|Publication number||US3236762 A|
|Publication date||Feb 22, 1966|
|Filing date||Feb 7, 1962|
|Priority date||Jan 28, 1951|
|Also published as||US3236761|
|Publication number||US 3236762 A, US 3236762A, US-A-3236762, US3236762 A, US3236762A|
|Inventors||James E Boyle, Paul E Pickert, Jule A Rabo|
|Original Assignee||Union Carbide Corp|
|Export Citation||BiBTeX, EndNote, RefMan|
|Patent Citations (5), Referenced by (95), Classifications (75)|
|External Links: USPTO, USPTO Assignment, Espacenet|
Feb. 22, 1966 J. A. RABO ET AL 3,236,762
HYDROCARBON CONVERSION PROCESS WITH THE USE OF A Y TYPE GRYSTALLINE ZEOLITE Filed Feb. 7, 1962 REACTION TEMPERATURE,C
YIELD OF ISOHEXANES,VOL. %OF FEED YIELD OF ISOHEXANE 39o O 10 2 0 3O 4O 5O 7O PERCENT Ca ION EXCHANGE BEST REACTION TEMPERATURE,C
MAXIMUM YIELD OF ISOHEXANES,VOL.%
INVENTORS W 9 JULE A RABO JAMES E. BOYLE PAUL E PICKERT A 7' TORNEV United States Patent 3,236,762 HYDROCARBON CGNVERSION PRDCESS WITH THE USE OF A 1 TYPE CRYSTALLINE ZEOLETE .linle A. Rabi), White Plains, Paul E. Pickert, North Tonawanda, and James E. Boyle, Wiliiarnsville, N.Y., assignors to Union Carbide Corporation, a corporation of New York Filed Feb. 7, 1962, Ser. No. 173,321 40 Claims. (Cl. 208-111) This is a continuation-inpart of copending application Serial No. 862,990, filed December 30, 1959 in the name of Jule A. Rabo et al., now abandoned.
This invention relates to a hydrocarbon conversion process and to a catalyst therefor. More particularly, this invention relates to a process for cracking, hydrocracking, polymerization, alkylation, dealkylation, reforming and isomerization of hydrocarbons using a zeolitic molecular sieve catalyst.
Hydrocarbon conversion and the isomerization of hydrocarbons in particular is of special importance to the petroleum industry. In recent years, with the advent of high horsepower gasoline-driven internal combustion motors, a need has arisen for higher octane number gasolines. Natural straight-run gasolines, i.e., naphthas, contain chiefly, normal parafiins, such as normal pentane and normal hexane, which have relatively low octane number, i.e., too low for modern high power requirements. It has become essential, therefore, to convert these low octane components to their higher octane counterparts. The isomerization of these hydrocarbon components accomplish this conversion, i.e., the isomers resulting have a much higher octane rating. Hence, the facility with which this isomerization is accomplished has become of prime importance.
Formerly, straight-run naphtha of low octane quality was used directly as motor gasoline. However, with the above-described need for higher-octane gasoline arising, attempts were made at thermally rearranging or reforming the naphtha molecular for octane number improvement. Reforming is the term employed by the petroleum industry to refer to the treatment of gasoline fractions having a boiling range above about 90 C. to obtain higher octane ratings and improved anti-knock characteristics through the formation of aromatic as well as branched chain hydrocarbons. The thermal reforming of gasoline proved to be inadequate and catalytic reforming in a hydrogenrich atmosphere, in large part, was substituted therefor by the gasoline industry.
In this regard, also, to permit full use to be made of tetraethyl lead (which is less effective with aromatics than with parafiins), high octane parafiins must be incorporated in gasoline blends. Such high octane parafiins can only be obtained from alkylation (which may require butane isomerization) or from the isomerization of pentanes, hexanes, or other light straight-chain hydrocarbons.
Among the isomerization processes known in the art, the most recent have dealt with converting normal paraffins, such as pentane and hexane, to their branch-chain counterparts by contacting, in the presence of hydrogen, the straight-chain hydrocarbons at an elevated temperature and pressure with a reforming type solid catalyst. US. Patent 2,831,908 and British Patent 788,588 relate to such processes. In each of the processes disclosed in these patents, however, a corrosive activator, such as a halide, is employed in the catalyst. Moreover, neither of these processes can be used for isomerizing a mixture of n-pentane and n-hex-ane with a high degree of efficiency.
The catalysts employed for the reforming of gasoline fractions boiling above 90 C., to higher octane products also employ acidic halide activators of objectionably corrosive nature.
It is known in the art to improve the quality of hydrocarbons, particularly petroleum hydrocarbons, by contacting them at various operating conditions with catalysts to eifect the above-mentioned hydrocarbon conversions. The conventional catalytic cracking process is carried out at almost atmospheric pressure (8-20 p.s.-i.g.) at 470- 510 C. on fluid or moving catalyst beds. In catalytic cracking, about 6% of the feed is converted to coke and will be deposited on the catalyst. Therefore, to enable continuous processing, the catalyst must be operated for about 10 minutes in cracking and subsequently for 20 minutes in regeneration by burn-oft". The catalyst used today is a SiO Al O composite with from 12 to 25% Al O content. About of the commercial cracking catalyst is synthetic SiO --Al O gel and 20% is properly processed montmorillonite clay. The only component of the final catalyst is SiO Al O Heretofore only strong mineral and Lewis-type acids have been found to be effective as catalysts for alkylation activity. Many difficulties have been encountered because of the corrosive nature of these strong acid catalysts thereby limiting the operating conditions of the conversion process.
An object of the present invention is to provide an improved process and catalyst for hydrocarbon conversion.
Another object of the present invention is to provide improved process and catalyst for isomerization, reforming, cracking, polymerization, alkylation, dealkylation, hydrogenation, dehydrogenation and hydrocracking of hydrocarbons.
Other objects and advantages of the present invention will be apparent from the ensuing description and appended claims.
By hydrocarbon conversion process is meant, in general, those processes for improving the octane number of gasoline or converting heavy hydrocarbons to light, low boiling hydrocarbons or converting hydrocarbons by hydrogenation or dehydrogenation to, for example, aromatics. Hence, among those processes included in the term hydrocarbon conversion are isomerization, reforming, cracking, polymerization, alkylation, dealkylation, hydrogenation, dehydrogenation and hydrocracking.
According to the present invention, a novel catalyst for the conversion of hydrocarbon comprises a crystalline zeol-itic molecular sieve having at least 40 percent of the aluminum tetrahedra satisfied by the presence of polyvalent meta-l cations and having a SiO /Al O molar ratio greater than 3, preferably greater than 3.3 with best results between about 3.5 and 6. It has been discovered that such a catalyst is effective for converting hydrocarbons in a process in which the mechanism of reaction of the hydrocarbons is postulated to pass through an ionic-type reaction, i.e., through the formation of carbonium ions. It has also been discovered that such a catalyst, when loaded with a catalytically active metal, especially metals of Group VIII of the Periodic Table, in an amount of at least about 0.05 weight percent, is also effective for converting hydrocarbons in a process in which the mechanism of reaction of the hydrocarbons is postulated to pass through an ionic-type reaction.
The term zeolite, in general, refers to a group of naturally occurring hydrated metal aluminosilicates, many of which are crystalline in structure. However, a number of synthetic crystalline zeolites have been prepared. They are distinguishable from each other and from the natur ally occurring material, on the basis of their composition, their crystal structure and their adsorption properties. A suitable method for describing the crystal structure, for example, is by their X-ray powder diffraction patterns.
Crystalline zeolites structurally consist basically of an open 3-dimensioned framework of SiO,, and A tetrahedra. The tetrahedra are cross-linked by the sharing of oxygen atoms, so that the ratio of oxygen atoms to the total of the aluminum and silicon atoms is equal to two, or O/(Al+Si)=2. The negative electrovalence of tetrahedra containing aluminum is balanced by the inclusion within the crystal, of cations, e.g., alkali metal or alkaline earth metal cations. This balance may be expressed by the formula 2Al/(2Na, 2K, 2Li, Ca, Ba, Sr, etc.)= 1:0.15. Moreover, it has been found that one cation may be replaced by another by suitable exchange techniques. Consequently, crystalline zeolites are often employed as ion-exchange agents. The cations are located in the vicinity of the A10 tetrahedra, but their exact location depends on the valency and the size of the cations. The replacement of the cations with other electropositive cations does not induce appreciable changes in the anionic framework. Therefore, any particular zeolitic molecular sieve can be identified independent of the type of cation it contains by the X-ray diffraction pattern and its other chemical components.
It is also known that the crystal structures of many zeolites exhibit interstices of molecular dimensions. The interstitial spaces are generally occupied by water of hydration. Under proper conditions, viz., after at least partial dehydration, these zeolites may be utilized as efiicient adsorbents whereby adsorbate molecules are retained within the interstitial spaces. Access to these channels is had by way of orifices in the crystal lattice. These openings limit the size and shape of the molecules that can be adsorbed. A separation of mixtures of foreign molecules based upon molecular dimensions, wherein certain molecules are adsorbed by the zeolite while others are refused, is therefore possible. It is this characteristic property of many crystalline zeolites that has led to their designation as molecular sieves. As stated heretofore, the novel hydrocarbon conversion processes with which this invention is concerned, has, as one of its essential features, the utilization therein of a novel zeolitic catalyst having the general structure also set forth above.
As previously indicated, a zeolitic molecular sieve is utilized as the novel catalyst for the improved hydrocarbon conversion process. It has been discovered that the catalytic activity of the zeolite employed is strongly dependent on (1) the pore size, (2) crystallinity, (3) the silica-to-alumina molar ratio and (4) the type of cation in the structure.
The pore size is important to the catalytic activity in that it must be larger than the molecules of the feed and the product. The molecules should be admitted and desorbed freely from the structure. Therefore, in hydrocarbon conversion processes only large pore size molecular sieves able to internally adsorb benzene are practical. The pore size can also be defined as one large enough to admit a substantial amount of branched-chain C to C hydrocarbons and to release their structurally rearranged counterparts or isomers.
The crystallinity of the zeolites strongly influences catalytic activity. Zeolite catalysts having crystalline structures are more active than the non-crystalline zeolites with the same chemical composition. The catalytic hydrocarbon reaction occurs at high temperatures. Hence, the crystal structure of the catalyst should be heat stable at the reaction temperature. In this regard, a higher silica-alumina ratio has been found to improve heat-stability.
The silica-to-alumina molar ratio, a measure of the distance between A10 tetrahedra, must be greater than 3 and preferably greater than 3.3 with best results in the range of 3.5 to about 6. This is of fundamental importance. Crystalline zeolites having silica-to-alumina molar ratios higher than about 6 are less desirable since the additional $0.; tetrahedra (lower A10 concentration, hence lower cation density) merely functions as a diluent. Calculations of the bond lengths, from which bond energies are calculated, show that in zeolites having a silicato-alumina molar ratio of 3 or less, such as that found in zeolite X, which is described in detail in U.S. Patent 2,882,244, the distance between the majority of nearest neighbor A10 tetrahedra is such that when the monovalent cations are replaced by divalent (polyvalent) cations by ion exchange, the cation will be located equidistantly between the A10 tetrahedra from which the monovalent metal cations have been replaced. As such, the negative charges on these A10 tetrahedra are equally balanced by the charges on the divalent (polyvalent) cations. On the other hand, in the crystalline zeolites with higher SiO /Al O molar ratios, where the A10 tetrahedra are separated by longer distances, calculations based on the potential energy of the electrostatic attraction of oppositely charged sites have shown that divalent (polyvalent) cations are located nearer to one of the two A10 tetrahedra from which it displaced two monovalent cations during the ion exchange. The bond length, hence bond strength, between the cation and the nearer A10 tetrahedron approaches that of a monovalent cation A10 bond. The remaining bond, however, will be long and of much lower energy. There exists then both positively and negatively charged sites throughout the structure separated by specific distances which are believed necessary to catalyze reactions proceeding through ionic reaction mechanisms.
