NO130394B - - Google Patents

Description

Method of manufacture

of aromatic hydrocarbons.

The present invention relates to a process for the production of polymethylbenzenes from alkylaromatic hydrocarbons which have at least one alkyl radical comprising more than one carbon atom.

Polymethylbenzenes are interesting materials used

in the production of aromatic polycarboxylic acids, e.g. fthai acids from xylenes, pyromallitic acid or anhydride from duren, trimesin-

acid from trimethylbenzene, etc. The commercial importance of these acids has been steadily increasing in recent years because they are useful raw materials for the production of synthetic resins.

On the other hand, large quantities of such alkylaromatic hydrocarbons are now available and are obtained, e.g. in the treatment of certain petroleum fractions or they are obtained as by-products in the production of other widely used compounds,

e.g. ethylbenzene, which is used in the production of styrene.

Consequently, it will be particularly advantageous to be able to

to convert these alkylaromatic hydrocarbons to polymethylbenzenes in an efficient manner.

Various processes are known to convert alkylaromatic compounds having an alkyl radical of more than 1 carbon atom to polymethylaromatic hydrocarbons. In such processes, the alkyl radicals with more than 1 carbon atom are transferred to methyl radicals, while the number of methyl radicals bound to the aromatic nucleus increases, e.g. conversion of ethylbenzene to dimethylbenzenes. The reaction can be carried out in one step (direct isomerization) or in 2 steps (hydroisomerization). In the latter case, the alkylaromatic hydrocarbons are initially subjected to hydrogenation and partial isomerization in one step and the resulting products are then subjected to further isomerization and dehydrogenation in another step.

The operating conditions of the process according to the state of the art are thus that the above-mentioned reactions are always impaired by secondary reactions, more particularly by the cracking reactions. Some of these processes are furthermore only available to one

single type of alkylaromatic hydrocarbons, so that the results obtained with homologous compounds are markedly less favorable.

One aspect of the present invention is to provide

a particularly selective process in the production of polymethylbenzenes with a minimum of secondary reactions. Another feature of the present invention is the production of methylbenzenes from fractions which have a high content of methylbenzene precursors. A further aspect of the present invention is to provide a method in which the polymethylbenzene foriopers are produced with good yields and in a form that can be easily separated. A further feature of the present invention is to produce a flexible process which allows the treatment of widely different alkylaromatic hydrocarbons as indicated above.

According to the present invention is provided

a process for the production of polymethylbenzenes from alkylaromatic hydrocarbons with the general formula:

where R 1 is an alkyl radical containing 2-6 carbon atoms, R 2 is an alkyl radical containing from 1-3 carbon atoms, R<3> is a methyl radical, n is an integer from 1 - 3 and m is 0 or 1

a) the alkylaromatic hydrocarbons are hydrogenated to the corresponding alkylcycloalkanes, the percentage being unreacted hydrocarbons

is not greater than 0.001,

b) the alkylcycloalkanes are isomerized in the presence of AICI3 at a temperature not higher than 150° C for a time sufficient to form a reaction product containing mainly gem-structured polymethylcyclohexanes with the same number of carbon atoms as in the starting material, c) the reaction product is brought into contact with an acid , difunctional dehydrogenation-isomerization catalyst in the presence of hydrogen in a reaction zone maintained at a temperature within the range of 200° - 600° C and a pressure not greater than 50 kg/cm<2>, and d) the reaction product containing polymethylbenzenes with the same number of carbon atoms as in the starting material is extracted.

The alkylaromatic hydrocarbons at the start can be added to the first stage of the process either alone or in admixture with another or with hydrogenated products of these. The feed is preferably mainly anhydrous if the product from the first stage is to be subjected to the second stage of isomerisation without further treatment, because the isomerisation reaction is impaired in the presence of water. Among the hydrocarbons having the above formula, ethylbenzene, cumene, diethylbenzenes, cymenes and tri-ethylbenzenes are preferably used, because these are easily available and they give the di-, tri-, tetra- and hexamethylbenzenes which are used in the preparation of the corresponding benzene-di -, tri-, tetra- and hexacarboxylic acids.

The hydrogenation of the feed according to the first step of the method according to the present invention is carried out in the presence of a catalyst containing a metal from group VIII in the periodic table, e.g. nickel alloy or a metal from the platinum group such as e.g. platinum, palladium, ruthenium, osmium, rhodium and the like, or a platinum alloy, e.g. an alloy with rhenium. The catalyst generally includes a carrier, a substrate, such as e.g. silicon dioxide, aluminum oxide, a mixture of silicon dioxide and aluminum oxide or a zeolitic compound. It can be used in any of different possible forms, i.e. in the form of granules or pellets or in finely powdered form. The catalyst can be used in a fixed layer or in a fluidized layer, or in another moving layer. In general, catalysts are used comprising nickel on silicon dioxide and containing from 5 to 60% by weight of nickel, and catalysts comprising platinum on a support and containing 0.1 to 1.5% by weight of platinum.

