DK163571B - CATALYST CONTAINING A TYPE OF L ZEOLITE AND AT LEAST ONE GROUP VIII METAL AND USE THEREOF - Google Patents

Description

DK 163571 B

The invention relates to a catalyst comprising a type L zeolite and at least one group VIII metal, as well as the use of such a catalyst.

Catalytic reforming is well known in the oil industry and indicates a treatment of the naphtha fractions - to improve the octane number in the production of aromatics. The more important hydrocarbon reactions that occur during the reforming process include dehydrogenation of cyclohexanes to aromatics, dehydroisomerization. of alkylcyclopentanes for aromatics and dehydrocyclization of acyclic hydrocarbons to aromatics.

There are also a number of other reactions including the following: dealkylation of alkyl benzenes, isomerization of paraffins, and hydrocracking reactions giving 15 gaseous hydrocarbons, e.g. methane, ethane, propane and butane. Hydrocracking reactions should be minimized especially during reform, as they reduce the yield of boiling products and hydrogen in the gasoline area.

Due to the demand for high-octane gasoline for use as motor fuel, etc., extensive research is being conducted to develop improved reforming catalysts and catalytic reforming processes. Catalysts for successful reform processes must have good selectivity, ie. be capable of producing large yields of liquid products in the benzine zinc boiling range, and containing high concentrations of high octane aromatic hydrocarbons and consequently small yields of light gaseous hydrocarbons. The catalysts must have good activity so that the temperature required to produce a certain quality product need not be too high. It is also necessary that the catalysts either have good stability to maintain the activity and selectivity properties for longer operating times, or are sufficiently regenerable to allow frequent regeneration without loss of performance.

Catalysts comprising platinum, e.g. platinum supported on alumina is well known and 2

DK 163571 B

widely used for reforming naphthas. The main products of catalytic reforming are benzene and alkylbenzenes. These aromatic hydrocarbons have great value as high octane number components in gasoline.

Catalytic reforming is also an important process for the chemical industry due to the large and growing demand for aromatic hydrocarbons for use in the manufacture of various chemical products such as synthetic fibers, insecticides, adhesives, 10 detergents, plastics, synthetic gums, pharmaceuticals, high octane gasoline, perfumes, drying oils, ion exchange resins and various other products well known to those skilled in the art. An example of this demand lies in the preparation of alkylated aromatics such as ethylbenzene, cumene and dodecylbenzene using the appropriate monoolefins to alkylate benzene. Another example of this demand lies in the area of chlorination of benzene to produce chlorobenzene, which is then used to produce phenol by hydrolysis with sodium hydroxide. The main use of phenol lies in the production of phenol-formaldehyde resins and plastics. Another phenol preparation pathway uses cumene as a starting material and involves oxidation of cumene with air 25 to cumene hydroperoxide which can then be degraded to phenol and acetone by action with appropriate acid. Demand for ethylbenzene is mainly due to its use in the production of styrene by selective dehydration and deification. The styrene, in turn, is used to make styrene-butadiene rubber and polystyrene. Orthoxylene is typically oxidized to phthalic anhydride by reaction with air in the vapor phase in the presence of a vanadium pentoxide catalyst. Phthalic anhydride, in turn, is used to prepare plasticizers, polyesters and resin types. Demand for paraxylene 3

DK 163571 B

is essentially due to its use in the preparation of terephthalic acid or dimethyl terephthalate, which in turn is reacted with ethylene glycol and polymerized to obtain polyester fibers. Substantial demand for benzene is also due to its use in the production of aniline, nylon, maleic anhydride, solvents and similar petrochemicals. Toluene, on the other hand, at least compared to benzene and Cg aromatics is not much in demand in the 10 petrochemical industry as a basic building block chemical. Accordingly, substantial amounts of toluene are hydrodalkylated to benzene or disproportionate to benzene and xylene. Another use of toluene is associated with the transalkylation of trimethylbenzene with toluene to give xylene.

In response to this demand for these aromatic products, the technique has evolved and the industry has used a variety of alternative methods to produce them in commercial quantities. One attempt has been the preparation of a considerable number of catalytic reformers for the production of aromatic hydrocarbons for use as starting materials in the preparation of chemicals. As is the case with most catalytic processes, the main goal of the efficiency of catalytic reforming involves the ability of the process to convert the starting materials into the desired products over extended periods of time with minimal impact from side reactions.

The dehydrogenation of cyclohexane and alkylcyclohexanes to benzene and alkylbenzenes is the thermodynamically most favorable type of aromatization reaction in catalytic reforming. This means that dehydrogenation of cyclohexanes can give a greater ratio of aromatic product to non-aromatic reactant than any of the other two types of aromatization reactions at a given reaction temperature and pressure. Furthermore, dehydrogenation of cyclohexanes is the fastest of the three aromatization reactions. As a result of these thermo- 4

DK 163571 B

For dynamic and kinetic considerations, the selectivity for dehydrogenation of cyclohexanes is higher than for dehydroisomer in series or dehydrocycl in series. Dehydroisomerization of alkylcyclopentanes is somewhat less favored both thermodynamically and kinetically. Its selectivity, although generally high, is lower than the selectivity for dehydrogenation. Dehydrocyclization of paraffins is much less favored both thermodynamically and kinetically. In conventional reforming 10, its selectivity is much lower than the selectivity of the other two aromatization reactions.

The disadvantage of selectivity in paraffin hydrocyclization is particularly great for the aromatization of compounds having a small number of carbon atoms per minute. molecule. The dehydrocyclization selectivity of conventional reforming is very low for C 6 hydrocarbons. It increases with the number of carbon atoms per but remains substantially lower than the aromatization selectivity for dehydrogenation or dehydroisomerization of 20 naphthenes with the same number of carbon atoms per molecule. Above all, a significant improvement in the catalytic reforming process will require a drastic improvement in dehydrocyclization selectivity while maintaining sufficient catalyst activity and stability.

In the dehydrocyclization reaction, both acyclic hydrocarbons are cyclized and dehydrogenated to produce aromatics. The conventional methods for carrying out these dehydrocyclization reactions are based on the use of catalysts comprising a noble metal on a support. Known catalysts of this type are based on alumina which supports 0.2 to 0.8% by weight of platinum and preferably another auxiliary metal.