Because of the three-dimensional character of the crystalline molecular sieve zeolites, it has been found that to be effective as ionic-type hydrocarbon conversion catalysts, at least 40%, preferably above 65%, of the monovalent metal cations should be replaced by polyvalent metal cations. Below this degree of ion-exchange, the ionic activity does not surpass the radical activity of the catalysts to an operable degree.
Therefore, the degree of ion-exchange of the alkali metal cations by polyvalent metal cations is critical and for the process of this invention it has been found that the equivalent amount of polyvalent cations contained on the zeolitic aluminosilicate should be at least 40 percent. Moreover, as more monovalent metal cations are removed, the activity of the catalyst increases. In other words, in the catalyst of this invention, less than 60 percent of the aluminum atoms will be associated with monovalent cations.
The polyvalent metal cations of the catalyst may be a mixture of two or more different metal cations to enhance the activity of the catalysts. Polyvalent metal cations preferred in the invention include aluminum, beryllium, calcium, cerium, chromium, magnesium, manganese, strontium, and zinc.
Additional alkali metal cations may be removed by ion-exchange with ammonium or alkyl-substituted ammonium cations. These may then be driven off during a later activation treatment and also provide additional active sites in the aluminosilicate structure. The production of such additional active sites is called decationization of the zeolite and does not cause destruction of the essential crystalline structure of zeolite Y. Decationization is more fully described in copending application, Serial No. 862,764, issued April 21, 1964, as US. Patent 3,130,006. The description thereof is incorporated herein by reference.
It has been discovered that when a catalytically active metal, especially metals of Group VIII such as platinum or palladium, is provided in finely-dispersed catalytic amounts, that is, 0.05 to 2.0 weight percent, on a zeolitic molecular sieve having at least 40 percent of the aluminum tetrahedr-a satisfied by the presence of polyvalent metal cations and having a silicon dioxide to aluminum trioxide molar ratio greater than 3, a novel catalyst results which, when coupled with the other essential reaction conditions of the present process, will provide a hydrocarbon conversion catalyst. It should be noted, however, that the presence of the metal in amounts higher than 2.0 percent will also catalyze the conversion of hydrocarbons. However, it has been found that the use of more than 2.0 percent of metals such as the noble metals does not substantially enhance catalytic activity and hence is superfluous as well as exorbitantly expensive. The catalytically active metals may be dispersed upon the molecular sieve in their elemental state or as oxides or compounds, such as sulfides, having catalytic properties. Among the metals and their oxides which have hydrocarbon conversion activity are copper, silver, gold, zinc, cadmium, titanium, tin, lead, vanadium, antimony, bismuth, chromium, molybdenum, tungsten, manganese, rhenium, iron, cobalt, nickel and the noble metals of the palladium and platinum groups.
It has been discovered that, although the polyvalent cationic metal containing and non-metal containing zeolite catalysts of the present invention exhibit catalytic activity for all hydrocarbon conversion processes, the polyvalent cationic metal containing zeolite catalysts show improved results in certain specific conversion processes. Included in such processes are: isomerization, reforming, hydrocracking, alkylation and dealkylation. The preferred polyvalent cations are magnesium and calcium, magnesium being particularly preferred, and the preferred loaded metals are palladium and platinum.
The catalytically active metal may be introduced to the crystalline aluminosilicate by any method which will result in the attainment of a highly dispersed catalytically active metal. Among the methods which have been successfully employed are (l) impregnation using an aqueous solution of a suitable metal compound followed by drying and thermal or chemical decomposition of the metal compound; (2) adsorption of a fluid decomposable compound of the metal followed by thermal or chemical decomposition of the metal compound; (3) cation exchange using an aqueous solution of a suitable metal salt followed by chemical reduction of the cation; (4) cation exchange using an aqueous solution of a suitable metal compound in which the metal is in the cationic state with coordination complexing agents followed by thermal or chemical decomposition of the cationic complex.
Methods 1), (2) and (3) are conveniently employed to introduce metals such as copper, silver, gold, cadmium, iron, cobalt and nickel while methods (1), (2) and (4) are suitable for introducing the platinum and palladium group metals. Method (2) is suitable for introducing metals such as titanium, chromium, molybdenum, tungsten, rhenium, manganese, zinc and vanadium. The metal loading techniques of methods (2), (3) and (4) are preferred, as the resulting products exhibit higher catalytic activity than those produced by method (1). The ion exchange techniques of methods (3) and (4) are particularly advantageous since their products have exhibited the highest catalytic activities. Methods (2), (3) and (4) are preferred because of the deposition of the active metal throughout the inner adsorption area of the molecular sieve, the most active dispersion being achieved by methods (3) and (4).
The impregnation method (1) may be practiced in any way that Will not destroy the essential structure of the crystalline zeolitic aluminosilicate. Impregnation differs from the other loading methods of this invention in that the metal is in the anionic part of a water soluble compound. The metal thus is only deposited on the external surfaces of the zeolite. In preparing the catalyst, a water soluble compound of the active met-a1, such as a Group VIII metal, in an amount sulficient to contain the quantity of metal desired in the finally prepared catalyst product is dissolved in water and mixed with the crystalline zeolite. The zeolite is then dried and heated to a temperature sufficient to thoroughly remove the water leaving the metal of the compound in a uniform deposit. Further heating may in some instances be required to convert the metal to its active state, such as heating in hydrogen or other reactive atmospheres.
Method (2) provides a means for depositing the active metals in the inner adsorption region of the molecular sieves. The zeolite is first activated to remove any adsorbed water and then contacted With a fluid decomposable compound of the metal thereby adsorbing the compound into the sieve. Typical of such compounds are the metal carbonyls, metal alkyls, volatile metal halides and the like. The internally adsorbed compound is then reduced thermally or chemically to its elemental metal thus leaving an active metal uniformly dispersed throughout the internal adsorption region of the molecular sieve.
It is preferred that the zeolitic crystalline aluminosilicate be polyvalent ion-exchanged to its desired degree prior to the impregnation step or adsorption step, depending on the method in use, for the reason that some removal of the already deposited metal compound would result if the polyvalent ion-exchange was accomplished after these steps. In order to effect the best distribution of the metal compound on the catalyst, the aqueous solution of the metal compound in the impregnation method should be as concentrated as practical. To this end the best results are obtained it at least some of the water contained in the inner pores of the polyvalent ion-exchanged zeolite has been removed prior to mixing with the impregnation solution. Such removal is effected by heating to about 125 C. Temperatures up to 200 C. may be employed for this drying and will effect a more complete removal of the water.
As stated above the solution should be as concentrated as practical commensurate with achieving a uniform distribution of the metal compound on the zeolite. Practical quantities of water are in the range by Weight of the zeolite from about 30 percent to percent. Less than 30 percent will not wet the zeolite thoroughly enough to distribute uniformly and over 100 percent will allow some solution to run-off with resultant loss of metal compound. About 60 weight percent has been found to produce good results. The thus impregnated zeolite may then be dried by heating to about C. to evaporate enough of the water so that a powder product is obtained which may easily be pelletized in a conventional pellet press. The product may be then stored and the activation process requiring higher temperature, and sometimes hydrogen, treatment may be effected when the pellets have been installed in the reaction chamber for use. The activation method is discussed hereinafter.
The ion-exchange methods (3) and (4) differ since (3) relates to the use of metal salts, such as the chlorides and nitrates of the iron group metals, wherein the metal itself is the cation, whereas (4) relates to the use of compounds of metals, such as the platinum and palladium group metals, in which the metal is contained in the cationic portion of the compound in coordination complex form.
The ion-exchange may be practiced in standard fashion, i.e., the metal compound is dissolved in an excess of water in an amount calculated to obtain the desired amount of metal in the catalyst product. This solution is preferably then added to the previously polyvalent metal cation-exchanged zeolite with stirring and after a sufiicient time has elapsed to allow the ion-exchange to take place the exchanged zeolite is separated by filtration. The ion-exchange of the active metal containing cations into the zeolite is substantially quantitative and the completeness of the exchange process can be detected by chemical tests for the metal in a sample of liquid from the exchanging solution. The filtered zeolite may then be washed to the extent necessary to remove any residual occluded salts followed by drying to produce a pelletizable powder. Decomposition of the active metal containing cation is eifected by heating to above 300 C. and preferably above 400 C. When the metal employed is of the iron-group, it is preferred to conduct this operation in a reducing atmosphere such as provided by hydrogen, methane or carbon monoxide while in the case of the noble metals air may be employed. This is preferably done after the powder has been pelletized, since if it is done beforehand it becomes necessary to perform pelletizing operation in a dry atmosphere to avoid rehydration beyond the extent that is preferred as discussed hereinafter.
The better catalytic activity shown by the metals introduced in the crystalline zeolite by ion-exchange rather than impregnation is believed due to the greater dispersion of the metal within the inner adsorption region of the crystalline zeolite achievable with ion-exchange techniques. It is believed that the metal introduced by ionexchange techniques is dispersed throughout the crystalline zeolite in essentially atomic dispersion. To more clearly show the improved activity of the metal loaded by ion-exchange techniques, two catalysts were prepared from the same zeolite and tested for hydrocracking activity. The two catalysts were prepared from magnesium exchanged (73%) zeolite Y by (1) ion-exchange with the Pd(NH cation, and (2) impregnation with the Pd CL,- anion. Both catalysts were prepared from the same magnesium exchanged zeolite Y preparation and were activated in the same manner, i.e., heated in air to 500520 C. prior to reduction in H at 500 C. The catalytic activity of these catalysts is set forth in the following table:
particular, noble metals such as platinum and palladium, in a range of 0.2-0.6 weight percent.
It should be emphasized that the present catalyst, unlike those of the prior art, does not employ the usual corrosive halide activators, i.e., such as chlorine, fluorine, etc. to enhance its activity. Moreover, the present catalyst is water-resistant under the reaction conditions set forth above. This feature is a direct result of the avoidance of halide activators. If halide activators were present in the catalyst, by adding water, corrosive hydrogen chloride or hydrogen fluoride would be formed and would leave the catalyst. Water amounts up to 1000 parts per million in the hydrocarbon feed, however, are tolerable for short periods of time to the catalyst of the present invention. In the prior art, on the other hand, the water in the feed had to be below 20 parts per million. Hence, the feed in the prior art had to be thoroughly dried before use. This feature is completely avoided by the process of the present invention. It should be emphasized, however, that under certain conditions activators may be employed in the process of this invention. However, even without the use of activators the equilibrium in the isomerization of the hexane and pentane fractions can be approached with facility.
Since the present catalyst containing active metal usually has a relatively high water content after introduction of the metal the present invention includes that it be activated before use since the catalyst is sensitive to the rate at which the water is desorbed. The activation procedure recommended for this catalyst involves the following steps:
(1) The catalyst should be heated slowly in air at 300 C. to 600 C. and preferably at 500 C.
(2) The catalyst should then be heated slowly from room temperature to about 500 C. in a stream of hydrogen gas at atmospheric pressure.
The temperature may be held at about 500 C. for several hours for maximum benefit; the temperature should then be reduced to the hydrocarbon process temperature, adjusting the pressure in the reactor to establish the conditions recommended for conversion. Hydrogen flow-rate during activation should be about 2 liters of gas Mg Y loaded with 1.0 wt.-percent Pd by ion exchange Catalyst Mg Y impregnated with 1.0 wt. percent Pd Process Conditions Weight-hour space velocity. b Volume percent 01 product distilling below 400 F.
The data clearly shows that at the same initial temperature (2800 Q), the catalyst loaded by the ion exchange at N.T.P., per cubic centimeters of catalyst per hour. A
method is more active than the catalyst loaded by impregnation. However, it is more significant to note, that although the catalyst loaded by impregnation had a respectable activity initially, the activity declined rapidly. The activity of the impregnated catalyst can be increased by increasing the reaction temperature but, as the data show, this activity declined at a rapid rate. The ion-exchange loaded catalyst, on the other hand, retained its activity after only a temperature increase of about 15 C. at least for 650 hours.
Moreover, while good results are obtained with catalysts containing the catalytically active metal, especially the metals of Group VIII, dispersed therein in amounts of from 0.05 to 2.0 weight percent, best results are obtained when employing an amount of the metals and in typical temperature program for activation is shown in Table 1.