Because the hydrogenation reaction is highly exothermic

it is preferred to dilute the feed with recycled products emerging from a previous hydrogenation and/or the second isomerization step. The weight ratio between fresh feed and recycled products can vary within wide limits, e.g. between 1 : 1 and 1 : 50, but this is preferably between 1 : 2 and 1 : 10. The reaction temperature is generally between 150 and 350°C and especially between 200 and 300<Q>C, to avoid cracking and deposition of soot on the catalyst.

To avoid deactivation of the isomerization catalyst during the second stage of the process, the hydrogenation of the starting hydrocarbons must be as complete as possible. For this purpose, the partial pressure of these hydrocarbons can be reduced by using a high ratio between hydrogen and hydrocarbons. However, the use of large amounts of hydrogen is not economical and in general the ratio between hydrogen and hydrocarbons (fresh feed plus recycled hydrocarbons) will vary from 2:1 to 30:1, and preferably between 3:1 and 20:1.

Hydrogenation yields can be increased by using higher pressures, but usually the pressure is kept between 20 and 70 kg per cm 2. In reality: higher pressures do not result in an improvement in the results that would justify additional costs for such higher pressures.

The space velocity before birth depends on many factors, e.g. dilution ratio, temperature and pressure. A space velocity for the liquid with regard to the total feeding, possibly of the order of magnitude 0.5 - 50, will prove to be appropriate, and in general the preferred residence time will vary between 1 and 12 volume parts per hour per volume fraction catalyst, but this time can be longer when the hydrogenation pressure is higher.

The relationship between the operating conditions for the hydrogenation reaction is shown in Table I, which indicates the results of hydrogenation experiments with a mixture comprising equal parts of diethylbenzenes and diethylcyclohexanes in the presence of a catalyst consisting of platinum (0.9% by weight) on a silicon dioxide-alumina support, at a pressure of 50 kg per cm 2. The investigated values were: temperature, molar ratio between hydrogen and hydrocarbons in the feed (H2/HC), and the space velocity (S.V.) expressed in liters per hour per liters of catalyst.

The operating conditions were as follows:

Experiment A : H2/HC = 3.5 : 1 S.V. = 2

.Experiment B : H2/HC =9 : 1 S.V. = 2

Experiment C : H2/HC =9:1 S.W. = 1

For each experiment, the percentage value of non-hydrogenated hydrocarbon in the reaction products is indicated in Table I.

A new set of experiments was carried out with a catalyst containing nickel (50% by weight) on silicon dioxide, at 240°C and 50 kg per cm<2>, with a H2/HC ratio of 6:1 and a space velocity S.V. of 4. Under these conditions, the percentage of unreacted hydrocarbon is less than 0.001. Practically the same results were obtained by using, under these conditions, a catalyst containing 0.7% by weight platinum on aluminum oxide.

The second step of the process involves the isomerization of the alkylcycloalkanes obtained from the hydrogenation reaction to polymethylcycloalkanes. For example, diethylcyclohexane is isomerized by breaking down the ethyl groups and forming methyl groups, a significant part of which are gem-methyl groups, according to the following reaction:

D-t has been shown- -that-this reaction gives unexpectedly high yields

while it is not effective with alkylaromatic hydrocarbons such as the output smat is in al er. To achieve a quick transfer by simply passing through

run, it is desirable from a practical point of view that the feed material is mainly free of aromatics, i.e. that it contains less than 1000 parts per million of aromatics. This condition is fulfilled when the first step or the hydrogenation step is carried out as described above.

To illustrate the formation of gem.-polymethylcycloalkanes,

namely, in the case of hydrocarbons containing at least 10 carbon atoms, an isomerization experiment was carried out with diethylcyclohexane in the presence of AlCl^, where the said diethylcyclohexane was obtained by hydrogenation of diethylbenzene as described above. The isomerization products were divided into 5 equal parts that had increasing boiling points. With these C-^Q hydrocarbons, it was found that the tetramethylbenzene precursors appeared to be primarily in the fractions which had boiling

point between 150 and 165°C. These fractions had a high content of gem.-tetramethylcyclohexanes. These geminated polymethylcyclohexanes have a lower boiling point than the corresponding non-geminated polymethylcyclohexanes, and consequently they can be easily separated by distillation.

The reaction is carried out according to the present invention under substantially anhydrous conditions in the presence of a Friedel-Crafts catalyst or a mixture of Friedel-Crafts catalysts. Any suitable Friedel-Crafts catalyst which is

known in the field can be used within the scope of the present invention. The preferred group of catalysts includes AlCl₂,

BF^, SnCl^, SbCl^, ZnCl9, TiCl)+, GaCl^, AlBr^, FeCl^, GaBr^ and

the like. The most useful catalysts are AlCl^ and mixtures of AlCl-, and SbCl^. The catalyst is generally used in amounts of from 1-60 percent by weight of the feed material in the reaction zone,

and preferably from 2-^-0 weight$ of the feed material in the reaction

the zone during batch reactions, and from 10 - U- 0% by weight of the feed material in the reaction zone during continuous reaction. A frequent addition of small amounts of catalyst during the continuous isomerization is advantageous for maintaining a high product yield, which will be further explained later in the following description.