A disadvantage of conventional naphtha resin catalysts is that they are generally more selective for Cg-Cg paraffins than other reactions (such as hydrocracking) than they are for dehydro cyclization. A significant advantage of the catalyst of the invention 5

DK 163571 B

is its high selectivity for dehydrocyclization.

The possibility of using carriers other than aluminum oxide has also been investigated and it has been proposed to use certain molecular sieves such as X and Y zeolites which have pores large enough for hydrocarbons in the gasoline boiling range can pass. However, catalysts based on these molecular sieves have not been commercially successful.

In the conventional process for carrying out the aforementioned dehydrocyclization, acyclic hydrocarbons to be converted are passed over the catalyst in the presence of hydrogen at temperatures of about 500 ° C and pressures from 0.5 to 3 MPa.

Part of the hydrocarbons are converted to aromatic hydrocarbons and the reaction is accompanied by isomerization and cracking reactions which also convert the paraffins to isoparaffins and lighter hydrocarbons.

The rate of conversion of the acyclic hydrocarbons to aromatic hydrocarbons varies by 20 the number of carbon atoms per reactant molecule, the reaction, the conditions and the nature of the catalyst.

The catalysts used so far have yielded evenly satisfactory results with heavy paraffins, but less satisfactory results with Cg-Cg paraffins, especially Cg paraffins. Catalysts based on a type L zeolite are more selective in the dehydro cyclization reaction and can be used to improve the rate of conversion to aromatic hydrocarbons without requiring higher temperatures than those required by thermodynamic considerations (higher temperatures generally have a significant negative effect on the stability of the catalyst), and they give excellent results with Cg-Cg paraffins, but type L zeolite catalysts have not achieved commercial use, apparently due to insufficient stability.

In another known method for dehydrocyclization of aliphatic hydrocarbons, hydrocarbons are introduced into 6

DK 163571 B

the presence of hydrogen in contact with a catalyst consisting essentially of a type L zeolite with exchangeable cations, at least 90% of which are alkali metal ions selected from the group consisting of lithium, sodium, potassium, rubidium and cesium ions; containing at least one metal selected from the group consisting of Group VIII metals in the periodic table, tin and germanium, said metal or metals including at least one Group VIII metal of the Periodic Table having a dehydrogenating effect, so that at least a portion of the starting material is converted to aromatic hydrocarbons.

In a particularly advantageous embodiment of this process, a platinum / alkali metal / type L-15 zeolite catalyst containing cesium or rubidium is used because of its excellent activity and selectivity to convert hexanes and heptanes to aromatics, but stability remains a problem.

The deficiencies of the prior art are overcome by the catalyst of the invention comprising a type L zeolite and at least one group VIII metal, the catalyst being characterized in that it further comprises an alkaline earth metal from the group barium, strontium and calcium.

This catalyst provides a superior selectivity for the conversion of acyclic hydrocarbons to aromatics compared to the selectivity obtained by known methods. The catalyst also provides a satisfactory service life.

Various suitable embodiments of the catalyst according to the invention can be used as set forth in claims 2-7.

The catalyst preferably contains (a) a type of L-zeolite containing from 0.1 to 5% by weight of platinum, preferably from 0.1 to 1.5% by weight of platinum, and 0.1 to 40% by weight of barium, preferably from 0.1 to 35 wt% barium, and more preferably 1 to 20 wt% barium, and (b) an inorganic binder. Over 50% of type L zeolite crystal 7

DK 163571 B

The clays are preferably greater than 500 Å, more preferably greater than 1000 Å. In the most preferred embodiment, at least 80% of the type L zeolite crystals are greater than 1000 Å. The inorganic binder is preferably either a silica, alumina, aluminosilicate or a clay. The hydrocarbons are contacted with the barium-exchanged type of zeolite at a temperature of from 400 to 600 ° C (preferably 430 to 550 ° C) an LHSV of from 0.1 to 10 (preferably from 0.3 to 5), a manometer-10 pressures of from 0 to 3.45 MPa (preferably from 0.35 to 2.07 MPa), and a t 2 / HC ratio of from 1: 1 to 10: 1 (preferably from 2: 1 to 6: 1) .

In general, the present invention relates to a catalyst comprising a type L zeolite, an alkaline earth metal as specified and a Group VIII metal, and the invention further relates to the use of the catalyst for reforming hydrocarbons, in particular dehydrocyclization of acyclic hydrocarbons having high selectivity, and further, dehydroisomerization of alkylcyclopentanes, and dealkylation of toluene to produce benzene.

The term "selectivity" as used in the present invention is defined as the percentage of moles of acyclic hydrocarbons converted to aromatics to the number of moles converted to 25 aromatics and cracked products, i.e.

100 x moles of acyclic hydrocarbons selectivity = converted to aromatics_ moles of acyclic hydrocarbons converted to aromatics and cracked products

Isomerization of paraffins and the mutual conversion of paraffins and alkylcyclopentanes with the same number of carbon atoms per molecule is not considered when determining the selectivity.

The term "n-hexane selectivity" as used in the present invention is defined as the percentage of moles of n-hexane converted to aromatics relative to the number of moles converted to aromatics and cracked products.

The selectivity for converting acyclic hydrocarbons to aromatics is a measure of process 8

DK 163571 B

efficiency of converting acyclic hydrocarbons into the desired and valuable products, i.e. aromatics and hydrogen, unlike the less desirable products of hydrocracking.

5 Extremely selective catalysts produce more hydrogen than less selective catalysts as hydrogen is formed when acyclic hydrocarbons are converted to aromatics and hydrogen is consumed when acyclic hydrocarbons are converted into cracked products.

An increase in the selectivity of the process increases the amount of hydrogen formed (more aromatization) and decreases the hydrogen consumed (less cracking).

Another advantage of using highly selective catalysts is that the hydrogen produced with highly selective catalysts is cleaner than that produced with less selective catalysts.

This greater purity results from more hydrogen being produced, while less boiling hydrocarbons (cracked products) are produced. The purity of the hydrogen produced by reforming is critical if the hydrogen, as is generally the case in an integrated refinery, is used in processes such as hydrotreating and hydrocracking which require at least some minimum partial pressure of the hydrogen. If the purity becomes too low, the hydrogen can no longer be used for this purpose and must be used in a less valuable way, e.g. such as fuel gas.