TABLE 1 Temperature, C.: Time (hours) Room temp. to 1 80 to 2 120 to 1 150 to 200 1 200 to 250 1 250 to 300 1 300 to 350 2 350 to 500 1 500 16 Total 26 9 Among the crystalline zeolites which have been found to be useful in the practice of the present invention, zeolite Y, zeolite L and faujasite are the most important.
The chemical formula for zeolite Y expressed in terms of mole oxides may be written as:
wherein x is a value greater than 3 up to about 6 and y may be a value up to about 9.
Zeolite Y has a characteristic X-ray powder diffraction pattern which may be employed to identify zeolite Y. The X-ray diffraction data are shown in Table A. The values for the interplanar spacing, d, are expressed in angstrom units. The relative intensity of the lines of the X-ray powder diffraction pattern are expressed as VS very strong, S strong, M medium, W weak, and VW very weak.
Zeolite Y is described in copending application Ser. No. 109,487 filed May 12, 1961, in the name of D. W. Breck, issued April 21, 1964, as US. Patent 3,130,007. The
description therein is incorporated herein by reference.
TABLE A h -I-Iz -H' d in A Intensity 14.3 -14. 4 VS. 8. 73- 8.80 M. 7. 45- 7. 50 M. 5. 67- 5. 71 S. 4. 75- 4. 79 M. 4.37- 4. 46 M. 3. 90- 3. 93 W. 3. 77- 3.79 S. 3. 57- 3.59 VW. 3. 46- 3. 48 VW. 3. 30- 3. 33 S. 3. 22- 3. 24 W. 3. 02- 3. 04 M. 2. 90- 2. 93 M. 2. 85- 2. 87 S. 2. 76- 2.78 M. 2. 71- 2.73 W. 2. 63- 2. 65 M. 2. 59- 2.61 M. 2. 52- 2. 54 VW. 2. 42- 2. 44 VW. 2. 38- 2. 39 M. 2. 22- 2. 24 VW. 2. 18- 2. 20 W. 2.16- 2.18 VW. 2. 10- 2. 11 W. 2. 06- 2.07 VW. 1. 93- 1. 94 VW. 1. 91- 1.92 VW. 1. 81- 1.82 VW. 1. 77- 1.78 VW. 1. 75- 1. 76 W. 1. 70- 1.71 W.
When an aqueous colloidal silica sol is employed as the major source of silica, zeolite Y may be prepared by preparing an aqueous sodium aluminosilicate mixture having a composition, expressed in terms of oxide-moleratios, which falls within one of the ranges shown in Table B.
TABLE B Range 1 Range 2 Range 3 NEtzO/SiOz 0.20 to 0.40 0.41 to 0.60 0.61 to 0.80 Si 10 to 40 10 to 30 7 to 30 HzOlNazo to 60 20 to 60 20 to 60 TABLE C Range 1 Range 2 Range 3 NagOlsiOg 0.6 to 1.0 1.5 to 1.7 1.9 to 2.1 Slog/A1203 8 to 30 10 to 30 about 10 H2O/NazO 12 to 20 to 90 40 to 90 The crystallization is conducted by holding the reaction mixture in the temperature range of 20 C. to 125 C. until the crystalline product is obtained. In this range it is preferred to use temperatures of from 80 C. to 125 C.
Zeolite Y may also be produced, when an aqueous colloidal silica sol is employed as the major source of silica, from the following reactant composition ranges which are expressed in terms of oxide-mole ratios:
The reactant mixture is first digested at ambient or roomtemperature and then heated to an elevated temperature and maintained at this temperature until sodium zeolite Y has crystallized. The ambient temperature step is preferably carried out for a 24 hour period and the elevated temperature is preferably C.
The composition of zeolite L may stoichiometrically be expressed in terms of mole ratios of oxides. Thus, a general formula for zeolite L may be represented as follows:
wherein M designates at least one exchangeable cation, as hereinbelow defined; n represents the valence of M; and y may be any value from 0 to about 7. Minor variations in the mole ratios of these oxides within the ranges indicated by the above formula do not significantly change the crystal structure or physical properties of the zeolite. Likewise, the value of Y is not necessarily an invariant for all samples of zeolite L. This is true because various exchangeable cations are of diiftferent size, and as no appreciable modification of the crystal lattice dimensions of the zeolite is effected by the exchange of these particular cations, more or less interstitial space should be available for the accommodation of water molecule. The value of Y therefore depends upon the identity of the exchangeable cation and also upon the degree of dehydration of the zeolite.
The exchangeable cations contemplated by the present invention include mono-, di-, tri-, and tetravalent metal ions, particularly those of Groups I, II, and III of the Periodic Table, as set forth in Websters New Collegiate Dictionary, 1956 edition, page 626, such as barium, calcium, cerium, lithium, magnesium, potassium, sodium, zinc ions etc. and the like, and other cations, for example, hydrogen and ammonium ions, which with zeolite L behave like the metal ions mentioned above in that they may be replaced for other exchangeable cations without causing a substantial alteration of the basic crystal structure of the zeolite.
In making zeolite L, the usual method comprises dissolving potassium or sodium aluminate and alkali, viz., potassium or sodium hydroxide, in water. This solution is admixed with a water solution of sodium silicate, or preferably with a Water-silicate mixture derived at least in part from an aqueous colloidal silica sol. The resultant reaction mixture is placed in a container made, for example, of metal or glass. The container should be closed to prevent loss of water. The reaction mixture is then stirred to insure homogeneity.
For best results, the crystallization procedure is carried out at a temperature of approximately 100 C. The zeolite may, however, be satisfactorily prepared at temperatures of from about 100 C. to about 120 C., the pressure being atmospheric or at least that corresponding to the vapor pressure of water in equilibrium with the mixture of reactants at the higher temperature.
In addition to composition, zeolite L may be identified and distinguished from other zeolites and other crystalline substances by its X-ray powder diffraction pattern, the data for which are set forth below in Tables D and E. In obtaining the X-ray powder diffraction patterns standard techniques were employed. The radiation was the K- alpha doublet of copper, and a Geiger counter spectrometer with a strip chart pen recorder was used. The peak heights, I, and the positions as a function of 20, where 6 is the Bragg angle, were read from the spectrometer chart. From these the relative intensities, 1001/ I where I is the intensity of the strongest line or peak, and d (A.) observed, the inter'planar spacing in Angstrom units, corresponding to the recorded lines were determined.
X-ray powder diffraction data for samples of the potassium form of zeolite L prepared from a potassium aluminosilicate reaction mixture (K L) and from a potassium-sodium aluminosilicate mixture (K-NaL) are given in Table D below. Also included in Table D are X-ray data for isomorphic forms of zeolite L in which varying proportions of the exchangeable cations originally present in the zeolite had been replaced by other exchangeable cations, viz., a 73.2 percent barium exchanged zeolite L (BaL) a 71.3 percent calcium exchanged zeolite L (CaL) a 28 percent cerium exchanged zeolite L (Ce L a 39.1 percent magnesium exchanged zeolite L (MgL), a 41.4 percent sodium exchanged zeolite L (Na L), a 48.3 percent strontium exchanged zeolite L (SrL) and a 22.8 percent zinc exchanged zeolite L (ZnL).
The more significant d(A.) values, i.e., interplanar spacings, for zeolite L are given below in Table E.
TABLE E In an example of the preparation of 8083% calcium exchanged zeolite Y, 1360.0 g. (anhydrous) of NaY having the following composition:
Wt. percent Moles SlO2ZAlzO3, molar ratio 5.08 NazOtAlzOa, molar ratio 0.93
TABLE D KgL KNaL BaL CaL C8314; MgL NazL SrL ZnL 15. 8 100 100 100 100 100 100 100 100 100 7. 89 14 6 38 10 38 12 9 12 15 7. 49 15 14 62 31 94 24 41 32 5. 98 25 16 56 33 94 29 21 44 38 5. 75 11 6 31 18 16 14 12 32 4. 57 32 69 37 75 33 34 32 G5 4. 39 13 13 38 16 63 12 13 32 18 4. 33 13 19 38 29 69 22 23 50 3. 91 30 35 56 33 81 39 34 63 47 3. 78 13 13 13 12 38 14 12 16 18 3. 66 19 18 22 56 2O 16 32 29 3. 48 23 21 62 22 50 24 25 41 38 3. 26 14 23 25 22 25 20 21 28 38 3. 17 34 48 100 47 88 51 46 56 56 3. 07 22 27 50 22 63 29 29 41 38 3. 02 15 14 38 10 25 12 11 31 12 2. 91 23 27 62 31 81 29 29 56 44 2. 65 19 18 44 16 69 22 21 31 32 2. 62 8 16 31 8 38 14 11 12 12 2. 53 8 6 25 4 38 6 5 12 6 2. 45 9 1O 19 6 44 6 9 22 12 2. 42 11 10 25 4 25 10 7 22 9 2. 19 11 10 25 10 56 12 11 28 12 The positions and relative intensities of the X-ray lines are only slightly different for the various cation forms of zeolite L. The patterns show substantially all of the same lines, and all meet the requirements of a unit cell of approximately the same size. The spatial arrangement of silicon, oxygen, and aluminum atoms, i.e., the arrangement of the AlO and SiO.; tetrahedra, are essentially identical in all forms of zeolite L. The appearance of a few minor X-ray lines and the disappearance of others from one cation form of zeolite L to another, as well as slight changes in positions and intensities of some of the X-ray lines, may be attributed to the different sizes and numbers of exchangeable cations present in the various forms of the zeolite.
calcium exchanges were made on the once-exchanged filter precipitate as described above. After the third calcium exchange, the filter precipitate was divided into 2 equal portions and each portion slurried in 2.5 liters of distilled H O. The resulting slurries were individually filtered and as the filter precipitates neared dryness, a hot solution of 1960 g. of calcium chloride in 3 liters of distilled H O was passed over each filter precipitate. Each of the two filter precipitates was then washed until the filtrate gave a negative test for 01- ion with AgNO solution. Three additional calcium exchanges were performed on the material as described above. The final filter precipitate was washed free of Clion. Drying at C. and granulating through a fine screen was completed on Wt. percent Moles/100 g.
In an example of the preparation of chromium exchanged zeolite Y, 533 g. of C1'Cl -6H O in 3.5 liters of distilled H O were charged to a 5-liter, three-necked flask equipped with a mechanical stirrer. fter the chromium salt was dissolved, 230 g. (anhydrous basis) of NaY having the following composition:
Wt., percent Na O 12.0 A1203 23.0 510 64 was added to the solution. The resulting slurry was allowed to stir for an additional /2 hour and then filtered with suction. The once-exchanged filter cake was reslurried in distilled H 0 and filtered with suction. As the filter precipitate neared dryness, 800 g. of
CrCl 6H O in 4 liters of distilled B 0 was passed over the precipitate in a continuous manner. The twice-exchanged material was then washed free of Cl ion until the filtrate washing gave a negative test with AgNO reagent. Analysis of final material calculated on an anhydrous basis was:
Wt., percent N320 Cr O 9.9
In an example of the preparation of magnesium exchanged zeolite Y, 500 g. (anhydrous basis) of NaY having the following composition:
Wt., percent N820 13.7 SiO 63.8 A1 0 22.1
in a 1 liter of distilled H O was charged to a 3-liter, threenecked flask equipped with mechanical stirrer, reflux condenser and heated by glas-col heating mantle.
To this slurry was added 265 g. of MgCl in 500 ml. of distilled H O. The resulting slurry was heated to reflux and stirred for an additional 3 hours. The slurry was filtered with suction and the filter precipitate washed with 2 liters of distilled H O. The once-exchanged material was exchanged two more times using the same procedure as described above. The three-times exchanged material was slurried in 2 liters of distilled H 0 and filtered with suction. As the filter precipitate neared dryness, a hot solution (80-85 C.) of 265 g. of MgCl in 1 liter of distilled B 0 was passed over the filter precipitate in a continuous manner. The four-times exchanged material was then washed with sufficient distilled H O to eliminate all soluble salts. The washed zeolite was then dried in an oven at -l25 C., granulated through a fine screen and tableted into ;-inch cylindrical tablets.