In addition to the Friedel-Crafts catalyst described above, it is desirable according to the present invention to include a promoter in the reaction zone, which promoter is selected from the group comprising anhydrous hydrogen halide, preferably hydrogen chloride, alkyl halides, olefins and mixtures thereof. Suitable alkyl halides include alkyl monohalides having from 1-1h carbon atoms, preferably 2-8 carbon atoms, while suitable olefins include monoacyclic and monoalicyclic olefins having from 2-1h carbon atoms, preferably 5-12 carbon atoms. Examples of suitable such materials are tert-butyl chloride, tert-butyl bromide, 2,2,^-trimethylpentene, 3,3,5-trimethylcyclohexene, iso-propylcyclohexene, sec-butyl chloride, 2,2, h-triethylpentene, 3,3, 5-triethylcyclohexene and the like .. When the reaction is carried out under batch conditions, especially when fresh catalyst is used, it has been found that the presence of both hydrogen halide and an alkyl halide, olefin or mixture thereof is desirable in order to achieve rapid conversion of the alkylcyclohexanes to the gem .-structured polymethylcyclohexanes. In addition, isomerization activity lost in the presence of significant amounts of aromatic compounds can be restored by such promoters. However, when the method according to the present invention is carried out with the active catalyst complex obtained under continuous process conditions, it has been shown that the presence of only the hydrogen halide gives satisfactory results.

The hydrogen halide is preferably fed to the alkylcyclohexane feed for contact with the Friedel-Crafts catalyst. In general, the amount of hydrogen halide used to promote the reaction according to the present invention is sufficient to saturate the feed under the conditions used, preferably is

the amount of hydrogen chloride used ranges from 0.00001 to 5 wt% of hydrogen halide. The promoter, e.g. The alkyl halide or olefin is preferably added to the catalyst in the reactor in the quantity? of less than approx. 5% by weight, calculated on the basis of the feed to the reactor,

and more preferably in the range from 0.01 - 1.0 by weight of the feed in the reactor.

Another reaction-promoting effect can be achieved by using

a mixture of 2 or more Friedel-Crafts catalysts, e.g. AICT^ and SbCT^. By using such mixtures, one also finds

high isomerization activities even in the absence of the above-mentioned olefin, alkyl halide and hydrogen halide promoters.

The isomerization reaction is carried out at a lower temperature

than 150°C. Preferably, the reaction is carried out at a temperature between 0 and 70° C to avoid secondary reactions, e.g. cracking reactions and disproportionation reactions. A more preferred reaction temperature range is between 10 and 50 C, e.g. used approx. 30°^ to ensure prolonged catalyst activity and a high conversion rate to gem.-polymethylcyclohexanes.

The pressure in the reaction zone can be atmospheric pressure or sub-atmospheric pressure, where the choice depends to a large extent on economic and practical considerations.

In the preferred embodiment with the most useful additional raw materials, the feed is led in series through two reactors of the "backmix" type, containing, for example, approximately equal amounts of the Friedel-Crafts catalyst. It has been shown that under such circumstances very high degrees of conversion to polymethylcyclohexanes are achieved. The use of two such reactors allows the use of a feed containing a higher percentage value of aromatic compounds than is usually desirable in the feed.

The contact time for the alkylcyclohexanes in the reactor depends on such factors as temperature, type of catalyst and the like. However, it is usually preferred to have a contact time between the alkylcyclohex:antilphbryl and the catalyst of between 0.1 and 10 hours. Typical values for the space velocities in continuous operation are from 0.1 - 10 parts by weight of fodder per weight part catalyst per hour. As

■ illustration of the fluctuation in the contact time according to the continuous process according to the present invention, a contact time of 6 hours can be used with 15% by weight of catalyst (calculated on the basis of the supply in the reactor) and a reaction temperature of 30°C, as well as when using 30% by weight of catalyst (calculated on the basis of the supply in the reactor) and a reaction temperature of 20°C. Another example includes a 2 hour contact time with 30% catalyst (calculated on the basis of the supply in the reactor) and a reaction temperature of 3°C.

In batch operation, the preferred embodiment is to mix the desired amount of HOI in the supply material at the beginning. Next, the desired amounts of Friedel-Crafts catalyst and alkyl halide or oleic promoter are added to the HC1-containing feed material. If a Friedel-Crafts catalyst mixture is used, e.g. AlCl^ and SbCl^ other promoters can be omitted.

The reactor can be any conventional reactor, e.g. a batch reactor system with rebar plant. After a reaction time of between 0.1 and 10 hours, the product from the batch reactor contains from 20 - ^ 0% polymethylcyclohexane product with gem. structure, when the starting material is diethylcyclohexane. This material can then be fed to a distillation zone and fractionated. They save. - structured polymethylcyclohexane products according to the present invention have a lower boiling point than the corresponding non-gerinated methylcyclohexanes, and therefore they can be easily separated by distillation. The preferred chemical structures, such as tetramethylcyclohexanes, also have atmospheric boiling points between 150 and 165°C and can be easily separated from the reactor's output stream by means of a conventional de-scaffolding column with several stages. The bottom fractions from such a separation are preferably recycled.

