (such as by a water jacket and heating coils).

The product recovery section (FIG. 5), comprises means for cooling the reaction mixture, means for separating gases (such as unreacted feed isoparaffin) from the desired liquid alkylate, and means for recycling unreacted feed hydrocarbons.

FURTHER DESCRIPTION OF THE INVENTION

Although our paraffin-olefin alkylation process requires the control of many inter-related process variables, such as reaction temperature, mixing rate, catalyst selection, concentration and preparation, the more critical conditions in our process are control of the maximum ratio of unreacted C.sub.3 -C.sub.9 olefin to C.sub.4 -C.sub.6 isoparaffin in the reaction mixture, that all the feed components (whether olefin or paraffin) must be well inter-mixed and kept at least partially (preferably predominantly) in the liquid phase, and most important, contact time of the olefin-containing reaction mixture with the catalyst must be controlled closely. A key measure of such contact is the mean residence time of the olefin-containing reaction mixture with the catalyst, in such units as mean hours per (gram of hydrocarbon per gram of catalyst). The criticality of these conditions with respect to obtaining a predominantly saturated alkylate has not been taught or suggested by the prior art.

The major reaction conditions which must be controlled in order to obtain a satisfactory product are inter-related in that they all influence the probability that a given molecule of reactant will collide with an active site of the zeolite catalyst and form a desirable carbonium ion. These major variables are the initial concentrations of isoparaffin and catalyst, the paraffin/olefin feed ratio, the intimacy of the paraffin-olefin premixing, the feed rate, agitation rate, temperature, pressure, catalyst type and particle size, and the contact time (or for a continuous stirred reactor, the mean residence time).

We have discovered that, within the operable ranges discussed herein, if the other variables are fixed (especially those which determine the probability that a given molecule of reactant will collide with an active site of the catalyst), a certain minimum contact time or induction period is required for substantial paraffin-olefin alkylation to occur. There is also a maximum contact time beyond which the quantity of saturated alkylate in the reaction mixture no longer increases, and unsaturated reaction products start to build up. We have found, in our process, that the production of saturated alkylation products is effected when the mean residence time of the reaction mixture with the catalyst is in the range of 0.05 to 0.5 hours.

The published art contains no disclosure, suggestion or even speculation of such criticality in a zeolite-catalyzed paraffin-olefin reaction system. The best mode of practice of our process utilizes this discovery to maximize conversion of our prarffin-olefin feed to desirable saturated products and to minimize, or effectively eliminate, the production of unsaturated hydrocarbons.

A preferred embodiment of our process for the preparation of an olefin-paraffin alkylate comprises the following steps:

1. contacting C.sub.3 -C.sub.9 monoolefin in admixture with C.sub.4 -C.sub.6 isoparaffin having a tertiary carbon atom, at a temperature below the critical temperature of the lowest boiling reactant and at a pressure such that the reactants are in liquid phase, with a substantially anhydrous acidic crystalline alumino-silicate zeolite;

2. controlling the addition of the said olefin reactant such that the amount of unreacted feed olefin in the reaction mixture is maintained at less than 12 mole percent and preferably less than about 7 percent, based on the total paraffin content and of the reaction mixture (or, more preferably, based on the unreacted C.sub.4 -C.sub.6 isoparaffin);

3. stopping such contacting after substantial alkylation has occurred but before the weight rate of formation of unsaturated hydrocarbon products exceeds the weight rate of formation of saturated hydrocarbons.

Preferably, in Step 3 above, the mean residence (or contact) time of the olefin-containing reaction mixture with the catalyst is in the range of 0.05 to 5 hours per (gram of hydrocarbon in the mixture per gram of catalyst). More preferred, is a mean residence time of 0.1 to 0.4 hours.

When our process is compared with the prior art processes for the alkylation of aromatic hydrocarbons with olefins, it becomes apparent that the prior art does not teach our process (particularly in view of the above-noted differences with regard to proton abstraction and hydride transfer between such reactions and paraffin-olefin reactions). In fact, it is evident that the prior art teachings relating to olefin-aromatic alkylation would lead the ordinary man skilled in the art (particularly one who attempted to substitute a C.sub.4 -C.sub.6 isoparaffin for the aromatic hydrocarbon in prior art examples) to conclude that paraffin-olefin "alkylation" catalyzed by crystalline zeolites results primarily in the production of unsaturated hydrocarbons.

Surprisingly, in view of the prior art reports of the ability of acidic crystalline zeolites, particularly the protonated or decationized zeolites, to strongly catalyze the polymerization of olefins, we have found a process, catalyzed by protonated and other acidic zeolites, whereby polymerization and/or other reactions which tend to produce unsaturate materials can be virtually eliminated, or held to a well-controlled minimum, while producing a saturated paraffin-olefin alkylate which is useful as a high octane component for gasoline.

We have further prefected our process so that with constant product removal and reactant recycle, saturated hydrocarbons may be produced continuously or continually in high yields. We can also control the molecular weight distribution of these hydrocarbons so that the desirable high anti-knock paraffins predominate. We have also discovered how to obtain a high degree of conversion of a C.sub.4 olefin-isobutane feed into alkylate in which the C.sub.8 paraffin portion contains 80 percent or more of trimethylpentanes and which has TMP/DMH.sub.x ratios greater than 5 and even as high as 11.

As is further illustrated herein, our process can, by appropriate selection of catalyst and conditions, be used to produce novel paraffin-olefin alkylates, useful as motor fuels and as gasoline blending components, comprising at least 60 mole percent C.sub.8 paraffins and less than 1 weight percent unsaturates and wherein the C.sub.8 paraffins consist of from 5 to 20 mole percent dimethylhexanes, from 0 to 1.5 mole percent methylheptanes, from 80 to 95 mole percent of trimethylpentanes, and wherein less than about 30 mole percent of the trimethylpentanes is 2,2,4-trimethylpentane.

Paraffin-olefin alkylates which contain a high proportion of trimethylpentanes have high octane ratings and good blending values and, therefore, are highly desirable components of gasoline. In contrast, olefinic hydrocarbons are less desirable as gasoline components because they are gum formers, are sensitive to oxidation, and when highly branched have poorer motor octanes than the corresponding paraffin hydrocarbons. Aromatic hydrocarbons, although they are generally good gasoline blending components with high octane ratings, are not desired in our process since they can decrease the activity of the catalyst.

The catalysts of our process are those substantially anhydrous acidic crystalline alumino-silicate zeolites which, in hydrated form, are chemically characterized by the empirical formula M.sub.x (A10.sub.2).sub.x (SiO.sub.2).sub.y.sup.. (H.sub.2 0).sub.z, where M is H.sup.+ and/or an equivalent valence of metal cations and x, y and z are integers, the ratio x/y being usually (but not necessarily from 1.0 to 0.2. A 10 percent aqueous suspension of the acidic zeolite catalyst will have a pH less than 7, preferably less than 5. For our process we prefer that the critical pore diameter of the zeolite be at least large enough to permit adsorption of benzene. We also prefer those acidic zeolites which contain both H.sup.+ and polyvalent metal cations (including metal cations in which part of the charge is balanced by oxide or hydroxyl groups), although catalysts having only H.sup.+ or polyvalent metal hydroxides (e.g., Ce(OH).sup.2 .sup.+or Ce(OH).sub.2 .sup.+ ) are effective in catalyzing paraffin-olefin alkylation.

These catalysts are normally prepared from alkali metal-containing zeolites (which in 10 percent aqueous suspension will have a pH greater than 7, and usually greater than 9) by ionexchanging the alkali metal ions for H.sup.+ and/or polyvalent metal cations. Hydorgen-ion (or proton) exchange can be effected by exchange from aqueous or non-aqueous medium with mineral acids, such as dilute aqueous HC1, or by exchange with solutions of acids and polyvalent metal ions (such as aqueous HN0.sub.3 and Ce(N0.sub.3).sub.3 ). For zeolites, such as the faujasites, which can be degraded by direct acid exchange, we prefer (as our exchange media) aqueous solutions containing, as at least one component, ammonium salts. Polyvalent metal exchange can be effected with solutions of salts of the metals, such as their nitrates.

Our preferred catalysts are prepared by such ammonium ion exchange, followed by polyvalent metal cation exchange, of an alkali metal faujasite (such as sodium type Y zeolite) having an Si0.sub.2 /A1.sub.2 0.sub.3 molar ratio in the range of 4.0 to 5.0.

ILLUSTRATIVE EXAMPLES

The present invention may be further illustrated by the following specific examples. Examples I-VI illustrate the preparation of acidic, or potentially acidic, solvated crystalline zeolites by aqueous cation exchange. Example VII illustrates the "activation" of the solvated zeolites by removing "solvent" from the zeolite.

The remaining examples, excepting Example XVIII, illustrate the use of such substantially anhydrous acidic crystalline alumino-silicate zeolites as alkylation catalysts. Of these, Examples VIII-XII show the effect of reaction temperature on the yield and product distribution. Examples X, XI and XIII show the unexpected, large increase in product yield effected by the use of various halide adjuvants. Example XII, when compared with Example IX, shows that at a given temperature the more highly exchanged catalyst does not necessarily produce the greatest product yield. When Example XII is compared with Examples X and XI, it is seen that, for a given catalyst and a given contact time, a large increase in reaction temperature will not necessarily give the large increase in yield which is obtainable by the use of a halide adjuvant.

Example XIV illustrates the influence of the contact time (or the time that the feed is in contact with the catalyst) and, in particular, shows the criticality, when maximizing saturate production based on olefin feed, of applicants' process step of stopping such contact after substantial alkylation has occurred but before the weight rate of production of unsaturated hydrocarbon becomes greater than the weight rate of production of saturated hydrocarbon. Further shown is the 100 percent consumption of feed olefin to produce an alkylate containing less than 0.5 percent of unsaturates. The importance of thorough premixing of the feed and incorporation of these steps into continuous process schemes is also disclosed.

Example XV illustrates the effect that catalyst composition has on the yield and product distribution. Example XVI shows the effect of various feed olefins on yield and product distribution. Example XVII shows the use of a fixed bed of catalyst in our process and compares the results with our preferred process using a catalyst slurry in a stirred reactor.

Example XVIII shows that highly unsaturated products are obtained when process conditions analogous to those of the prior art (in particular, the process conditions in U.S. Pat No. 3,251,902) are used to attempt to react butene-2 with isobutane in liquid phase in the presence of a substantially anhydrous acidic zeolite (similar to that of Example III). Example XVIII, when compared with other examples such as Example VIII or Example XIX, shows the importance of our requirement that the addition of the feed olefin be controlled such that the amount of unreacted olefin in the reaction mixture is maintained at less than 12 mole percent (preferably less than 7 percent) based on the unreacted C.sub.4 -C.sub.6 isoparaffin.

Example XIX illustrates a preferred embodiment of our invention wherein we produce a highly saturated, novel alkylate, highly desirable as a solvent and as a component of blended fuel for high compression automobile engines, containing less than 1 percent of unsaturates and comprising at least 60 mole percent of C.sub.8 paraffins, said C.sub.8 paraffins consisting of less than 1 mole percent methylheptanes, 5-10 percent dimethylhexanes, and at least 90 percent trimethylpentanes, said trimethylpentanes comprising less than 20 percent 2,2,4-trimethylpentane.

Example XX illustrates the effect of the gas used in catalyst activation on the paraffin yield, per weight of olefin charged, obtained from the resulting catalyst.

Example XXI illustrates the practice of our process in a continuous production operation wherein fresh portions of the hydrocarbon reactants (feed paraffin and feed olefin) are constantly added to the reaction mixture and a catalyst-free alkylate is constantly separated from the reaction mixture and withdrawn from the reactor.

Example XXII illustrates the effect, on yield and product quality, of the catalyst/olefin ratio.

Example XXIII illustrates the correlation between alkylate yield and the ESR measurements of total spin count when aromatic hydrocarbons are absorbed on the CeHY catalyst.

EXAMPLE I

This example describes the ammonium exchange of a crystalline, alkali metal alumino-silicate zeolite which can be heated to remove "loosely bound" water and to decompose the ammonium ion to produce a substantially anhydrous acidic crystalline alumino-silicate zeolite which can be used as a catalyst in our process. Preferably, before such decomposition or "decationizing," such ammonium-exchanged zeolites are further exchanged with polyvalent metal cations, as is shown in EXAMPLE III hereinafter.

A kilogram of a commercially available hydrated crystalline alumino-silicate zeolite, identified as sodium zeolite Y, was dried in air at 125.degree.C. for 18 hrs., broken up into particles of 100 mesh or less, redried in air at 125.degree.C. for 18 hrs., and suspended with stirring, in 1.7 liters of a 9.1% by weight aqueous solution of ammonium chloride at 80.degree.C. After 30 minutes the resulting ammonium-exhanged Y zeolite was separated from the liquid by filtration and recontacted at 80.degree.C. in a similar manner with a second 1.7 liter portion of fresh NH.sub.4 C1 solution.

After 6 more such 30-minute exchange cycles, the filtered zeolite was washed with distilled water (pH 6.5) at 20.degree.C. until no chloride ion could be detected in the spent wash liquid with acidic silver nitrate reagent.

The washed ammonium-exchanged zeolite was dried for about 18 hours in air at 125.degree.C., then ground to about 200 mesh and stored. The dried ammonium-exchanged zeolite produced by the above series of eight ammonium-exchanges analyzed 1.34 percent Na and 4.6 percent N, and had a loss on ignition of 26.5 percent. After the first ammonium-exchange cycle, a similarly washed and dried portion of the zeolite analyzed 5.5 percent Na and 2.3 percent N, and had a loss on ignition of 25.6 percent.

The sodium Y zeolite before this ammonium exchange had a pore size sufficiently large to enable it to absorb benzene and analyzed 7.5% sodium and 8.86 percent aluminum, and had an A1/Si atomic ratio of 0.40. The sieve had a loss on ignition at 1,800.degree.F. of 23.8%. All ignition losses referred to hereinafter were run at 1800.degree.F.

EXAMPLE II

This example illustrates the preparation of more highly ammonium-exchanged zeolites than that of Example I. Example I was repeated except that the sodium Y zeolite was subjected to 8 additional hot NH.sub.4 C1 exchange cycles before it was washed chloride free. The washed, dried ammonium-exchanged zeolite, resulting from this total of 16 ammonium-exchange cycles, contained 0.77 percent sodium and 4.14 percent nitrogen, and had 29.8 percent loss on ignition.

A similar exchange for a total of 32 cycles produced a washed, dried zeolite containing 0.21 percent Na and 4.64 percent N and having 28.7 percent loss on ignition.

Ammonium exchange of alkali metal zeolites can also be accomplished by suspending the zeolite in a vessel containing the exchange solution and maintaining a flow of fresh exchange solution into the vessel while withdrawing an equal volume of catalyst-free liquid from the vessel. Removal of catalyst-free liquid from the vessel can be effected by forcing the liquid with pressure or suction through a pleated microporous, woven stainless steel screen "10 " filter. In such continuous flow processing, the flow rate is preferably regulated so as to maintain a relatively constant pH in the exchange vessel. Hydrochloric acid ir nitric acid addition can also be used for pH control. With 10 percent ammonium chloride solutions it is preferred to maintain a pH of about 4.5 .+-. 0.3 (at 80.degree.C.). Ammonium exchange can also be effected by percolating the exchange solution through a fixed bed of zeolite.

EXAMPLE III

This example illustrates the further exchange of an ammonium-exchanged zeolite with a solution containing polyvalent metal ions in order to produce a zeolite containing both polyvalent metal ions and ammonium ions. A portion of the dried, ammonium-exchanged zeolite of Example I was contacted, with stirring, for 30 minutes at 80.degree.C. with 1.7 parts by weight of a 1.3 percent solution Ce(NO.sub.3).sub.3.sup.. 6H.sub.2 O, then separated from the exchange solution by filtration and recontacted for 30 minutes at 80.degree.C. with 1.7 parts by weight of fresh cerium nitrate solution. After 6 more such exchange cycles (or a total of 8 exchanges), the filtered Ce.sup..sup.+3 -exchanged/ammonium-exchanged zeolite was washed with water until no nitrate ion could be detected in the spent liquor by diphenylamine reagent. The washed Ce-NH.sub.4 .sup.+ Y zeolite was dried for 0b 18 hours at 125.degree.C., ground, redried for 18 hours at 125.degree.C., and stored in a moisture-tight container. The dried Ce.sup..sup.+ 3 -NH.sub.4 .sup.+-exchanged zeolite analyzed 6.18% Ce, 1.25% Na, and 1.43% N. It had a 25.6 percent weight loss on ignition.

EXAMPLE IV

This example illustrates the use of additional cerium exchange cycles and a more highly ammonium-exchanged "base" zeolite in order to obtain zeolites with a greater cerium content and a lower sodium content than the zeolite of Example III. A portion of the washed, dried "16 cycle" NH.sub.4 .sup.+-exchanged zeolite of Example II was contacted according to the procedure of Example III for a total of 16 Ce(NO.sub.3).sub.3 exchange cycles, then similarly washed and dried. The resulting Ce.sup..sup.+3 -NH.sub.4 .sup.+ -exchanged zeolite analyzed 10.1% Ce, 0.69% N, 0.68% Na, and had a loss on ignition of 24.4%.

A similar series of 16 cerium exchanges performed on the "32 cycle" ammonium-exchanged zeolite of Example II produced a washed, dried Ce.sup..sup.+3 -NH.sub.4 .sup.+ -exchanged zeolite which analyzed 0.23% Na, 10.3% Ce, 0.8% N, and had a loss on ignition of 24.7%.

Hereinafter, sometimes, a catalyst will be identified according to the number and type of such exchange cycles according to the code: number of ammonium exchange cycles/number of polyvalent metal exchange cycles. That is, the above zeolite propared by 6 cerium exchange cycles of a 32 cycle ammonium is, by this code, a 32/16 zeolite, (or, after activation, a 32/16 catalyst).

EXAMPLE V

This example illustrates the preparation of a Ce.sup..sup.+3 -exchanged sodium Y zeolite. A portion of the commercial sodium Y zeolite of Example I was ground and exchanged for 16 exchange cycles with Ce(NO.sub.3).sub.3 solution in a manner similar to the exchange of Example III, then washed and dried. The resulting Ce.sup..sup.+3 -exchanged Na Y zeolite containted 9.6 percent certium, 1.69 percent sodium, and had a loss on ignition of 25.1 percent.

The cerium exchanges of Examples III, IV and this example can be effected in a continuous manner, similar to that described in Example II for ammonium exchanges. Preferably the pH should be about 4.5. The particular polyvalent metal salt chosen and the pH of the exchange solution will determine whether the cationic exchange species is the metal or a hydroxylated complex ion of the metal. Other polyvalent metal ions, such as those referred to hereinafter, and in particular cations of the polyvalent rare earth metals and mixtures thereof, may be similarly exchanged with alkali metal-containing and/or ammonium containing crystalline zeolites. Especially preferred catalysts can be obtained from crystalline alumino-silicate zeolites which have been so exchanged with aqueous solutions of salts of gadolinium, such as Gd(NO.sub.3).sub.3, or with mixtures of salts of Gd and Ce. In this specifiction the term "rare earth metals" includes lanthanum, that is, the term "rare earth" herein is used as a synonym for "lanthanon." The lanthonons include La, Ce, Pr, Nd Sm, Eu, Gd, Tb, Dy, Ho, Er, Tm, Yb and Lu.

EXAMPLE VI

This example illustrates the ammonium exchange of a Ce.sup..sup.+3 -exchanged sodium Y zeolite. A portion of a washed, dried cerium-exchanged zeolite prepared similar to that of Example V and containing 7.7% Ce, 0.63% Na and with 26.4% loss on ignition was exchanged with hot aqueous ammonium chloride, for a total of 8 cycles, using the procedure of Example I. The washed, dried NH.sub.4.sup.+ -Ce.sup..sup.+3 -exchanged Y zeolite analyzed 0.22% Na, 2.8 N, 5.3 Ce.sup..sup.+3 and had a loss on ignition of 27.6%. Therefore, about 30 percent of the cerium was removed from the cerium-exchanged Na Y zeolite during the ammonium-exchange cycle.

An alkali metal zeolite which was exchanged by the reverse procedure, that is, 8 ammonium exchanges followed by 16 cerium exchanges, contained 87 percent more cerium (it analyzed 9.9% Ce, 1.3% Na and 0.26% N, and had a loss on ignition of 22.5 percent).

EXAMPLE VII

This example illustrates a preferred method of "activation" of hydrous crystalline alumino-silicate zeolites prior to their use as catalysts in our paraffin-olefin alkylation process. In general, hydrous crystalline zeolites are activated by controlled heating under vacuum or in a stream of a gas, such as air, hydrogen, nitrogen, helium or oxygen, to remove water. In the case of ammonium-exchanged zeolites, not only is loosely bound water removed but also the ammonium ion is decomposed to obtain a substantially anhydrous, "decationized" or "protonated" zeolite. Such zeolites are highly acidic and are similar catalytically to those prepared by direct exchange with an aqueous acid.

When the hydrous ammonium zeolite also contains polyvalent metal ions, the resulting activated zeolite will be partially protonated or "cation deficient." Such zeolites are not only highly acidic, but are more resistant to the detrimental effects of the activation procedure.