Analysis of the final material calculated to an anhydrous basis was:
Wt., percent N320 4.1 MgO 6.3
In another example of the preparation of calcium exchanged zeolite Y a slurry was formed in a 1-liter beaker of 500 mls. of distilled water and 350 grams of a NaY zeolite having the following composition:
Calculated to Anhy- Found, Wt.-percent drous Basis, Wt.
percent LOI (loss on ignition).
The slurry was then transferred to a 2-liter pressure filter and formed into a filter precipitate. Over this filter precipitate there was passed a solution of 450 grains of CaCl dissolved in 2 liters of distilled water. This was then followed by 2 liters of distilled water, and by another solution of 450 grams of CaCl in 2 liters of distilled water. A final 2 liters of distilled water was thereupon added. The Ca+ exchanged filter precipitate was removed and washed with additional distilled water until the washings gave a negative test for chloride ion with AgNO reagent. The washed material was thereupon dried in an oven at C., and re-equilibrated with water vapor. The material was then found, by percent loss on ignition to contain 25 wt. percent water.
The following example shows a method of preparation of the preferred form of noble metal-loaded zeolite Y by the ion-exchange technique. A 3-liter, S-necked flask equipped with a mechanical stirrer and dropping tunnel was charged with 280 grams of the zeolite prepared in the above example and 500 mls. of distilled water. To the stirred suspension there was added dropwise from the dropping funnel over a period of between one and one-half hours, a solution of 1.97 grams Pt(NH Cl -H O dissolved in 400 mls. of distilled water.
When addition of the Pt(NH Cl -H O solution was complete, stirring was continued for an additional 3 hours and the slurry filtered with suction. The filter precipitate was then washed twice by re-slurrying in fresh 400 ml. portions of distilled water and filtered. The washed filter precipitate was dried in an oven at 125 C., granulated through a 20 mesh screen, and converted into tablets in a pelletizer. The tablets were analyzed and were found to have the following metal and cation composition:
Calculated to Found, Wt. percent Anhydrous Basis,
Wt. percent The tablets as prepared above were activated by heating the tablets in an electric oven to 350 C. to 510 C. for 7 hours.
As has been pointed out heretofore, the catalytic activity of the catalysts of this invention for hydrocarbon converting reactions is superior to the catalytic activity of the non-crystalline aluminosilicates and the crystalline metal aluminosilicates having a silica-to-alumina molar ratio of 3 or less. The superior activity in most hydrocarbon converting reactions can be beneficially utilized to conduct the reaction at a lower temperature than heretofore except in some instances wherein the prior art employed relatively large amounts of corrosive, acidic activators. It
should not be inferred from this that the new catalysts must not under any circumstances have added thereto or to the reactant feed some Lewis acid type halide containing activator. When desired for special effect, activators may be employed. The benefit achieved through the addition of an activator will vary with changes in feed compositions, temperature of reaction, moisture or other impurity in the feed, and the like. In some instances such as in hydrocracking, hydroisomerization and hydrodealkylation the activation may be employed to assist in reaching stable operating conditions more quickly and easily.
The unique activity of the crystalline polyvalent cationic zeolites for hydrocarbon conversion processes is illustrated by the following examples:
CATALYTIC CRACKING The aim of the cracking processes including catalytic cracking and hydrocracking is to produce gasoline from hydrocarbon fractions boiling above the gasoline range. In a few cases the product desired is a specific gaseous hydrocarbon compound such as ethylene, propylene, and the like.
The conventional catalytic cracking process is carried out at almost atmospheric pressure (8-20 p.s.i.g.) at 470- 510 C. on fluid or moving catalyst beds. In oncethrough operation only 55-60 percent gasoline-containing product is produced. Therefore, the unconverted fraction boiling above 200 C. has to be recycled.
The prior art catalytic cracking catalysts split the large hydrocarbon molecules and also isomerize the split small molecules. They do not isomerize, however, the feed components. The lack of activity in isomerizing the paraflin hydrocarbons without splitting the CC bond is characteristic of all known catalytic cracking catalysts. The property is disadvantageous, however, since the isoparafiins can be cracked more easily. Due to the lack of isomerization activity for the feed components, the feed is degraded by dehydrogenation and thermal cracking during the cracking process. Since, in once-through operation, only 55-60% of the heavy boiling feed is converted, recycle operation is used to increase the gasoline yield. Because of the degradation of the feed and the build-up of the aromatic content, however, a substantial Cir 16 quantity of the recycled oil cannot be further cracked and has to be disposed of as cheap by-product.
Table II below, shows the effect of the degree of polyvalent cation exchange on catalytic activity of zeolite Y in catalytic cracking. Similar data is shown for the performance of a representative commercially available cracking catalyst having a SiO/Al O molar ratio of approximately 11. Test procedures were similar to those described by Conn and Connolly (Ind. Eng. Chem., 39, 1138 (1947)). The reactor was of conventional design and was loaded with 170 cc.s of activated catalyst in a 1 inch diameter x 12 inch long bed. The temperatures in the catalyst bed were monitored throughout the run by thermocouples placed at the top, 3 inches from the top and the bottom of the bed in a thermowell located axially in the center of the reactor. The feed stock was a heavy virgin gas oil with approximately a 450 C. end point (Sp. Gr. 60/60 F.=0.864). It was introduced at atmospheric pressure at a weight-hourly-space-velocity of 2.0 for a period of 60.0 minutes. At the end of the run, the reactor was purged with N at 2 s.c.f.h. for a period of 1.0 minutes. The gaseous hydrocarbons released during the purge were not collected. The cracked liquid product was collected at atmospheric pressure and room temperature. All of the noncondensed gases were collected by water displacement and analyzed by gas chromatography. The weight and volume of the liquid product where measured and a portion distilled using a modified ASTM (D-86) distillation procedure. The distillation was stopped when the vapor temperature reached 400 F. The weights and volumes of the initial boiling point (IBP) to 180 F., 180-400 F. and 400 F. (gas oil) fractions were obtained. The IBP-180 F. cut was analyzed by gas chromatography and combined with the analysis of the non-condensed gases to give total C -180 F. product composition. The catalyst was removed from the reactor and the weight of the product retained on the catalyst was computed. Material balances, which ranged from 90-97 wt.-percent were drawn on the unit for each run. Yields were calculated on a no-loss basis.
Table III, below, shows the eflect of the crystalline zeolites SiO /Al O ratio on catalytic cracking activity. The test procedures were identical to those described for the tests of Table II.
TABLE II Efiect of degree 0 polyvalent cation exchange on catalytic cracking activity of zeolite Y Temp, C, 375 375 375 375 480 350 350 350 350 350 Catalyst Composition:
Percent Mg+ Exchange 0 24 40 d 71 Amorphous 40 52 64 d 71 Percent Cation Deficiency. 15 12 8 5 SiOzAl O 8 0 8 5 Percent Na+ Not Exchanged... 85 64 52 24 Reference 48 28 24 SlOzZAlzOz Molar Ratio- 5.0 5. 0 5. 0 5.0 Catalyst 5.0 5. 0 5.0 Gasoline Yields 11.9 25.4 36.0 34. 6 34. 5 29. 1 25.0 26. 3 Selectivity 97. 7 97. 5 92. 9 90. 9 84. 3 96. 5 96. 7 93. 3 D L c 15.0 31. 5 48. 5 46. 7 45.0 39. 5 32.0 35. 5 Yields, Wt. Percent of Feed:
z 0. 01 0. 01 0. 03 0.14 0. 01 0. 01 0. 01 0. 02 0. 11 0. 12 l. 22 0. 02 0. 01 0.02 0.07 0. 5 0. 4 1. 57 0. 12 0.09 0. 16 0.5 2.6 3.6 5.14 .0 1.1 0.8 2.1 1.2 5.0 7.0 7.96 .8 3.4 2.5 4.5 1.4 4.0 4.4 4.49 .0 3.2 2.2 3.4 2.1 4.5 4.3 6.54 .9 4.0 2.8 3.4 21. 9 27. 6 25.9 23. 51 .2 21.8 19. 9 19. 5 64. 2 42. 3 42. 2 44. 31 .4 55. 2 60.8 54. 5 8. 7 13. 6 12.1 5. 12 6 11. 1 10. 8 12. 1 77.2 79. 1 79. 8 73. 5 .0 81.9 82.4 81. 8 82. 5 91. 5 89.1 86.6 91. 4 91. 7 91. 7 91. 2 Sp. Gr. 400 F.+, 60/GO F 0.858 0.861 0.862 0.885 0.854 0.856 0.855 0.853 0.856
' Wt. percent 05-400" F. gasoline, based on feed.
11 Wt. percent C4-400 F. product in C1-400 F. product. 0 Vol. percent of cracked liquid product distilling below 400 F.+gas losses.
d 64% Mg+ exchange, 7% Ca+ exchange.
Aerocat 75/85 sold by the American Cyanamid Co.
TABLE III E act zeolite SiO :Al 0 ratio on catalytic cracking 2 2 activity Temp., C 430 430 415 330 350 375 400 425 375 Catal st com os't'o P rcent hil g f l lii change Amorphous f 77 75 77 79 d 71 71 d 71 83 5253335 132 130 332332.9311:3:31: 43 23 23 43 21 2.? 2? 2i 19 210 2 11201, Molar Ratio g g g fi 2. 3 2 4. 4% 5. 1 5. 5. 5%
e0 1 Q Gasoline Y1eld 34. 5 40. 0 35. 4 32. 9 23. 3 34. 5 40. 0 30. 9 23. 4 Selectivity 34.3 39.7 95.2 95.3 94.2 00.9 91.7 35.7 33.5 ga u d 45. 0 50. 0 43. 5 41. 5 33.0 45. 7 53.5 54.0 27.0
ie t s percent 0 e5 H2 0.14 0.10 0. 03 0. 01 0. 01 0. 03 0. 03 0. 05 0. 01 1. 2 0. 30 0. 13 0. 00 0.05 0.12 0. 23 0. 25 0. 01 1.5 1. 2 0. 0. 21 0.15 0.44 0. 04 1.2 0.31 5.1 3.3 1.4 1.1 1.3 3.0 3.4 5.9 3.5 3.0 5.2 3.1 2.5 4.5 7.0 7.3 11.2 7.4 4.5 4.0 2.0 2.1 3.5 4.4 5.7 0.7 2.7 0.5 3.3 3.4 3.1 3.9 4.3 5.4 5.0 2.3 23. 5 32. 1 29. 4 27. 5 21. 4 25. 9 27.9 24. 2 13. 5 400 F.+ 44. 3 43. 5 52. 9 53.5 53.1 42. 2 33. 2 33. 0 53.4 Catalyst Depos1t. 5.1 4. 9 5. 5 9. 5 11.5 12.1 10. 2 11.4 11.9 i-C in 04, Wt. percent 73. 5 52. 2 75 79. s 31. 9 79. 3 79. 0 75. 3 79. 2 M in 05, Wt. pereent 35. 5 57. 2 35 90. 4 91. 3 39.1 90. 5 3s. 5 03. 3 Sp. Gr. 400 F., /00 F 0. 335 0.304 0. 353 0.353 0. 355 0.302 0. 359 0. 374 0.350
9 Wt. percent C5-400 F. gasoline, based on feed.
b Wt. percent (Jr-400 F. product in C1-400 F. product.
9 Vol. percent of cracked liquid product distilling below 400 F. d 64% Mg+ exchange, 7% 03+ exchange.
3 Calcium cation exchanged.
f Aerocat /85 sold by the American Cyanamid C0.
The data of Table II shows that the polyvalent exchanged zeolite Y catalyst is considerably more active at 350 C. than the commercial amorphous reference catalyst at 350 C. Furthermore, these data show that a zeolite Y catalyst having at least 40% of its cation sites satisfied by polyvalent cations is as active at 375 C. as the commercial reference catalyst is at its optimum operating temperature of 480 C.