When operating the continuous liquid phase reaction according to

for the present invention, the continuous reaction can be initiated with another Friedel-Crafts catalyst or with the alkyl halide or the olefin promoter in combination with the hydrogen halide as indicated above. Immediately when stable catalyst activity is achieved, however, the presence of the second Friedel-Crafts catalyst or the promoter based on alkyl halide and/or olefin is unnecessary, but the presence of hydrogen halide is still desirable. Alternatively, the continuous reaction can be initiated longitudinally without the other Friedel-Crafts catalyst, when the promoter based on alkyl halide or olefin is present in the reaction zone.

In general, the catalyst activity is a function of its concentration in relation to the feed, as well as the temperature, where the initiation time decreases when. the temperature rises. When approx. 50 parts by weight for seals per weight part AlCl^ sr obtained, the catalyst will become liquid. Fresh catalyst is then added in an amount depending on the operating conditions. Such an amount will usually be from 0.1 -

5.0% by weight, preferably 0.2 - 2% by weight, with respect to the amount of hydrocarbons fed into the reactor. An equivalent amount of the catalyst complex is withdrawn to maintain a constant catalyst concentration.

According to a preferred embodiment of the process according to the present invention, diethylcylohexane is reacted in a continuous process in the presence of 10-^-0% by weight of AlCl^, calculated on the basis of the feed in the reactor and at a catalyst temperature between 20 and 30° C. Diethylcyelohexane is initially saturated with HCl (dry) and is maintained saturated with HOI during the contact time, usually 1-1} hours. • High yields of gem.-structured tetramethylcyclohexanes are obtained by this method.

According to another embodiment thereof. present invention, diethylcylohexane is reacted in a batch operation in the presence of 5 - 30% by weight fresh AlCl.,, calculated on the basis of the supply in the reactor, and at a catalyst temperature of between 20 and 30 0. The diethylcylo-hexane is maintained for a reaction period of 3 - 6 hours saturated with hydrogen chloride. An alkyl halide with 3 - 6 carbon atoms, e.g. sec.- or tert.-butyl chloride, is present as another promoter in an amount of 0.01 - 1.0% by weight, calculated on the basis of the feed.

The polymethylcyclohexanes from the second stage of the process are then subjected to dehydrogenation and isomerization to polymethylbenzenes. In the case of the previous example, this reaction is:

The polymethylcyclohexanes from the second step can be sent directly to the third step of the process, and such a method is usually used when the hydrocarbons contain 8 or 9 carbon atoms. In the case of hydrocarbons having more than 9 carbon atoms, however, the polymethylcycloalkanes are advantageously distilled to separate the gem compounds from the other products.

The catalyst used in the third step of the method according to the present invention is difunctional with acidic points which promote the isomerization action and which also have dehydrogenation points." The catalyst preferably comprises a metal from the platinum group or an alloy with platinum, for example with rhenium, on a support such as silicon dioxide and/or aluminum oxide and/or zeolitic compounds A difunctional catalyst preferably comprises between 0.1 and 5% by weight, and more particularly between 0.2 and 3 wt.% platinum on an amorphous or crystalline silicon dioxide-aluminum oxide carrier consisting of 5-95 wt.% silicon dioxide, preferably 10-50 wt.% silicon dioxide for amorphous silicon-aluminum oxide and 30-90 wt.% silicon dioxide for crystalline silicon dioxide-aluminum oxide.

The reaction involving dehydrogenation-isomerization is endothermic and thus high temperatures are advantageous. In order to prevent cracking reactions, however, it is preferred to use low temperatures. Thus, a compromise must be sought between temperature and reaction rate. In general, the temperature is kept between 200 and 600 o 0, preferably between 300 and 500<o>"C. On the other hand, the pressure will usually be less than 50 kg per cm 2 and preferably kept between atmospheric pressure and 35 kg per cm 2 , while the space velocity for the treated hydrocarbons will be kept between 0.5 and 30 liters of hydrocarbons per hour and per liter of catalyst.

The polyalkylcyclohexane feed must be mixed with the appropriate ratio of aromatic compounds in order to achieve good temperature regulation of the reaction. The additional aromatics can be mixed with the semi-structured polyalkylcyclohexane binder for this feed to the dehydrogenation reactor, or they can be fed separately to the reactor. Preferably, the two materials are mixed before they are fed to the reactor. It is preferred that the molar ratio between gem.-structured polyalkylcylohexanes and aromatic compounds in the reactor is maintained within the range 0.1 : 1 to : 1. On the basis of factors such as product yield and the availability of appropriate aromatic compounds, the gem.-structured polyalkylcylohexanes with aromatics in a molar ratio of 0.1:1 to 1:1. The preferred aromatic compounds which can be mixed with the feed are alkylbenzenes. When, for example, the durene is produced from gem.-tetramethylcyclohexanes, tetramethylbenzenes such as e.g. the isodure and the prehnite preferred. Other suitable aromatic compounds include benzene, toluene, xylenes, trimethylbenzene, ethyltoluene, and the like.