The heating rate and temperatures of such "activation" will depend to a great extent on the type of zeolite, that is, an Al/Si atomic ration, and the type and percent of polyvalent cations and monovalent ions such as hydrogen or ammonium ion. In any event the hydrated zeolite is first heated at a temperature sufficiently high to remove the bulk of the "uncombined" or "uncomplexed" water from the pores of the zeolite. At atmospheric pressures this temperature is preferably from 125.degree.-300.degree.C., most preferably from 125.degree.-240.degree.C.

In the case of an ammonium-exchanged zeolite the temperature is then raised to a higher temperature than that used for such water removal and such temperature is maintained for a sufficient time to remove a substantial amount of the ammonium ion from the zeolite as NH.sub.3. This removal may also involve decomposition of the ammonium ion by such reactions as oxidation of ammonia to nitrogen oxides or nitrogen and water.

At atmospheric pressure, with ammonium-exchanged zeolites which also contain appreciable quanitites of exchanged polyvalent metal cations, this higher temperature is preferably 320 -500.degree.C.

With ammonium-exchanged zeolites which contain no polyvalent metal cations or have a low content of polyvalent cations, it is important that the activation temperature be kept below about 400.degree.C., since at higher temperatures the intensity of the X-ray diffraction peaks of the zeolite decreases greatly (due to a degradation of crystalline structure) and the resulting catalyst is less active for paraffin-olefin alkylation. In U.S. Pat. No. 3,130,007 a similar intensity measurement is used to determine the "percent zeolite," and appears to relate to crystallinity of the zeolite.

We have also found that, if an ammonium-exchanged crystalline alkali metal zeolite is further exchanged with polyvalent metal cations, the resulting polyvalent metal NH.sub.4.sup.+ -exchanged zeolite retains a much greater proportion of its X-ray peak intensity after activation than does the base NH.sub.4.sup.+ -exchanged zeolite. Although small quantities of polyvalent cations will be of some benefit in this respect, for our catalysts it is preferable that the zeolite contain at least the following quantity of polyvalent metal cations (or a combination thereof of equivalent valence):

1. at least one tetravalent metal, metal oxide or metal hydroxide for every 16 atoms of aluminum in the alumino-silicate tetrahedra of the zeolite, or

2. at least one trivalent metal, metal oxide or metal hydroxide for every 12 atoms of aluminum in the alumino-silicate tetrahedra, or

3. at least one divalent metal, metal oxide or metal hydroxide for every 8 atoms of aluminum in the alumino-silicate tetrahedra.

In addition, for optimum activity, the polyvalent cation should be selected from classes 1, 2 and 3 above (and mixtures thereof) when the atomic ratio Al/Si of the alumino-silicate tetrahedra comprising the zeolite is greater than 0.65, or from classes 2 and 3 above (and mixtures thereof) when the atomic ratio Al/Si is from 0.65 to 0.35, or from class 3 above when the atomic ration Al/Si is less than 0.35. For example, the cation of our zeolite catalyst is preferably selected from the following:

1. at least one cation selected from the class consisting of V.sup..sup.+4, Mo.sup..sup.+4, W.sup..sup.+4, Pa.sup..sup.+4, U.sup.115 4, VOH.sup..sup.+4, Cr(OH).sub.2.sup..sup.+4, CrO.sup..sup.+4, MnO.sup..sup.+4, Mn(OH).sup..sup.+4, NbOH.sup..sup.+4, MoOH.sup..sup.+4, Mo(OH).sub.2.sup..sup.+4, MoO.sup..sup.+4, RuO.sub.2.sup..sup.+4, Ru(OH).sub.4.sup..sup.+4, RuO.sup..sup.+4, Ru(OH).sub.2.sup..sup.+4, SbOH.sup..sup.+4, OW.sup..sup.+4, W(OH).sub.2.sup..sup.+4, WOH.sup..sup.+4, Re(OH).sub.3.sup..sup.+4, Re(OH).sub.2.sup..sup.+4, ReO.sup..sup.+4, Os(OH).sub.4.sup..sup.+4, OsO.sub.2.sup..sup.+4, OOs.sup. .sup.+4, Os(OH).sub.2.sup..sup.+4, IrO.sup..sup.+4, Ir(OH).sub.2.sup..sup.+4, BiOH.sup..sup.+4, BiOH.sup..sup.+4, PaOH.sup..sup.+4, UO.sup..sup.+4, U(OH).sub.2.sup..sup.+4, and UOH.sup..sup.+4, when Al/Si is from 1.0 to 0.65,

2. at least one cation selected from the group consisting of Al.sup..sup.+3, Ni.sup..sup.+3, TiOH.sup..sup.+3, V.sup..sup.+3, VOH.sup..sup.+3, VO.sup..sup.+3, V(OH).sub.2.sup..sup.+3, Cr(OH).sub.3.sup..sup.+3, Mn(OH).sub.4.sup..sup.+3, MnO.sub.2.sup..sup.+3, Mn(OH).sub.3.sup..sup.+3, Mn(OH).sup..sup.+3, Mn.sup..sup.+3, GeOH.sup..sup.+3, ZrOH.sup..sup.+3, Nb(OH).sub.2.sup..sup.+3, NbO.sup..sup.+3, Mo(OH).sub.3.sup..sup.+3, Mo(OH).sub.2.sup..sup.+3, MoO .sup..sup.+3, MoOH.sup..sup.+3, Mo.sup..sup.+3, Ru.sup..sup.+3 , RuOH.sup..sup.+3, Ru(OH).sub.3.sup..sup.+3, Ru(OH).sub.5.sup..sup.+3, Rh.sup..sup.+3, RhOH.sup..sup.+3, PdOH.sup..sup.+3, SnOH.sup..sup.+3, Sb.sup..sup.+3, Sb(OH).sub.2.sup..sup.+3, SbO.sup..sup.+3, La.sup..sup.+3, HfOH.sup..sup.+3, Ta(OH).sub.2.sup..sup.+3, TaO.sup..sup.+3, W(OH).sub.3, WO.sup..sup.+3, W(OH).sub.2.sup..sup.+3, WOH.sup..sup.+3 W.sup..sup.+3, Re(OH).sub.4.sup. .sup.+3, ReO.sub.2.sup..sup.+3, Re(OH).sub.3.sup..sup.+3 ReOH.sup..sup.+3, Os.sup..sup.+3, OsOH.sup..sup.+3, Os(OH).sub.3.sup..sup.+3, Os(OH).sub.5.sup..sup.+3, Ir.sup..sup.+3, IrOH.sup..sup.+3, Ir(OH).sub.3.sup..sup.+3, PtOH.sup..sup.+3, PbOH.sup..sup.+3, Bi.sup..sup.+3, Bi(OH).sub.2.sup..sup.+3, Bio.sup..sup.+3, PoOH.sup..sup.+3, Ce.sup..sup.+3, CeOH.sup..sup.+3, Pr.sup..sup.+3, PrOH.sup..sup.+3, Sm.sup..sup.+3, Gd.sup..sup.+3, Tb.sup..sup.+3, TbOH.sup..sup.+3, Dy.sup..sup.+3, ThOH.sup..sup.+3, PaO.sup..sup.+3, Pa(OH).sub.2.sup..sup.+3, PaOH.sup..sup.+3, U(OH).sub.3.sup..sup.+3, U(OH).sub.2.sup..sup.+3 UO.sup..sup.+3, UOH.sup..sup.+3, U.sup..sup.+3, when Al/Si is from 0.65 to 0.35, and

3. at least cation selected from the class consisting of Mg.sup..sup.+2, Ca.sup..sup.+2, Ba.sup..sup.+2, Sr.sup..sup.+2, ScOH.sup..sup.+2, TiO.sup..sup.+2, Ti(OH).sub.2.sup..sup.+2, TiOH.sup..sup.+2, V(OH).sub.3.sup..sup.+2, V(OH).sub.2.sup..sup.+2, VO.sup..sup.+2 VOH.sup..sup.+2, V.sup..sup.+2, Cr(OH).sub.4.sup..sup.+2, CrO.sub.2.sup..sup.+2, CrOH.sup..sup.+2, Cr.sup..sup.+2, Mn(OH).sub.5.sup..sup.+2, Mn(OH).sub.4.sup..sup.+2, MnO.sub.2.sup..sup.+2, Mn(OH).sub.2.sup..sup.+2, MnO.sup..sup.+2, Mn.sup..sup.+2, MnOH.sup..sup.+2, Fe.sup..sup.+2, FeOH .sup..sup.+2, Co.sup..sup.+2, CoOH.sup..sup.+2, Ni.sup..sup.+2, NiOH.sup..sup.+2, GaOH.sup..sup.+2, Ge(OH).sub.2.sup..sup.+2, GeO.sup..sup.+2, YOH.sup..sup.+2, Zr(OH).sub.2.sup..sup.+2, ZrO.sup..sup.+2, Nb(OH).sub.3.sup..sup.+2, NbOH.sup..sup.+2, Mo(OH).sub.4.sup..sup.+2, MoO.sub.2.sup..sup.+2, Mo(OH).sub.3.sup..sup.+2, Mo(OH).sub.2.sup..sup.+2, MoO.sup..sup.+2, MoOH.sup..sup.+2, Mo.sup..sup.+2, Ru.sup..sup.+2, RuOH.sup..sup.+2, Ru(OH).sub.2.sup..sup.+2, RuO.sup..sup.+2, Ru(OH).sub.4.sup..sup.+2, Ru(OH).sub.6.sup..sup.+2, RuO.sub.3.sup..sup.+2, Rh.sup..sup.+2, RhOH.sup..sup.+2, Rh(OH).sub.2.sup..sup.+2, RhO.sup..sup.+2, Pd.sup..sup.+2, Pd(OH).sub.2.sup..sup.+2, InOH.sup..sup.+2, RuO.sub.2.sup..sup.+2, Sn(OH).sub.2.sup..sup.+2, SnO.sup..sup.+2, Sn.sup..sup.+2, SbOH.sup..sup.+2, Sb(OH).sub.3.sup..sup.+2, LaOH.sup..sup.+2, Hf(OH).sub.2.sup..sup.+2, HfO.sup..sup.+2, Ta(OH).sub.3.sup..sup.+2, W(OH).sub.4.sup..sup.+2, WO.sub.2.sup..sup.+2,W(HO).sub.3.sup..sup.+2, W(OH).sub.2.sup..sup.+2, WO.sup..sup.+2, WOH.sup..sup.+2, W.sup..sup.+2, Re(OH).sub.5.sup..sup.+2, Re(OH).sub.4.sup..sup.+2, ReO.sub.2.sup..sup.+2, Re(OH).sub.2.sup..sup.+2, ReO.sup..sup.+2, Re.sup..sup.+2, Os.sup..sup.+2, OsOH.sup..sup.+2 Os(OH).sub.2.sup..sup.+2, OsO.sup..sup.+2, Os(OH).sub.4.sup..sup.+2, OsO.sub.2.sup..sup.+2, Os(OH).sub.6.sup..sup.+2, OsO.sub.3.sup..sup.+2, Ir.sup..sup.+2, IrOH.sup..sup.+2, Ir(OH).sub.2.sup..sup.+2, IrO.sup..sup.+2, Ir(OH).sub.4.sup..sup.+2, IrO.sub.2.sup..sup.+2, Pt.sup..sup.+2, Pt(OH).sub.2.sup..sup.+2, PtO.sup..sup.+2, AuOH.sup..sup.+2, TlOH.sup..sup.+2, Pb(OH).sub.2.sup..sup.+2, PbO.sup..sup.+2, Pb.sup..sup.+2 BiOH.sup..sup.+2, Bi(OH).sub.3.sup..sup.+2 Po(OH).sub.2.sup..sup.+2, PoO.sup..sup.+2, Po.sup..sup.+2, AcOH.sup..sup.+2, CeOH.sup..sup.+2, Ce(OH).sub.2.sup..sup.+2, CeO.sup..sup.+2, PrOH.sup..sup.+2, Pr(OH).sub.2.sup..sup.+2, PrO.sup..sup.+2, NdOH.sup..sup.+2, Eu.sup..sup.+2, PmOH.sup..sup.+2, SmOH.sup..sup.+2, Sm.sup..sup.+2, EuOH.sup..sup.+2, Gd.sup..sup.+2, GdOH.sup..sup.+2, TbOH.sup..sup.+2 Tb(OH).sub.2.sup..sup.+2, Dy.sup..sup.+2, DyOH.sup..sup.+2, HoOH.sup..sup.+2, EROH.sup..sup.+2, TmOH.sup..sup.+2, Tm.sup..sup.+2, YbOH.sup..sup.+2, Yb.sup..sup.+2, LuOH.sup..sup.+2, Th(OH).sub.2.sup..sup.+2, ThO.sup..sup.+2, Pa(OH).sub.3.sup..sup.+2, Pa(OH).sub.2.sup..sup.+2, PaO.sup..sup.+2, U(OH).sub.4.sup..sup.+2, UO.sub.2.sup..sup.+2, U(OH).sub.3.sup..sup.+ 2, UO.sup..sup.+2, U(OH).sub.2.sup..sup.+2, UOH.sup..sup.+2, when Al/Si is less than 0.35.

Thomsonite, levynite, and the Type X zeolite of U.S. Pat. No. 2,822,244 are crystalline zeolites having an Al/Si atomic ratio greater than 0.65. Analcite, chabazite, phillipsite, and the Type Y zeolite of U.S. Pat. No. 3,130,007 have Al/Si ratios between 0.65 and 0.35. Heulandite and the Type L zeolite of U.S. Pat. No. 3,013,984 have Al/Si ratios less than 0.35. Mordenite has an Al/Si ratio in the range of 0.2 and some mordenites have been reported to have an Al/Si ratio appreciably less than 0.2 (e.g., 0.13). Such low Al content mordenites, when exchanged and activated by the procedures taught herein, have some catalytic activity in our process but are not among our preferred catalysts.

As catalysts in our process we further prefer substantially anhydrous protonated alumino-silicates which are capable of adsorbing benzene, wherein the ratio Al/Si in the tetrahedra is from 0.65 to 0.35 and which contain at least one rare earth metal cation for every 9 aluminum atoms in the tetrahedra since such catalysts have high alkylation activity and retain a high degree of X-ray peak intensity on activation or regeneration.

For example, in illustration of our preferred method of activation of a preferred species of hydrous zeolite, the 16-cycle Ce.sup..sup.+3 -exchanged/16-cycle NH.sub.4 -exchanged zeolite of Example IV was heated at 230.degree.C. in a rotating kiln in a stream of flowing air for about one hour to remove water. No loss of ammonium ions was detected during this heating period. The temperature of the kiln was then raised at the rate of about 10.degree.C. per minute to a temperature of 400.degree.C. During this heating, ammonia could be detected, by MnSo.sub.4 -AgNO.sub.3 reagent, in the exhaust gases from the kiln. The kiln was maintained at 400.degree.C. for 2 hours, at which point no ammonia could be detected in the exhaust gases. The heat was then removed from the kiln and the kiln was cooled rapidly in a flowing stream of dry air. The activated catalyst was maintained overnight in a slowly flowing stream of dry air. The resulting, substantially anhydrous, protonated crystalline alumino-silicate had a loss on ignition of 3.7 percent.

Summation of the intensity of the significant X-ray diffraction peaks of the hydrous zeolite before activation and of an activated sample showed no decrease in intensity for the activated zeolite. In contrast, a similarly activated portion of the base 16-cycle ammonium-exchanged zeolite showed an intensity decrease of 64 percent.

To illustrate the stabilizing effect of even small quantities of polyvalent metal ions, a sample of the base 16-cycle ammonium-exchanged zeolite was submitted to a 16-cycle Ce.sup..sup.+3 exchange using one-tenth the usual cerium salt concentration to produce a dried, washed zeolite which analyzed 1.23% Ce (ignited basis). After activation according to the above procedure, the activated zeolite showed an intensity decrease of 47.4 percent.

The bulk density in g/ml of the "dry" (at 125.degree.C.) hydrated zeolite is about 0.71 for sodium Y zeolite, 0.78 for highly ammonium-exchanged sodium Y zeolite (NH.sub.4 Y), 0.90 for highly cerium-exchanged NH.sub.4 Y (CeNH.sub.4 Y) and 0.89 for highly Gd exchanged NH.sub.4 Y (GdNH.sub.4 Y). If one assumes no significant volume change in activation, the calculated bulk density of the corresponding activated or substantially anhydrous zeolite would be in the range of 0.6 g/ml for the NaY and 0.75 g/ml for the CeHY.

Quantitative studies of the activation of "equilibrated" highly ammonium-exchanged sodium Y zeolite (hereinafter, sometimes NH.sub.4 Y) and cerium exchanged, NH.sub.4 Y (hereinafter, sometimes CeNH.sub.4 Y) have shown that, in our preferred catalysts, even after our optimum activation, water can be evolved from the catalyst upon ignition at 1,800.degree.F. This water is called hereinafter sometimes, "bound," or "combined" or "complexed" water to distinguish it from that water which is readily evolved from the exchanged zeolite below 300.degree.C. Equilibrated zeolite is a zeolite which has been exposed to air of about 50 percent relative humidity, at about 68.degree.C. for about 12 hours.

We have further established that our preferred substantially anhydrous, acidic crystalline zeolite catalyst, containing polyvalent metal ions and, more preferably, having some degree of protonation sometimes termed "cation deficiency", will evolve substantially no bound water when heated for about 1 hour at 300.degree.C. but when ignited at 1800.degree.F. will evolve about 1/4 to 2 mole of water for each atom of exchanged polyvalent metal. In particular, in our novel, activiated, cerium-containing catalysts, for each atom of cerium in the catalyst, 0.8 to 1.2 molecules of water will be evolved upon ignition at 1,800.degree. F.

We have concluded that in the catlyst this water is present, in the form Ce(OH).sup.2.sup.+. To understand the basis for this finding, one must first consider the behavior on activation of the hydrated NaY and NH.sub.4 Y zeolites. Behavior of NH.sub.4 Y catalyst during activation at temperatures from 150.degree.to 1292.degree.F. (65.degree. to 700.degree.C.) and for times up to 4 hours, as is summarized in Table 1. Two series of experiments were performed. Experiment A being at different times at constant temperature and Experiment B being at different temperatures at constant time. Total water (that is, sorbed and combined) appeared to be retained by this catalyst more firmly than ammonia.

Table 1 shows that about two-thirds of the total water present on "dried " NH.sub.4 Y zeolite had been removed after 1 hour at 450.degree.F. This water removal is an endothermic reaction and probably represents "loosely" held water that is in molecular form when sorbed. Data from DTA-EGA measurements agree with this observation.

Water removable only at 750.degree.F. and higher temperatures is more firmly bound and is chemically combined in a form other than molecular H.sub.2 O. Ammonia, which is largely in the form of NH.sub.4.sup.+ ions, was removed more readily than the water which remains after 1 hour at 450.degree.F. Furthermore, ammonia removal releases protons (NH.sub.4.sup.+ .fwdarw. NH.sub.3 + H.sup.+ ). When an activated NH.sub.4 Y zeolite is ignited at 1800.degree.F., OH groups are destroyed and H.sub.2 O is evolved in an amount equivalent to one molecule of H.sub.2 O for every two OH groups.

Uytterhoeven, Christner and Hall, J. PHYS. CHEM. 69, 2117-26 (1965), have proposed the following stoichiometry to account for protonation and dehydroxylation: ##SPC1##

In an analogous manner, residual NH.sub.4.sup.+ ions can be expected to undergo similar changes when an activated zeolite is ignited at 1800.degree.F. Therefore, whether NH.sub.4.sup.+ groups are intact or have been deaminated, one H.sub.2 O should be detected upon ignition of activated catalyst for every two NH.sub.4.sup.+ ions originally in the zeolite.

Comparing H.sub.2 O contents of activated NH.sub.4.sup.+ zeolite with half the lattice NH.sub.4.sup.+.sup.+ -- on the basis of Uytterhoeven's stoichiometry -- reveals the following differences between calculated and measured H.sub.2 O: Experiment A - activation for zero time at each temperature ______________________________________ H.sub.2 O Difference (g.mole) Temp.,.degree.F. (Experimental-Calculated) ______________________________________ 572 0.107 752 0.071 932 0.046 1112 0.056 Experiment B - activation at 750.degree.F. H.sub.2 O Difference (g.mole) Time, min. (Experimental-Calculated) 0 0.022 60 -0.038 120 0.052 180 0.047 240 0.042 ______________________________________

It is interesting to note that in most instances the excess H.sub.2 O in the preceding table is about numerically equivalent to the residual sodium value of about 0.044 g. ion -- all on the basis of 100 g. anhydrous base. Carter, Lucchesi and Yates, J. PHYS. CHEM., 68, 1385-1391 (1964), described IR bands on NaX zeolite at 3400 and 1655 cm.sup..sup.-1 which presisted up to 450.degree.C. (842.degree.F.) and which they concluded were "due to residual hydrogenbonded `polymeric` water." It is probable that the H.sub.2 O measured over and above that produced from intact and deaminated NH.sub.4.sup.+ sites is this hydrogen-bonded water structurally related to residual Na.sup.+ species in the lattice.

In fact, the NaY zeolite itself may contain more than one kind of H.sub.2 O. For example, if it is assumed that the 0.414 g. ion Na.sup.+ had one mole H.sub.2 O associated - and that this represents strongly bound H.sub.2 O - then the H.sub.2 O equivalent is 7.46 wt. % H.sub.2 O. This firmly bound H.sub.2 O would represent 31 percent of the total 24.32 wt. % H.sub.2 O found (Table 1). With NH.sub.4 Y, about 33 percent of the H.sub.2 O was firmly bound enough to remain after one hour at 450.degree.F.