The data of Table III shows the polyvalent zeolite Y having a SiO /Al O molar ratio greater than 3 is as active, at considerably lower temperatures, as the commercial reference catalyst and zeolite X. At 350 C., it has been found that magnesium exchanged zeolite X has no catalytic cracking activity. At 350 C., and 375 C. on the other hand, magnesium exchanged zeolite Y exhibits approximately the same activity and considerably more selectivity (less C -C products) than the commercial reference catalyst at 480 C. With higher SiO /Al O molar ratios, progressively lower temperatures are necessary to achieve comparable activities with the commercial reference catalyst at 480 C. At 375 C., magnesium exchanged zeolite Y (SiO /Al O =5) had the same activity as the commercial reference catalyst at a 100 C. higher temperature. In general, because high conversions are obtained at lower temperatures with the polyvalent cation exchanged zeolite Y catalyst, better quality products are obtained than the commercial reference catalyst or zeolite X. The zeolite Y catalyst produces C and C products containing substantially higher quantities of the higher octane number branched-chained isomers.
The process conditions for the cracking reaction using polyvalent metal cationic forms of zeolite Y are:
The major difference between catalytic cracking and hydrocracking is in the application of a considerably higher hydrogen partial pressure in the hydrocracking process. Generally, metals with catalytic hydrogenation activity are preferably incorporated into the hy-drocracking catalysts. These components reduce the tendency for coke to accumulate on the catalyst, hence increase the catalyst life.
The performance data of the polyvalent cationic zeolite Y catalysts for hydrocracking n-hexane, are shown in Table IV. Similar data obtained with an amorphous commercial cracking catalyst are also shown. The hydrocracking experiments were run in the same reactors used for catalytic cracking while operating at a total hydrogen plus hydrocarbon pressure of 450 p.s.i.g. and a hydrogen-to-hydrocarbon molar ratio of 35 1.
Comparison of the hydrocrackzng activity of zeolite catalyst with amorphous commerical SiO A1 O catalyst TABLE IV Catalyst Ca(82%)Y Commercial Amorphous Mg(72%)Y- Cracking Catalyst SiO /Al,0
Reaction Temp, C 375 400 425 425 450 475 345 390 Pressure, p.s.i.g 400 450 450 450 450 450 450 450 W.H.S. 1.0 1.0 1.0 1.0 1.0 1.0 2.0 2.0 3:1 3:1 3:1 3:1 3:1 3:1 3:1 3:1 17. 5 19.3 24. 5 6.0 8. 5 11. 5 4. 0 9. 0 ee n-Hexane n-Hexane n-Hexane Mel Percent Conversion 25. 0 35. 9 39. 4 36. 6 45. 49. 6 25.4 Analysis, Mol Percent:
n-Hexane 75. 0 04. 1 60. 6 63. 4 54. 5 50. 4 74. 6 51. 9 g-gngthyitpleiigareu 7. 8 9. 8 9. 1 4. 7 3. 1 2. 6 8. 1 10. 0
, lme y u ane" 9. 9 11.3 11.3 0. 3 4. 7 2,2-dimethylbutane 0. 2 0.4 0. 3 0. 1 0. 1 n-Pentane 0. 7 0. 8 iso-Pentane 2.8 3. 3 n-Butanc 2. 2 3.0 iso-Butane 1. 6 3. 3 4. 0 5. 6 7. 9 Propane 2. 3 5. 7 7. 7 13. 2 21.2 Ethane. 0. 1 0.3 0. 6 0. 5 1. 0 Total iso-C 17. 9 22.0 21.2 11.1 7. 9 Volume of Catalyst Used, cc 80 80 80 80 The performance data show that the CaY and MgY zeolite catalysts convert as much of the n-hexane feed at a 25-50" C. lower temperature as the amorphous cracking catalyst. At the same level of conversion, the production of branched-chain isomers of the feed was over twice as much (even when operating at twice the weight-hourlyspace velocity) with the zeolite catalyst and with a corresponding lower production of less desirable propane. This isomerization of the feed leads to a higher octane product boiling in the useful gasoline range.
A comparison of the hydrocracking activity of an amorphous aluminosilicate, a zeolitic molecular sieve having a SiO /Al O ratio of less than 3, and a catalyst of the present invention, all having a polyvalent cation and 1.0 wt. percent palladium loaded thereon is shown in Table V. The feed stock for these runs was a No. 2 fuel oil having a distillation end point of 590 F. This data clearly shows the improved results obtainable with the catalysts of this invention. Metal-loaded zeolite Y operating at a temperature of 200 C. less than the amorphous zeolite and 100 C. less than the low SiO /Al O ratio molecular sieve converts a substantially greater amount of C product distilling below 400 F.
TABLE V Catalytic hydrocracking activity of polyvalent cation (Mg+ exchanged supports loaded with 1.0 wt. percent Pd Zeolex 238D, manufactured by J. M. Huber Corp.: Amorphous aluminosilicate with base exchange capacity: SiOzZAlzOa molar ratio=10.
Vol. Percent of C4+Produet distilling below 400 F. minus volume present (20%) below 400 F. in feed stock.
Feed stock for above run-No. 2 fuel oil:
Distillation range, F. 250-590 Sp. Gr. 60/60 0.843 Gasoline content, vol. percent 20 The hydrocracking is preferably carried out at temperatures of 150400 C., particularly 250-350 C., at pressures of 300-2000 p.s.i.g., particularly 400-1000 p.s.i.g., at weight-hour space velocities of 0.55.5, particularly 1-3, and at a hydrogen to hydrocarbon mole ratio of 540, particularly -20, for the metal loaded catalyst of the present invention. When the non-metal loaded catalyst of the present invention is employed, the hydrocracking is preferably carried out at temperatures of 300-600 C., particularly 400-500 C., at pressures of 3002000 p.s.i.g., particularly 400-1000 p.s.i.g., at weighthour space velocities of 0.5-1, and at a hydrogen to hydrocarbon mole ratio of 15, particularly 3 Magnesium and calcium are the preferred polyvalent cations and palladium and platinum the preferred loaded metals for the hydrocracking process, magnesium and palladium being particularly preferred.
A'LKYLATI ON The ability of crystalline polyvalent cationic zeolite Y to catalyze alkylation reactions is particularly unique. It
is generally an accepted fact that this type reaction, which may be illustrated by the addition of olefins to aromaticring containing compounds, occurs through an ionic-type reaction mechanism. Heretofore, only strong mineral and Lewis-type acids have been found to be eflective catalysts. The ability of our zeolite catalysts is not only unique, but advantageous, since many of the diffioulties encountered in the use of the corrosive strong acid catalysts are not prevalent with the non-corrosive zeolite catalysts. For example, the zeolite catalysts are readily removed from the reaction mixture in liquid-phase alkylations by simple filtration or centrifugation. Furthermore, they are readily adaptable to either liquid-phase or vaporphase alkylations, thereby increasing process flexibility.
In this embodiment of the present invention, isoparaflins and aromatics can be alkylated. Typical of the feed stocks for such a conversion process are iso C -C paraffins plus gaseous C -C olefins or aromatic hydrocarbons such as benzene and substituted benzenes such as phenol and chlorobenzene plus gaseous and liquid C -C olefins. The alkylation process is preferably carried out at temperatures of 20300 C. at pressures of atmospheric- 1000 p.s.i.g., particularly atmospheric 700 p.s.i.g. for the non-metal loaded catalyst of the present invention. The metal-loaded catalyst of the present invention will generally operate at lower temperatures than the non-metal loaded catalyst and yield a greater amount of polyalkylated products.
The examples described below illustrate the catalytic activity of the crystalline polyvalent cationic zeolite catalysts in alkyaltion reactions.
To 156 g. (2.0 moles) of benzene was charged 10.0 g. of activated Ca (82%) Y powder. Propylene gas was then bubbled into the slurry, at atmospheric pressure, at a rate of approximately cc. per minute. The temperature of the slurry rose, due to the exothermic reaction, from 27 C. to 35 C. and remained at this temperature for approximately 30 minutes. A sample of this slurry at this point was analyzed by vapor chromatography. The analysis showed that the reaction mixture contained approximately 3 mole percent isopropylbenzene. The slurry was then heated to reflux temperature as the flow of propylene gas was continued and maintained at this temperature for 2-3 hours. Analyses showed the reaction mixture contained 7.2 mole percent isopropylbenzene. No other products were formed.
The following examples illustrate the alkylation activity of polyvalent cation exchanged zeolite Y with and without incorporated catalytically active metals. The data show that although the polyvalent cation exchanged zeolite Y catalysts alone possess unique catalytic activity in ionic type hydrocarbon conversion reactions, the incorporation of small quantities of catalytically active metals, particularly those of the platinum group of Group VIII, into the zeolite, preferably by the ion exchange techniques previously described further enhances the basic activity of the zeolite. The same procedures were used for the metal-loaded and non-metal-loaded catalysts.
Fifteen (15) grams of catalyst (in the powder form), activated by heating to 500 C. in an oven purged with air, was slurried in 39 g. (0.5 mole) of benzene and charged to a 300 ml. pressure bomb. Ethylene was then charged to the bomb until the pressure was approximately 750 p.s.i.g. at room temperature (-25 C.). Simple gaslaw relationships (PV=nRT) indicated benzene to ethylene ratio of approximately 1 at this pressure. The bomb was then heated in a rocker-type apparatus until reaction commenced, as indicated by a drop in pressure. It was. held at temperature until the pressure was constant. With each catalyst, the final gage pressure was zero when the bomb was cooled to room temperature. This showed substantially complete conversion of all the ethylene. The contents of the bomb was removed, filtered to remove the catalysts and the liquid filtrate distilled. Pertinent process conditions and product distribution data are summarized as follows:
1 Temperature at which reaction commenced, i.e., pressure started to a Time until pressure, at; reaction temperature, was constant.
ISOMERIZATION The catalytic activity of zeolite Y, having at least 40 percent of the alumina tetrahedra satisfied by polyvalent metal cations, for the isomerization of parafiinic hydrocarbons is demonstrated by the unexpectedly high yield of isohexanes in the hydrocracking process. When an active metal such as one of the platinum group of Group VIII, is present finely dispersed in the pore systems of the catalyst, it is a preferred catalyst for the isomerization process. The preferred isomerization process contemplated for the present invention is not only dependent on the catalyst hereinbefore described, but, is also dependent on such features as (1) reaction temperature, (2) space velocity, (3) hydrogen-to-hydrocarbon ratio and, (4). reaction pressure.
With respect to temperature, the isomerization process should be carried out at a range of between 300 C. and 425 C. It is preferred, however, when isomerizing a pentane fraction, to carry out the reaction at a temperature of between 350 C. and 380 C. It is also preferred, however, when isomerizing a hexane fraction to carry out. the reaction at a temperature of between 340 and 370 C. As a consequence of the closeness of the optimum isomerizing temperatures for pentane and hexane fractions, another surprising advance represcnted by the invention has also been found, i.e., it is now possible to isomerize mixtures of normal pentane and normal hexane fractions. In this regard, it has been discovered that a mixture of pentane and hexane can be effectively carried out employing the Pt or Pd catalysts at a temperature of between 350 to 375 C. In this range, the effective isomerization of both the pentane and hexane fractions will result. However, it is to be understood that the higher temperature limit of the just-mentioned temperature range will more effectively isomerize the pentane fraction, but would crack a higher proportion of the hexane to gas products. While, conversely, the lower temperature limit of the temperature range will isomerize the hexane efliciently but will result in the conversion of the pentane fraction being less. Moreover, with the addition of activators of Lewis acid type halides or other type hydrocarbons may be isomerized at temperatures substantially lower than 300 C.
It is to be clearly understood, therefore, that temperature is an exceedingly critical factor in the process of the invention. It is essential that the temperature of isomerization not be carried out above 425 C. since undue cracking will occur. Indeed, even above 400 C. hydrocracking becomes significant and decreases the net yield of liquid product. However, butane can be effectively isomerized even at about 425 C.