The total mole ratio between hydrogen and hydrocarbon in the components entering the dehydrogenation reactor is preferably in the range 3 : 1 to <>>+0 : 1, and more preferably from 6 : 1 to 25 r i. The method according to the present invention produces more than sufficient hydrogen for recirculation to the reaction zone to provide appropriate hydrogen to hydrocarbon ratios.

Comparative examples have shown that in the method according to the present invention, alkylaromatic hydrocarbons as indicated above are transferred to polymethylbenzenes with yields and selectivities that are significantly higher than what is achieved by processes according to the state of the art.

For example, a feed containing 100% ethylbenzene was hydrogenated and the ethylcyclohexane obtained was isornerized in the presence of AlCl₂. The isornerization products were then dehydrogenated without prior distillation. The analysis of the products gave the following results:

In the process according to the present invention, the conversion yield of ethylbenzene to xylenes was thus 80.0% by weight, while processes according to the state of the art give yields which are generally in the range of 0 - 5% by weight.

In another trial, bee cumen was treated according to the method according to the invention. On analysis, it was found that the products contained 68.6% by weight of trimethylbenzenes. A dividend on

about. <*>+5% by weight is usually achieved by conventional known processes.

■ Other features and characteristics of the method according to the present invention will be further described and exemplified by the following examples, which are indicated with reference to the attached drawings, where: Fig. 1 shows a schematic flow diagram of a embodiment of the present invention. Fig. 2 shows a schematic flow diagram of another specific embodiment of the present invention.

Example- 1

A feedstock consisting mainly of diethylbenzenes was treated in accordance with the method according to the present invention, as shown in fig. 1. This alkylaromatic feed was previously used. The feed mixture, 100 parts by weight, was led through pipe line 10 in combination with a recirculation stream containing material from the hydrogenation step, from the isomerization step and from the dehydrogenation-isomerization step in the process according to the invention. The resulting mixture amounted to 576.8 parts by weight and was quickly sent via pipeline 11 to the hydrogenation reactor 12.

In the hydrogenation reactor 12, the mixture of fresh alkylaromatic feed and recirculate is passed over a catalyst containing 50% by weight of nickel and the remaining part is a silicon dioxide carrier on which the nickel was distributed. Hydrogen was supplied to the hydrogenation reactor 12 by means of pipeline 13 at the same time as the hydrocarbon feed was supplied via pipeline 11. The molar ratio between hydrogen supply and hydrocarbon feed was 6:1. The space velocity of the liquid per hour was ^f, and the pressure was 50 kg per cm", while the temperature was maintained at approx. 2<1>+0°G. -

The hydrogenated product quickly left the hydrogenation reactor 1? by means of piping lh- and was separated into a recirculation stream (301 parts by weight) and an isomerization supply stream (275.8 parts by weight). The recirculate stream is led by means of pipe line 15 into a common recirculation pipe line 16, through which the combined recirculate stream from the hydrogenation step, the isomerization step and the dehydrogenation-isomerization step is led into pipe line 10 and mixed with fresh feed to the hydrogenation reactor 12. The isomerization feed is led from pipe line l<*>f via rudder line 17 into an isomerization reaction zone 18.

The hydrogenated hydrocarbons are brought into the reaction zone 18

in contact with an AlCl^ catalyst. The concentration of the AlCl₂ catalyst in the resulting reaction mass was 30% by weight, calculated on the basis of the weight of the liquid. Simultaneously with the supply of the hydrogenated hydrocarbons in the isomerization reaction zone 18, anhydrous HC1 is supplied to the reaction zone 18 via pipeline 19. In an alternative design, the anhydrous HC1 is mixed with the hydrogenated hydrocarbons in pipeline 17 and the resulting mixture is supplied to the reaction zone 18.

The operating conditions inside reaction zone 18 were a temperature of 30 o C 5 o C, atmospheric pressure and. a contact time of 2 hours. The amount of added HCl was sufficient to keep the hydrogenated hydrocarbon feed substantially saturated under the conditions used.

Spent AlCl^ was periodically withdrawn and fresh AlCl^ was added

to ensure that the catalyst activity remained essentially constant.

The products from the Isomerization reaction zone 18 were removed

and led by means of pipe line 20 to a first fractionation column 21. The most volatile components which boiled below 1'50°C were taken as top fraction from column 21 via pipe line 22. This fraction constituted 1<;>+.1 parts by weight and can be sent to any application (due to its heat-producing value) or for further processing such as reforming or it can be fed to the removal for waste products. The bottom fraction in column 21

was led via rudder line 23 to another de scaffolding column 2h.

From the second distillation column 2^f, a fraction which boiled

in the range 150 - 165°C and which amounted to 155.5 parts by weight as top fraction quickly out via pipe 26. The heavy-boiling bottom fractions which amounted to 106.3 parts by weight were removed from the bottom of column 2k by means of pipe 25 and recycled, mixed with fresh feed to the hydrogenation reactor 12.