Activation studies of two ammonium-, cerium-exchanged catalysts were made in a manner similar to those for NH.sub.4 Y.

Cerous nitrate exchange of NH.sub.4 Y catalysts replaced most of the NH.sub.4 .sup.+ ions with cermium but removed only 20-25% of the small residual sodium (Table 2). The Ce.sup.3.sup.+ -exchanged product, therefore, contained residues of NH.sub.4 .sup.+ and Na.sup.+. In contrast to NH.sub.4 Y catalyst, the subsequently Ce.sup.3.sup.+ -exchanged material was able to lose NH.sub.4 .sup.+ during activation down to a level of 0.01 mole/100 g. or less at 750.degree.F. The NH.sub.4 Y had required temperatures above 750.degree.F. to accomplish this degree of removal.

The sum of chemical equivalents for Na.sup.+, NH.sub.4 .sup.+ and Ce.sup.3.sup.+ after exchange was always noticeably less than the 0.414 g. ion Na.sup.+/100 g. anhydrous base found with the original NaY zeolite. One explanation for this cation deficiency of the exchanged catalyst is that some protons are structurally incorporated during exchange but not directly measured by analysis. Chemically this incorporation is possible because the pH of the cerous nitrate solution was about 4.5, and favored cerous salt hydrolysis.

An increase in SiOH groups and a growing cation deficiency was observed as the catalyst became progressively deaminated during activation. However, the SiOH groups and intact NH.sub.4 .sup.+ ions were not enough to account for all of the H.sub.2 O measured by ignition loss at 1800.degree.F. Total H.sub.2 O measured was 0.16 to 0.10 mole/100 g. anhydrous base, but SiOH and NH.sub.4 .sup.+ could not have produced more than 0.05 mole H.sub.2 O on ignition.

The NH.sub.4 Y study showed that residual sodium ions were complexed at a H.sub.2 O/Na.sup.+ ratio of about 1. Continuation of this behavior in the Ce.sup.3.sup.+ -exchanged catalyst could at most produce 0.04 mole H.sub.2 O on ignition. Therefore, the water not derivable from SiOH, NH.sub.4 .sup.+ and Na.sup.+ amounted to nearly half the measured H.sub.2 O. Another source was obviously contributing to the total.

Calculation of (H.sub.2 O)/Ce.sup.3.sup.+ ratios, are shown for two 750.degree.F. activations in Table 3. Calculation of H.sub.2 O measured at 1800.degree.F. does not imply the existence of associated molecular water but is equally capable of interpretation as ce(OH).sup.2.sup.+ species according to the equation: ##SPC2##

This correspondence of measured water to the amount needed for Ce(OH).sup.2.sup.+ formation indicates that substantially all of the cerium in Type Y zeolite activated at 750.degree.F. (400.degree.C.) is in this form or in an equivalent combination of such forms as Ce.sup..sup.+3, Ce(OH).sup.2.sup.+, Ce(OH).sub.2 .sup.+ , Ce.sup..sup.+4 and Ce(OH).sup..sup.+3.

A 750.degree.F. activation with a sample of the same hydrated CeNH.sub.4 Y zeolite in another experiment revealed in H.sub.2 O/Ce.sup.3.sup.+ ratio of 1.041 (Table 4). Continuing this experiment at a series of temperatures up to 1292.degree.F. (700.degree.C.) showed a steady decline of the ratio to 0.370 with increasing temperature. Heating for up to 4 hr. at 750.degree.F. did not lower the H.sub.2 O/Ce.sup.3.sup.+ (or Ce(OH).sup.2.sup.+ /total Ce.sup.3.sup.+) ratio effectively below 1. NH.sub.4 .sup.+ removal occurred during this time, and, thus SiOH increased accordingly - as evidenced by the growing cation deficiency. Ce(OH).sup.2.sup.+, however, was far more stable, and that behavior in itself is more indicative of Ce(OH).sup.2.sup.+ than of Ce(H.sub.2 O).sup.3.sup.+. Only by increasing temperature above 750.degree.F. could the fraction of total Ce.sup.3.sup.+ in the Ce(OH).sup.2.sup.+ state be reduced.

Isoparaffin-olefin alkylations with NH.sub.4 .sup.+ -, Ce.sup.3.sup.+ - exchanged Type Y gave maximum alkylate yields and selectivity when the catalyst had been activated at about 750.degree. F. rather than at lower or higher temperatures, Possibly, maximal Ce(OH).sup.2.sup.+ establishes the sites needed for isoparaffin-olefin alkylation. Also, Ce.sup.3.sup.+ -exchanged Type Y catalysts have been far more stable toward temperatures above 400.degree.C. than NH.sub.4 Y, as measured by X-ray diffraction. When Ce(OH).sup.2.sup.+ sites become dehydroxylated, Ce.sup.3.sup.+ ##SPC3##

When we state that Ce(OH).sup.2.sup.+ sites are preferred for alkylation, we do not means that these sites per se are the sole locus of activity. Rather, these sites form an essential part of a newtowrk or complex of sites, including ##SPC4##

species. Very possibly the entire complex is required to achieve the carbonium ion-olefin combinations accompanied by good hydride transfer which are vital for a highly paraffinic alkylate. Structural considerations further suggest that only a portion of these complexes may be effective for alkylation, even though enough H.sub.2 O for total Ce(OH).sup.2.sup.+ formation is a necessary ingredient of catalyst composition.

As is shown hereinafter, catalytic activity for paraffin-olefin alkylation is related to cerium content. ESR total spin counts of these catalysts with aromatics (such as benzene, perylene, anthracene, etc.) sorbed on them revealed a dependence of electron withdrawl ability upon cerium content. Calculations reveal about 5 percent of the cerium ions to be on the external catalyst surface and the data indicated a 1:1 numerical correspondence between these ceriums and the total spin count.

As is illustrated hereinafter in Example XXIII, there is a nearly linear relation between the total ESR spin count of adsorbed aromatics on our cerium catalyst (when activated at temperatures below about 450.degree.C.) and the alkylate yield under a given set of reaction conditions. In general the more preferred cerium-containing catalysts have total ESR spin counts above 3 .times. 10.sup.19 /g. with anthracene. Thus, there is a correlation between alkylation yield - with its essential dependence upon hydride transfer - and the electron withdrawal from anthrancene by our catalyst.

EXAMPLE VIII

This example illustrates the use of substantially anhydrous acidic crystalline alumino-silicate zeolite as a paraffin-olefin alkylation catalyst. The activated 16-cycle Ce.sup..sup.+3 /16 -cycle NH.sub.4 .sup.+ -exchanged zeolite of Example VII was charged in amount of 23.3 g. into a 1-liter, stirred autoclave containing a four-member baffle to diminish vortex formation. Then 444 milliliters of liquid isobutane was added. The stirring rate (of a six-member, flat-blade turbine) was adjusted such that substantially all of the zeolite was suspended in the liquid isobutane (about 550 rpm). The temperature in the reactor was raised to 80.degree.C. using sufficient nitrogen to produce a total pressure of 250 p.s.i.g. Under these conditions nearly all of the hydrocarbon is in the liquid phase. Then a liquid mixture of one part by volume of butene-2 and five volumes of isobutane was charged from a Jerguson gauge via a needle valve and dip tube into the isobutane-catalyst slurry (and near the bottom of the reactor) at the rate of one milliliter of mixture per minute for a period of 220 minutes. Nearly all of the hydrocarbon was maintained in liquid phase. At the end of this time the reaction was stopped by cooling the reactor to 17.degree.C., then separating the reaction mixture from the catalyst by first removing the normally gaseous hydrocarbons at room temperature and atmospheric pressure, and then separating the liquid product from the catalyst by filtration. The used catalyst analyzed 0.9% coke (non-volatile residue). Some propane and n-butane but no methane, ethane, ethylene or propylene were found in the normally gaseous hydrocarbons. The C.sub.5 .sup.+ paraffin yield of the reaction mixture, based on the weight of olefin charged, was 71.4% and the C.sub.5 .sup.+ unsaturate yield was 0.24 percent on the same basis. Hereinafter all yield data are reported as based on the weight of olefin charged.

EXAMPLE IX

when the reaction of Example VIII was repeated except that the temperature was 120.degree.C. and the pressure 475 p.s.i.g., the C.sub.5 .sup.+ paraffin yield was 129.4% and the unsaturated C.sub.5 .sup.+ hydrocarbon yield was 4.3%.

Table 5 further characterizes the C.sub.5 .sup.+ product obtained in the reactions of Examples VIII and IX.

EXAMPLE X

This example illustrates the unexpectedly large increase in degree of conversion of olefin reactant to saturated C.sub.5 .sup.+ hydrocarbon product when a small amount of a halide adjuvant is present in the reaction mixture. The reaction of Example VIII was repeated at 60.degree.C. except that the catalyst used was the 32/16 zeolite of Example IV which had been prepared by 32 NH.sub.4 .sup.+ -exchange cycles followed by 16 Ce.sup..sup.+3 -exchange cycles. The catalyst was activated by the procedure of Example VII. Without halide addition at 60.degree.C., the yield of C.sub.5 .sup.+ paraffins was 90 percent and the yield of C.sub.5 .sup.+ unsaturates was 11.5 percent. On a mole basis this amounted to 0.44 mole of C.sub.5 .sup.+ paraffins per mole of C.sub.4 olefin charged. In contrast, when 2.4 .times. 10.sup..sup.-3 mole of tertiary butyl chloride (hereinafter, sometimes, TBC) was added to the reactor for each mole of initial isobutane, the yield of C.sub.5 .sup.+ paraffins was 120 percent and C.sub.5 .sup.+ unsaturates 6 percent. The TMP/DMH.sub.x ratio was 6.6. The used catalyst had no measurable coke content.

EXAMPLE XI

With the same proportion of t-butyl chloride the reaction of Example X was repeated at 40.degree.C. (125 p.s.i.g.) and at 25.degree.C. (125 p.s.i.g.). At 25.degree.C. only 12 percent of C.sub.5 .sup.+ paraffins was produced, and 0.49 percent of C.sub.5 .sup.+ unsaturates. At 40.degree.C. 120 percent of C.sub.5 .sup.+ paraffins was produced and 6.5 percent of C.sub.5 .sup.+ unsaturates. The TMP/DMH.sub.x ratio was 4.10 at 25.degree.C. and 7.86 at 40.degree.C.

EXAMPLE XII

When Example X was repeated at 120.degree.C. (484 p.s.i.g.) without halide addition, 75 percent of C.sub.5 .sup.+ paraffins and 1.1 percent of C.sub.5 .sup.+ unsaturates were produced. The TMP/DMH.sub.x ratio was 3.16.

Of commercial importance is the finding that, by practice of our invention, we cannot only obtain good yields of alkylate which has a high TMP/DMH.sub.x ratio and is high in trimethylpentanes but that in these trimethylpentanes there is a low proportion of the less desirable 2,2,4-trimethylpentane (regarding this undesirability, see U.S. Pat. No. 2,646,453). Table 6 compares the percent of the total trimethylpentanes which is 2,2,4-trimethylpentane (percent 2,2,4 in TMP) in the products of Examples X, XI and XII with similar distributions reported by Cupit, C.R. (Ibid. p. 211) for H.sub.2 SO.sub.4 alkylation and Kennedy, R.M. (Ibid. p. 30) for HF or AlCl.sub.3 alkylation of isobutane with butene-2. For the compositions of the products obtained with other olefins (e.g., propylene, cyclohexane) and H.sub.2 SO.sub.4 catalyst see J. E. Hofmann, J. ORG. CHEM., 29 (Part II), 1497-1499 (1964).

As is further illustrated herein, our process can be used to directly product novel paraffin-olefin alkylates, useful in gasoline blending, comprising at least 60 mole percent C.sub.8 paraffins and less thann one weight percent unsaturates and wherein the C.sub.8 paraffins consist of from 5-20 mole percent dimethylhexanes, from 0-1.5 mole percent methylheptanes, from 80-95 mole percent trimethylpentanes, and wherein less than 30 mole percent of the trimethylpentanes is 2,2,4-trimethylpentane. Such novel alkylates can also be produced using hydrogenation and/or adsorbents to reduce the unsaturates in such products of our process as that of Example X which is cited in Table 6.

In general, in the temperature range from about 25.degree.C. to 120.degree.C., the proportion of 2,2,4-TMP in the total TMP's decreases (and usually the proportion of 2,3,3-TMP increases) as the reaction temperature decreases.

Product distributions in the alkylates produced by our 32/16 catalyst (and in virtually all of the alkylates reported herein as produced by our process) are far removed from calculated equilibrium values. For example, the calculated equilibrium TMP among C.sub.8 paraffins at about 60.degree.C. is only 12%l whereas, in our process TMP usually constitute more than 80 percent of the C.sub.8 paraffins. Among the TMP there is a similar departure from calculated equilibrium, as shown by the following mole percent data at 60.degree.C.:

Experimental Calculated TMP With TBC No TBC Equilibrium ______________________________________ 2,2,4- 14.8 40.4 65.4 2,2,3- 5.0 10.7 15.0 2,3,3- 44.3 27.3 8.3 2,3,4- 35.9 21.7 11.3 ______________________________________

These departures from equilibrium suggest that either the published free energies for the equilibrium calculation are seriously in error or else that our alkylation reactions are kinetically controlled under the conditions of our process.

Kinetic control of product distribution through specific reactions is considered more likely for two reasons:

1. TBC promoter with solid catalyst had an evident influence upon TMP distribution which is not characteristic of equilibrium control.

2. Equilibrium calculations predict that the fraction of 2,2,4-TMP decreases with rising temperature. With solid catalyst, it is increased as temperature went up.

Alkylation with solid zeolite catalyst appears to be the result of a sensitive balance among competing kinetic paths. Our halide adjuvant, such as TBC, favorably alters one or more of those paths.

Surveying temperatures with a 16/16 catalyst (under conditions as in Examples VIII but with improved feed premixing) similarly points out the alkylate yield gain to be realized by operating at an intermediate temperature of about 80.degree.C. (Table 7). In a continuous reactor residence time must be taken into account; the same temperature with a given catalyst may not be preferred if residence time is changed.

It is interesting to note that 60.degree.C. with the 16/16 catalyst produced no more than half as much alkylate as 80.degree.C. with this catalyst. The 32/16 catalyst was not so responsive to temperature changes above 40.degree.C. Again, as temperature decreased from 80.degree. to 60.degree.C., a shift toward a heavier product occurred (C.sub.9 .sup.+), but the 2,2,4-TMP content was desirably low. When this isomer decreased, the largest gain was in 2,3,3-TMP, as it has been with the 32/16 catalyst.

Temperature is a useful device in elucidating catalyst differences. When the catalyst exchanged only with NH.sub.4 .sup.+ (32/0) was tested at 120.degree.C. without TBC promoter, it was less active than a 16/16 catalyst (Table 8). The C.sub.5 .sup.+ paraffin yields were 109.5%, based on olefin charge, with 32/0 and 129.4 percent with 16/16.

Evaluating an NH.sub.4 .sup.+-exchanged catalyst (16/0) with TBC relative to 16/16 at 80.degree.C. revealed a more dramatic difference in alkylate yield and product distribution.

The 16/0 catalyst produced too little C.sub.8 paraffin and too much C.sub.9 .sup.+ and C.sub.5. These factors could also be used to understand the importance of a polyvalent metal, such as cerium, on the catalyst. But testing catalysts at milder conditions is even more effective in uncovering differences between them, as shown by the data from 32/0, 16/16, and 16/0. Therefore, low operating tempertures can be used as a research tool in distinguishing among alkylation catalysts that appear to be more similar at relatively high temperatures.

EXAMPLE XIII

Table 5 reports the products obtained from similar runs at 8.degree.C. using the activated catalyst of Example VIII with t-butyl, chloride, n-propyl chloride or n-butyl chloride as adjuvants (at a level of 2.4 .times. 10.sup..sup.+3 mole of adjuvant per mole of initial isobutane).

Table 9 reports the products obtained from similar runs (but with more intimate premixing of the feed olefin and feed paraffin) using CCl.sub.4, TBC and various other adjuvants and using either continuous or "pulsed" addition of the adjuvant to the reaction mixture. In this table, the amount of adjuvant is reported as millimoles per mole of feed olefin charged (m.mole/m. OC).

In run 606 of Table 9, the catalyst was preconditioned by contact with a solution of perylene in CCl.sub.4. The perylene was quantitatively adsorbed by the catalyst along with some CCl.sub.4. The catalyst developed a dark, intense, blue color upon contact with the perylene solution. Removal of residual CCl.sub.4 by vacuum-pumping at ambient temperature caused the catalyst color to turn to black. This black catalyst was the catalyst used in run 606.

In run 600, the catalyst was preconditioned with carbon tetrachloride as a control experiment for 606. The catalyst developed an intense red color on contact with the CCl.sub.4. Upon vacuum pumping, the red color disappeared. It is this "pumped" catalyst which was used in run 600.

Potential catalyst adjuvants are those halides, both organic and inorganic (e.g., AlBr.sub.3, BF.sub.3, HBCl.sub.2, AsCl.sub.3), which are capable, under the reaction conditions, of sufficient polarization to promote carbonium ion reactions or to have carboniogenic properties. For precise control of the reaction product distribution (or alkylate quality) and to prolong catalyst life, we prefer to avoid adjuvants which contain atoms other than hydrogen, carbon, bromine, fluorine and chlorine (although as seen in run 632, oxyen, as in the form of alcoholic OH groups, can be present in reaction mixture. Water, C.sub.1 to C.sub.10 saturated alcohols (e.g., tertiary butyl alcohol, cyclohexanol) or mixtures thereof can be used, per se, as adjuvants or in combination with halides. To avoid accumulation of large organic molecules at the catalyst surface, we prefer to avoid those organic halides wherein the organic radical has a critical diameter greater than about 9A, such as the chlorinated naphthenic waxes. Note, however, in Table 9, that perylene presorbed on the catalyst from CCl.sub.4 solution did not act as a "poison" but allowed about 10 relative percent more C.sub.5 .sup.+ paraffin yield than a control experiment with carbon tetrachloride alone. This carbon tetrachloride control experiment itself produced a better than 10 relative percent increase in C.sub.5 .sup.+ paraffin yield over a similar experiment with tertiary butyl chloride and without CCl.sub.4. In contrast, NH.sub.3 presorbed on the catalyst acted as a poison, even when TBC was added continuously to the reactor.

Our preferred halide adjuvants, when present in solution in the reaction mixture at a level of from 1 .times. 10.sup..sup.-5 to 1 .times. 10.sup..sup.-1 mole per mole of C.sub.4 -C.sub.6 isoparaffin reactant, are HF, HC1, HBr and the saturated halohydrocarbons containing at least one atom per molecule of bromine, chlorine or fluorine. Mixtures of these substances can also be used as adjuvants. Of these adjuvants we prefer carbon tetrachloride and the aliphatic saturated monochlorides having no more than six carbon atoms. When the isoparaffin reactant is predominantly isobutane, we prefer to use, as halide adjuvants, the aliphatic saturated monochlorides having 3 or 4 carbon atoms.

The adjuvant can also be added to the catalyst after the final washing, in the exchange procedure but, more preferably, is added to catalyst after activation, as by passing gaseous HC1 through the catalyst at the final stage of activation (or while cooling catalyst after activation). It can be important, especially in our continuous process, to control the amount of adjuvant present in the reactor vapor space since the vapor pressure of such adjuvant affects the adjuvant concentration in the reaction mixture.

EXAMPLE XIV

This example illustrates, at a given feed rate, the influence of "reaction time" or, more precisely, a combination of catalyst/feed contact time and residence time (of the paraffin product) on the yield of C.sub.5 .sup.+ saturates and C.sub.5 .sup.+ unsaturates. This reaction time, which combines residence and contact time, reflects both kinetic influence and the age of the catalyst-reactant system.

The process of Example VIII was repeated at 80.degree.C., 250 p.s.i.g., using the activated catalyst of Example VIII, with 2.4 .times. 10.sup..sup.+3 mole of tertiary butyl chloride present in the reactor initially per mole of initial isobutane (444 ml.). As in Example VIII the rate of addition of the butene-2/isobutane feed was 1 milliliter per minute. Four runs were made at various contact times (60, 120, 220 and 280minutes). In each run contacting was stopped by rapidly cooling the samples to 15-20.degree.C., then slowly reducing the pressure to atmospheric and simultaneously distilling off lighter gases. The remainder of the reaction mixture, which was liquid at that temperature, was separated from the solid catalyst by filtration.

FIG. 1 illustrates the variation in the yield of C.sub.5 .sup.+ paraffins based on the olefin reactant as the reaction time increased.

FIG. 2 illustrates, by the solid curve, the weight percent of C.sub.5 .sup.+ unsaturates produced, based on the olefin reactant, as the reaction time increased. The broken curve, of FIG. 2, show the weight percent of n-butane produced per mole of olefin converted, as the reaction time increased.

From FIGS. 1 and 2, it can be seen that after 2 hours only negligible amounts of C.sub.5 .sup.+ unsaturates were found in the reaction mixture and the yield of C.sub.5 .sup.+ paraffins was 96 percent of the weight of olefin charged. Of the C.sub.5 .sup.+ paraffins, 60 mole percent was C.sub.8 and of the C.sub.8 paraffins there was 0 mole percent methylheptanes. The ratio TMP/DMH.sub.x was 7.18. Of the trimethylpentanes, 24.7 percent was 2,2,4-TMP. At this point a total of 20 milliliters of butene-2 had been added to the reactor along with an additional 100 milliliters of isobutane. When added to the original 444 milliliters of isobutane this amounted to a total of 544 milliliters of isobutane which had been charged to the reactor at that time along with 20 milliliters of butene-2 (density 0.60), which produced 11.5 grams of C.sub.5 .sup.+ paraffin.