Regarding the acyclic hydrocarbons which can be dehydrocyclized by the use of the invention, they are mostly paraffins, but can generally be any acyclic hydrocarbon capable of cyclization to form. of an aromatic hydrocarbon. That is, the present invention also aims to dehydrocyclize any acyclic hydrocarbon capable of undergoing cyclization to produce an aromatic hydrocarbon capable of being evaporated at the dehydrocyclization temperatures used herein.

9

DK 163571 B

More particularly, suitable acyclic hydrocarbons comprise acyclic hydrocarbons containing 6 or more carbon atoms per liter. molecules such as Cg-C2Q paraffins and ε6-Ο20 · “θ1θ £ iner. Specific examples of their own acyclic hydrocarbons are: (1) paraffins such as n-hexane, 2-methylpentane, 3-methylpentane, n-heptane, 2-methylhexane, 3-methylhexane, 3-ethylpentane, 2,5-dimethane thylhexane, n-octane, 2-methylheptane, 3-methylheptane, 4-methylheptane, 3-ethylhexane, n-nonane, 2-methyloctane, 3-methyloctane, n-decane and similar compounds as well as (2) olefins such as 1-hexene , 2-methyl-1-pentene, 1-heptene, 1-octene, 1-nonene and similar compounds.

In a preferred embodiment, the acyclic hydrocarbon is a paraffinic hydrocarbon having 15 to 10 carbon atoms per minute. molecule. It should be noted that the above-mentioned acyclic hydrocarbons may be applied to the present application individually, in admixture with one or more of the other acyclic hydrocarbons, or in admixture with other hydrocarbons such as naphthenes, aromatics and the like. Thus, mixed hydrocarbon fractions containing significant amounts of acyclic hydrocarbons and commonly available in a typical refinery are suitable additives. sealing materials for the catalyst in question, e.g. highly paraffinic "straight-run" naphthas, paraffinic refiners from aromatic extraction or adsorption, Cg-Cg paraffin-rich streams and similar refinery streams. A particularly preferred embodiment relates to the use of a starting material which is a paraffin-rich naphtha fraction boiling in the range of 60 to 177 ° C. Generally, the best results are obtained with a starting material consisting of a mixture of Cg-C20 paraffins, especially Cg-Cg paraffins.

The starting material is preferably substantially free of sulfur, nitrogen, metals and other known toxins for reforming catalysts. This catalyst is particularly sensitive to sulfur. The starting material can be made substantially free of sulfur, nitrogen, metals

DK 163571B

10 and other known poisons by conventional hydro-purification methods and using sorbents which remove sulfur compounds.

In the case of a starting material that does not already have low sulfur content, acceptable levels can be obtained by hydrating the starting material in a pretreatment zone where naphtha is contacted with a hydrogen purification catalyst resistant to sulfur poisoning. A suitable catalyst for this hydrodesulfurization process is e.g. a catalyst with an alumina-containing support and a minor proportion of molybdenum oxide, cobalt oxide and / or nickel oxide. Hydro desulphurization is generally carried out at 315 to 455 ° C at a pressure gauge of from 1.38 to 13.8 MPa and an LHSV of 1 to 5.

The sulfur and nitrogen contained in the naphtha are converted into hydrogen sulfide and ammonia, respectively, which can be removed prior to reforming by appropriate conventional processes.

According to one aspect of the present invention, the acyclic hydrocarbon is contacted with the catalyst in a dehydrocyclization zone. which are kept under dehydrocyclization conditions. This contacting can be accomplished by using the catalyst in a solid-layer, moving-layer system, a fluidized system, or in batch operation. However, given the danger of loss of the valuable catalyst and well-known operating advantages, it is preferred to use either a a solid-layer system or a high-density moving layer system. It is also contemplated that the contacting step may be carried out in the presence of a physical mixture which comprises particles of a conventional dual-function catalyst of known type. In a solid-layer system, the acyclic hydrocarbon-containing starting material is preheated by any suitable heating method to the desired reaction temperature and then introduced into the dehydrocyclization zone containing a solid layer of the catalyst. It should be noted that the dehydrocyclization zone may, of course, be one or

DK 163571 B

1 1 several separate reactors with suitable means therebetween, ensuring that the desired conversion temperature is maintained at the entrance to each reactor. It is also important to note that the reactants can be contacted with the catalyst layer in either upward, downward or radial flow modes. In addition, the reactants may be in liquid phase, mixed liquid phase or vapor phase when contacted with the catalyst, with the best results being obtained in the vapor phase. The dehydrocyclization system then preferably comprises a dehydrocyclization zone containing one or more solid or high density moving layers of the catalyst.

Of course, in a multi-layer system, it is within the scope of the present invention to use the present catalyst in fewer than all the layers, with a conventional dual-function catalyst being used in the rest of the layers. The dehydrocyclization zone may be one or more separate reactors with suitable heaters therebetween to compensate for the heat loss occurring in each catalyst layer due to the endothermic nature of the dehydrocyclization reaction.

Although hydrogen is the preferred diluent for use in the dehydrocyclization method described, in some cases advantageous other diluents known in the art may be used either alone or in admixture with hydrogen such as C C-C₂ paraffins such as methane, ethane. , propane, butane and pentane and similar diluents and mixtures thereof. Hydrogen is preferred because it serves the dual function of not only lowering the partial pressure of the acyclic hydrocarbon, but also suppressing the formation of hydrogen-poor carbonaceous deposits (usually referred to as coke) on the catalyst composite.

In general, hydrogen is used in amounts sufficient to ensure a hydrogen to hydrocarbon molar ratio of from 1: 1 to 10: 1, with the best results being obtained in the range of 2: 1 to 6: 1. The hydrogen supplied to the dehydrocyclization zone will typically be in

DK 163571B

12 is a hydrogen-rich gas stream which is recycled from the outlet stream from this zone after a suitable gas / liquid seep stage.

The hydrocarbon hydrocyclization conditions used in carrying out the present application include a reactor pressure selected in the range of 0 to 3.45 MPa overpressure, the preferred pressure being 0.35 to 2.07 MPa overpressure. The temperature of the dehydrocyclization is preferably from 430 to 550 ° C. As is well known to those skilled in the art, the first choice of temperature within this wide range is performed primarily as a function of the desired level of conversion of the acyclic hydrocarbon, taking into account the properties of the starting material and the catalyst. In general, the temperature is then slowly increased during the run to compensate for the inevitable deactivation that occurs so as to obtain a relatively constant value for the conversion.