The top fraction from column 2h was led through the rudder line

26 to a dehydrogenation-isomerization zone 27. The remaining liquid from a crystallization step following the dehydrogenation-isomerization step was fed via recirculation pipe 28 to pipe 26 and mixed with the top fraction from column 2h. The amount of such remaining liquid amounted to 228.5 parts by weight.

Inside the dehydrogenation-isomerization zone 27, the hydrocarbon reactant mixture was brought into contact with a catalyst containing 0.5% by weight of platinum deposited on a silicon dioxide-aluminum oxide carrier. The operating conditions inside the dehydrogenation-isomerization zone 27

was temperature of <1>+25°C, pressure of 15 kg per cm^ and a space velocity for the liquid per hour of 6. A ratio between hydrogen and hydrocarbon of 9:1 was used inside the dehydrogenation-isomerization zone 27-

The products from the dehydrogenation-isomerization zone 27 were removed via pipeline 29 - and thereby added to a fractionation column 30. From the fractionation column 30, the lightest components were

(69.5 parts by weight) with a boiling point lower than 192°C taken out as top fraction via pipeline 31 and fed into pipeline 16, after which these components were recycled to the hydrogenation reactor 12. The bottom fraction (311+.3 parts by weight) from the fractional ionization column 30 was led by means of a rudder line 32 to a crystallization zone 33.

In the crystallization zone 33, the dura (85.8 parts by weight) was crystallized from the liquid. The dura was removed from the crystallization zone 33 by rudder line 3<!>+. The remaining mother liquor inside the crystallization zone 33 was removed via pipe line 28 as indicated above for recirculation to the dehydrogenation-isomerization zone 27.

Example 2

Another embodiment of the method according to the present invention was carried out by using the scheme indicated in fig. 2. According to this embodiment, a supply raw material containing approx. 80% by weight diethylbenzenes, and where the remaining part was approx. equal parts with respect to lower-boiling and higher-boiling aromatics, were added by means of conduit 101 to admixture with a recirculation stream (specified more fully in the following part of the example) from conduit 102, and the resulting mixture was rapidly introduced into the hydrogenation reactor 103. Inside the hydrogenation reactor 103, the mixture of feed and recirculate treated with hydrogen was supplied to the reactor 103 via pipe line 10k. The hydrogen used included fresh hydrogen as well as recycled hydrogen supplied to pipeline 10lf by pipeline 105 from separation tank 106, as well as recycled hydrogen supplied by pipeline 108. The catalyst and operating conditions used in the hydrogenation reactor 103 were essentially the same as those in

.used in the same steps of Example 1 given above.

The products from the hydrogenation reactor 103 were drawn out through pipeline 107 and were sent via pipeline 109 to a separation tank 106. Unreacted hydrogen is separated and recycled using pipeline 105 to supply pipeline 10^- for hydrogen. The remaining liquid product was led via rudder line 110 and was separated into 2 parts. One part was led to the recirculation tube line 102 for renewed supply to the hydrogenation reactor 103. The remaining part was led to a first distillation column 111.

From the scaffold column 111, the compounds boiling below 150°C as top fraction were quickly removed via pipe line 112 for further processing as such to a refarming unit. The heavier bottom fraction was led from distillation column 111 by means of pipe line 113 to another distillation column 114-. Together with the bottom fraction extracted from column 111, a recirculate stream 115 from the isomerization step according to the present method was supplied.

A fraction which boiled within the range 150 - 130°C was removed as the top fraction from the stilling column 114- by means of pipe line 117 through which it was led to the isomerization reactor 116. The compounds boiling above 180°C were removed from the distillation column 114- by of rudder cable 11.8.

Inside the isomerization reactor 116, the catalyst and operating conditions were mainly as described in the same steps according to example 1, and HC1 was supplied to the isomerization reactor 116 via pipe line 119.

The products from the isomerization reactor were removed via pipe line 120 and quickly to a stripping column 121. HC1 was removed as an overhead fraction from the stripping column 121 and recycled by means of pipe line 122 in admixture with fresh HC1 in pipe line 119 for supply to the isomerization reactor 116. From stripping column 121, stripped isomerization product by means of pipe line 123 quickly to a washing tank 12<*>+ to which soda was added by means of pipe line 125 °g of water was added by means of pipe line 126.

In the new washing tank 12<>>+, the isomerised product is first washed with soda and then washed with water. The washing liquids were quickly sent from the washing tank 124 by means of pipe line 127 to waste removal or other treatment. The washed isomerization product was led from washing tank 124 by means of pipe line 128 through which it was supplied to distillation column 129.

From the distillation column 129, compounds boiling below 150°C were removed as top fraction by means of pipe line 130

and quickly to further treatment such as reform!ag or for removal. The remaining bottom fraction was removed from distillation column 129 by means of pipe line 131 through which it was charged to another distillation column 132. From distillation column 132 a fraction with a boiling point within the range 150 - 16 5°C was taken out as top fraction through pipe line 133 °g thereby fed into a dehydrogenation-i-somerization reactor 13<!>+. Bottom "raks ion from distillation column 132 was removed and recycled via pipe 115 to pipe 113 and then to distillation column 11:+.