After 220 minutes, 37 ml. of butene-2 had been charged to the reactor along with 183 ml. of isobutane to give a total hydrocarbon charge of 664 ml. The C.sub.5 .sup.+ paraffin yield was 142 percent based on the weight of olefin charged (37 ml.) or a total C.sub.5 .sup.+ paraffin yield of 31.6 grams. Of the C.sub.5 .sup.+ paraffins, 60 mole percent was C.sub.8. Of the C.sub.8 paraffins, 1 mole percent was methylheptane. The ratio TMP/DMH.sub.x was 4.98. Of the trimethylpentanes, 20.9 percent was 2,2,4-TMP.

After 280 minutes a total of 724 ml. of hydrocarbon had been charged to the reactor of which 47 ml. was butene-2. The C.sub.5 .sup.+ paraffin yield was 110 percent of the weight of olefin charged or 31.1 grams. The ratio TMP/DMH.sub.x was 5.44 and 21.0 percent of the trimethylpentanes was 2,2,4-TMP. At 220 minutes 0.84 grams of C.sub.5 .sup.+ unsaturates had been detected in the reaction mixture or 3.75 percent based on the olefin charged. At 280 minutes 1.6 grams of C.sub.5 .sup.+ unsaturates were detected in the reaction mixture, which amounted to 5.5 percent of C.sub.5 .sup.+ unsaturates based on the weight of olefin charged.

An inspection of FIGS. 1 and 2 shows that at point B of FIG. 1 and point B' of FIG. 2, the net rate of production of unsaturated hydrocarbon is about 2 weight percent C.sub.5 .sup.+ unsaturates/olefin charged/hour but the rate of production of saturated hydrocarbon is about 100 weight percent C.sub.5 .sup.+ saturates/olefin charged/hour. At point C in FIG. 1 and C' in FIG. 2, the net weight rate of production of unsaturated hydrocarbon becomes greater than the net weight rate of production of saturated hydrocarbon per weight of olefin charge.

The broken line in FIG. 2 shows that there is a high initial production of n-butane but that by point A" (corresponding in time to point A' of the solid curve) the proportion of n-butane as a function of time had become nearly constant. This behavior is understandable if n-butane is formed as a result of olefin protonation and subsequent hydride transfer (which is considered necessary for initiating alkylation).

If one wishes to maximize the production of C.sub.5 .sup.+ paraffin per olefin charged the reaction should be stopped at the point corresponding to the letter A in FIG. 1, at which point the C.sub.5 .sup.+ product of the reaction will contain about 3.6 percent of unsaturated hydrocarbons (see point A' of FIG. 2).

However, if one wishes to have a substantially olefin-free C.sub.5 .sup.+ product, the reaction would be stopped in the vicinity of point B' in FIG. 2. In the latter case, the catalyst life can be greatly prolonged in comparison with operation between points B and A of FIG. 1 or B' and A' of FIG. 2.

Similar relations were observed at 60.degree.C. and 120.degree.C. At 120.degree.C. (without halide addition) 28.7 grams of C.sub.5 .sup.+ paraffin were obtained after 220 minutes (129.4 weight percent yield/olefin charged) while only 23.3 grams were obtained after 384 minutes of reaction time (104.8% yield). When the weight of catalyst used at 120.degree.C. (without adjuvant was doubled, 22.5 grams of C.sub.5 .sup.+ paraffin were obtained after 220 minutes (101.4% yield).

Catalyst life can be greatly prolonged at conversion ratios approaching that at point B by constantly separating a catalyst-free alkylate and a concomitant amount of unreacted feed from the reaction zone while constantly adding an approximately equal volume of fresh portions of the hydrocarbon reactants. This constant separation and withdrawal of alkylate and unreacted feed in conjunction with the addition of fresh reactants (including recycle of unreacted feed) can be accomplished continuously by adding a steady stream of reactants and withdrawing a steady stream of the mixture, as by utilizing a continuous stirred reactor system such that of FIGS. 3, 4 and 5.

Surprisingly, in view of U.S. Pat. No. 3,251,902, the degree of conversion of butene-2 to paraffins is very high in our process. For example, in the above runs, analysis of the reaction mixture for unreacted butene-2 showed that in the 60-minute and 120-minute runs the butene-2 conversion was 100%. In the 220-minute run 92.6 percent of the feed butene-2 was converted and 88.5% was converted in the 280-minute run.

The most important consideration in continuous operation is to coordinate the rate of olefin addition with the rates of feed olefin consumption and removal in order that the amount of unreacted olefin in the reaction mixture is maintained at less than 12 mole percent based on the unreacted C.sub.4 -C.sub.5 isoparaffin, and preferably less than about 7 percent. Also, the preferred mean residence time of the hydrocarbons in the reaction mixture, with the catalyst, is in the range of 0.05-0.5 hour per gram of catalyst per gram of hydrocarbon in the reaction mixture.

In this respect, the preferred procedure is to thoroughly premix the feed olefin and feed paraffin. The uniformity and intimacy of such premixing can greatly influence the character of the C.sub.5 .sup.+ product. In Table 10, for example, two runs are shown which were identical except for the feed premixing technique. In one run the feed was introduced through the bottom of the Jerguson gauge. This increased the uniformity and intimacy of the premixing (by reducing charge segregation or layering). The resulting C.sub.5 .sup.+ product contained only 0.26 percent unsaturates and was also lower in C.sub.9 .sup.+ paraffins, higher in pentanes and had a higher TMP/DMH.sub.x ratio than the product from the corresponding run where the feed was introduced at the top of the mixing buret (which, unless otherwise noted, was the technique used in all the other examples reported herein). Note that the product from the run with the feed introduced at the top of the mixing buret had a 685 percent greater concentration of unsaturates than the product from the former run where the premixing was more uniform and more complete. In these examples the runs utilizing the improved feed premixing (from the bottom of the buret) are so noted or are numbered in the range of 502-698 and 804-898.

For any given type of activated catalyst and feed hydrocarbon, the rate of olefin consumption and, correspondingly, the rate of olefin addition, will be a function of the reaction temperature, the mean retention time of feed olefin in the reactor, the mixing rate, the particle size and concentration of catalyst, and the rate of product removal. One method of controlling such a "continuous" process, under conditions of good feed olefin mixing, is to control the rate of feed addition and the rate of reaction mixture withdrawal such that the C.sub.5 .sup.+ component of the reaction mixture contains substantially no C.sub.5 .sup.+ unsaturates and, preferably, such that there is little unreacted feed olefin in the withdrawn portion.

Another important variable to be considered in our process is the proportion of the reactants (particularly of the feed olefin) which can be present in the vapor space of the reactor. The proportion of olefin in the vapor space is a function of the vapor pressure of the olefin at the reaction temperature and of the degree of reactor filling (that is, the vapor space in the reactor). In our continuous stirred reactor system of FIGS. 3, 4 and 5 the degree of reactor filling can be precisely controlled, as by means of the differential pressure cell.

One convenient means of stopping the contact of the olefin-isoparaffin feed with the zeolite catalyst and effecting the concomitant removal of a portion of the C.sub.5 .sup.+ product from the reaction mixture is to submerse a line, the opening of which is covered by a filter device, such as a very fine screen, into the reaction mixture and to constantly withdraw a catalyst-free portion of the reaction mixture from the reactor to a zone where unreacted feed olefin (if present) and feed isoparaffin are separated by distillation from the C.sub.5 .sup.+ hydrocarbons, and recycled to the reactor. This means can be utilized in the reactor section (FIG. 4) of our continuous stirred reactor system (of FIGS. 3, 4 and 5) for producing an olefin-paraffin alkylate.

In the reaction illustrated by FIGS. 1 and 2, the mean residence time in hours (per gram of catalyst per gram of hydrocarbon) of the hydrocarbons in the reaction mixture with the catalyst was 0.089 after the first 60 minutes, 0.167 after 120 minutes, 0.236 after 180 minutes and 0.297 after 240 minutes. This is to be contrasted with 0.62 hours in Example II, Table II, Column 3, 1.25 hours in Example VII and 1.47 hours in Example I, Table I, Column 1 of U.S. Pat. No. 3,251,902, previously cited herein.

In our process the preferred mean residence time is in the range of 0.05-0.45 hour, more preferably 0.1 to 0.4 hour.

An illustration of the calculation of mean residence time, for the first 60 minutes in the reaction illustrated by FIGS. 1 and 2 herein, follows:

(444 ml. i-butane)(0.5543) = 246.11 g. isobutane for entire time

23.3 g. of catalyst

Change 1 vol. butene-2 (density = 0.5988 g./ml.)

5 vol. isobutane (density = 0.5543 g./ml.)

(6)(D) = (5)(0.5543) + (1)(0.5988)

= 2.7715 + 0.5988 = 3.3703

d* = density of hydrocarbon mixture = 0.5617

For 60 min.

(1 hour)(23.3 grams catalyst 246.11 + (60 min.)(1 ml./min.)(0.5617 g.ml.) =

0.08861 hr./(g. Hydrocarbon)

(g. Catalyst)

EXAMPLE XV

This example illustrates the effect that catalyst composition has on the yield of C.sub.5 .sup.+ reaction product and on the product distribution, in particular with regard to the proportion of C.sub.8 paraffins and the distribution of these C.sub.8 paraffins into trimethylpentanes and dimethylhexanes.

The process of Example VIII was repeated except that the reaction temperature was 120.degree.C. (which was close to the critical temperature of the reaction mixture), the reaction pressure was 500 p.s.i.g., and the reaction time was 3.67 hours. Separate runs were made with equal weights (activated basis) of zeolites of varied Na, H and polyvalent metal contents, which were prepared similarly to the catalysts of Examples II, III, IV and V.

Runs were also made, at 80.degree.C., 250 p.s.i.g., and 2.4 .times. 10.sup.-.sup.3 moles t-butyl chloride per mole of initial i-butane, with catalysts prepared from the following: the 1.72 percent (ignited) Ce zeolite of Example VII; a 16-cycle ammonium-exchanged NaY zeolite which was further exchanged with 16 cycles of a 13.3 g./1. aqueous solution of La(NO.sub.3).sub.3.sup.. 6H.sub.2 O; a 16-cycle ammoniumexchanged NaY zeolite which was further exchanged with 16 cycles of a 13.3 g./l. aqueous solution of hydrated mixed rare earth nitrates (approximate salt analysis, 48% Ce.sub.2 O.sub.3, 24% La.sub.2 O.sub.3, 17% Nd.sub.2 O.sub.3, 5% Pr.sub.2 O.sub.3, 3% Sm.sub.2 O.sub.3, 2% Gd.sub.2 O.sub.3); and, a 16-cycle ammoniumexchanged NaY zeolite which was further exchanged with 16 cycles of aqueous Ce(NO.sub.3).sub.3.sup.. 6H.sub.2 O (as in Example IV).

All of these catalysts were activated by the procedure of Example VII.

The yields of the C.sub.5 .sup.+ paraffin and C.sub.5 .sup.+ unsaturates, based on the weight percent of olefin charged, the C.sub.5 .sup.+ paraffin distribution and the C.sub.8 paraffin distribution of the products are shown in Table 11.

The yields and product distributions shown in Table 11 indicate that, in substantially anhydrous acidic crystalline alumino-silicate zeolites which have been prepared by ammonium exchange of sodium zeolites with ammonium ions and polyvalent metal ions, the catalytic activity and selectivity in paraffinolefin alkylation are dependent upon the amount and type of exchanged polyvalent metal and the degree of "protonation" or "cationic deficiency" (which is related to the nitrogen content before activation). Therefore, when other reaction variables are fixed, an appropriate selection of the catalyst can be used to vary the yield and product distribution in our process.

Table 12 and Table 13 illustrate the effect on the ultimate catalyst of the type of salt used in the exchange solution.

It is evident from Table 13 that the yield differences are not determined only by the total amount of rare earth metal present. Therefore, it appears that different cations and their accompanying anions can have pronounced effects on catalyst performance. Other desirable catalysts can be prepared by exchanging NH.sub.4 Y zeolite with salts of Gd.sup..sup.+3, Dy.sup..sup.+3 and Sm.sup..sup.+3.

One precaution to be taken with data from Table 13 concerns the apparent gain in selectivity for C.sub.8 paraffins with the La(NO.sub.3).sub.3 and CeCl.sub.3 catalysts. In fact, this gain is more in line with the selectivity gain which is typical when our process is operated at a relatively low degree of reactant conversion or product yield. In other words, if the Ce(NO.sub.3).sub.3 catalyst had been used to produce only 68 to 73% C.sub.5 .sup.+ paraffin yield (the range for La(NO.sub.3).sub.3 and CeCl.sub.3), the molar C.sub.8 paraffin content of the C.sub.5 .sup.+ paraffins would have increased to about 80 percent instead of remaining at the 69.0 percent actually observed at 132.0 percent C.sub.5 .sup.+ paraffin yield.

As shown by these data, the anion in the exchange solution exerts an influence on catalyst performance. The effect is related to the condition of metal cations in aqueous solution as a function of anion, cation concentration, pH and temperature. An effect such as the following is the probable cause:

[RE(H.sub.2 O).sub.n ].sup.3 .sup.+ + X.sup.(m).sup.- .revreaction. [RE(H.sub.2 O).sub.n-1 X].sup.(3.sup.-m).sup.+

Other cations which can affect the catalyst are the alkali metals, such as lithium, sodium, potassium and cesium. As shown in Table 14, at comparable sodium levels, C.sub.5 .sup.+ paraffin yield progressed from 26 to 132 wt. % olefin charge for an increase of cerium from 2.0 to 13.5%. Even at 8.3% cerium, the C.sub.5 .sup.+ paraffin yield was only 62.7% on the same basis. The probability that the 1.68% sodium content did not have the principal deleterious effect upon this 62.7 % yield is supported by the 118.9% yield for a catalyst containing 2.8% sodium but 13.0% cerium and by the 115.4% yield for another catalyst with a 1.68% sodium and a 12.8% cerium content.

Some gain in C.sub.5 .sup.+ paraffin yield (115.4 to 132.0) can be inferred for a reduction in sodium content from 1.68 to 0.76%.

Selectivity effects of cerium are illustrated by the relatively high C.sub.5 .sup.+ unsaturate production with catalysts containing less than about 12 percent cerium. Trimethylpentane/dimethylhexane (TMP/DMX.sub.x) ratios were also comparatively low for those catalysts, and relatively undesirable C.sub.9 .sup.+ paraffins constituted as much as 27.2 mole % of the total C.sub.5 .sup.+ paraffins for the lowest cerium catalyst. These data show that with less than about 12 percent cerium, alkylate will be not only lower in yield but also poorer in quality.

A series of NH.sub.4 .sup.+-, Ce.sup.3 .sup.+-exchanged catalysts having very similar sodium levels clarified the essential role of cerium in producing favorable yields of high quality alkylate.

When cerium replaced ammonium on a Type Y zeolite at constant sodium level, the following effects were observed:

1. Appreciable gains were realized in C.sub.5 .sup.+ paraffin yield, in relative proportion of C.sub.8 paraffins, and in selectivity for trimethylpentanes (TMP/DMH.sub.x ratio).

2. Simultaneously, desirable decreases were found in C.sub.5 .sup.+ unsaturates and in the relative proportion of C.sub.9 .sup.+ paraffins.

3. The only undesirable trned was an increase in the relative amount of 2,2,4-TMP up to 26.4 mole % of the total TMP. However, typical sulfuric acid alkylates have 2,2,4-TMP contents above 40 percent. This isomer has the lowest F-1 octane number of all the TMP.

4. the largest gains in yield and selectivities occurred at values of (Ce.sup.3 .sup.+/NH.sub.4 .sup.+) equivalent ratio below about 2.5. Higher ratios are desirable, but corresponding product improvements become smaller.

These catalysts were prepared from the same common lot of NH.sub.4 .sup.+-exchanged Type Y zeolite. The following analytical data establish that Ce.sup.3 .sup.+ was exchanging for NH.sub.4 .sup.+ and that no net loss of Na.sup.+ occurred from the NH.sub.4 .sup.+-exchanged zeolite:

Analysis Catalyst No. (g. equivalent/100 g. anhydrous residue) ______________________________________ Na NH.sub.4.sup.+ Ce.sup.3.sup.+ * .SIGMA. ______________________________________ FX10 0.064 0.389 -- 0.543 FX10-1-2 0.054 0.190 0.175 0.419 FX10-1-3 0.047 0.102 0.253 0.402 FX10-1-4 0.047 0.049 0.299 0.395 ______________________________________ * Average of 3 analyses

The original zeolite had a sodium content of 0.426 equiv./100 g. anhydrous residue after correction for 1,800.degree.F. ignition loss. Residual sodium content was thus 11-13 percent of the original.

Another interesting but undecided aspect of these catalysts is their growing cation deficiency as cerium exchange increases. A deficiency is said to occur when the sum of residual sodium, ammonium and rare earth does not equal the positive charged initial sodium. The presence of protons-bound or "solvated" -can account for the apparent deficiency.

As has been shown in Examples I to VII, we prefer to prepare the substantially anhydrous acidic alumino-silicate zeolites by controlled activation of zeolites which are prepared from crystalline sodium zeolites by first exchanging the bulk of the sodium with ammonium ions and then exchanging the resulting zeolite, which is low in sodium and high in ammonium ions, with solutions of polyvalent metal cations. When the base zeolite is sodium Y, the ammonium-exchanged zeolite should contain, on an ignited basis, less than 3% Na and preferably less than 1.0% Na.

In our ammonium exchange we also prefer that the sodium content of the exchange solution be kept as low as is practicable. One means of removing sodium ions from ammonium salt solutions is by a separate cation exchange of the solution with a bed of ammonium-containing ion-exchange resins or noncrystalline ammonium zeolites. In this sodium-ion removal step, which is particularly advantageous in continuous ammonium exchange (as in the procedures of Example II), the sodium ion in the ammonium-ion exchange solution exchanges with the ammonium ion in the resin and the resulting ammonium-rich solution is recycled to the vessel containing the crystalline zeolite for additional exchange with the sodium in the zeolite. The ion-exchange resin bed (or noncrystalline zeolite bed) can be regenerated by contacting the ammonium-sodium equilibrium resin with an ammonium-rich stripping stream. The sodium-rich effluent from the regeneration is discarded after, if desired, residual ammonia has been recovered by flash distillation.

Products obtained from a preferred Gd catalyst and from two other, less preferred, catalyst types are shown in Table 15. One of the two less preferred catalysts was obtained by activation (as in Example VII but with 8 hours at 400.degree.C. to insure good NH.sub.3 removal) of a highly (16 cycles) ammonium-exchanged type Y zeolite (to produce HY catalyst). The HY catalyst produced only about one-fourth as much alkylate, together with more C.sub.9 .sup.+ and C.sub.5 and less C.sub.8, as its cerium counterpart.

The other less preferred catalyst was prepared by activation of a 16-cycle cerium exchanged, 16-cycle ammonium-exchanged sodium X zeolite (to produce CeHX catalyst). In comparison with CeHY catalyst (run 664) the CeHX catalyst produced an appreciably smaller C.sub.5 .sup.+ paraffin yield. An Analysis of this paraffin product showed 23.9 mole % to be isopentane (which is 2 to 4 times isopentane usually found in alkylate produced by CeHY catalyst). Accordingly, the C.sub.8 paraffin in the alkylate produced by the CeHX was only 59 mole % compared wth about 70% for CeHY.

Runs 628 and 674 were made with catalysts prepared by an exchange procedure similar to that of Example IV and activated as in Example VII (except that for the run 674 catalyst helium was substituted for air), but wherein gadolinium nitrate was used instead of cerium nitrate in the exchange solution. The resulting novel Gd-alumino-silicate, upon activation, produced a novel catalyst which is very useful for hydrocarbon conversion reactions, particularly in our process for paraffin-olefin alkylation.

EXAMPLE XVI

This example shows the effect of feed olefins other than butene-2 on the yield of C.sub.5 .sup.+ products and their distribution. Example VIII was repeated, with a similarly prepared catalyst, except that the feed olefin was butene-1. The C.sub.5.sup.+ paraffin yield and the C.sub.5.sup.+ unsaturate yield were about the same as those obtained in Example VIII with butene-2 and the distribution of C.sub.5.sup.+ paraffins and the C.sub.8 distributions (see Table 10) were similar to those obtained in Example IX with butene-2.

We have found that butene-1, in the presence of n-butane, is readily isomerized to cis and trans butene-2 under our alkylation conditions with acidic zeolite catalysts. This ready isomerization provides the explanation for the similarity between the products obtained when isobutane is alkylated with butene-2 and the products obtained when butene-1 is the feed olefin. Surprisingly, a significant quantity of highly saturated C.sub.5.sup.+ liquid product is also obtained from this lilquid phase isomerization of butene-1 in the presence of n-butane. A very small amount of isobutane was also detected. An analysis of the C.sub.5.sup.+ liquid product of one such run is shown in Table 10. At least some of this C.sub.5.sup.+ liquid appears to be the result of the combination of the n-butane and the C.sub.4 olefin.