The specific volumetric liquid delivery rate (LHSV) used in the present dehydrocyclization method is selected in the range of 0.1 to 10 hours, with a value in the range of 0.3 to 5 hours -1 being preferred.

Reformation generally results in the formation of hydrogen. Thus, exogenous hydrogen need not necessarily be added to the reforming system except for pre-reduction of the catalyst and once the feed stream is introduced. Once the reforming is underway, a portion of the hydrogen produced is generally recycled across the catalyst. The presence of hydrogen serves to reduce the formation of coke that tends to poison the catalyst. Hydrogen is preferably supplied to the reforming reactor in an amount ranging from 1 to 10 35 moles of hydrogen per liter. moles of feed material. The hydrogen may be in admixture with light gaseous hydrocarbons.

If, after a certain period of operation, the catalyst has been deactivated by the presence of carbon

DK 163571 B

containing deposits, these deposits can be removed from the catalyst by passing an oxygen-containing gas such as diluted air into contact with the catalyst at elevated temperature to burn away the carbonaceous deposits from the catalyst. The regeneration can either be carried out in the semi-regenerative manner in which the reforming process is interrupted after a shorter or longer period of time and the catalyst regeneration is carried out, or in the continuous regenerative manner in which part 10 of the catalyst is regenerated while the reforming process is continued over the rest of the catalyst. There are two types of continuous regeneration known in the art, viz. cyclical and continuous reform. In cyclic reforming, the catalyst is regenerated in one of a number of reactors, while the reforming is continued in the rest of the plant.

On continuous: reforming, part of the deactivated catalyst is removed from the plant and regenerated in a separate regeneration system while the reforming unit is continued in the plant and the regenerated catalyst is returned to the annulus. The manner in which the catalyst is regenerated will depend on whether it is operating with a solid layer, moving layer or fluidized layer. Regeneration methods and conditions are well known in the art.

The catalyst according to the invention is a type L-25 zeolite which is supplied with a Group VIII metal and an alkaline earth metal.

Type L zeolites are synthetic zeolites. A theoretical formula is Mg ^ n [(Alt ^) g (SiC ^^ yl / wherein M is a cation with the valence n.

The real formula may vary without changing the crystalline structure, e.g. For example, the molar ratio of silicon to aluminum (Si / Al) may range from 1.0 to 3.5.

Although there are a number of cations that may be present in the L zeolite, it is preferred in one embodiment to synthesize the potassium form of the zeolite, ie. the form in which the interchangeable cations present are essentially all potassium ions. The reactants, 14

DK 163571 B

as a result, are readily available and generally water soluble. The interchangeable cations present in the zeolite can then conveniently be replaced by other interchangeable cations, as will be seen below, thereby providing isomorphic forms of zeolite L.

In a process for preparing zeolite L, the potassium form of zeolite L is prepared by appropriately heating an aqueous metal-aluminum silicate mixture, the composition of which is expressed in the molar ratios of the oxides falling within the range: 0.33 to approx. 1 / SiC> 2 ..... from approx. 0.35 to approx. 0.5

SiC ^ / A ^ And ......... from approx. 10 to approx. 28 H2O / ...... from approx. 15 to approx. 41 15 The desired product is thereby crystallized relatively free of zeolites of different crystal structure.

The potassium form of zeolite L may also be prepared by another process together with other zeolite compounds by using a reaction mixture whose composition, expressed by the molar ratios of oxides, is within the following range: ^ 0 / (^ 0 + ^ 20) ... ..from approx. 0.26 to approx. 1 (K 2 O + Na 2 O) / SiO 2 .... from ca. 0.34 to approx. 0.5

SiC ^ / A ^ O ^ ......... from ca. 15 to approx. 28 25 H2O / (^ OHS ^ O) ..... from ca. 15 to approx. 51

It is noted that the presence of sodium in the reaction mixture is not critical to the present invention.

When the zeolite is prepared from sodium-containing reaction mixtures, sodium ions are generally included in the product as part of the replaceable cations along with the potassium ions. The product obtained from the above molar ratio ranges has a composition expressed as moles of oxides 35 corresponding to the formula: 0.9-1.3 [(1x) K 2 O, xNa 2 O]: Al 2 O 3: 5.2-6.9 SiC ^ ryl ^ O in which x may have any value from 0 to 0.75 and y may have any value from 0 to 9.

15

DK 163571 B

In the preparation of zeolite L, representative reactants are activated alumina, γ-alumina, alumina trihydrate and sodium aluminate as the source of alumina. Silica can be obtained from sodium or potassium silicate, silica gels, silicic acid, aqueous colloidal silica sols and reactive amorphous solid silica. The preparation of typical silica sols suitable for use in the present process is described in U.S. Pat.

10 2,574,890 and US Patent No. 2,597,872. Typical of the group of reactive amorphous solid silica, which preferably has a larger particle size of less than 1 µm, are such materials as silica fumes and chemically precipitated and precipitated silica sols. Potassium-15 and sodium hydroxide can supply the metal cation and help regulate the pH.

In the preparation of zeolite L, the usual process consists in the fact that potassium or sodium aluminate and alkali, e.g. potassium or sodium hydroxide, 20 dissolved in water. This solution is mixed with an aqueous solution of sodium silicate or preferably with a water-silicate mixture which is at least partially derived from an aqueous colloidal silica sol. The resulting reaction mixture is placed in a container which e.g.

25 is made of metal or glass. The container should be closed to prevent water loss. The reaction mixture is then stirred to ensure homogeneity.

The zeolite can be prepared satisfactorily at temperatures of from. 90 to 200 ° C, the pressure 30 being atmospheric pressure or at least a pressure corresponding to the vapor pressure of water in equilibrium with the mixture of reactants at the higher temperature. Any suitable heater may be used, e.g. an oven, a sand bath, an oil bath, or a sheathed autoclave. The heating is continued until the desired crystalline zeolite product is formed. The zeolite crystals are then filtered off and washed to separate them from the reactant mother liquor. The zeolite crystals should be washed with preferably distilled 16

DK 163571 B

water until the outgoing wash water, in equilibrium with the product, has a pH of between 9 and 12. While the zeolite crystals are washed, the interchangeable cation in the zeolite may be partially removed and it is assumed to be replaced by hydrogen cations. If washing is interrupted when the run-off wash water has a pH of between 10 and 11, the (K 2 O + Na 2 O) / Al 2 O 3 molar ratio of the crystalline product will be about 1.0. Then the zeolite crystals can be dried, conveniently in a ventilated oven.