A particularly good optimum temperature for isomerizing a pentane fraction has been found to be 380 C. A particularly effective isomerization temperature for hexane fraction has been found to be 360 C. A particularly effective temperature for the isomerization of a mixture of hexane and pentane fractions has been found to be 370 C. While isomerization of other parafiins besides normal pentane and normal hexane can be accomplished by the process of our invention, it is to be understood that isomerization of these lastnamed paraffins is of the utmost importance, because of their importance in gasoline upgrading.
With respect to the space velocity, the reaction should be carried out at a velocity of about 1 .to 10 grams of feed per gram catalyst per hour. It is preferred, however, to carry the reaction out at a velocity of 2 to 5 grams feed per gram catalyst per hour. It has been found that as the space velocity is increased the yield of isoparafiins is decreased at a given temperature. It should be noted, however, that when the space velocity is kept constant, an increase in temperature will result in the yield of isoparafiins increasing steadily to a maximum. However, as aforementioned, after passing above the optimum temperatures set forth above, hyd-rocracking of the feed will increase. In this regard, the selectivity of the isomerization reaction is very high up to the optimum temperature. Above this temperature, however, the selectivity decreases.
The hydrogen-to-hydrocarbon ratio should extend from 0.5 to 1 to 10 to 1. It is preferred, however, that the hydrogento-hydrocarbon ratio extend from 2:1 to 5:1.
The pressure at which the isomerization reaction of this invention is carried out should extend from about 100. to about 1000 p.s.i.g. It is preferred, however, that the pressure of the reaction range from about 300 to about 600 p.s.i. At constant contact time the reaction appears to be favored by lower pressures. At low reaction temperatures the selectivity of the catalyst does not appear to be affected by the total reaction pressure. However, if the reaction temperature is increased above the optimum some hydrocracking will result. In this case higher operating pressures have been found to reduce the amount of this hydrocracking.
As indicated, heretofore, the invention has severaladvantageous features over those processes found in the prior art. Foremost among the advantages is the quality of the product itself. In this regard, the amount of isomers produced by the process has been found to closely approach the equilibrium in all normal paraflin to isomeric-paraflin relationships. The amount of 2,2- dimethylbutane contained in the hexane isomerizate formed by the process has been found to be more than 12 mol percent. This isomer, i.e., 2,2-diemethylbu'tane, is among the most important of all the isomers formed in the reforming of straight-run gasoline fractions. Hence, a 12 and greater percent yield of this isomer represents a real contribution to the art.
Another advantageous feature resulting from the process is the fact that the corrosive activators formerly employed in all the known prior art processes need not now be employed.
A third advantageous feature of our process is that as a result of the fact that corrosive activators are not required, the feed of hydrocarbons need not be specially dried prior to contacting the same with the novel catalyst of our invention.
A still further advantage, is that the optimum isomerization temperatures are generally lower than those temperatures found in the prior art. In addition, the optimal isomerization temperature for the pentane and hexane fraction is very close i.e., only 10 to 15 C. apart. As aforementioned, this factor enables a mixture of the pentane and hexane fraction to be isomerized simultaneously.
As aforementioned, it has been found that (1) the type of the cation, (2) the crystallinity, (3) the silicaalumina ratio and, (4) the pore size, influence the activity of metal-loaded zeolitic catalysts. To establish the importance of these effects, different type zeolites were loaded with 0.5 weight percent platinum and after subsequent activation were tested for their isomerization activity using n-hexane as feed with 100 cu. cm. volume of catalyst in a high pressure testing unit.
Among the zeolites prepared and tested were an amorphous zeolite of high silica-alumina molar ratio and certain crystalline zeolites. One of the crystalline zeolites had a small pore size and the other, the metal-loaded zeolite Y of the invention, had a large pore size and a high SiO /Al O molar ratio. All the zeolites selected were prepared in two diflFerent cationic forms, one of each of the zeolites containing sodium or calcium. The characterization of the metal-loaded zeolitic catalysts used and the data for the activity are collected in Table VI which follows. The yield of the isohexane and the amount of 2,2-dimethylbutane present in the product indicates the isomerization activity. The amount of dimethylbutane in the product indicates high activity of the catalyst because the production of this compound is much more difiicult than the production of the monomethyl pentanes.
TABLE VI 24 all for isomerization. As can be seen from Table X, the calcium ion-exchanged zeolite Y with platinum as the active metal has far more activity for isomerization than the other zeolites.
The catalytic properties of the noble metal-loaded zeolites indicate that zeolite Y is superior, in catalytic application to hydrocarbon reactions involving ionic reaction mechanisms, to any other known catalyst.
A series of 0.5 percent platinum-loaded polyvalent metal ion-exchanged zeolite Y catalysts were also prepared and tested. The degree of equivalent ion-exchange for these zeolites was between 60-80 percent. In the selection of the cations used for the ion exchange of the original sodium cations, mono-, di-, tri-, and tetravalency cations were used. Each group contained cations of different electropositive character. All ion-exchanged zeolite Y catalysts retained the original crystal structure. The ion-exchanged and platinum-loaded samples were first activated at 350 C. in air and then in hydrogen prior to being tested for isomerization activity using n-hexane feed, and for reforming activity using gasoline feed. The experiments were carried out in a 100 cubic centimeter volume high-pressure experimental unit.
The experimental data collected. in Table VII also shows that zeolites having alkali metal cations have lsomerization activity of platinum loaded zeolites isomerization of n-hexane The reaction conditions for the above tests were:
Space Velocity: 2.0 g. hydrocarbon teed/g. catalyst/hour. Pressure: 450 p.s.i.g. Hydrogen to Hydrocarbon Mole Ratio 5:1.
Mol per- Pt Metal Liquid Yield cent 2,2- Reaction Type of Zcolite Sim/A1 0 Cation Distribution Content, wt. Recovery, iso-Ca, dimethyl Tempera- Molar Ratio equiv. in percent percent percent M01 Butane in ture, C.
percent Liquid Product Amorphous Zcolitc 5. 3 100 Na 0.50 96. 7 1 4 0. 5 450 Amorphous Zeolite (not a 5. 3 Na, Ca 0.42 90+ 5 0 0.2 400 molecular sieve).
Erionitc- 6. 6 100 (Na-l-K) 0.70 91. 6 6. 6 0.1 400 Erionite (Pore Size. 4.8 11.). 6. 6 21 (Na-l-K). 79 Ca 0.47 71.0 3.4 0.1 400 Metal-loaded Zcolite X 2. 4 100 Na 0. 49 99. 0 1. 0 0.9 450 (Pore Size, 9-1011.) 2. 4 13 Na, 87 Ca 0.46 91.0 9.6 0.2 400 Metal-loaded Zeolite Y 4. 7 100 Ne 0.50 95.1 0. 4 0.2 400 (Pore Size. 10A.) 4. 7 20 Na, 80 Ca 0. 43 97. 0 66. 5 10. 4 400 These tests revealed that platinum loaded alkali 5 no, or very poor, activity and that only the lithium metal cationic zeolites do not have appreciable activity as isomerization catalysts. The alkali metal cationic erionite zeolite contained potassium and sodium cations and had a pore size of about 4.8 angstrom units. It adsorbs only the n-paraffin molecules but the larger isoparafiins are not admitted. The relatively low yield indicates that a small percent of the adsorbed n-hexane may have been converted to iso-paraflins, but these are trapped and decomposed. The other alkali metal cationcation provided a small degree of catalytic activity. All of our polyvalent cationic catalysts have an appreciable degree of activity. In the divalent cationic zeolites, all cations like Mg++, Ca++, Sr++, Zn++ and Mn++ provided very high isomerization activities. The tests show, however, that calcium exhibits catalytic activity to a greater degree than the other polyvalent cations. Moreover, calcium not only provides the highest activity, but is also able to retain the activity for a longer period of ic zeolites containing sodium cations have no activity at time.
TABLE v11 [Feed=n-hexane] Mol Percent Distribution of Pt Wt. Per- Liquid Re- Mol Percent ,--di- Reaction Catalyst Type Cations in Percent cent Content covery Yield iso-Ca methyl Temp, C.
butane in Product 33 Na, 67 L1 0. 28 91.6 9. 2 0. 2 400 100 Na- 0.5 95. l 0.4 0. 2 400 2 Na, 98 0. 45 100 0.0 0.0 450 19 Na, 81 Mg- 0.39 94. 2 55. 4 6. 7 400 20 Na, 80 Ca 0.43 97.0 '66. 10. 4 400 17 Na, 83 Sr 0.43 78. 6 32. 1 4.1 400 37 Na, 63 Zn 0. 48 97 60 6. 8 400 6 Na, 74 Mn 0.47 92.6 62 10.2 400 34 Na, 66 Ce 0. 45 99 66 10. 5 400 32 Na, 68 AL- 0.43 87.7 50. 3 6. 9 400 32 Na, 68 Ce 0.34 100 42. 5 2. 3 400 1 Reaction conditions were: 7
Space velocity: 2.0 g. feed/g. catalyst/hour.
Pressure: 450 p.s.i.g.
Hydrogen to Hydrocarbon Mole Ratio: 5:1.
The catalytic potency of the metal-loaded zeolitic cathydrocarbons. An example of such a mixture a pealyst of our invention for cracking, reforming and other ionic-type hydrocarbon reactions is based, in large part, on its active aluminum site. In order to stabilize the aluminum in its most active form, the SiO /Al O molar ratio of our catalyst should reach a-considerably high value so that substantially all of the'available aluminum oxides will be in a highly active form. For this reason, the SiO /Al O ratio should be greater than 3 and preferably greater than 3.3.
It has been discovered therefore, that a polyvalentexchanged zeolitic molecular sieve having preferably a metal of Group VIII of the Periodic Table finely-dispersed thereon in an amount of at least about 0.05 weight percent, having at least 40 percent of the aluminum oxygen tetrahedra associated withpolyvalent metal cations, having a silica to alumina ratio of greater than 3 and having a pore size sufiicient to release the hydrocarbon isomer products is an excellent, superior catalyst under the reaction conditions previously set forth.
In this regard, as aforementioned, the metal of Group VlIIshould be added to the zeolite and dispersed thereon by ion-exchanging procedures previously described, though other procedures may be employed. Platinum troleum light gasoline fraction. Moreover, our isomerization process will also isomerize branched chain hydrocarbons to their further branched chain isomers, e.g. monomethyl pentane can be isomerized or converted to dimethylbutane.
The isomerization of substantially pure normal pentane and normal hexane is also of great use in the chemical industry, however. Suitable charge stocks for our process, however, also include paraffinic fractions rich in normal pentane and normal hexane which are separated from the products of conversion processes. For example, a suitable charge is a light parafiinic fraction, rich in n-pentane and n-hexane which is obtained by distilling the reformate from a naphtha reforming process into a light and heavy fraction. Suitable paraifinic fractions can also be obtained from the reforming products of other separation methods such as solvent extraction, molecular sieve adsorption separations etc.
To exemplify the superior performance of the catalysts of this invention, a calcium exchanged zeolite Y molecular sieve loaded with platinum was used to isomerize a stream of normal hexane. The results of the ten runs of this test are shown in Table VIII below.
TABLE XIII Run No 1 2 3 4 5 6 7 8 9 10 Reaction Temp, C 350 375 375 400 375 400 425 375 400 400 Pressure p.s.l.g 450 450 450 450 450 450 450 450 450 460 g./fecd/g. catalyst/hour. 2. 0 2. 0 2. 0 2. 0 2. 0 2. 0 2. 0 2. 0 1. 0 2. 0 H /Hydr0carb0n 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 Hours on Stream 27 30 57 5O 6 27 30 25 104 126 Liquid Yield, V01. percent 94. 7 97.1 96. 7 84.1 99. 4 94. 2 87. 1 100 91. 4 96. 7 Liquid Analysis, Vol. percent:
n-Hexane 19. 8 17. 2 22.4 21. 4 32. 1 22. 8 22.7 27. 0 21.9 24. 2 3-meitlliylpengane. 18. 1 15. 2 19. U 18. 0 18. 5 19. 1 18. 7 19. 5 20. 3 19. 3 2-1ne y pen ane. zygmmethylbutanen 36. i 31.0 38. 2 35.6 34. 9 37. 3 36. 4 37.1 36. 9 36. 9 2,2-dlmethylbutane- 12.0 10. 1 12. 4 10.8 8. 7 11. 2 10.4 9. 6 10. 6 9. 7 Total hexanes 86. 6 73. 5 92. 0 85. 8 94. 2 90. 4 88.2 93.2 89. 7 90. 1 Yield of iso-hexanes 1- 63. 3 55.0 67. 3 54. 2 61. 7 63. 7 50.3 66.2 62.0 63. 7
and palladium are those metals of Group VIII par- To further exemplify the superior performance of a ticularly preferred in the practice of this invention, especially where they are present in amounts of from 0.2 to 0.6 weight percent and more particularly where the equivalent polyvalent metal cation content of the catalyst is preferably at least 65 percent.