A similar run was made using 2-methylbutene-2 as the feed olefin and a catalyst, prepared in a manner similar to the catalyst of Example III, prepared from a zeolite which before activation analyzed 5.0 percent cerium, 1.19 percent sodium, and had a loss on ignition of 25.65%. The catalyst was activated (final temperature 400.degree.C.) as in Example VII. The C.sub.5.sup.+ paraffin yield was 28.6 percent and the C.sub.5.sup.+ unsaturate yield was 31.2%, based on the weight of olefin charged. The molar ratio C.sub.8 /C.sub.9 of the C.sub.5.sup.+ paraffins was 1.00.

A similar run with a somewhat more acidic catalyst produced a C.sub.5.sup.+ paraffin yield of 49.0 percent and 15.4 percent C.sub.5.sup.+ unsaturates based on the weight of olefin charged. Of the C.sub.5.sup.+ paraffins 29 mole percent were C.sub.9.sup.+ paraffins and 36.6 mole percent were C.sub.8 paraffins (molar ratio C.sub.8 /C.sub.9 was 1.25). The presence of C.sub.8 paraffins indicates that self-alkylation of isobutane occurred during the reaction.

A similar run using a portion of the same catalyst (activated at 500.degree.C.) and a butene-2 feed resulted in a C.sub.5.sup.+ paraffin yield of 51.8 percent of which 82.8 mole percent was C.sub.8 and 4.2 mole percent C.sub.9.sup.+ paraffins. The distributions of the C.sub.5.sup.+ paraffins in these two products are shown in Table 10.

Similarly when isobutylene or propylene or a "B-B" refinery stream (i.e., a mixture of butanes and butenes containing a minor amount of propylene) is the feed olefin, good yields can be obtained (of a product in which the C.sub.5.sup.+ saturates predominate over the C.sub.5.sup.+ unsaturates) by stopping contact of the catalyst with the reaction mixture after substantial alkylation has occurred but before the weight rate of production of C.sub.5.sup.+ olefins becomes greater than the weight rate of production of C.sub.5.sup.+ paraffins.

In general, our process can be used to produce (either by alkylation, self-alkylation or both) good yields of C.sub.5.sup.+ saturated hydrocarbons from C.sub.4 -C.sub.6 isoparaffin (or mixtures thereof) and any monoolefin having from 3 to 9 carbon atoms, including the cyclic olefins, such as cyclohexene, and mixtures of such monoolefins.

Table 16 reports a typical product obtained from isobutane at 60.degree.C. in our process under reaction conditions similar to those of Example VII when propylene is the feed olefin, with and without halide adjuvant. Table 17 details similar products obtained in our process from isobutane and propylene under varied reaction conditions.

Note that under the conditions of Run 544 propylene-isobutane alkylation produced a C.sub.5.sup.+ alkylate yield of 123.3 wt. % based on olefin charge. This number corresponds to 0.513 mole of C.sub.5.sup.+ paraffins per mole of olefin charged to the reactor. Close examination of the product revealed 66.0 mole % C.sub.7 among the C.sub.5.sup.+ paraffins, and 97.2% of the C.sub.7 was 2,3-dimethylpentane.

From these results and other such studies the following observations can be made:

1. In such batch reactions alkylation at 60.degree.C. for 2 hr., with attendant increases in relative amounts of promoter and catalyst, produced the highest alkylate yield.

2. Raising or lowering temperature to 80.degree. or 40.degree.C., decreasing isoparaffin/olefin ratio (15/1 to 10 or 5/1), increasing contact time at 60.degree.C. (2 to 4 hr.), and allowing more olefin to be in the vapor phase-all had adverse effects upon yield and selectivity. Low alkylate yields invariably were accompanied by relatively high amounts of C.sub.9.sup.+ paraffin and low quantities of C.sub.7.

3. Although isobutane self-alkylation is strongly indicated by C.sub.8 paraffin in the product, the distribution of trimethylpentanes is even more convincing. The 2,2,4-TMP content of the C.sub.8 fraction was 61 to 70 percent for all propylene-isobutane experiments; with butene-1 or 2 and isobutane it is normally 15 to 25 percent for exchanged Type Y catalysts. Mechanistic considerations show 2,2,4-IMP as a likely initial product of self-alkylation. The reaction sequence can be formulated in the following way: ##SPC5##

4. Comparison of Run 544 with published data on continuous H.sub.2 SO.sub.4 alkylation of propylene with isobutane shows the much higher 2,3-DMP content in C.sub.7 product from the CeHy catalyst. The F-1 clear value of 2,3-DMP is 91.1, while the same octane number of 2,4-DMP is only 83.1.

Although the alkylate yield per olefin charged decreased as the total residence time went from 2 to 4 hr., the cumulative yield of C.sub.5.sup.+ paraffins continued to increase from 13.27 to 15.06 g. This behavior strongly indicates that the smaller rate of alkylate production was not the result of a dead catalyst but was more probably brought on by more product cracking and net degradation. We have found that C.sub.6 and C.sub.8 isoparaffins can crack to some degree under alkylation conditions and that reaction conditions can be chosen which avoid or minimize such cracking.

Lowering the isobutane/propylene minimum molar ratio at 60.degree.C. from 15/1 to 10/1 and 5/1 apparently decreased the C.sub.5.sup.+ paraffin yield from 89.1 wt. percent of olefin charge to 49.8 and 5.4 percent, respectively. This apparent effect is also complicated by other factors. Amounts of catalyst and promoter were chosen so that their volumetric concentration in the reaction remained constant. Then, however, the promoter/olefin and catalyst/olefin ratios necessarily changed. C.sub.5.sup.+ paraffin yields were again high when these ratios were high.

Note the virtual absence of olefin polymerization as the relative amount of olefin charge was increased. Only at a 5/1 isoparaffin/olefin ratio was a measurable amount of C.sub.5 -C.sub.8 unsaturate observed-0.07 wt. percent of olefin charged or slightly over 1 percent of the C.sub.5.sup.+ product. This low degree of polymerization and the implied degree of hydride transfer are the result of our novel CeHY catalyst especially in combination with a halide adjuvant. With a less highly exchanged (8/8) CeHY catalyst and with no promoter significant quantities of C.sub.5.sup.+ unsaturates will be in the product when the isoparaffin/olefin ratio is 5.

Table 18 shows typical products obtained at 80.degree.C. when isobutane is contacted, in our process, with various pentenes. A report from the literature of a pentene-isobutane alkylation with a sulfuric acid catalyst is also shown in Table 18.

Table 19 shows typical products obtained with isobutane when isobutylene, 2,3-dimethylbutene-1 or diisobutylene is the feed olefin. Note that diisobutylene acts like two moles of isobutylene.

Over a 220-minute period at 500 p.s.i.g. at 120.degree.C., 9.9 liters of gaseous ethylene were introduced into a stirred slurry of 627 ml. of liquid isobutane and 23 g. of a catalyst containing (ignited basis) 1.30 percent Na and 12.6% Ce and which was prepared similarly to that of Example VIII. No liquid product (i.e., C.sub.5.sup.+ hydrocarbon) was obtained.

Similarly, when ethylene was charged to an isobutane-catalyst slurry under pressure at 80.degree.C. no C.sub.5.sup.+ product was found. The isobutane/ethylene molar ratio was 12.6. Ethylene feed was cut off after 2 hr. on stream, when the reactor reached a predetermined pressure limit of 510 p.s.i.g. However, the reactor contents were kept at run conditions for an additional 2 hr. No C.sub.5.sup.+ product was found. Gas analyses revealed only ethylene and isobutane contaminated with the small amount of n-butane originally present.

P. E. Eberly, Jr., has reported (J. PHYS. CHEM., 71, 1717-22 (1967)) that, among C.sub.2 -C.sub.6 olefins, only ethylene fails to produce conjugated polyene structures on HY catalyst. Eberly's report and the present example show that there is an inherent difference in ethylene's behavior relative to that of the other olefins when contacted in the presence of acidic alumino-silicate zeolites.

If ethylene is charged to a reservoir of carbonium ions and active catalysts, it is probable that ethylene reactivity for alkylation can be enhanced; therefore, butene-2 was charged to an isobutane-catalyst slurry for 2 hr. at 80.degree.C., at the end of which ethylene introduction was begun and continued for an additional 2 hr.

No sign of excess C.sub.6 paraffins was apparent in the liquid product, and the total alkylate yield was only 51.3 percent based on olefin charge instead of the expected 95 to 100 percent paraffin yield. Distribution of paraffins in this alkylate an in another from a typical 3.67 hr. run shows good resemblance between them. The chief difference was a greater proportion of C.sub.9.sup.+ paraffins in the longer run. The relatively low alkylate yield suggested that some cracking had occurred during the 2 hr. period of ethylene charging.

To test the possibility of considerable alkylate cracking, representative products were exposed to typical alkylation conditions for 2 hr.: 12.3 g. 2,2-DMB and 8.7 g. of a C.sub.8 paraffin mixture with a high 2,2,4-TMP content. These amounts corresponded to a 50 percent theoretical yield of 2,2-DMB and a 100 percent theoretical yield of C.sub.8 alkylate on the basis of the preceding run. Isobutane was then charged in the same amount as in the preceding run, and ethylene was introduced during a 2 hr. time period in the same amount. It could not be determined whether the isobutane had more than a diluent action and whether the ethylene had any effect at all.

Results of this experiment, showed 77.4 percent conversion of the 2,2-DMB and 45.7% conversion of 2,2,4-TMP at only 80.degree.C. Net destruction of C.sub.6 paraffin was 72.0 percent because 5.4 percent of the 2,2-DMB was converted to 2,3-DMB and 3-MP. Other C.sub.8 paraffins present to the extent of 2 percent or more of the C.sub.5.sup.+ charge were 49-75 percent cracked, and 52.3 percent of the 2,3-DMP present was also degraded. These results suggest that, in our process, if ethylene in the presence of isobutane is to produce any C.sub.5.sup.+ paraffin, it will be at a temperature in the range of 0.degree.-60.degree.C.

EXAMPLE XVII

This example illustrates use of a fixed bed of substantially anhydrous acid zeolite to catalyze the liquid phase paraffin-olefin alkylation. A bed of 20/60 mesh zeolite catalyst was set up in a vertical column. The catalyst used was an acidic Y zeolite containing 7.8 percent of mixed rare earth ions (from exchange with a crude didymium salt) of which 41 percent was lanthanum. The catlayst contained less than 0.6 percent of cerium, 1.6 percent of sodium, and had an atomic ratio Al/Si of 0.44.

On an ignited basis, the rare earth content (by X-ray fluorescence analysis) was 4.21 percent lanthanum, 0.58 percent cerium, 1.25 percent praseodymium, 3.64 percent neodymium and 0.6 percent samarium.

The catalyst bed was activated in situ and in a manner similar to the activation of Example VII. The catalyst was then pre-wet with isobutane. A feed containing 1 part by weight of butene-2 and 15 parts by weight of isobutane was passed once through the column at a weight hourly space velocity (WHSV) of 8.3 (WHSV of olefin = 0.5), at 440 p.s.i.g. and 120.degree.C., in a period of 80 minutes. The C.sub.5.sup.+ paraffin yield was 88.5 percent. The C.sub.5.sup.+ naphthene yield was 0.6 percent, and the C.sub.5.sup.+ unsaturate yield was 4.5 percent. Of the C.sub.5.sup.+ paraffins, over 78% was pentanes, and 85% of these pentanes was n-pentane. The trimethylpentane to dimethylhexane ratio in the C.sub.8 fraction was 1.58. Note that the maximum possible concentration of unreacted olefin in the reaction mixture is 6.1 mole percent.

A control reaction with the same catalyst and a similar activation in a stirred reactor, run similar to Example VIII, at 120.degree.C. and 480-565 p.s.i.g. for 220 minutes produced 88.0% C.sub.5.sup.+ paraffins and 6.9% C.sub.5.sup.+ unsaturates. Of the C.sub.5.sup.+ paraffins, less than 10 percent was pentanes and only 15 percent of these pentanes was n-pentane. The trimethylpentane to dimethylhexane weight ratio was 2.72 in the C.sub.8 paraffin fraction.

Since the performance of a solid catalyst is in part dependent upon reactant adsorption, we prefer the stirred slurry reactor since the hydrocarbon to catalyst ratio can be higher than that which can be obtained in a fixed catalyst bed of practical dimensions and catalyst packing. However, this deficiency of a fixed bed reactor can to some extent be overcome by using a lower olefin/isoparaffin ratio in the feed, multiplie feed injection, specialized catalyst distribtuion (including dilution with a relatively inactive solid, such as acid clay), high space rates and pulsed flow.

EXAMPLE XVIII

This example illustrates the highly unsaturated products obtained by the use of our acid Y zeolite catalysts and procedures which would be suggested to the person having ordinary skill in the pertinent art and who studied the previously-mentioned prior art relating to paraffin-olefin alkylation with zeolite catalysts. That is, the process conditions in the examples are a fair combination of conditions taught by the prior art wich are applicable to liquid phase reaction of butene-2 and isobutane using an acidic zeolite catalyst. In particular, the example shows a fair combination of an activated cerium-exchanged, ammonium-exchanged, sodium Y zeolite and process conditions analogous to those of U.S. Pat. No. 3,251,902.

A comparison of this example with the previous examples illustrating our process shows the undesirable results of not controlling the addition of the olefin such that the amount of unreacted olefin in the reaction mixture is less than 12 mole percent (preferably less than 7 mole percent) based on the unreacted C.sub.4 to C.sub.6 isoparaffin.

A portion of an 8-cycle cerium-exchanged/8-cycle ammonium-exchanged-sodium Y zeolite (prepared similarly to that of Example III) was activated by the process of Example VII except that the maximum activation temperature was 650.degree.C. 16.8 g. of the activated zeolite, which analyzed 1.7% Na (ignited) and 6.7% Ce (ignited) were charged into a 1-liter stirred autoclave, to which was added 6.1 moles of liquid isobutane. The temperature was raised to 80.degree.C. and the pressure adjusted to 260 p.s.i.g. with nitrogen, then 1.2 moles of liquid butene-2 was gradually charged to the reactor, with stirring, over a period of 2.2 hours. The total C.sub.5 .sup.+ hydrocarbon yield was 10.8 percent of the weight of olefin charged and contained 47.3 percent saturated and 52.7 percent unsaturated hydrocarbons. Note that in this procedure the maximum probable concentration of unreacted olefin in the reaction mixture is 16.4 mole percent.

When the above reaction was repeated with a 22.4 gram portion of the same zeolite which was similarly activated except that the maximum activation temperature was 500.degree.C., the total C.sub.5 .sup.+ yield was 19.7% of the weight of olefin charged and was 47.5 volume percent saturated and 51.4 volume percent unsaturated.

With 17.0 grams of this activated zeolite and a reaction temperature of 40.degree.C. (100 p.s.i.g.), the C.sub.5 .sup.+ yield, based on weight of olefin charged, was 4.6 percent and was 41 volume percent saturated and 59 volume percent unsaturated.

With 23.0 grams of the activated zeolite and a reaction temperature of 120.degree.C. (450 p.s.i.g.), the C.sub.5 .sup.+ yield was 45.2% based on the weight of olefifn charged and was 61 volume percent saturated and 39 volume percent unsaturated.

The high degree of unsaturation in the products of this example indicates that a major reaction was olefin polymerization as in the general reaction

2 (gas olefins) .sup.acid zeolite 1 (liquid olefin).

EXAMPLE XIX

Example VIII was repeated except that the run was for 120 minutes at 60.degree.C., 200 p.s.i.g., and the catalyst was prepared with 32 ammonium and 16 cerium exchanges, similar to that of Example IV. Before activation the exchanged zeolite had a 24.74 ignition loss and, on an ignited basis, analyzed 0.31% Na and 14.3% Ce. The C.sub. 5.sup.+ paraffin yield was 50.1 percent and the C.sub.5 .sup.+ unsaturate yield was 0.00%. Of the C.sub.5 .sup.+ paraffins 80.5 mole percent was C.sub.8 and 0 mole percent of the C.sub.8 paraffins was methylheptanes.

This example illustrates one preferred embodiment of our invention wherein a highly acidic Y zeolite with a halide adjuvant is used at low temperature and with a short retention time to catalyze the reaction of isobutane and butene-2 to produce a high yield of a novel C.sub.5 .sup.+ alkylate containing very little C.sub.5 .sup.+ unsaturates and wherein the ratio TMP/DMH.sub.x is desirably high and the proportion of the less desirable 2,2,4-trimethylpentane is low. In particular, the character of this novel alkylte (which is a highly desirable motor fuel or gasoline blending component) should be contrasted with the undesirable products of Example XVIII (which are illustrative of practice of the prior art).

This example, when compared with Example XVIII and Example XV (see Table 11) shows the superiority in our process of the more highly exchanged and highly acid catalysts (that is, those with the lower Na and the higher polyvalent metal (contents).

Our invention can also be practiced in conjunction with the dehydrogenation of lower paraffins to produce olefins useful as feeds for our alkylation process. That is, the lower paraffins in the C.sub.5 .sup.+ liquid (particularly C.sub.5 and C.sub.6 paraffins) can be separated from the higher paraffins (such as octanes) and passed to a reactor (such as the fixed bed type) containing a dehydrogenation catalyst at conditions favoring monoolefin rather than diolefin formation (500.degree.-1200.degree.F., 0.5-1.5 atm., H.sub.2 /hydrocarbon molar ratio 0.1 to 10 and LHSV of 0.5 to 5). The resulting monoolefins (which need not be separated from unreacted paraffins) can be then transported to the premixing zone, diluted with isoparaffin (if necessary) and utilized as a feed in our paraffin-olefin alkylation process. Operable dehydrogenation catalysts are platinum-alumina, nickel-silica, nickel-magnesia, nickel-alumina, copper-alumina, chromia-alumina and the like.

When our fixed bed alkylation process is operated under conditions which produce high yield of n-pentane (as in Example XIII), it is preferred that at least some of the n-pentane is so dehydrogenated and at least some of the resulting olefins are used as feed in the alkylation stage.

When practiced in combination with such a dehydrogenation stage, this process is preferably operated under conditions which favor paraffin self-alkylation reactions, such as

2 i-C.sub.4 H.sub.10 + C.sub.5 H.sub.10 .fwdarw. i-C.sub.8 C.sub.18 + C.sub.5 C.sub.12.

Such self-alkylation reactions are well known in sulfuric acid catalyzed paraffin-olefin reactions (see Hofmann, J. E. and Schriesheim, A., JACS. 84, 955 (1962)). The present process wherein self-alkylation is combined with the dehydrogenation of C.sub.5 .sup.+ paraffins can be of great economic value to the refiner who has a shortage of butenes and a good supply of isobutane.

Pentanes and hexenes for such self-alkylation can also be derived from relatively inexpensive sources, as dehydrogenated natural gasoline (see U.S. Pat. No. 3,016,344). An especially preferred olefin is 2,3-dimethylbutene, which can be obtained from the dimerization of propylene, since, in this isobutane self-alkylation, it is converted to 2,3-dimethylbutane which has an excellent blending octane number.

Although the previous examples are illustrative of practice of our invention, the yields of many of these examples based on the weight of olefin charged, can be improved upon since it is probable that some of the olefin feed was not consumed in the reaction, but was lost because of slight leakages from the reactor system. For example, Table 20 reports two runs, at 80.degree.C., 250 p.s.i.g. with a catalyst similar to that of Example VIII, which were identical except that Run A utilized the same reactor system as in the previous examples and Run B utilized the same system but greater care was used to prevent loss of feed olefin from the system. It can be seen that the yield in Run B was 12 relative percent greater than Run A.

The calculated F-1 clear octane number of the alkylate of Run B (excluding materials boiling higher than 2,2,4-trimethylhexane) was 98.0. This high octane number, in combination with the reported yield, is one indication of the commercial promise of our invention. Note that a similar calculation shows the novel alkylate of Example XIX (see Table 20) to have a 99.9 octane number.

EXAMPLE XX

This example illustrates, in Table 21, the effect of the gas used in catalyst activation on the C.sub.5 .sup.+ paraffin yield, obtained from the resulting catalyst. Also shown below is a brief summary of the effect on C.sub.5 .sup.+ yield of the final activation temperature.

Activation technique of a CeNH.sub.4 zeolite to produce CeH zeolite can be divided into three distinct parts, during each of which the catalyst is maintained at a definite temperature for a fixed time. First is a preliminary drying, for example, at about 65.degree.C. Second is a dehydration, for example, at about 230.degree.C., which removes virtually all of the adsorbed water but probably does not affect hydroxyls or other waterforming entities more firmly incorporated into the structure. Third is the final activation stage, whch is characterized by NH.sub.4 .sup.+ decomposition and a relatively small removal of water. The activated catalyst, however, contains a definite and reproducible amount of water.

Early in the solid alkylation catalyst research program it had been observed that 400.degree.C. appeared to be a preferred temperature for catalyst activation in air. That observation has been verified with current techniques and more highly exchanged catalysts. The following yields are illustrative:

Wt. % C.sub. 5.sup.+ Yield Based Temperature of Run on Olefin Charged Catalyst Base Air Activation No. Paraffins Unsaturates __________________________________________________________________________ 32NH.sub.4.sup..sup.+, 16Ce.sup.3.sup.+ Y 325.degree.C. 842* 29.4 0.99 do. do. 400.degree.C. 782* 119.0 5.85** do. do. 500.degree.C. 844* 113.2 1.83 16NH.sub.4.sup.+, 16Ce.sup.3.sup.+ Y 400.degree.C. 868*** 125.8 0.24 do. do. 400.degree.C. 880*** 133.8 0.36 do. do. 500.degree.C. 860*** 107.8 1.96 Catalyst Compositions (wt. %, ignited basis) Run No. wt. % cerium wt. % sodium 782 14.0 0.31 842, 844 13.7 0.23 860, 868, 880 14.1 0.98 __________________________________________________________________________ *Operating Conditions: 60.degree.C., 200 psig., 220 min., i-C.sub.4 /C.sub.4 -ene = 14.9 min. **Use of less preferred charge stock preparative technique for this run should not have affected C.sub.5.sup.+ paraffin yield but probably increased the unsaturate yield. ***Operating Conditions: 80.degree.C., 250 psig., 250 min., i-C.sub.4 /C.sub.4 -ene = 14.9 min.