10 Zeolite L has been described in "Zeolite Mole cular Sieves" by Donald W. Breck, John Wiley & Sons, 1974, as having a lattice consisting of 18 cancrinite-type tetrahedral cages connected to double 6-column columns and cross-linked with single-15 oxygen bridges to form planar 12-membered rings.

These 12-membered rings produce wide channels parallel to the c-axis without stacking errors. Unlike erionite and cancrinite, the cancrinite cages are arranged symmetrically across the double 6-ring units. There are four types of cation positions: A) in the double-6-ring gene, B) in the cancrinite-type cages, C) between the cancrinite-type cages, and D) on the duct wall. The cations at position D appear to be the only interchangeable cations at room temperature. During dehydration, cations in position D probably move from the channel walls to a fifth position, position E, found between the A positions. The hydrocarbon sorption pores have a diameter of approx. 7 to 8 Å.

A more complete description of these zeolites is given in e.g. U.S. Pat.

3,216,789, which provides a more detailed description of these zeolites in more detail.

Various factors influence the X-ray diffraction pattern of a zeolite. Such factors include temperature, pressure, crystal size, impurities and the type of cations present. Eg. the X-ray diffraction pattern becomes wider and less accurate 17

DK 163571 B

thick as the crystal size of the type L zeolite becomes smaller. Thus, the term "zeolite L" includes any zeolite consisting of cancrinite cages having an X-ray diffraction pattern substantially similar to the 5 X-ray diffraction patterns set forth in U.S. Patent No. 3,216,789.

The crystal size also has an effect on the stability of the catalyst. For reasons not yet fully understood, catalysts with at least 10 80% of the type L zeolite crystals greater than 1000 Å give longer operating times than catalysts with substantially all type L zeolite crystals between 200 and 500 Å.

The larger among these crystal sizes of type L zeolite is thus the preferred support material.

Type L zeolites are conventionally synthesized mainly in the potassium form, ie. that in the previously stated theoretical formula most of the M cations are potassium. The M-Cat ions are interchangeable, so that a given type of L zeolite, e.g. a potassium type L zeolite, may be used to obtain type L zeolites containing other cations by subjecting the type L zeolite ion exchange treatment to an aqueous solution of appropriate salts. However, it is difficult to replace all the original cations, e.g. potassium, since some interchangeable cations in the zeolite are found at positions that are difficult to reach for the reagents.

Other zeolites which could be used according to the invention are "zeolites of the L family". The term "zeolite of the L family" is defined as any zeolite consisting of cancrinite cages to obtain a pore structure whose pores are limited by a 12-sided ring wherein the zeolite has dehydrocyclization activity if a group VIII metal is introduced into the zeolite. Zeolites representative of a "zeolite of the L family" include type L zeolite as described in U.S. Patent No. 3,216,789, AGI as disclosed in GB Patent No. 1,393,365, AG4 as described in GB Patent No. 1,394,163, AG5, as described in U.S. Pat

DK 163571 B

tent publication no. 3,298,780 and K, Ba-G zeolite.

An essential feature of the present invention is the presence of an alkaline earth metal in the Type L zeolite. This alkaline earth metal must be either 5 barium, strontium or calcium. The alkaline earth metal is preferably barium. The alkaline earth metal can be introduced into the zeolite by synthesis, impregnation or ion exchange.

Barium is preferred over the other ground-eel potassium metals because the resulting catalyst has high activity, high selectivity and high stability.

In one embodiment, at least a portion of the alkali metal is replaced by barium using methods known for ion exchange of zeolites.

This involves contacting the zeolite with a solution containing an excess of Ba ++ ions. The barium should preferably be from 0.1 to 35% by weight of the zeolite, more preferably from 1 to 20% by weight.

The catalyst of the invention contains one or more Group VIII metals, e.g. nickel, ruthenium, rhodium, palladium, iridium or platinum.

The preferred Group VIII metals are iridium, palladium and especially platinum as these are more selective in dehydrocyclization and are also more stable under the dehydrocyclic reaction conditions than other Group VIII metals.

The preferred percentage of platinum in the catalyst is between 0.1 and 5%, more preferably from 0.1 to 1.5%.

Group VIII metals are introduced into the L zeolite by synthesis, impregnation or exchange in an aqueous solution of a suitable salt. When it is desired to introduce two Group VIII metals into the zeolite, this can be done simultaneously or sequentially.

Eg. platinum can be introduced by impregnating the zeolite with an aqueous solution of tetramine platinum.

DK 163571 B

(II) nitrate, tetraamine platinum (II) hydroxide, dinitrodiami noplatin or tetramine platinum (II) chloride. In an ion exchange process, platinum can be introduced using cationic platinum complexes such as tetraamine platinum (II) -5 nitrate.

An inorganic oxide can be used as a carrier to bind the type L zeolite containing the Group VIII metal and the alkaline earth metal / And give the catalyst additional strength. The carrier may be a natural or synthetic inorganic oxide or combination of inorganic oxides. Preferred content of inorganic oxide is from 5 to 25% by weight of the catalyst. Typical inorganic oxide carriers which may be used include aluminosilicates (such as clays), alumina and silica, wherein acidic positions are preferably exchanged with cations which do not provide strong acidity.

A preferred inorganic oxide carrier is "Ludox", which is a colloidal suspension of silica 20 in water, stabilized with a small amount of alkali.

When a urogenic oxide is used as a carrier, there are two preferred methods by which the catalyst can be prepared, although other embodiments may also be used.

In the first preferred embodiment, the Type L zeolite is prepared, then the Type L zeolite is ion exchanged with a barium solution, separated from the barium solution, dried and calcined, impregnated with platinum, calcined and then mixed with the inorganic oxide and extruded through a nozzle to form cylindrical pellets. Advantageous methods for separating the type L zeolite from the barium and platinum solutions are with a batch centrifuge or a press filter. This embodiment has the advantage that 35 al of barium and platinum are taken up in the type L zeolite and that nothing is absorbed in the inorganic oxide. It has the disadvantage that the type L zeolite has a small size and is difficult to separate from the barium solution and

DK 163571 B

platinum solution.