The feed of hydrocarbons may comprise straight-run gasoline fractions consisting essentially of pentanes and hexane in substantially pure state, separately, or in a mixture of the gasoline boiling range. The latter will calcium exchanged zeolite Y molecular sieve loaded with platinum in accordance with the invention, three runs were made employing a stream of normal pentane as feed. The results are included in Table IX below as runs 1 through 3 inclusive. Also indicated in Table IX below as exemplifying the catalytic action of the metal loaded calcium-exchanged zeolite Y molecular sieve are the results of two runs employing a mixed hexane-pentane feed.
comprise normal pentane and normal hexane and other These are represented as runs 4 and 5.
TABLE IX Run No 1 2 3 4 5 Reaction Temp., C 350 400 450 400 42 5 Pressure p.s.i.g 450 450 450 450 450 W.H.S.V. g./g./hr 2.0 2.0 2.0 2.0 2.0 II /Hydrearb0n :1 5:1 5: l 5: 1 5 :1 Hours on Stream 131 149 152 174 177 Liquid Yield Vol. Percent 97.6 95. 2 95. 5 91.3 Liquid Analysis Vol. Percent n-hexane 12.0 10. 9 3met}l:ylp 8. 3 8. 3 2met y pentane 2,3-dimethylbutane 6 8 2,2-dimethylbutane 3. 1 3. 7 Iso-pentane 9. 9 49. 8 50.4 22. 2 26. 4 n-Pentane 87. 1 48. 5 37. 5 33. 4 30. 4 Octane Number F-l Clear 72. 3 73. 7 Octane Number F-l L aded 90. 7 92. 0
To investigate the catalytic effects of changing the degree of ion-exchange of polyvalent cation in an 0.5 weight percent platinum-loaded zeolite Y catalyst, a series of isomerization runs on n-hexane were conducted. The results of the runs are tabulated in Table X below:
In FIGURE 2 are plotted the optimum isomerization temperature and the volume percent yield of isohexanes obtained at that temperature versus the degree of calcium ion-exchange in the zeolite Y catalyst with 0.5 weight percent platinum metal loading. Examination of the curves TABLE X [Feed=n-l1exane] Catalyst Na (100%) Y Ca (45%) Y Run. N0 1 2 3 4 5 6 7 8 9 10 11 12 Reaction Temp., C 375 400 425 450 475 500 370 385 410 425 440 450 Pressure p.s.i.g 450 450 450 450 450 450 450 450 450 450 450 450 W.H.S.V. g./g./hr 2.0 2.0 2.0 2.0 2. 0 2.0 2. 0 2.0 2.0 2.0 2.0 2.0 H /HydrocarbonM:M 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 Liquid Yield Vol. Percent 98. 7 99.0 99.1 98. 7 98. 7 94. 7 100.0 100.0 100. 0 109.0 95. 5 96.9 Product Analysis:
n-Hexane. 96.8 97. 0 95. 7 93. 6 89. 0 71. 3 100.0 91. 0 76. 5 70. 6 57. 4 54. 0
3-metgylpeng 7. l 8.8 11. 8 14. 8 15. 9
2 met y pen anezadimethylbuta 0.1 1.0 2. 4 5. 2 9. 5 5. 6 11.2 14.0 21. 3 23. 5
2,2-dimethylbutane- 0. 2 0. 0 0. 5 0. 8 1. 8 2. 3 Yield iso-C Vol. Percent 0. 0 U. 1 1. 0 2. 4 5. 1 15. 7 0. 0 5. 6 20. 0 25. 8 36. 2 40. 4 Propane M01. Percent." N0 propane analysis (or t iis series 0. 2 0.3 1. 2 1. 1
Catalyst Ca (45%) Y Ca (65%) Y Ca (85%) Y Run No 13 I 14 l 15 i 16 17 18 i 19 l 20 i 21 22 23 24 Reaction Temp., C 460 470 480 500 350 400 410 425 435 450 350 360 370 Pressure p.s.i.g.... 450 450 450 450 450 450 450 450 450 450 450 450 450 W,H S.V g./g./hr 2.0 2.0 2.0 2.0 2.0 2.0 2. 0 2.0 2. 9 2.0 2.0 2.0 2.0 Ii /Hydrocarbon M:M 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 5:1 Liquid Yield Vol. percent 96. 2 96. 2 97. 4 88. 8 100. 0 99. 3 99. 4 98. 7 99. 4 97. 4 97.4 97. 4 Product Analysis:
n-Hexane 47. 0 41. 4 40. 5 27. 9 67. 0 35. 7 34. 0 32. 3 32. 2 25. 9 27. 0 23. 2 20.3
3-methylpentane 17. 4 18. 4 l8. 1 17. 4 11.6 19. 1 19. 7 20. 2 20.6 18. 3 19. 4 18. 9 17. 4
2methylpentane- 2Y3 dlmethy1butane 26.9 29. 1 28. 2 29. 7 17.3 32. 5 34. 1 34. 0 33. 1 33. 6 36. 2 26. 4 34. 3
2,2-dimethylbutane 3. 3 4. 2 4. 3 6.0 1. 2 6. 7 7. 0 7. 4 7. 7 8. 5 10. 7 11. 7 11. 4 Yield ISO-C5 Vol. percent of feed" 45. 8 50. 0 49. 3 47. 2 30. l 57. 9 60. 4 60. 8 61. 4 60. 4 64. 6 65. 3 59. 8 Propane M01. percent 2. 1 2. 9 3. 9 7. 3 0. 7 2.0 2.0 2. 0 2. 6 7. 7 2. 2 3.0 5. 3
The results of Table X are plotted in FIGURE 1 wherein the effect of increasing the degree of polyvalent-ion exchange is indicated. As can be seen from this figure, as the degree of ion-exchange increases, i.e., from to 100 percent, the degree of isomerization similarly increases.
In FIGURE 1, runs 1 through 6, representing a nonion exchanged catalyst, are represented as curve A. Runs 7 through 16 representing a 45 percent calcium exchanged catalyst are represented as curve B. Runs 17 through 22 representing a percent calcium exchange catalyst of this invention are represented as curve C. Runs 23 through 25, representing an percent calcium exchanged zeolite Y catalyst of this invention are represented as curve D.
REFORMING The reforming activity of the metal loaded catalyst of the present invention is shown by the following performance data. In the reforming reaction shown, a light gasoline feed was used. The reforming tests were carried out at a temperature of between 450500 C. and 450 p.s.i.g. pressure. The hydrogen to hydrocarbon molar 7 ratio in the feed was 5.
Distribution of Pt Wt. Octane No. at Catalyst Type Cations in perpercent 85 Vol. percent Content cent, liquid yield According to the performance data shown, the alkaline earth-metal cationic zeolite catalysts produced a higher octane number gasoline product than the other diand higher valent cationic zeolite catalysts and are preferred.
The following table shows a comparison of the reforming activity of various platinum loaded aluminosilicate.
TABLE XI Comparison of catalytic reforming activity 09 polyvalent cation exchanged aluminosilicates loaded with 0.5 wt. percent Pt SiO :Alz0 =5. The same process conditions were used with each catalyst:
Equilibrium Temp, C 500 Pressure, p.s.i.g. W.H.S.V., g./g./hr H :H.C. Molar Ratio Feed Stock, Light naphtha, 350 0. end piont. Composition:
Paraifins95 Vol. percent; Olefins-IO Vol. percent; Aromatics5Vol. percent.
Table XI shows the improved results of the combination of metal and zeolite Y in a reforming process. The comparison of the results of Ca+ X and Ca+ Y shows the importance of the SiO /Al O ratio. The comparison of the amorphous aluminosilicate and Ca+ Y shows the importance of crystallinity to the present invention.
The reforming process employing the polyvalent metal cation exchanged zeolite Y catalyst containing within the pore system from 0.01 to 5.0 wt. percent of active noble metal from the platinum group of Group VIII is preferably carried out within the following ranges of process conditions. The broad temperature range is 300 to 600 C. and more particularly 400 to 525 C. The pressure should be in the range of 100 to 1200 p.s.i.g. more particularly 300-600 p.s.i.g., and the hydrogen to hydrocarbon feed mole ratio should be from 1:1 to 2:1 and more particularly 2:1 to 5:1. The contact time expressed as weight hourly space velocity WHSV, should be 0.1 to about 7 and more particularly from 0.5 to 3.
DEALKYLATION The preferred catalyst for dealkylation activity in the present invention is the metal loaded catalyst in a hydrogen atmosphere. The usual feeds to a catalytic dealkylation unit are alkyl-substituted aromatics such as toluene. The present hydro-dealkylation process is preferably carried out at a temperature of 400600 C., particularly 450-550, at a pressure of -1000 p.s.i.g. particularly 50-500 p.s.i.g. at a weight-hour space velocity of .5-5 particularly .5-2 and at a hydrogen to hydrocarbon molar ratio of 320, particularly 5-10.
The following table shows a comparison of the hydrodealkylation activity of different platinum loaded aluminosilicates.
TABLE XII Catalytic hydrodcalkylation (dealkylation) activity of polyvalent cation exchanged aluminosilicates-comparison of difierent aluminosilicates and different polyvalent cations Amorphous Zeolite Zeolite Zeolite Zeolite Catalyst (Aluminosilicate) 1 Alumino- T 1 Y Y Y silicate Cation Ca Ca+ Ca+ Mn" 2 AW Degree of Exch 40 8 85 75 Composition of Liq. Prod, Mol
Benzene 3. 6 4. 2 22. 2 58. 2 33. 3 Toluene. 95. 0 93. 2 63. 7 34. 3 51. 3 Xylenes 0. 0 0. 0 11.3 4.1 10.2 Cracked Products 1. 3 2.7 2.8 3. 4 5. 1 Stability of Catalyst Good Fair Good Fair Poor 1 Each catalyst contained 0.5 wt. percent Pt-loaded by ion-exchanged Pt (NHa)4 cation and decompose by heat in an oxygen containing atmosphere prior to conditioning in H before reaction.
1 Described in U.S. Patent No. 2,950,952 issued to D. W. Brock et al., August 30, 1960.
Process Conditions: Each catalyst was evaluated under the same process conditions "h .21 The Table XII comparison shows the importance of crystallinity, pore size, and cation to the hydrodealkyla- 32 product solution yielded the following fractions after removal of the solvent:
The dimer fraction consisted of approximately 90-91 wt.- 2.4,4-trimcthylpentenc-1 a-olefin) and 9-10 wt. percent; of the less desirable 2,4,4-trimethylpentene-2 (ti-olefin) by comparing the observed refractive index to the retraetivc index of standard mixtures of aand fl-olefin. The usual ratio produced by Bronsted and Lewis acid catalysts is 75% a-olefin: 25% B-olefin.
tion catalyst. The amorphous aluminosilicate and zeolite T are respectively an amorphous zeolite having an SiO2/Al203 5 HYDRODEALKYLATION OF TOLUENE WITH 1.0 WT. PERCENT CuMg+ Y Run N 1 2 Feed Temperature, 0., Inlet Average Temperature for Run Pressure, p.s.i.g
H,. H.C W.H.S.V Hours on Stream at End of Rum... Duration of Run to Nearest Hour..- Liquid Yield Vol. Percent Liquid Yield Wt. Percent Conversion Molar Selectivity to:
Bemene Xylene Non-aromatics Toluene 550 POLYMERIZATION The preferred catalyst for polymerization for the present invention is the non-metal loaded catalyst Low molecular weight gaseous and liquid olefins including C -C olefins are polymerized to low molecular weight products boiling in the gasoline range and useful as high octane number gasoline and as petrochemical intermediates. The present polymerization process is preferably carried out at a temperature of 0-300 C., particularly 20-200 C., at a pressure of atmospheric-10,000 p.s.i.g. particularly atmospheric-1500 p.s.i.g. and at a weighthour space velocity of 0.01-10, particularly .5-2.