With the 32NH.sub.4 .sup.+ -, 16Ce.sup.3.sup.+ Y-catalyst, a 5.8 percent loss of C.sub.5 .sup.+ paraffin yield was obtained after activation at 500.degree.C. relative to 400.degree.C. With the 16NH.sub.4 .sup.+ -, 16Ce.sup.3.sup.+ Y-catalyst, a 22.0 percent means loss of C.sub.5 .sup.+ paraffin yield was obtained by activation at 500.degree.C. instead of 400.degree.C. The poor result following 325.degree.C. activation may be a result of i incompletely developed acidity in the solid or of residual ammonium, even though a negative test for ammonia evolution had been observed at the end of this activation. DTA and EGA experiments have shown that ammonium decomposition occurs at 300.degree.-320.degree.C. with NH.sub.4 Y zeolite.

To some extent the bound water lost on activation at higher than optimum temperature can be re-introduced to the catalyst. A hydrated CeNH.sub.4 Y zeolite was activated by the procedure of Example VII, the final heating stage being at 400.degree.C. The resulting activated catalyst was contacted with an isobutane-butene-2 feed to produce a 141.5 percent yield of C.sub.5 .sup.+ paraffin.

A similar alkylation with a similar catalyst which had been activated at 600.degree.C. in the final step produced only 128.5 percent of C.sub.5 .sup.+ paraffin.

A catalyst from a similar 600.degree.C. activation was allowed to rehydroate (by exposure to humid air) until the rehydrated zeolite had reached equilbrium. This equilibrated zeolite was then reactivated at 400.degree.C. When an isobutane-butene-2 feed was contacted with this rehydrated, reactivated catalyst, at 140.5% yield of C.sub.5 .sup.+ paraffin was obtained.

EXAMPLE XXI

FIGS. 3, 4 and 5 illustrate the three basic sections which comprise a continuous stirred reactor system for producing an olefin-paraffin alkylate by our process. FIG. 3 illustrates the feed section. The valving arrangement at the top of the mixing vessels 17 and 18 allows feed paraffin 1 or feed olefin 3 to be placed in either vessel 17 or vessel 18 from either the top or the bottom of the vessel. For example, paraffin can be introduced through the bottom of vessel 17 closing valve 9, 12, 19 and 23 and opening valves 4, 8 and 21. Then, feed olefin 3 is transported to vessel 17 by closing valves 4, 11, 12, 14, 19 and 23 and opening valves 7, 8, 9, 10 and 21. Alternately, the mixing of the incoming feed olefin and feed paraffin can be effected by means of an inline mixer; however, for precise control of the reactant proportions and to insure intimate admixing of paraffin and olefin, we prefer that a substantial amount of paraffin admixed with olefin be maintained in a stirred mixing vessel as vessels 17 and 18.

Similarly, by sequencing the position of the valves, the feed paraffin and the feed olefin can be introduced in any desired pattern. One sequence of placing feed paraffin and feed olefin in vessel 17 is to allow feed paraffin to enter vessel 17 toa level a. sufficient feed olefin is then brought into vessel 17 to produce a volume of paraffin-olefin admixture represented by level b. The remainder of the required paraffin feed is added to vessel 17 until the level of the total feed mixture is at c. Such a sequence of paraffin-olefin-paraffin addition allows for better internal mixing of the reactants in vessel 17 (n addition, uniform mixing is insured by mixing devices 15, 16, such as, turbine blade rotary mixers). We have also found that additional mixing can be accomplished by brining the inert gas heat into the bottom of the mixing vessel as through valve 21, rather than into the top of the vessel, as through valve 12.

It is generally preferable to introduce feed components to the mixing vessels in a number of alternate portions (except unless when the feed components are simultaneously proportioned into an inline mixer) to insure uniform mixing. Similarly, sequencing of valves can be used to fill vessel 18 while the mixture in vessel 17 is being fed to the reactor 43 (of FIG. 4). In order to use vessel 17 regardless of whether vessel 18 is being filled or not, pressure, as by an inert gas 2 (e.g., nitrogen) is imposed upon the liquid in vessel 17, as by closing valves 12, 19 and 21 and opening valve 5. Normally, the nitrogen head is allowed to build up until the pressure in the mixing vessel is about 50 p.s.i. less than the pressure in the reactor 43.

In order to allow a mixed paraffin-olefin feed to enter the reactor 43, the nitrogen head is imposed upon vessel 17 and valve 19 is opened. A constant head pressure on vessel 17 allows the pump 31 to pump the mixed feed through line 33 to the reactor at a constant rate.

When valve 19 is open, the feed mixture passes through a microfilter 24 (which protects the pump and meters from damage caused by foreign particles), then through a high pressure rotometer 29 (which serves as a flow indicator). The feed can then enter the pump 31 when valve 30 is open and, when valve 32 is open and 34 closed, the feed is pumped into the reactor 43. In the event of a pump failure, valve 30 and valve 32 may be closed, needle valve 34 opened and the nitrogen head increased sufficiently to allow the feed to flow through line 35 to line 33 and, thence, to the reactor 43.

FIG. 4 illustrates the reactor section, comprising a continuous stirred reactor vessel and the associated lines and valving required for introducing feed, removing reaction products, and for operation of the differential pressure cell which is used for liquid level control. The reactor also contains heat transfer and control means (as a water jacket and electrical heaters, not shown) for maintaining the desired reaction temperature. The paraffin-olefin feed from line 33 enters the reactor 43 through valve 34 and line 35. To insure maximum olefin dilution, we prefer that the liquid feed be allowed to enter the reactor 43 below the reactor liquid level 42 and in the vicinity of the mixing means 36.

The liquid level is controlled by a differential pressure cell, hereinafter DP cell, having a high pressure section 61 and a low pressure section 60, the differential pressure being in the range of 5-50 inches of water column. Inert gas 49 enters the DP cell through an inline filter 50 from which it diverges through meter 52 to the high pressure section 61 and through meter 57 to the low pressure section 60. That is, for the high pressure section, valve 51 is open allowing the inert gas stream to flow through the high pressure meter 52 through open valve needle 53 and valve 64 into the high pressure side 61 of the DP cell, then through valve 62 into a line 67 which leads below the liquid level 42 in the reactor. Pressure gauges, 54a and 54b, indicate the pressure in the high pressure side and the lower pressure side, respectively. Other pressure gauges, thermometers, and analytical devices, can be advantageously incorporated into the three sections comprising the apparatus of FIGS. 3, 4 and 5; however, for simplicity such devices are not shown in the figures.

The inert gas can also be diverted through valve 56 to the low pressure meter 57 through needle valve 58 to the low pressure side 60 of the DP cell and then through valve 63 and line 68 to the vapor space above the liquid level 42 in the reactor. In operation, the DP cell senses a differential pressure which is equal to the height of liquid through which the inert gas from the higher pressure side of the cell must travel from the bottom of line 67 (which must be below the liquid level) to the vapor space 39. The difference between the pressure of the high side 61 and the pressure of the low side 60 of the DP cell is equal to the pressure required to push a bubble of gas through the height of the liquid. Since the DP cell measures the mass of a column of fluid (pressure) and not the volume, its measurement is independent of temperature and, although at a given temperature the actual level of the liquid will vary somewhat, the mass of the volume of liquid above the opening line 67 can be maintained at a constant value regardless of the temperature and pressure of the reactor.

The nitrogen (or other inert gas) which is introduced through the DP cell can be vented through a valve system 40, which can consist of an Annin control valve (a spline-type, highly sensitive metering device) and a block valve ahead of the Annin valve. The Annin valve can be actuated by a pressure transmitter (and gas meter) 38 in order to maintain a constant pressure in the vapor space of the reactor.

Catalyst-free reaction mixture is removed from the reactor via line 37 through valve system 69 (which can consist of an Annin valve and a hand-block valve ahead of the Annin valve). The liquid reaction mixture is separated from the suspended catalyst particles by means of a submerged screen 44 and is withdrawn from the reactor through line 37. Although screen plugging is not a frequent occurrence, if plugging occurs the screen can be back-flushed with nitrogen. This nitrogen back-flush can enter the reactor through line 37 and the excess nitrogen vented through valve system 40 in order that the reaction pressure is maintained constant. This flushing can be effected while the catalyst particles are maintained in suspension and the reaction mixture is maintained in contact with the catalyst particles. In the event that catalyst must be added or removed from the reactor, it may be accomplished through line 48, flush valve 47 and line 46. Similarly, the entire contents of the reactor can be drained through the flush valve 47, if the reaction mixture becomes contaminated or if for any other reason it is desired to drain the reactor contents. For example, if scale builds up on the reactor baffles 45, the reaction mixture can be dumped by opening valve 47, then cleaning materials can be pumped into (and removed from) the reactor through the same valves and lines.

The gases removed via valve 40 can be sent to a gas meter (and pressure transmitter) 38 which can also contain devices for chemical analysis or sampling. For example, the gases so removed can contain HCl, from the halide adjuvant. The HCl concentration in the vapor space is preferably maintained at a constant partial pressure, as by adjusting the quantity of adjuvant which enters the reactor via line 65, valve 66 and line 67. Such adjuvants can also be introduced into the reactor if they are directly added to the paraffin olefin feed in the mixing vessels 17 and 18.

The catalyst-free liquid reaction product (comprising C.sub.5 .sup.+ "alkylate", unreacted feed isoparaffin, some C.sub.5 - paraffin product and, usually, a small amount of unreacted feed olefin) which is removed from the reactor via line 37, passes through valve 69 (where the pressure is reduced from reactor pressure to 25 psig or less) and to condenser 72. The condensed liquid and noncondensed gas (e.g., feed isoparaffin) and inert gas (e.g., nitrogen) pass through valve 86 into vessel 92 or, alternately, through valve 81 to vessel 95. We prefer to have two such collecting vessels in order that product can be collected in one vessel while product is removed from the other vessel. The liquid product removed from these vessels can be transported to product storage tanks or to a means for blending the alkylate with other gasoline components in order to make a blended gasoline product which can be transported to a stabilizer and then to storage area or to tank trucks, etc.

In the product recovery section illustrated in FIG. 5, the liquid produced by condensation of gaseous products in condenser 72 and uncondensed gases pass through line 77 and valve 86 (valve 81 is closed) and enter vessel 92, which is maintained at a temperature and pressure such that liquid alkylate can be removed via valve 96 through line 98 (valve 97 is closed) to tank trucks, a blending area, storage tanks, etc. Uncondensed gases (which consist primarily of unreacted feed isoparaffin) leave vessel 92 via valves 93, 90 and 74 (valves 94, 89 and 73 being closed) and can be passed to means for gas purification and separation 71 or, under some conditions, can be recycled to the reactor or to the mixing vessels via line 76, or via valve 73 and line 70. Minor amounts of the halide promotor which may be present in the reaction product can be removed as by means of an adsorbent which can be between valve 69 and the condenser 72 or at any other appropriate location in the product recovery section. When the halide promotor is a readily distillable gas such as HCl or methyl chloride, it can be removed from the reaction product by an intermediate condensation.

Adjuvants, such as tertiary butyl chloride, can be added directly to the reactor as by line 33, valve 34 and line 35, or can be added to the one of the feed components, such as the isoparaffin, or can be added to the paraffin-olefin mixer in the mixing vessel 17 or 18; however, in the event that the promoter can react with the microfilter 24, or cause corrosion in the pump 31, it is preferred that the adjuvant be added at some point after the pump as by line 65, valve 66 and line 67 (the high pressure side of the DP cell), thus the promoter becomes dispersed in the flowing nitrogen from the high pressure side of the DP cell and passes into the reactor below the liquid level and bubbles up through the reactor contents.

A liquid mixture of 17 volumes of isobutane and 1 volume of butene-2, having a density of 0.557 g./ml., was fed, from the mixing section, and continuously contacted, in the reactor, in the presence of a halide adjuvant, with a CeHy catalyst (16/16) which had been activated by the procedure of Example VII (maximum activation temperature 400.degree.C.). The liquid level in the reactor was maintained constant by the DP cell so that the concentration of the catalyst in the reaction mixture was about 10 weight percent.

The feed rate and the rate of withdrawal of catalyst free reaction mixture were continuous and controlled, such that the mean retention time was 0.3112 hour per gram of hydrocarbon per gram of catalyst. The catalyst was maintained primarily in suspension by continuous stirring.

The halide adjuvant was 0.13 g. t-butyl chloride and 0.22 g. CCl.sub.4 per gram of olefin charged. The temperature in the reaction was maintained at 178.degree.F. and the pressure was autogenous at the temperature (the hydrocarbon being maintained primarily in the liquid phase).

The C.sub.5.sup.+ paraffin product recovered (continuously) from the so withdrawn catalyst-free reaction mixture was about 100 wt. percent based on the weight of olefin charged. The proportion of 2,2,4-trimethylpentane was about 27.5 mole percent of the trimethylpentane fraction, and relatively constant, indicating a continuing self-alkylation function of the catalyst. The total trimethylpentane content in the C.sub.8 fraction of the liquid product was also relatively constant at about 89 to 85 mole percent of the C.sub.8 fraction of the continuously produced alkylate. Less than 1 percent of the product was methylheptane. C.sub.5.sup.+ unsaturates in the liquid product were negligible.

EXAMPLE XXII

This example, by Table 22, illustrates the effect on yield and product quality of the catalyst/olefin ratio.

EXAMPLE XXIII

This example illustrates the correlation between alkylate yield and Electron Spin Resonance (ESR) measurements of total spin count when aromatic hydrocarbons are adsorbed on the CeHY zeolite catalyst.

Several kinds of aromatic hydrocarbons (benzene, p-xylene, naphthene, anthracene, perylene, etc.) were adsorbed upon CeHY catalysts prepared by varied numbers of Ce.sup.+ and NH.sub.4.sup.+ exchange cycles and with various types of activation (e.g., temperature, type of gas). The total ESR spin count of the adsorbed hydrocarbon was then measured. The sorption of the aromatics corresponded to the order of decreasing ionization potential (benzene being first). Catalysts with a higher degree of exchange produced larger spin counts with any particular aromatic. When C.sub.5 .sup.+ paraffin yields obtained with similar catalysts were plotted versus the spin count with a particular aromatic on each catalyst, a good correlation was achieved.

The most satisfactory correlations are for compounds having ionization potentials equal to or larger than that of naphthalene (about 8 ev.). Spin counts of compounds with lower ionization potential changed less than one order of magnitude while relatively large differences in alkylate yield were being observed. These correlations imply a relationship between radigenic nature of a catalyst and its performance in an alkylation reaction (which is highly dependent upon hydride transfer).

A far more exact relation between alkylate yield and anthracene spin count was realized when a series of catalysts of increasing cerium but constant sodium content was used. Anthracene spin count increased less after a (Ce.sup.3.sup.+ /NH.sub.4.sup.+) equivalent ratio of about 2.5 had been reached.

Note, for example, a nearly linear relation between alkylate yield and anthracene spin count can be seen by plotting the data below:

Wt. % C.sub.5.sup.+ paraffin Sorbed based on wt. of Spin Count Hydrocarbon olefin charged (Spins/g.Cat.) .times. 10.sup..sup.-19 ______________________________________ Anthracene Cat. A 91 1.2 do B 126 2.4 do. C 150 3.8 ______________________________________ Note: Reactions at 80.degree.C., Example VIII conditions with "improved feed premixing", isobutane-butene-2 feed. Sodium in catalyst 11 to 13% of cation capacity of zeolite.

Poor hydride transfer, as represented by C.sub.5.sup.+ unsaturate formation, was intensified at low (Ce.sup.3.sup.+ /NH.sub.4.sup.+) equivalent ratio. Other aspects of product quality --low C.sub.9.sup.+ high C.sub.8, and high TMP in the C.sub.8 -- also improved when this composition ratio increased.

These data offer excellent support for the importance of the cerium in CeHY catalyst for alkylation and strongly imply a relation between hydride transfer facility of a catalyst and its electron withdrawal ability.

TABLE 1 __________________________________________________________________________ Exchanged Zeolite Catalysts Chemical Composition of Activated Equilibrated Ammonium-Only __________________________________________________________________________ Catalyst Catalyst: 16 Exchange Cycles with NH.sub.4 NO.sub.3 solutes, Dry Air Medium, Rotary Kiln __________________________________________________________________________ Activation Catalyst No. Conditions Weight Percent TIL* Na-Equivalent Moles** __________________________________________________________________________ Maximum Temperature .degree.F Time**** Na.sup.+ N H.sub.2 O Na.sup.+ NH.sub.4.sup.+ .SIGMA. __________________________________________________________________________ ***Base NaY Zeolite -- -- 9.51 -- 24.32 0.414 -- 0.414 Base after NH.sub.4.sup.+ Exchange -- -- 1.04 5.50 23.76 0.045 0.393 0.438 Experiment A Run No. A-1 150 30 1.01 5.51 24.63 0.044 0.394 0.438 A-2 450 0 1.03 5.28 24.88 0.045 0.377 0.422 A-3 450 60 1.07 4.72 25.39 0.047 0.337 0.384 A-4 750 0 1.00 2.88 26.14 0.043 0.206 0.249 A-5 750 60 1.07 1.98 26.27 0.047 0.141 0.188 A-6 750 120 1.13 0.84 27.26 0.049 0.060 0.109 A-7 750 180 1.05 0.57 27.58 0.046 0.041 0.087 A-8 750 240 1.09 0.54 27.62 0.047 0.039 0.086 Experiment B B-1 150 30 1.03 5.42 24.54 0.045 0.387 0.432 B-2 450 0 1.03 5.04 25.07 0.045 0.360 0.405 B-3 450 60 1.16 4.74 25.17 0.050 0.339 0.389 B-4 572 0 1.14 4.41 25.44 0.050 0.315 0.365 B-5 752 0 1.15 3.33 26.17 0.050 0.238 0.288 B-6 932 0 1.17 0.48 27.43 0.051 0.034 0.085 B-7 1112 0 1.17 0.14 27.65 0.051 <0.010 <0.061 B-8 1292 0 1.14 0.13 25.19 0.050 <0.009 <0.059 B-9 1292 120 1.28 0.13 25.57 0.056 <0.009 <0.065 __________________________________________________________________________ *TIL -- True ignition loss corrected for ammonium. **Na -- Equivalent moles, moles/100 g. ignited catalyst (TIL). ***Molar ratio Na.sub.2 O/Al.sub.2 O.sub.3 was 0.98. Molar ratio SiO.sub. /Al.sub.2 O.sub.3 was 4.70. ****Time, in minutes, at indicated maximum temperature.

TABLE 2 __________________________________________________________________________ Exchanged Zeolite Catalysts Chemical Composition of Activated Ammonium Cerium Catlaysts __________________________________________________________________________ Activation Conditions: Rotary Kiln, Dry Air, Ambient pressure, Programed Temperatures, Air rate = 0.6SCFM __________________________________________________________________________ Activation Conditions LOI* Moles** __________________________________________________________________________ Run No. Max. Temp. .degree.F Time*** H.sub.2 O Na.sup.+ Ce.sup.3.sup.+ NH.sub.4 .sup.+ H.sub.2 O __________________________________________________________________________ Experiment C C-1 150 30 17.21 0.037 0.097 0.025 0.354 0.937 C-2 450 0 10.25 0.037 0.096 0.054 0.379 0.523 C-3 450 60 5.21 0.038 0.096 0.068 0.394 0.228 C-4 750 0 4.05 0.038 0.097 0.043 0.371 0.186 C-5 750 60 3.38 0.034 0.098 0.020 0.350 0.169 C-6 750 120 3.56 0.036 0.098 0.014 0.345 0.186 C-7 750 180 3.09 0.033 0.098 <0.009 <0.336 0.163 C-8 750 240 3.40 0.035 0.098 <0.009 <0.337 0.180 Experiment D D-1 150 30 19.58 0.038 0.096 0.056 0.381 1.044 D-2 450 0 13.18 0.042 0.092 0.056 0.373 0.721 D-3 450 60 6.89 0.038 0.094 0.56 0.375 0.333 D-4 750 0 4.89 0.039 0.098 0.054 0.388 0.222 D-5 750 60 3.92 0.038 0.099 0.031 0.365 0.189 D-6 750 120 3.80 0.038 0.099 0.041 0.375 0.174 D-7 750 180 3.76 0.037 0.100 0.014 0.352 0.197 D-8 750 240 3.47 0.040 0.101 0.010 0.352 0.183 Experiment E E-1 150 30 15.09 0.037 0.096 0.051 0.376 0.796 E-2 450 0 9.76 0.043 0.098 0.051 0.387 0.497 E-3 450 60 5.08 0.042 0.090 0.057 0.370 0.230 E-4 572 -- 4.84 0.043 0.099 0.057 0.396 0.216 E-5 752 -- 4.08 0.041 0.098 0.049 0.384 0.182 E-6 932 -- 3.04 0.040 0.097 0.012 0.341 0.157 E-7 112 -- 2.70 0.043 0.098 <0.009 <0.347 0.127 E-8 1292 0 2.24 0.045 0.100 <0.009 <0.355 0.116 E-9 1292 120 1.66 0.034 0.099 <0.009 <0.340 0.083 __________________________________________________________________________ *LOI - Loss on ignition (includes ammonium) **Moles/100 g ignited catalyst (TIL) ***Time, in minutes at indicated maximum temperature.