In the second preferred embodiment, the type L zeolite is mixed with the inorganic oxide and extruded through the nozzle to form cylindrical pellets, and then these pellets are ion exchanged with a barium solution, separated from the barium solution, impregnated with platinum, separated from the platinum solution and separated. calcined. This embodiment has the advantage that the pellets are easily separable from the barium and platinum solutions, but it has the disadvantage that barium and platinum can also be deposited on the inorganic oxide support material and can catalyze unwanted reactions. Thus, the choice of the embodiment used depends on a balance between the catalyst selectivity and the ease with which the catalyst is separated from the barium and platinum solutions.

In a third possible embodiment, the Type L zeolite is ion exchanged with a barium solution, dried and calcined, mixed with the inorganic oxide and extruded through the nozzle to form cylindrical pellets, then impregnated with platinum, separated from the platinum solution and calcined.

In the extrusion of type L zeolite various extrusion aids and 25 pore forming agents can be added. Examples of suitable extrusion aids are ethylene glycol and stearic acid. Examples of suitable pore forming agents are wood flour, cellulose and polyethylene fibers.

After the desired Group VIII metal or 30 metals has been introduced, the catalyst is treated in air or diluted C> 2 at 260 to 500 ° C and then reduced in hydrogen at temperatures of 200 to 700 ° C, preferably 480 to 620 ° C. , and more preferably 530 to 620 ° C.

At this stage, the catalyst is ready for use in the dehydrocyclization process. In some cases, e.g. however, when the metal or metals have been introduced by an ion exchange process, 21 is preferred

DK 163571 B

removing any residual acidity in the zeolite by treating the catalyst with an aqueous solution of a salt or hydroxide of an appropriate alkali or alkaline earth metal to neutralize any hydrogen ions formed during the reduction of the metal ions with hydrogen.

In order to achieve optimum selectivity, the temperature should be regulated so that the reaction rate is significant, but the conversion is less than 98%, as too high a temperature and too far - reaction can have a detrimental effect on the selectivity. The pressure should also be regulated within an appropriate range. Excessive pressure will result in a thermodynamic (equilibrium) constraint. for the desired reaction especially for hexane aromatization, and too low a pressure can result in coking and deactivation.

Although the primary advantage of the invention is an improvement in the selectivity for converting acyclic hydrocarbons (especially Cg-Cg paraffins) to flavorings, it has also surprisingly been found that the selectivity for conversion of methylcyclopentane to benzene is excellent. This reaction, involving conventional reformed catalysts based on chlorinated alumina, involves an acid-catalyzed isomerization step, occurring on the catalyst of the invention with a selectivity as good as or better than on the chlorinated alumina-based catalysts known. Thus, the present invention can also be used to catalyze the conversion of high starting materials of 5-membered alkyl naphthene to aromatics.

The dehydroisomerization of alkylcyclopentanes is preferably carried out at a temperature of from 426 to 538 ° C, more preferably from 437 to 510 ° C, an LHSV of 35 from 0.1 to 20, more preferably from 0.3 to 10, a pressure of from 0 to 0. to 3.45 MPa overpressure, more preferably from 1 atmosphere to 2.07 MPa overpressure, and a K

DK 163571 B

of from 0 to 20: 1, more preferably from 1: 1 to 10: 1.

It has also been surprisingly found that the selectivity for converting toluene to benzene is also excellent. The dealkylation of toluene is preferably carried out at a temperature of from 426 to 649 ° C, more preferably from 454 to 593 ° C, an LHSV of from 0.1 to 20, more preferably from 0.3 to 10, a pressure of from 0 to 20.7 MPa overpressure, more preferably from 1.38 to 13.8 MPa overpressure, and an I / HC ratio of from 0 to 20: 1, 10 more preferably from 1: 1 to 10: 1.

Another advantage of the invention is that the catalyst of the invention is more stable than known zeolite catalysts. The stability or resistance of the catalyst to deactivation determines its useful life. Longer service life results in less downtime and less cost of regenerating or replacing the catalyst charge.

The invention is described in more detail in the following examples which indicate particularly advantageous methods and compositions.

Example 1

An Arabian Light so-called "straight run" naphtha, which was hydro-purified to remove sulfur, oxygen, and nitrogen was reformed at 0.69 MPa overpressure, a LHSV of 2 and an H / HC ratio of 6 with three different catalysts. The feed stream contained 80.2 volume percent paraffins, 16.7 volume percent naphthenes and 3.1 volume percent aromatics, and it contained 21.8 volume percent C 2, 30 52.9 volume percent C 9, 21.3 volume percent and 3.2 volume percent C 9.

In the first, the naphtha was reformed at 499 ° C using a commercial sulphided platinum-rhenium alumina catalyst prepared as described in U.S. Patent No. 3,415,737.

In the second experiment, the naphtha was reformed at 493 ° C using a platinum-potassium-type L zeolite catalyst prepared by: (1) impregnating a potassium-type L zeolite with crystal sizes from ca.

23

DK 163571 B

1000 to 4000 Å containing 0.8% platinum impregnated as tetramine platinum (II) nitrate; (2) drying the catalyst; (3) calcining the catalyst at 260 ° C and (4) reducing the catalyst in hydrogen at 480 ° C 5 to 500 ° C for 1 hour.

In the third experiment performed according to the present invention, the naphtha was reformed at 493 ° C using a platinum-barium-type L zeolite catalyst prepared by: (1) ion exchange of a potassium-type L zeolite with crystal sizes of 1000 to 4000 Å with a sufficient volume of 0.17 molar barium nitrate solution to obtain an excess of barium relative to the ion exchange capacity of the zeolite; (2) drying the resulting barium-exchanged type L zeolite catalyst; (3) calcining the catalyst at 590 ° C; (4) impregnating the catalyst with 0.8% platinum using tetramine platinum (II) nitrate; (5) drying the catalyst; (6) calcining the catalyst at 260 ° C and (7) reducing the catalyst in hydrogen at 20 480 to 500 ° C for 1 hour.

The results of these three experiments are listed in Table I.