Together with the top fraction from the distillation column 132, the mother liquor from the crystallization zone 135 was supplied by means of pipe line 136 and pipe line 133 to reactor 13<4>-, The catalyst and the operating conditions of reactor 13<4>- were essentially the same as the conditions used in the same step according to example 1. The products from reactor 134 were sent via pipe line 137 to a separation tank I38 to recover hydrogen, part of which was recycled through pipe line 139 to reactor 131 and the remaining part was recycled via pipe line 108 to pipe line 104- and the hydrogenation reactor 103.

The liquid products from the separation tank 138 flow via pipe line 14-0 to a distillation column 14-1. In the distillation column 14-1, the fraction was separated into a lighter fraction which boiled below 192°C and which passed quickly via pipe line 14-2 to pipe line 101 for recirculation to the hydrogenation reaction step according to the present method. The heavier fraction or the bottom fraction was sent from distillation column 14-1 to distillation column 14-4 by means of pipe line 14-3. From this distillation column 14-4-, a heavier fraction that boiled above 210°C quickly came out via pipe line 14-5. This fraction can be further processed as desired. The top fraction from column 14-4 is led via pipe line 14-6 as a fraction with a boiling point in the range 192 - 210°C and fed to a crystallization zone 13? in which the dure was crystallized. The dura was removed via pipe line 14-7 while the mother liquor was recycled by means of pipe line 136 to pipe line 133 to be re-fed to the reactor Ijh.

Oak sample 3

Three isomerization experiments with batch operation (second step of the process) were carried out, each time using a total feed of 100 parts by weight of diethylcyclohexanes together with 10 parts by weight of A1C1-.

(10% by weight of the feed) and a reaction temperature of o 30 OC. All experiments were carried out under anhydrous conditions. The feed was a dry mixture of diethylcyclohexanes containing less than 10 parts by weight of aromatics per million.

First, the feed was saturated at 0°C with dry HCl and it was then fed into the batch reaction chamber and reacted for 6 hours. In the second experiment, the feed was led into the dry reaction chamber containing the AlCly catalyst and 0.1 wt.

(of the feed) tert-butyl chloride. In the third trial, the feed at 0°C was saturated with dry HOI and then passed through the reactor

containing AlCl₂ catalyst and 0.1% by weight (of the feed) tert-butyl chloride. The results obtained are indicated in the following table II:

It appears from Table II" that the presence of both HC1 and an alkyl halide is desirable in a batch reaction to achieve rapid conversion to tetramethylcyclohexane product. Similar results were obtained by replacing the diethylcyclohexane with l-methyl-M-isopropylcyclohexane.

Example U -

Subsequently, 5 isomerization trials were carried out in the batch reactor used in example 3, using a different promoter system. In each of the experiments, 100 parts by weight of dry diethylcyclohexane containing less than 10 parts by weight of aromatics per million saturated with dry HC1 at ambient temperature (22-23°C) led to the reactor containing 10 parts by weight of AlCl^ (10% by weight calculated on the basis of the feed) together with 0.1 part by weight of an alkyl halide or an olefin promoter for the AlCl^ catalyst (0 .1% by weight calculated on the basis of the weight of the fodder). In the last experiment, no promoter was present, but instead 0.1% by weight of diethylbenzene was placed in the reactor. The results are indicated in the following table III:

It appears from Table III that the best results were obtained with 2,2,<4>--trimethylpentene. The effects of tert-butyl chloride and tert-butyl bromide were comparable. The effect of 3,3,5-trimethyl-eyelohexene is quite significant compared to AlCl^ without promoter. The presence of diethylbenzene is found to suppress the Isomerization activity under the conditions used.

Example *?

Four isomerization reaction experiments were carried out by using batch operation, and by using a total feed of 100 parts by weight of diethylcylohexane together with 10 parts by weight of A1C1-, (10% by weight of the feed) and a reaction temperature of o 30 oC. All experiments were carried out under difficult conditions. The feed contained less than 10 parts of aromatics per million.

In the first trial, the feed was saturated at 0 QC with dry HCl and it was then fed into the batch reactor and maintained there for 6 hours. In the second experiment, 5 parts by weight became SbCl^

(5 wt. of the feed) was added to the AlCl^ catalyst. All other conditions were identical to the conditions used in the first experiment. In a third experiment, the feed was fed into the dry reaction chamber containing 10 wt% (of the feed) AlCl^ and 5 wt#

(of the feed) SbCl^ and then 0.1 voct$ (of the feed) tert-butyl chloride. In a fourth attempt, the feed was saturated at 0°C with dry HC1 and then quickly into the reactor containing 10% by weight Alcl3°S

5" wt% SbCl^ and 0.1 wt% tert-butyl chloride. The results are given

in Table IV:

It appears from Table IV that a decrease in the isomerization activity is achieved by the addition of 5% by weight (calculated on the basis of the feed) of SbCl^. Both hydrogen chloride saturation and the presence of liquid promoter have a beneficial effect on the isomerization activity.