TABLE 4 __________________________________________________________________________ Exchanged Zeolite Catalysts Water-cerium Ratio for Ammonium-Cerium Exchanges Zeolite Catalysts Activated at Different Temperatures __________________________________________________________________________ Basis of Data: Moles ion/100 g anhydrous catalyst Catalyst: 16NH.sub.4 .sup.+ exchanges followed by 16Ce.sup.3.sup.+ exchanges Activation Conditions: Rotary kiln, Dry air, Ambient pressure, Programed temperatures, Air Rate = 0.6 __________________________________________________________________________ SCFM Time at Total H.sub.2 O Temp Temp. Non-Ce H.sub.2 O.sup.(2) Non-Ce Total for Ce H.sub.2 O/Ce Run No. (.degree.F) (min) .DELTA..sup.(1) SiOH Na.sup.+ NH.sub.4 .sup.+ H.sub.2 O H.sub.2 O Ce (moles) Ratio __________________________________________________________________________ E-3 450 60 0.044 0.022 0.042 0.028 0.092 0.230 0.138 0.090 1.533 E-4 572 0 0.018 0.009 0.043 0.028 0.080 0.216 0.136 0.099 1.375 E-5 752 0 0.030 0.015 0.041 0.024 0.080 0.182 0.102 0.098 1.041 E-6 932 0 0.073 0.036 0.040 0.006 0.082 0.157 0.075 0.097 0.733 E-7 1112 0 0.072 0.036 0.043 0.002 0.081 0.127 0.046 0.098 0.470 E-8 1292 0 0.069 0.034 0.045 0.000 0.079 0.116 0.037 0.100 0.370 E-9 1292 120 0.083 0.042 0.034 0.000 0.076 0.083 0.007 0.009 0.077 __________________________________________________________________________ .sup.(1) .DELTA. = [(Molar equivalent Na.sup.+ in Na-Base)-(Molar equivalents of measured ions in catalyst)] /100 g anhydrous base .sup.(2) Non-Ce H.sub.2 O calculated as 2SiOH.fwdarw.1H.sub.2 O & 2NH.sub.4 .sup.+ .fwdarw.1H.sub.2 O. .sup.(3) Anhydrous base = ignited catalyst to which is added, as NH.sub.4 .sup.+, the NH.sub.3 evolved on ignition at 1800.degree. F.

TABLE 5 __________________________________________________________________________ ISOBUTANE-BUTENE-2 ALKYLATION WITH ZEOLITE CATALYST __________________________________________________________________________ Example No. VIII IX XIII XIII XIII* t-Butyl n-Propyl n-Butyl Adjuvant None None Chloride Chloride Chloride __________________________________________________________________________ Reaction Temp. 80.degree.C. 120.degree.C. 80.degree.C. 80.degree.C. 80.degree.C. Reaction Press. 250 475 250 250 250 (psig.) Wt.% C.sub.5 .sup.+ Paraffin 71.4 129.4 142.8 129.6 117.2* Yield, based on olefin charged C.sub.5 .sup.+ Paraffin Dist., Mole % C.sub.9 .sup.+ 4.5 10.7 14.1 5.8 10.1 C.sub.8 67.3 54.7 60.1 63.3 71.5 C.sub.7 5.1 12.0 5.9 5.7 7.1 C.sub.6 4.2 8.9 4.6 4.2 5.9 C.sub.5 19.0 13.7 15.3 21.0 5.4* Wt.% C.sub.5 .sup.+ Unsatu- 0.24 4.34 3.54 0.24 0.49 rate Yield based on olefin charged C.sub.8 Paraffin Dist., Mole % Trimethyl- 88.2 74.0 82.4 88.1 85.0 pentanes Dimethyl- 11.8 24.6 16.6 11.9 14.9 hexanes Methylhep- 0.0 1.5 1.1 0.0 0.1 tanes TMP/DMH.sub.x Ratio** 7.47 3.01 4.98 7.38 5.71 % 2,2,4 in TMP*** 26.9 -- 21.0 25.0 23.0 __________________________________________________________________________ In Ex. XIII halide concentration = 1.6 millimole/mole total hydrocarbon charged. 80.degree.C., 250 psig., i-C.sub.4 -ane/C.sub.4 -ene = 14.9 mola (min.), 3.67 hr. 16NH.sub.4 .sup.+ -, 16Ce.sup.3.sup.+ -Catalyst (10.13% Ce, 0.68% Na-before 400.degree. C max. activation) with an ignition loss 24.37% at 1800.degree.F. Feed introduced into bottom of Jerguson guage. *One product gas sample was lost. Some isopentane thereby not accounted for, and the C.sub.5 .sup.+ yield and distribution are affected. **Mole ratio trimethylpentanes to dimethylhexanes. ***Mole percent 2,2,4-trimethylpentane in total trimetylpentanes.

TABLE 6 __________________________________________________________________________ Source Cupit Kennedy Kennedy Ex. X Ex. XI Ex. XI EX. XII __________________________________________________________________________ Cata- 96% HF AlCl.sub.3 CeHY* CeHy* CeHY* CeHY lyst H.sub.2 SO.sub.4 Reac. 7.2 20 30 60 40 25 120 Temp., .degree.C. % 2,2,4 43.9 54.3 65.3 14.8 9.1 4.0 32.3 in TMP __________________________________________________________________________ *2.4 .times. 10.sup..sup.-3 mole t-butyl chloride adjuvant/mole initial i-butane

TABLE 7 __________________________________________________________________________ ISOPARAFFIN-OLEFIN ALKYLATION WITH ZEOLITE CATALYSTS __________________________________________________________________________ Temperature Effects upon Isobutane-Butene-2 Feed with a 16NH.sub.4 .sup.+ /16Ce.sup..sup.+3 Catalyst i-C.sub.4 -ane/C.sub.4 -ene-2 = 14.9 (molar min.), 220-min. time TBC = 1.6 mmole/mole hydrocarbon Catalyst Air-activated at 750.degree.F. __________________________________________________________________________ Temperature, .degree.C. 60 80 80 100 Run Number 560 552 566 558 C.sub.5 .sup.+ Paraffin Yield, wt. % OC 61.3 161.8 151.9 127.7 C.sub.5 .sup.+ Unsaturates, wt. % OC 0.00 0.00 0.00 0.00 C.sub.5 .sup.+ Paraffin Distribution, mole % C.sub. 9.sup.+ 12.1 6.0 7.7 6.8 C.sub.8 65.1 63.8 68.2 60.1 C.sub.7 5.1 6.8 7.7 9.6 C.sub.6 6.6 6.2 6.1 8.5 C.sub.5 11.0 17.3 10.3 15.0 C.sub.8 Paraffin Distribution TMP 80.6 86.4 86.6 80.9 DMH.sub.x 19.4 13.6 13.4 19.1 MH.sub.p 0.0 0.0 0.0 0.0 TMP/DMH.sub.x 4.16 6.37 6.45 4.23 TMP Distribution 2,2,4 10.3 22.3 20.5 27.0 2,2,3 5.0 5.0 5.0 5.0 2,3,4 36.6 33.5 34.0 32.9 2,3,3 48.2 39.3 40.5 35.1 __________________________________________________________________________

TABLE 8 __________________________________________________________________________ ISOPARAFFIN-OLEFIN ALKYLATION WITH ZEOLITE CATALYSTS __________________________________________________________________________ Temperature Effects upon Isobutane-Butene-2 with NH.sub.4 .sup.+ Catalyst vs NH.sub.4 .sup.+ /Ce.sup..sup.+3 Catalyst i-C.sub.4 -ane/C.sub.4 -ene-2 = 15 (molar min.), 220-min. time Catalysts Air-activated at 750.degree.F. __________________________________________________________________________ Catalyst Exchange 16NH.sub.4 .sup.+ / 16NH.sub.4 .sup.+ / Cycles 32NH.sub.4 .sup.+ /O 16Ce.sup..sup.+3 16NH.sub.4 .sup.+ /O 16Ce.sup..sup.+3 TBC Adjuvant No NO Yes Yes Run Number 742 738 596 578 Temperature, .degree.C. 120 120 80 80 C.sub.5 .sup.+ Paraffin Yield, wt. % OC 109.5 129.4 43.5 159.1 C.sub.5 .sup.+ Unsaturates, wt. % OC 2.8 4.3 0.20 0.00 C.sub.5 .sup.+ Paraffin Distribu- tion, mole % C.sub. 9.sup.+ 19.7 10.7 18.9 5.3 C.sub.8 42.5 54.7 54.3 69.5 C.sub.7 9.5 12.0 6.2 7.1 C.sub.6 9.7 8.9 7.0 5.9 C.sub.5 18.6 13.7 13.6 12.2 C.sub.8 Paraffin Dsitribution TMP 72.0 74.0 59.3 87.2 DMH.sub.x 27.2 24.6 38.7 12.8 MH.sub.p 0.7 1.5 2.0 0.0 TMP/DMH.sub.x 2.64 3.01 1.53 6.83 TMP Distribution 2,2,4 24.5 31.5 13.6 22.1 2,2,3 5.5 6.4 5.1 5.1 2,3,4 37.6 31.3 39.6 32.9 2,3,3 32.5 30.8 41.7 39.9 __________________________________________________________________________

TABLE 9 __________________________________________________________________________ ISOPARAFFIN-OLEFIN ALKYLATION WITH ZEOLITE CATALYSTS __________________________________________________________________________ CCl.sub.4 as Adjuvant -80.degree.C., i-C.sub.4 -ane/C.sub.4 -ene-2 = 15 (min.), 3.67 hr., NH.sub.4 .sup.+ -, Ce.sup.3.sup.+ -Type __________________________________________________________________________ Adjuvant Type TBC TBC TBC* TBC** CCl.sub.4 & TBC TBC & TBA*** CCl.sub.4 Technique Continuous Pulse Pulse Pulse Continuous Continuous Continuous Amount, mmole/m OC 42.6 28.4 30.1 30.1 30.1 CCl.sub.4 30.1 TBC 30.1 20.2 TBC 30.1 TBA Run Number 570 614 606 600 630 632 688 C.sub.5 .sup.+ Paraffin Yield, Wt.% OC 162.1 150.0 193.4 176.3 141.1 150.1 178.5 C.sub.5 .sup.+ Unsaturates, wt.% OC 0.00 0.00 0.04 C.sub.5 .sup.+ Paraffin Distribn., mole% C.sub.9.sup.+ 9.1 4.6 7.3 5.8 4.1 5.0 5.8 C.sub.8 73.2 76.3 71.1 72.3 71.5 72.2 74.9 C.sub.7 7.3 7.3 7.1 7.1 6.8 7.0 7.5 C.sub.6 6.4 5.9 5.8 5.7 5.2 5.7 5.4 C.sub.5 4.0 5.9 12.4 10.2 6.4 C.sub.8 Paraffin Distribution TMP 86.0 88.4 87.1 87.8 88.0 88.2 87.9 DMH.sub.x 13.7 11.6 12.9 12.0 11.9 11.8 12.1 MH.sub.p 0.3 0.0 0.0 0.2 0.0 0.0 0.1 TMP Distribution 2,2,4- 19.5 24.5 22.9 27.2 29.2 28.3 26.4 2,2,3- 5.3 5.4 5.6 5.9 6.5 5.0 5.1 2,3,4- 34.2 30.9 31.8 29.1 27.1 28.7 30.9 2,3,3- 41.0 39.2 39.7 37.8 37.2 38.0 37.6 __________________________________________________________________________ *Perylene (from CCl.sub.4) presorbed on catalyst, TBC added continuously **CCl.sub.4 presorbed on catalyst, TBC added continuously ***TBA = t-butyl alcohol, TBC = t-butyl chloride

TABLE 10 __________________________________________________________________________ Paraffin Isobutane Isobutane Isobutane Isobutane n-Butane* Isobutane Isobutane* *** *** 2-Methyl- 2-Methyl- *** *** **** *** **** Olefin butene-2 butene-2 Butene-2 Butene-1 Butene-1 Butene-2 Butene-2 __________________________________________________________________________ Temperature .degree.C. 120 120 120 120 80 80 80 Pressure, psig. 460 485 455 455 250 250 250 Catalyst Wt.% Na (ignited) 1.68 1.11 1.11 1.38 0.76 0.76 0.76 Wt.% Ce (ignited) 6.8 8.7** 8.7** 12.4 13.5 13.5 13.5 C.sub.5 .sup.+ Paraffin Yield 28.6 49.0 51.8 119.7 25.6 135.0 132.0 Wt.% Olefin Chg. C.sub.5 .sup.+ Unsaturate 31.2 15.4 0.5 2.8 1.6 1.88 0.26 Yield, Wt.% Olefin Chg. C.sub.5 .sup.+ Paraffin Dist., Mole % C.sub.9.sup.+ 32.5 29.2 4.2 9.7 26.6 14.6 8.0 C.sub.8 32.6 36.6 82.8 55.0 25.6 69.3 69.0 C.sub.7 15.6 10.3 6.2 12.0 2.0 6.7 6.1 C.sub.6 15.2 10.3 5.2 11.4 2.0 5.5 5.2 C.sub.5 4.0 13.6 1.6 11.9 43.8 4.0 11.7 C.sub.8 Paraffin Dist., Mole % Trimethyl- 71.2 74.7 63.8 75.7 33.7 85.4 88.0 pentanes Dimethyl- 26.4 24.3 36.4 23.1 65.6 14.6 12.0 hexanes Methylhep- 2.4 1.0 0.7 1.2 0.7 0.0 0.0 tanes __________________________________________________________________________ *2.4 .times. 10.sup..sup.-3 mole tertiary butyl chloride used as adjuvant per mole of n-butane. **Catalyst activated at 500.degree.C. (all other runs at 400.degree.C.) ***Feed introduced at top of Jerguson gauge. ****Feed introduced at bottom of Jerguson gauge.

TABLE 11 __________________________________________________________________________ Ex. Ex. Ex. Ex. Ex. Ex. Ex. Catalyst Prep. II IV IV VII** XV** XV** IV** __________________________________________________________________________ Temperature .degree.C. 120 120 120 80 80 80 80 Pressure, psig. 500 500 500 250 250 250 Wt.% Na (ignited) 0.26 1.38 0.3 0.6 0.82 0.9 0.76 Wt.% Ce (ignited) -- 12.4 13.5 1.72 -- -- 13.5 *** Wt.% La (ignited) -- -- -- -- 12.3 **** -- Wt.% N (before 6.42 0.98 0.66 5.20 1.18 0.57 0.86 activation) Wt.% Ignition 30.25 24.24 25.84 28.41 25.25 24.95 24.70 Loss Wt.% C .sup.+ Paraf- 109.5 119.3 75.3 26.0 68.4 142.4 132.0 fin.sup.5 Yield* Wt.% C.sub.5.sup.+ Unsatu- 2.8 5.6 1.1 10.9 0.13 0.13 0.26 rate Yield* C.sub.5.sup.+ Paraffin Dist. Mole % C.sub.9.sup.+ 19.7 11.5 11.2 27.2 5.4 3.6 8.0 C.sub.8 42.5 50.3 72.0 57.9 81.0 60.0 69.0 C.sub.7 9.5 9.0 8.6 5.8 5.9 4.1 6.1 C.sub.6 9.7 8.4 2.2 5.8 4.4 3.1 5.2 C.sub.5 18.6 20.7 6.0 3.3 3.2 29.2 11.7 C.sub.8 Paraffin Dist. Mole % Trimethyl- 72.0 73.1 75.0 55.1 89.0 88.4 88.0 pentanes do. Dimethyl- 27.2 25.7 23.8 41.4 11.0 11.6 12.0 hexanes do. Methyl- 0.7 1.2 1.2 3.5 0.0 0.0 0.0 heptanes do. TMP/DMH.sub.x 2.64 2.84 3.15 1.33 8.06 7.65 7.30 __________________________________________________________________________ *Based on weight of olefin charged. **t-Butyl chloride adjuvant. Olefin and paraffin entered bottom of Jerguson gauge. ***Catalyst prepared from La(NO.sub.3) solution. ****Catalyst prepared from mixed rare earth nitrate solution and analyzed 13.8% total rare earth metals (ignited).

TABLE 12 ______________________________________ Catalyst Content Salt Used of Rare Earth Metals C.sub.5.sup.+ Paraffin for Exchange (g. ion/100 g. Yield (wt. % Solution anhydrous cat.) olefin charge) ______________________________________ Ce(NO.sub.3).sub.3 0.286 132.0 CeCl.sub.3 0.272 73.1 La(NO.sub.3).sub.3 0.275 68.4 LaCl.sub.3 0.250 112.0.sup.(a) RE(NO.sub.3).sub.3 0.301 142.4 RECl.sub.3 0.236 103.1 Gd(NO.sub.3).sub.3 0.305 163.0 ______________________________________ .sup.(a) Use of the magnetic drive on the reactor possibly increased this yield as much as 15% over what it would have been with the same packed drive used for the other runs. Evan at (112-115 = 98%), its yield vastly exceeds that from La(NO.sub.3).sub.3.

TABLE 13 __________________________________________________________________________ ISOPARAFFIN-OLEFIN ALKYLATION WITH ZEOLITE CATALYSTS __________________________________________________________________________ Rare Earth Cation and Anion Effects on Catalysts 80.degree.C., 250 psig., i-C.sub.4 -ane/C.sub.4 -ene-2 = 14.9 (min.), 3.67 hr. 1.0 g. t-Butyl Chloride __________________________________________________________________________ Salt for Exchange Ce(NO.sub.3).sub.3 CeCl.sub.3 RE(NO.sub.3).sub.3 RECl.sub.3 La(NO.sub.3).sub.3 LaCl.sub.3 Catalyst Composition Sodium, wt.%(ignited residue basis) 0.76 0.78 1.17 0.89 0.82 1.09 Rare Earth, wt.% (ignited residue basis) 13.5 12.9 14.2 11.2 13.1 11.9 Run Number (467-) 830 822 820 854 818 858 C.sub.5 .sup.+ Paraffin, wt.% chg. 132.0 73.1 142.4 103.1 68.4 112.0 C.sub.5 .sup.+ Unsaturates, wt. % olefin chg. 0.26 0.15 0.13 0.77 0.13 0.16 C.sub.5 .sup.+ Paraffin Distribution C.sub.9 .sup.+ , mole % 8.0 6.5 3.6 10.6 5.4 8.1 C.sub.8 , do. 69.0 78.1 60.0 66.9 81.0 61.7 C.sub.7 , do. 6.1 5.7 4.1 6.7 5.9 6.4 C.sub.6 , do. 5.2 4.1 3.1 3.9 4.4 4.4 C.sub.5 , do. 11.7 5.5 29.2 11.9 3.2 19.5 C.sub.8 Paraffin Distribution TMP, mole % 88.0 88.6 88.4 84.8 89.0 86.2 DMH.sub.x, do. 12.0 11.4 16.6 15.2 11.0 13.8 MH.sub.p , do. 0.0 0.0 0.0 0.0 0.0 .0 TMP/DMH.sub.x Ratio 7.30 7.74 7.65 5.09 8.06 6.25 __________________________________________________________________________

TABLE 14 __________________________________________________________________________ ISOPARAFFIN-OLEFIN ALKYLATION WITH ZEOLITE CATALYSTS __________________________________________________________________________ Catalyst Cerium-Sodium Effects on Alkylation 80.degree.C., 250 psig., i-C.sub.4 -ane/C.sub.4 -ene-2 = 14.9 (min.), 3.67 hr. 1.0 g. t-Butyl Chloride __________________________________________________________________________ Catalyst Composition Sodium, wt. % (ignited residue basis ) 0.23 0.76 1.68 2.76 1.68 1.24 Cerium, wt. % ( do. ) 13.7.sup.(a) 13.5.sup.(a) 12.8.sup.(a) 13.0 8.3.sup.(a) 2.0 Run Number (467-) 852 830 848 850 846 828 C.sub.5 .sup.+ Paraffin, wt. % olefin charge 135.2 132.0 115.4 118.9 62.7 26.0 C.sub.5 .sup.+ Unsaturates, wt. % olefin charge 0.32 0.26 0.42 0.77 4.52 10.85 C.sub.5 .sup.+ Paraffin Distribution C.sub.9 .sup.+ , mole % 8.4 8.0 10.6 10.1 19.1 27.2 C.sub.8 , do. 66.0 69.0 63.4 72.8 64.1 57.9 C.sub.7 , do. 6.6 6.1 6.7 7.2 7.4 5.8 C.sub.6 , do. 4.9 5.2 5.5 5.7 6.0 5.8 C.sub.5 , do. 14.1 11.7 13.9 4.1 3.4 3.3 C.sub.8 Paraffin Distribution TMP, mole % 87.5 88.0 85.7 87.2 82.9 55.1 DMH.sub.x, do. 12.4 12.0 14.3 12.8 16.2 41.4 MH.sub.p , do. 0.1 0.0 0.0 0.0 0.9 3.5 TMP/DMH.sub.x Ratio 7.08 7.30 6.00 6.80 5.10 1.33 __________________________________________________________________________ .sup.(a) By X-ray fluorescence. Others were by gravimetry.