DK 163571 B

24

Table I

499 ° 493 ° 493 °

Pt / Re / Pt / K / L Pt / Ba / L

Feed stream alumina 5% by weight feed stream 2.8 5.5 3.6 C2 6.6 2.5 1.3 C3 9.3 3.2 1.5 iC4 0.1 5.8 0.9 0.5 KC4 0.5 6.8 · 3.8 2.4 10 iC5 5.1 13.6 6.7 5.6 NC5 11.3 9.8 12.6 12.6 C6 + P + N 81.3 13.4 7.8 9.3

Benzene 1.5 15.1 40.6 43.8 C + Aromatics 0.8 15.8 12.7 15.0 15 C5 + LV% yield 63 69.9 74.4

Hydrogen, SCF / B 470 1660 2050

Selectivity, mole% 20 72 87

Cg + P -> aromatics

This series of experiments shows that the use of a platinum-barium type L zeolite catalyst in reforming provides a selectivity for converting hexanes to benzene which is clearly better than the selectivity of the prior art. It should be noted that, together with this better selectivity, an increased production of 25 hydrogen gas can be obtained which can be used in other processes. It is also noted that the hydrogen purity is greater by. The Pt / Ba / L test, as more hydrogen and less plus C 2 are produced. Example 2

A second series of experiments was carried out using hydro-purified n-hexane as feed stream. All the experiments in this series were performed at 490 ° C, 0.69 MPa overpressure, a LHSV of 3 and a H2 / HC ratio of 3.

In the first experiment, a platinum-potassium type L zeolite was used which was prepared following the procedure set forth in the second experiment in Example 1.

In the second experiment, a platinum-barium 25 was used

DK 163571 B

type L zeolite prepared following the procedure set forth in the third experiment of Example 1 except that the barium nitrate solution was 0.3 molar instead of 0.17 molar. The results of these experiments are given below in Table II.

Table II

Conversion rate Selectivity for _ n-hexane_ 5 hours 20 hours 5 hours 20 hours

Pt / K / L 70 59 76 79 10 Pt / Ba / L 85 85 89 92

Thus, in operation, the introduction of barium into the type L zeolite causes a dramatic improvement in the selectivity for n-hexane. Note that the stability of the platinum-barium type L zeolite is excellent. After 15 h, there was no decrease in conversion rate using the platinum-barium type L zeolite.

Example 3 In a third series of experiments, various 20 cation replacements were used. All experiments in this series were performed at 490 ° C, 0.69 MPa overpressure and a H2 / HC ratio of 6. The feed stream was hydro-purified, containing 80.9% by volume paraffins, 16.8% by volume naphthenes, 1.7% by volume aromatics , 0.4 vol% olefins, and contained 2.6 vol% Cg, 47.6 vol% Cg, 43.4 vol% C 2 and 6.3 vol% C 6 carbohydrates.

In the first experiment, a platinum-barium type L zeolite prepared according to the procedure used in the second experiment of Example 2. was used. It was tested with an LHSV of 2.0.

In the second experiment, a platinum-calcium type L zeolite was prepared following the same procedure except for a 0.3 molar calcium nitrate solution. It was tested at a LHSV of 2.0.

In the third experiment, a platinum-strontium-type L zeolite was used which was made after the same

DK 163571 B

except a 0.3 molar strontium nitrate solution was used in the exchange. It was tested at a LHSV of 2.0.

In the fourth experiment, a platinum-cesium 5 type L zeolite prepared according to the same procedure was used except that in the exchange a 0.3 molar cesium nitrate solution was used. It was tested at a LHSV of 2.0.

In the fifth experiment, a platinum-barium-10 type L zeolite was used which was prepared by the same procedure as in the first experiment. It was tested at an LHSV of 6.0.

In the sixth experiment, a platinum-potassium-type L zeolite was prepared following the procedure set forth in the second experiment in Example 1. It was tested at an LHSV of 6.0.

In the seventh experiment, a platinum-rubidium-type L zeolite was prepared following the procedure of the first experiment, except that a 0.3 molar rubidium nitrate solution was used in the exchange. It was tested at an LHSV of 6.0.

In the eighth experiment, a platinum-lanthan-type L zeolite prepared according to the procedure of the first experiment was used, except that in yield 25 a 0.3 molar lanthanum nitrate solution was used. It was tested at an LHSV of 6.0.

In the ninth experiment, a platinum-magnesium-type L zeolite was prepared following the procedure of the first experiment, except that at the replacement a 0.3 molar magnesium nitrate solution was used. It was tested at a LHSV of 6.0.

In the tenth experiment, a platinum-lithium-type L zeolite was prepared following the procedure of the first experiment, except that in yield 35 a 0.3 molar lithium nitrate solution was used. It was tested at an LHSV of 6.0.

In the eleventh experiment, a platinum-sodium type L zeolite was used which was prepared according to the procedure of 27

DK 163571 B

the first experiment, except that a 0.3 molar sodium nitrate solution was used in the replacement.

It was tested at a LHSV of 6.0. The results from all eleven trials are given in Table III.

5 'Table III

Paraffinoma Formation - Selection Aroma degree of mole% mole% feed stream 3 hours 20 hours 3 hours 20 hours 3 hours 20 hours

2.0 LHSV

Pt / Ba / L 87 87 83 85 75 76

Pt / Ca / L 96 93 63 70 65 70 10 Pt / Sr / L 92 83 61 69 62 63

Pt / Cs / L 84 73 73 75 66 61

6.0 LHSV

Pt / Ba / L 73 63 80 83 64 59

Pt / K / L 75 66 71 75 60 57

Pt / Rb / L 84 74 72 77 65 63 15 Pt / La / L 63 55 63 58 50 48

Pt / Mg / L 23 <5 23 - 22 14

Pt / Li / L 77 72 70 74 61 60

Pt / Na / L 74 67 72 75 60 58

Thus, in operation, the introduction of barium into type L zeolites results in a dramatic improvement in selectivity over type L zeolites with other cations (more than 25% reduction in the amount of manufactured cracked products for an L zeolite with any other cation ).

Example 4 A platinum-barium type L zeolite prepared following the procedure set forth in the second experiment in Example 2 was used in two experiments. In the first experiment, the L-zeolite crystallite size was 1000 to 2000 Å according to transmission electron microscopy. In the second experiment, the L-zeolite crystallite size was approx. 400 Å. The feed stream in both experiments contained 70.2% by volume paraffins, 24.6% by volume naphthenes, 5.0% by volume aromatics, and 29.7 by volume C1, 43.4% by volume C1, 21.2% by volume, 5.0% by volume. 35 octane percent Cg and 0.6 ounce C3. "Research octane clear" in the feed stream was 71.4. The catalyst for

DK 163571B

28 the first experiment was reduced in hydrogen for 20 hours at 566 ° C. The catalyst for the second experiment was reduced in hydrogen for 2 hours at 1050 ° F. The test conditions were 0.69 MPa overpressure, a LHSV of 1.5 and 1 H / HC 5 recirculation of 6.0. The temperature was adjusted to give 50% by weight of aromatics in the liquid C2 + product, which corresponds to 89 "research octane clear". The results of the first experiment are listed in Table IV.