The following example describes the polymerization of isobutylene to low molecular weight products distilling in the gasoline boiling range and also useful as valuable mono-olefinic petrochemicals, with a polyvalent cation exchanged zeolite Y catalyst in the presence of an inert solvent to dissolve the polymer product as it is formed.
Thirty grams (30) of magnesium exchanged (73%) zeolite Y previously activated (dehydrated) by heating to 500 C. in an oven purged with air, was slurried in 150 ml. (98 g.) of n-hexane. Isobutylene, at atmospheric pressure, was bubbled into this slurry at a rate of approximately 1 l. per minute for 90 minutes. Conversion of the isobutylene was complete and was as rapid at the end of the run as at the start. Some cooling of the slurry was necessary to keep the temperature between 20-35 C. The gain in weight was 247.4 g. (4.43 moles of isobutylene). The catalyst was removed from the slurry by filtration and all but 2 grams of adsorbed solvent and product were removed by heating the catalyst to 235 C. under a vacuum of 1.0 mm. Hg. (Note: The recovered catalyst was again used in a similar run. The activity was the same as the original.) DlSIillation of the solvent What is claimed is:
1. A process for the conversion of hydrocarbons, which comprises contacting said hydrocarbons with a zeolitic molecular sieve having at least 40 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure capable of internally adsorbing benzene and a silicon dioxide to alumium trioxide molar ratio greater than 3, under hydrocarbon converting conditions.
2. A process for the conversion of hydrocarbons, which comprises contacting said hydrocarbons with a zeolitic molecular sieve having at least 40 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure capable of internally adsorbing benzene, a silicon dioxide to aluminum trioxide molar ratio greater than 3 and up to about 6, under hydrocarbon converting conditions.
3. A process for the conversion of hydrocarbons, which comprises contacting said hydrocarbons with a zeolitic molecular sieve having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure capable of internally adsorbing benzene, a silicon dioxide to aluminum trioxide molar ratio between about 3.5-6, under hydrocarbon converting conditions.
4. A process as described in claim 1 wherein said zeolitic molecular sieve is zeolite Y.
5. A process as described in claim 3 wherein said zeolitic molecular sieve is zeolite Y.
6. A process as described in claim 1 wherein said zeolitic molecular sieve is selected from the group consisting of zeolite Y, zeolite L and faujasite.
7. A process as described in claim 4 wherein said bydrocarbon conversion process is cracking.
8. A process as described in claim 4 wherein said hydrocarbon conversion process is alkylation.
9. A process as described in claim 4 wherein said hydrocarbon conversion process is polymerization.
10. A process for the cracking of hydrocarbons which comprises contacting said hydrocarbons with zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure, a silicon dioxide to aluminum trioxide molar ratio between about 3.5-6, said polyvalent cations selected from the group consisting of magnesium and calcium, under cracking conditions.
11. A process for the alkylation of hydrocarbons which comprises contacting said hydrocarbons with zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure, a silicon dioxide to aluminum trioxide molar ratio between about 3.5-6, said polyvalent cations selected from the group consisting of magnesium and calcium, under alkylating conditions.
12. A process for the polymerization of hydrocarbons which comprises contacting said hydrocarbons with zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure, a silicon dioxide to aluminum trioxide molar ratio between about 3.5-6, said polyvalent cations selected from the group consisting of magnesium, calcium and chromium, under polymerizing conditions.
13. A process for the conversion of hydrocarbons which comprises contacting said hydrocarbons with a zeolitic molecular sieve having at least 40 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure capable of internally adsorbing benzene, a silicon dioxide to aluminum trioxide molar ratio greater than 3, and containing a catalytically active elemental metal in an amount of at least about 0.05 weight percent; under hydrocarbon converting conditions.
14. A process for the conversion of hydrocarbons, which comprises contacting said hydrocarbons with a zeolitic molecular sieve having at least 40 percent of the alu-minum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure capable of internally adsorbing benzene, a silicon dioxide to aluminum trioxide molar ratio greater than 3, and containing an elemental metal of Group VIII of the Periodic Table in an amount of at least about 0.05 weight percent; under hydrocarbon converting conditions.
15. A process for the conversion of hydrocarbons which comprises contacting said hydrocarbons with zeolite Y having at least 40 percent of the aluminum tetrahedra satisfied by the presence of polyvalent cations and containing a catalytically active elemental metal in an amount of at least about 0.05 weight percent; under hydrocarbon converting conditions.
16. A process for the conversion of hydrocarbons which comprises contacting said hydrocarbons with zeolite Y having at least 40 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, and containing an elemental metal of Group VIII of the Periodic Table in an amount of at least from about 0.05 weight percent, under hydrocarbon conditions.
17. A process for the conversion of hydrocarbons which comprises contacting said hydrocarbons in a stream of hydrogen gas with zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a silicon dioxide to aluminum trioxide molar ratio between about 3.5-6, and containing an elemental metal of Group VIII of the Periodic Table in an amount of from about 0.05-2.0 weight percent; under hydrocarbon converting conditions.
18. A process as described in claim 17 wherein said hydrocarbon conversion process is isomerization.
19. A process for the hydrocracking of hydrocarbons which comprises contacting said hydrocarbons in a stream of hydrogen gas with zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a silicon dioxide to aluminum trioxide molar ratio between about 3.5-6, and containing an elemental metal of Group VIII of the Periodic Table in an amount of from about 0.05-2.0 weight percent; said polyvalent cations selected from the group consisting of magnesium and calcium and said metal selected from the group consisting of platinum and palladium; under hydrocracking conditions.
20. A process for the hydrodealkylation of hydrocarbons which comprises contacting said hydrocarbons in a stream of hydrogen gas with zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the pres ence of polyvalent metal cations, a silicon dioxide to aluminum trioxide molar ratio between about 3.5-6, and containing an elemental metal of Group VIII of the Periodic Table in an amount of from about 0.05-2.0 weight percent, said polyvalent cations selected from the group consisting of magnesium and calcium; under hydrodealkylation conditions.
21. A process for the reforming of hydrocarbons which comprises contacting said hydrocarbons in a stream of hydrogen gas with zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a silicon dioxide to aluminum trioxide molar ratio between about 3.5-6, and containing an elemental metal of Group VIII of the Periodic Table in an amount of from about 0.05-2.0 weight percent, said 34' polyvalent cations selected from the group consisting of magnesium and calcium and said metal selected from the group consisting of palladium and platinum, under reforming conditons.
22. A process for the alkylation of hydrocarbons which comprises contacting said hydrocarbons in a stream of hydrogen gas with zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a silicon dioxide to aluminum trioxide molar ratio between about 3.5-6, and containing an elemental metal of Group VIII of the Periodic Table in an amount of from about 0.05-2.0 weight percent, said polyvalent cations selected from the group consisting of magnesium and calcium and said metal selected from the group consisting of palladium and platinum; under alkylating conditions.
23. A process as claimed in claim 16 for the isomerization of parafiinic hydrocarbons containing from five to six carbon atoms.
24. A hydrocarbon conversion catalyst, which comprises a zeolitic molecular sieve having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure capable of internally adsorbing benzene and a silicon dioxide to aluminum trioxide molar ratio greater than 3.
25. A hydrocarbon conversion catalyst, which comprises zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations.
26. A catalyst as claimed in claim 25, in which the polyvalent metal cations are selected from the group consisting of aluminum, beryllium, calcium, cerium, chromium, magnesium, manganese, strontium, and zinc.
27. A hydrocarbon conversion catalyst which comprises a zeolitic molecular sieve having at least 40 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure capable of internally adsorbing benzene, a silicon dioxide to aluminum trioxide molar ratio greater than 3, and containing a catalytically active elemental metal in an amount of at least about 0.05 weight percent.
28. A hydrocarbon conversion catalyst which comprises a zeolitic molecular sieve having at least 40* percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a crystalline structure capable of internally adsorbing benzene, a silicon dioxide to alumninum trioxide molar ratio greater than 3, and containing an elemental metal of Group VIII of the Periodic Table in an amount of at least about 0.05 weight percent and in the inner adsorption region of the zeolitic molecular sieve.
29. A hydrocarbon conversion catalyst which comprises zeolite Y in the partly decationzed state but having at least 40 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations and a silicon dioxide to aluminum trioxide molar ratio betwen about 3.5-6.
30. A hydrocarbon conversion catalyst which comprises zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, and a silicon dioxide to aluminum trioxide molar ratio between about 3.5-6.
31. A hydrocarbon conversion catalyst which comprises zeolite Y having at least 40 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a silicon dioxide to aluminum trioxide molar ratio greater than 3, and containing an elemental metal of Group VIII of the Periodic Table in an amount of at least about 0.05 weight percent and in the inner adsorption region of the zeolite Y.
32. A hydrocarbon conversion catalyst which comprises zeolite Y having at least 65 percent of the aluminum tetrahedra satisfied by the presence of polyvalent metal cations, a silicon dioxide to aluminum trioxide mo- 35 Iar ratio between about 3.56, and containing an elemental noble metal of Group VIII of the Periodic Table in an amount of from 0.2 to 0.6 weight percent and in the inner adsorption region of the zeolite Y.
33. A catalyst as claimed in claim 32 in which the Group VIII metal is platinum.
34. A catalyst as claimed in claim 32, in which the Group VIII metal is palladium.
35. A catalyst as claimed in claim 32, in which the polyvalent metal cations are selected from the group consisting of aluminum, beryllium, calcium, cerium, chromium, magnesium, manganese, strontium, and zinc.
36. A process as described in claim 13 wherein the metal is contained in the inner adsorption region of the zeolitic molecular sieve.
37. A process as described in claim 14 wherein the metal is contained in the inner adsorption region of the zeolitic molecular sieve.
38. A process as claimed in claim 1, in which the zeolitic molecular sieve is in the partly decationized state.
39. A process as claimed in claim 16, in which the zeolitic molecular sieve is in the partly decationized state.
40. A catalyst as defined in claim 31 in which the zeolite Y is in the partly decationized state.
References Cited by the Examiner UNITED STATES PATENTS DELBERT E. GANTZ, Primary Examiner.
MILTON STERMAN, ALPHONSO D. SULLIVAN,
Examiners. A. RIMENS, Assistant Examiner.
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|U.S. Classification||208/111.25, 585/533, 585/486, 208/138, 585/467, 502/60, 585/722, 585/433, 585/739, 585/489, 585/666, 585/475, 585/646, 585/434, 585/650, 585/468, 585/752, 585/654, 585/419, 585/653, 208/2, 585/481, 208/111.35, 585/751, 585/482, 502/74, 502/64|
|International Classification||C07C2/12, C07C5/27, C10G35/095, C07C2/66, C07C5/22, C10G47/16, B01J20/18, C07C4/08, C07C2/54, C10G11/05, B01J29/06, C07C4/06, C07C4/18, B01J29/08|
|Cooperative Classification||C07C2/12, C07C2/54, C07C5/222, B01J29/06, C10G11/05, C07C5/226, C07C2/66, C10G47/16, B01J20/186, C10G35/095, C07C5/2724, C07C4/18, C07C4/06, C07C4/08, C07C2529/60, C07C2529/08, B01J29/084, C07C5/2791|
|European Classification||C07C5/27A8, C07C5/22B4, C07C2/12, C07C4/18, C07C5/22B8, C10G11/05, B01J29/06, C07C5/27D2J, C07C2/54, B01J20/18D, B01J29/08Y, C10G47/16, C10G35/095, C07C4/08, C07C4/06, C07C2/66|