TABLE 15 __________________________________________________________________________ Liquid Phase Isoparaffin-Olefin Alkylation with Solid Zeolite __________________________________________________________________________ Catalysts Gadolinium versus Ammonium versus Cerium and Type X versus Type Y Zeolite Autogenous pressure, 80.degree.C., i-C.sub.4 -ane/C.sub.4 -ene-2 = 15 (min.), 3.67 hr., 1.0 g. tertiary butyl chloride adjuvant __________________________________________________________________________ Catalyst Zeolite before activation GdNH.sub.4 Y GdNH.sub.4 Y NH.sub.4 Y CeNH.sub.4 X CeNH.sub.4 Y Activation (400.degree.C.) Gas Air He Air Air Air Run No. 628 674 596 622 642 C.sub.5 .sup.+ Paraffin Yield, wt.% OC 163.0 169.8 43.5 130.0 161.6 C.sub.5 .sup.+ Unsaturates, wt.% OC 0.00 0.05 0.2 0.0 0.00 C.sub.5 .sup.+ Paraffin Distribution, mole % C.sub.9.sup.+ 3.7 5.8 18.9 8.3 5.4 C.sub.8 67.6 71.4 54.3 59.0 71.2 C.sub.7 5.7 7.7 6.2 4.8 7.4 C.sub.6 4.4 5.7 7.0 4.0 5.9 C.sub.5 18.5 9.3 13.6 23.9 10.0 C.sub.8 Paraffin Distribution TMP 88.1 88.2 59.3 85.7 85.9 DMH.sub.x 11.9 11.8 38.7 14.2 14.1 MH.sub.p 0.0 0.0 2.0 0.1 0.0 TMP Distribution 2,2,4- 27.4 28.4 13.6 15.5 24.4 2,2,3- 5.9 5.4 5.1 4.1 5.6 2,3,4- 28.8 29.1 39.6 34.0 32.0 2,3,3- 37.9 37.1 41.7 46.4 38.0 Catalyst Analysis (ignited basis, before activation) wt. % Na 9.97 0.97 1.05 0.93 0.97 wt. % Ce or Gd 14.39 Gd 14.39 Gd -- 15.5 Ce 13.99 Ce wt. % N 1.09 1.09 5.86 1.85 0.84 wt. % loss on ignition 25.35 25.35 29.67 25.47 26.28 Analysis of Base Na zeolite (before exchange, ignited base) * * wt. % Na 9.51 9.51 -- -- 9.42 wt. % Al.sub.2 O.sub.3 16.56 16.56 -- -- 16.32 wt. % SiO.sub.2 45.29 45.29 -- -- 47.87 wt. % loss on ignition 24.32 24.32 -- -- 25.05 __________________________________________________________________________ *Al/Si atomic ratio is 0.69 for the CeNH.sub.4 X zeolite before activation.

TABLE 16 ______________________________________ Isoparaffin-Olefin Alkylation with Solid Zeolite Catalysts Isobutane and Propylene with and without Promoter ______________________________________ 60.degree.C., i-C.sub.4 -ane/C.sub.3 -ene = 20 (min.), 120 min. CeHY catalyst from 16NH.sub.4 /16 Ce Base ______________________________________ 42.3 mmole Promoter None TBC/m OC 96% Run No., 690 544 H.sub.2 SO.sub.4 * C.sub.5 .sup.+ Paraffins, 66.1 123.3 Wt % OC C.sub.5 .sup.+ Unsaturates, 0.00 0.00 Wt. % OC C.sub.5 .sup.+ Paraffin Distribn., mole % Vol.% C.sub.9.sup.+ 5.2 7.2 11.0 C.sub.8 12.1 11.2 9.8 C.sub.7 63.8 66.0 71.1 C.sub.6 6.9 5.9 4.2 C.sub.5 11.9 9.7 3.8 C.sub.8 Paraffin Distribn. TMP 82.5 86.5 DMH.sub.x 16.1 13.5 MHp 1.4 0.0 TMP Distribn. 2,2,4- 67.8 70.0 54.6 2,2,3- 1.7 0.4 0.0 2,3,4- 12.7 13.4} (45.4)** 2,3,3- 17.8 16.3 C.sub.7 Paraffin Distribn. 2,3-DMP 95.3 97.2 70.9 Other DMP 3.2 2.0 29.1 MH.sub.x 1.5 0.8 0.0 ______________________________________ *At 7.degree.C., 47 Vol. % emulsion, from Cupit, C.R.,et al, Petro Chem. Eng., p. 204 (December, 1961). **Total of 2,3,4- and 2,3,3-TMP.

TABLE 17 __________________________________________________________________________ ISOMERIZATION-OLEFIN ALKYLATION WITH SOLID ZEOLITE CATALYSTS __________________________________________________________________________ Isobutane and Propylene Catalyst: CeHY (Base zeolite 0.35% Na, 14.19%Ce, 1.34% N) .sup.(a) Feed rate 1 ml/min. Time and Isoparaffin/Olefin Effects __________________________________________________________________________ Run No. 534 544 542 550 548 Temp. .degree.C. 60 60 60 60 60 Time, hr. 3 2 4 3 3 Iso/Olefin Molar Ratio (min.) 15 20 15 10 5 Wt. TBC, g. 1.0068 1.0018 1.0000 0.9006 0.6998 Wt. Catalyst, g. 21.99 22.54 22.04 19.95 15.83 C.sub.5 .sup.+ Paraffin Yield, wt. % OC 89.1 123.3 70.0 49.8 5.4 C.sub.5 .sup.+ Unsaturates, wt. % OC 0.41 0.00 0.00 0.00 0.07 C.sub.5 .sup.+ Paraffin Distri- bution, mole % C.sub.9.sup.+ 16.2 7.2 22.6 21.8 28.4 C.sub.8 9.1 11.2 10.2 9.8 23.1 C.sub.7 56.3 66.0 52.4 53.9 38.6 C.sub.6 4.8 5.9 6.6 4.9 8.1 C.sub.5 13.6 9.7 8.2 9.6 1.7 TMP Distribution, mole % 2,2,4- 69.3 70.0 62.1 69.0 66.7 2,2,3- 0.3 0.4 0.6 0.0 1.1 2,3,4- 14.1 13.4 12.2 15.1 15.6 2,3,3- 16.3 16.3 15.2 15.9 16.7 C.sub.7 Paraffin Distri- bution, mole % 2,3-DMP 98.2 97.2 95.9 97.2 92.9 Other DMP + TMB 1.4 2.0 1.8 1.2 1.6 MH 0.4 0.8 2.3 1.6 1.5 __________________________________________________________________________ .sup.(a) Analysis on unactivated catalyst

TABLE 18 __________________________________________________________________________ Isoparaffin-Olefin Alkylation with Solid Catalyst __________________________________________________________________________ Isobutanes and Pentenes Isopar./Olef. = 15 (min.) CeHy catalyst Autogeneous pressure __________________________________________________________________________ Run No. 584 582 590 586 H.sub.2 SO.sub.4 Olefin C.sub.5 -ene-1 C.sub.5 -ene-1 C.sub.5 -ene-1 3MB-ene-1 C.sub.5 -ene-1 Olefin Vaporization Capacity (% OC) 10 10 10 10 Catalyst/Olefin (g/mole) 57.3 57.3 57.3 57.3 TBC/Olefin (mmole/mole) 28.4 28.4 28.4 28.4 Temperature (.degree.C) 60 80 100 80 Time (min.) 120 220 220 220 C.sub.5 .sup.+ Paraffin Yield, wt.% OC 57.6 67.5 64.2 115.1 C.sub.5 .sup.+ Unsaturates, wt.% OC 0.07 4.77 7.43 0.00 C.sub.5 .sup.+ Paraffin Distribution, mole % C.sub.9.sup.+ 61.3 47.8 40.6 11.1 38.2 C.sub.8 12.3 14.8 17.8 27.4 25.6 C.sub.7 4.9 3.6 6.3 4.4 0.7 C.sub.6 3.3 3.7 5.7 4.2 1.3 C.sub.5 18.1 30.0 29.6 52.8 33.9 C.sub.8 /C.sub.5 0.680 0.493 0.602 0.520 0.745 C.sub.8 Paraffin Distribution TMP 78.7 85.2 78.2 89.1 DMH.sub.x 20.8 14.8 20.2 10.9 MHp 0.5 0.0 1.6 0.0 TMP Distribution 2,2,4- 62.6 55.9 56.3 58.4 2,2,3- 0.3 1.3 1.6 2.5 2,3,4- 17.3 20.2 22.0 18.4 2,3,3- 19.8 22.6 20.1 20.7 __________________________________________________________________________ * From J. E. Hofman, J. Orgn. Chem., 29, 1497-99 (1964)

TABLE 19 __________________________________________________________________________ Isoparaffin-Olefin Alkylation with Solid Zeolite Catalysts Isobutane with Butene-1 and 2, Isobutylene, and Diisobutylene __________________________________________________________________________ Isopar./Olef. = 15 (min.), CeHY (16NH.sub.4 /16Ce) Catalyst, 80.degree.C., 220 minutes Catalyst/Olefin = 57.3 g./mole, TBC/Olefin = 28.4 mmole/mole __________________________________________________________________________ Run No. 578 594 592 588 666 Olefin C.sub.4 -ene-2 C.sub.4 -ene-1 i-C.sub.4 -ene 2,4,4-TMP-ene-1 2,3-Dimethylbutenol Olefin Vaporization Capacity (% OC) 22.3 22.3 22.3 1 C.sub.5 .sup.+ Paraffin Yield, Wt.% OC 159.1 168.3 123.9 120.5 87.0 C.sub.5 .sup.+ Unsaturates, Wt.% OC 0.00 0.00 {0.58 } {29.2 as C.sub.8} 0.08 {as C.sub.8} {0.21 as C.sub.5} C.sub.5 .sup.+ Paraffin Distribn., mole % C.sub.9.sup.+ 5.3 8.2 18.5 28.2 24.6 C.sub.8 69.5 69.2 49.3 49.5 23.0 C.sub.7 7.1 6.8 8.5 7.2 10.3 C.sub.6 5.9 5.7 7.6 6.5 31.7 C.sub.5 12.2 10.1 16.0 8.6 10.4 C.sub.8 Paraffin Distribn. TMP 87.2 85.4 88.7 87.0 83.9 DMH.sub.x 12.8 14.6 11.1 10.2 13.8 MHp 0.0 0.0 0.2 2.8 2.2 TMP Distribn. 2,2,4- 22.1 21.2 60.1 64.6 58.3 2,2,3- 5.1 5.0 2.6 1.9 2.2 2,3,4- 32.9 33.8 17.7 16.7 19.3 2,3,3- 39.9 40.0 19.6 16.8 20.2 __________________________________________________________________________ *C.sub.6 paraffin distribution, for Run 666 was 97% 2,3-DMB, 1.3% 2,2-DMB and 1.7% MP.

TABLE 20 ______________________________________ Run A* Run B* Ex. XIX* ______________________________________ % Yield C.sub.5 .sup.+ Paraffin** 132 148 50.1 % Yield C.sub.5 Unsaturates** 0.26 0.19 0.00 C.sub.5 .sup.+ Paraffin Distribution C.sub.9 .sup.+ , Mole % 8.0 8.8 5.3 - C.sub.8 , Mole % 69.0 71.8 80.5 C.sub.7 , Mole % 6.1 6.3 3.8 C.sub.6 , Mole % 5.2 5.2 2.1 C.sub.5 , Mole % 11.7 7.9 8.4 C.sub.8 Paraffin Distribution TMP, Mole % 88.0 88.8 92.0 DMH.sub.x, Mole % 12.0 11.2 8.0 MHp, Mole % 0.0 0.0 0.0 TMP/DMH.sub.x Ratio 7.30 7.91 11.44 Trimethylpentane Distribution % 2,2,4- 25.2 17.1 % 2,2,3- 3.7 5.0 % 2,3,4- 31.8 33.3 % 2,3,3- 39.3 44.7 ______________________________________ *Feed olefin introduced at bottom of Jerguson gauge. **Based on weight of olefin charged.

TABLE 21 __________________________________________________________________________ Helium and Hydrogen versus Air Activation at 400.degree.C. NH.sub.4 .sup.+ -, Ce.sup.3 - Type Y Base 80.degree.C, autogeneous pressure, i-C.sub.4 -ane/C.sub.4 -ene-2 = 15 (min.), 3.67 hr., 1.0 g. TBC __________________________________________________________________________ Activation Gas Air Air H.sub.2 He Run No. 654 656 658 660 C.sub.5 .sup.+ Paraffin Yield, wt.% OC 142.6 139.7 148.8 160.6 C.sub.5 .sup.+ Unsaturates, wt.% OC 0.00 0.05 0.00 0.14 C.sub.5 .sup.+ Paraffin Distribn., mole % C.sub.9.sup.+ 9.2 9.1 8.1 7.8 C.sub.8 66.8 66.1 69.2 70.5 C.sub.7 6.4 6.2 6.7 5.8 C.sub.6 6.2 5.8 5.7 5.9 C.sub.5 11.4 12.9 10.3 9.9 C.sub.8 Paraffin Distribn. TMP 85.6 86.4 86.7 87.5 DMH.sub.x 14.4 13.5 13.1 12.4 MHp 0.0 0.2 0.2 0.0 TMP Distribn. 2,2,4- 22.3 23.8 25.8 26.8 2,2,3- 5.0 5.3 5.3 5.4 2,3,4- 33.9 32.3 31.1 30.4 2,3,3- 38.9 38.7 37.8 37.3 __________________________________________________________________________

TABLE 22 __________________________________________________________________________ Isoparaffin-Olefin Alkylation with Solid Catalysts __________________________________________________________________________ Amount of Catalyst and Performance i-C.sub.4 -ane/C.sub.4 -ene-2 = 15 (molar min.) 220-min. Time 16/16 -type Catalyst __________________________________________________________________________ Catalyst 1F7 1F1 1F8 1F9 Relative Amt. 1.0 2.0 0.5 1.0 1.0 1.5 Run Number 738 744 564 552 566 568 TBC Promoter No No Yes Yes Yes Yes Temperature, .degree.C 120 120 80 80 80 80 C.sub.5 .sup.+ Paraffins, wt.% OC 129.4 101.4 98.5 161.8 151.9 140.8 C.sub.5 .sup.+ Unsaturates, wt.% OC 4.3 0.81 0.00 0.00 0.00 0.00 C.sub.5 .sup.+ Paraffin Distribn., mole % C.sub.9.sup.+ 10.7 6.9 8.1 6.0 7.7 3.4 C.sub.8 54.7 43.1 70.0 63.8 68.2 66.2 C.sub.7 12.0 9.9 6.9 6.8 7.7 6.8 C.sub.6 8.9 11.7 6.5 6.2 6.1 6.0 C.sub.5 13.7 28.4 8.5 17.3 10.3 17.5 C.sub.8 Paraffin Distribn. TMP 74.0 76.1 81.6 86.4 86.6 87.5 DMH.sub.x 24.6 22.8 17.6 13.6 13.4 12.5 MHp 1.5 1.1 0.8 0.0 0.0 0.0 TMP-DMH.sub.x 3.01 3.34 4.65 6.37 6.45 7.03 TMP Distribn. 2,2,4- 31.5 40.4 17.9 22.3 20.5 27.3 2,2,3- 6.4 10.7 4.8 5.0 5.0 5.5 2,3,4- 31.3 21.7 36.0 33.4 34.0 38.1 Grams of Catalyst 24.53 48.43 10.97 22.79 22.78 34.37 __________________________________________________________________________

Claims

We claim:

1. A paraffin-olefin alkylation process which comprises contacting monoolefin of the C.sub.2 -C.sub.9 range in admixture with paraffin of the C.sub.4 -C.sub.6 range and with a substantially anhydrous acidic crystalline alumino-silicate zeolite under alkylating conditions and wherein there is present in solution in the reaction mixture from 10.sup..sup.-5 to 10.sup..sup.-1 mole per mole of C.sub.4 -C.sub.6 paraffin of a halide, said halide being selected from the group consisting of HCl, carbon tetrachloride and the aliphatic saturated monochlorides having no more than 6 carbon atoms.

2. An isoparaffin-olefin alkylation process wherein C.sub.3 -C.sub.9 monoolefin in admixture with C.sub.4 -C.sub.6 isoparaffin having a tertiary carbon atom is contacted with a substantially anhydrous, acidic crystalline alumino-silicate zeolite under alkylation conditions at a temperature below the critical temperature of the lowest boiling hydrocarbon reactant and at a pressure such that the reactants are substantially in liquid phase, and wherein there is present in solution in the reaction mixture from 10.sup..sup.-5 to 10.sup..sup.-1 mole per mole of C.sub.4 -C.sub.6 isoparaffin of a halide, said halide being selected from the group consisting of HCl, carbon tetrachloride and the aliphatic saturated monochlorides having 1 to 4 carbon atoms.

3. A process for the preparation of an isoparaffin-olefin alkylate comprising contacting isobutane with monoolefin selected from the group consisting of isobutylene, butene 2 and butene-1 and with a substantially anhydrous acidic crystalline alumino-silicate zeolite, at a temperature in the range of 25-120.degree.C. and at a pressure such that each of the reactants is substantially in liquid phase,

i. said contacting being effected utilizing sufficient agitation so that substantially all of said zeolite is maintained in suspension in the liquid reaction mixture,

the amount of unreacted olefin in the reaction mixture being maintained at less than 7 mole percent based on the unreacted isobutane, and

iii. wherein the mean weight hourly space velocity of the hydrocarbons in the reaction mixture is in the range of 2-20 gram hydrocarbons per hour-gram catalyst; and wherein there is present in solution in the reaction mixture from 10.sup..sup.-5 to 10.sup..sup.-1 mole per mole of isobutane of a halide adjuvant selected from the group consisting of HCl, carbon tetrachloride and the aliphatic saturated monochlorides having 1 to 4 carbon atoms.

4. Process for the preparation of an isoparaffin-olefin alkylate comprising contacting isobutane with monoolefin selected from the group consisting of isobutylene, butene-2 and butene-1, and with a substantially anhydrous acidic crystalline alumino-silicate zeolite, at a temperature in the range of 40.degree.-80.degree.C. and at a pressure such that each of the reactants is substantially in liquid phase,

i. said contacting being effected utilizing sufficient agitation so that substantially all of said zeolite is maintained in suspension in the liquid reaction mixture,

ii. the amount of unreacted olefin in the reaction mixture being maintained at less than 7 mole percent based on the unreacted isobutane,

iii. wherein the mean weight hourly space velocity of the hydrocarbons in the reaction mixture is in the range of 2-20 gram hydrocarbons per hour-gram catalyst; and

iv. wherein said contacting is in the presence of an alkylation-promoting amount of a halide containing chlorine, bromine or fluorine.

5. Process according to claim 4 wherein there is present in solution in the reaction mixture from 10.sup..sup.-5 to 10.sup..sup.-1 mole per mole of isobutane of a halide, said halide being selected from the group consisting of HCl, carbon tetrachloride and the aliphatic saturated monochlorides having 1 to 4 carbon atoms.

6. An isoparaffin-olefin alkylation process which comprises contacting 2,3-dimethylbutene in admixture with C.sub.4 -C.sub.6 isoparaffin having a tertiary carbon atom in the presence of an alkylation-promoting amount of a halide containing chlorine, bromine or fluorine with a substantially anhydrous acidic crystalline alumino-silicate zeolite under alkylating conditions,

i. wherein said contacting is at a temperature below the critical temperature of the lowest boiling hydrocarbon reactant and is at a pressure such that the reactants are substantially in liquid phase, and

ii. wherein said product of said contacting comprises 2,3-dimethylbutane.

7. Process for the preparation of an olefin-paraffin alkylate comprising contacting butene-1 with n-butane and with substantially anhydrous acidic crystalline alumino-silicate zeolite, at a temperature in the range of 40.degree.-80.degree.C. and at a pressure such that each of the reactants is substantially in liquid phase,

i. said contacting being effected utilizing sufficient agitation so that substantially all of said zeolite is maintained in suspension in the liquid reaction mixture,

ii. the amount of unreacted olefin in the reaction mixture being maintained at less than 7 mole percent based on the unreacted n-butane,

iii. wherein the mean weight hourly space velocity of the hydrocarbons in the reaction mixture with the catalyst is in the range of 2-20 gram hydrocarbon per hour-gram catalyst; and

iv. wherein said contacting of step (i) is in the presence of an alkylation-promoting amount of a halide containing chlorine, bromine or fluorine.

8. Process for the preparation of an olefinparaffin alkylate comprising contacting butene-1 with n-butane and with a substantially anhydrous acidic crystalline alumino-silicate zeolite, at a temperature of 25-120.degree.C. and at a pressure such that each of the reactants is substantially in liquid phase,

i. said contacting being effected utilizing sufficient agitation so that substantially all of said zeolite is maintained in suspension in the liquid reaction mixture,

ii. the amount of unreacted olefin in the reaction mixture being maintained at less than 7 mole percent based on the unreacted n-butane,

iii. wherein the mean weight hourly space velocity of the hydrocarbons in reaction mixture with the catalyst is in the range of 2-20 gram hydrocarbon per hour-gram catalyst; and,

iv. wherein said contacting of step (i) is in the presence of an alkylation-promoting amount of a halide containing chlorine, bromine or fluorine.