Table IV

Test time, hours To 50 wt.% Aromatics C + + yield - temperature ° C_ liquid vol% 500 459 86.4 1000 464 86.2 2000 469 86.1 15 2500 471 86.2

The results of the second experiment are listed in Table V.

20

Table V

Test time, hours To 50 wt.% Aromatics C + yield - temperature ° C_ liquid vol% 100 466 87 200 472 86 25 400 478 85

In operation, the large crystal size catalyst of the invention provides an unusually long test time, whereas the small crystal size 30 catalyst is deactivated much faster.

Example 5

A fifth series of experiments was performed to show the effects of the reduction temperature. The 35 feed stream used was the same as in Example 1. The test conditions were 490 ° C, 0.69 MPa overpressure, a LHSV of 2.0, and a H2 / HC ratio of 6.0.

29

DK 163571 B

The catalyst used in the first two experiments was a Pt / Ba / K / L prepared as in the third experiment of Example 1, except that the barium replacement was performed three times with a 1 molar barium nitrate solution; This gave only a slight increase in the barium content.

The catalyst in the last two experiments was a Pt / K / L prepared as in the second experiment in Example 1.

The results obtained after three hours of operation are listed in Table VI.

10

Table VI

Reducing conditions Arcmater Selectivity _ mol% feed stream mol% _

Pt / Ba / K / L 490 ° C, 1 hour 59 78

Pt / Ba / K / L 620 ° C, 20 hours 60 89

ID

Pt / K / L 490 C, 1 hour 55 66

Pt / K / L 620 ° C, 20 hours 23 23

The results show a large and surprising difference in the effect of the preparation on the two catalysts. The selectivity of the barium-containing catalyst improved significantly, while the selectivity of the barium-free catalyst was greatly reduced.

Example 6 A methylcyclopentane feed stream which had been hydro-purified to remove sulfur, oxygen and nitrogen was dehydroisomerized at 493 ° C, 0.69 MPa overpressure, a LHSV of 2 and a H2 / HC ratio of 6 with a dehydroisomeric acid. seringskatalysator. This catalyst was prepared by 30 (1) ion exchange of a potassium barium type L zeolite with a sufficient volume of 0.17 molar barium nitrate solution to contain an excess of barium relative to the ion exchange capacity of the zeolite; (2) drying the resulting barium-exchanged type L zeolite catalyst; (3) calcining the catalyst at 590 ° C; (4) impregnating the catalyst with 0.8% platinum using tetramine-platinum (II) nitrate; (5) drying the catalyst; (6) calcining the catalyst at 260 ° C and (7) reducing the catalyst in hydrogen at 480 to 500 ° C.

30

DK 163571 B

The feed stream contained 73% by weight of methylcyclopentan. The feed flow distillation data was start: 64 ° C, 5%: 68 ° C, 10%: 69 ° C, 30%: 70.6 ° C, 50%: 71 ° C, 70%: 72 ° C, 90%: 74 ° C, 95%: 78 ° C and end point: 106 ° C.

5 The feed stream contained 10% by volume paraffins, 80% by volume naphthenes and 10% by volume aro-mater. The conversion rate was 78% after 3 hours and 32% after 20 hours. The selectivity for dehydroisomerization was 82 mol% at 3 hours and 86% at 20 hours.

10

Example 7

An Arabian so-called "straight run" feed stream which had been hydro-purified to remove sulfur, oxygen and nitrogen was dealkylated at 566 ° C, 0.69 MPa overpressure, an LHSV of 4 and an I / HC ratio of 3 with a dealkylation catalyst. This catalyst was prepared by (1) ion exchange of a potassium barium type L zeolite with a sufficient volume of 0.17 molar barium solution to contain an excess of barium relative to the ion exchange capacity of the zeolite; (2) drying the resulting barium-exchanged type L zeolite catalyst; (3) calcining the catalyst at 590 ° C; (4) impregnating the catalyst with 0.8% platinum using tetraminplatin (II) nitrate; (5) drying the catalyst; (6) calcination of the catalyst at 260 ° C and (7) reduction of the catalyst in hydrogen at 480 ° C to 500 ° C.

The feed stream contained 1 volume percent C 2, 22 volume percent C 2, 53 volume percent C 2, 21 volume 30 percent C 2 The results of this experiment are given in the table below (all percentages are by weight%). The conversion of Cg + paraffins 35 was 96.04 mol% and the selectivity for aromatization was 84.02.

DK 163571 B

31

Table VII

Feed stream Product

Methane 0. 7.66

Ethane 0 5.66 5 Propane 0 6.49 iC4 0.08 1.55 nC4 0.52 6.49 iC5 5.88 3.26 nC5 12.61 2.81 10 C5 + 80.91 70.34 H2 0 2 , 91

Benzene 1.69 51.10 C + Aromatics 1.62 10.92

Claims (9)

Catalyst according to claim 1, characterized in that the alkaline earth metal is barium and the group VIII metal is platinum. Catalyst according to claim 2, characterized in that it contains 0.1 to 35% by weight of barium 10 and 0.1 to 5% by weight of platinum. Catalyst according to claim 2 or 3, characterized in that it contains 1 to 20 wt% barium and 0.1 to 1.5 wt% platinum. Catalyst according to claims 1-4, characterized in that more than 50% of the type L zeolite crystals are greater than 500 Å. Catalyst according to claims 1-5, characterized in that more than 50% of the type L zeolite crystals are greater than 1000 Å. Catalyst according to claims 1-6, characterized in that at least 80% of the type L zeolite crystals are greater than 1000 Å. Use of the catalyst according to claims 1-7 for hydrocarbon reforming. Use according to claim 8 for dehydrocyclization of acyclic hydrocarbons. Use according to claim 8 for dehydroisomerization of alkylcyclopentanes. II. Use according to claim 8 for dealkylation of toluene for the production of benzene.