Example 6

A continuous liquid phase isomerization was carried out according to the method according to the invention. At the start, 100 parts by weight of diethylcyclohexanes per hour containing no more than 10 parts by weight aromatics per million added to a continuous reactor with recirculation. 2 parts by weight per hour of dry HC1 were added to the reactor. The reactor originally contained 10 parts by weight of A1C1-, with regard to the amount of hydrocarbons present in the reaction zone. The temperature was kept at !+0 C

for 2<!>+ hours, and then the temperature was lowered to 30°C. The contact time for hydrocarbons with the catalyst was 6 hours. Under these conditions, a practically constant activity in the vicinity of ^.0% yield of tetramethylcyclohexane was obtained. After 4-8 hours of operation, 19 parts by weight of AlCl^ were added to the reactor (12% by weight of the amount of hydrocarbons present in the reactor zone). This resulted in a clear decrease in the isomerization activity (a decrease of up to h3% conversion). The AlCl^ catalyst had been converted to liquid after approx.

99 hours of operating time. Subsequently, approx. 1% by weight (calculated on the basis of the amount of hydrocarbons fed to the reactor) of catalyst complex extracted from the reactor and a corresponding amount of 1% by weight fresh AlCl^ catalyst was fed to the reactor every 6 hours. This procedure resulted in a constant yield of tetramethylcyclohexanes of approx. 38 - 4-0% by using only HC1 as promoter. For continuous liquid phase isomerization, the presence of HC1 alone as promoter thus proves to be satisfactory.

Example 7

A continuous liquid phase isomerization was carried out according to the method according to the invention. Diethylcyclohexanes (100 parts) were passed through 2 reactors 1 series each containing 20% by weight active catalyst complex, calculated on the basis of hydrocarbons present in the reactor. A stream of dry HC1 which amounted to 2 parts by weight per hour was also led in series through the 2 reactors. The temperature in the 2 reactors was kept at 30°C. The contact time in each reactor was 2 hours. The extraction of catalyst complex and the additions of fresh AlCl3 catalyst to each reactor amounted to 1.0% by weight, calculated in relation to the hydrocarbon supply to the reactors. At the outlet of the first reactor, 33% gem.-tetramethylcyclohexanes were obtained. At the outlet of the second reactor, the percentage value of average water was 4-4%. The same results were obtained with a feed consisting of a bottom product after removal of the gem.-tetramethylcyclohexanes by fractionation.

Example 8

A continuous liquid phase isomerization was carried out using the liquid active AlCl₂ complex. One hundred parts by weight per hour of triethylcyclohexanes containing not more than 4-0 parts by weight per million of aromatics was added to a continuous reactor with stirring. 2 parts by weight were added to the reactor. hour dry HC1. The reactor contained 4-0 parts by weight AlCl^ as active, complex,

in relation to the hydrocarbons present in the reactor. The temperature was kept at 20°C and the contact time was 6 hours. The reaction product contained gem.-hexamethylcyclohexanes.

Example 9

An experiment with dehydrogenation-isomerization (3rd stage of the process) was carried out using a system as illustrated in fig.2.

100 parts by weight of a naphthenic feed stream containing gem.-polymethylcyclohexane and with a composition as indicated in the following table V was supplied via pipeline 133 °g mixed with 186 parts by weight of filtrate from crystallization zone 13 5. The composition of the filtrate is shown in table V.

75.6 parts by weight of hydrogen were additionally added to pipeline 133

and the resulting mixture was fed into reactor 13<4>-, the Reactor

contained 31.9 parts by weight of a catalyst comprising 0.8% by weight of platinum on a silicon dioxide-alumina support containing 30% by weight of silicon dioxide<y>d. The reactor operated at 425°C and a pressure of 10.5 kg per cm<2>. The feed was passed through the reactor "at a volume rate pS -'H- volume-vK-ler per hour per volume part of catalyst. The exit stream from the reactor was passed through the separator 138 and the distillation columns 1<4>-1 and 14-4- , and the product obtained from distillation column 14-4- was passed through a crystallizer in separation zone 135 via pipe line 14-6.

Example 10

A further dehydrogenation-isomerisation experiment was carried out, using a single reactor system. 100 parts by weight of a naphthenic feed stream containing gem.-polymethylcyclohexane, and which had the composition indicated in the following table VI, was mixed with 125 parts by weight of diethylbenzene and fed to the reactor I34-. 29.6 parts by weight of hydrogen were additionally passed to the reactor 13<4>-, which contained 22.6 parts by weight of a catalyst comprising 0.8% by weight of platinum on a molecular sieve NaY. The reactor operated at +25 o C and a pressure of 14 kg per cm 2. The feed was passed through the reactor at a space velocity of 6 vo^am parts per hour per volume fraction catalyst. The reaction product contained 57 parts by weight of tetramethylbenzene and 18 parts by weight of unreacted products which can be recycled to the reactor. The composition of the tetramethylbenzene was as follows:

Claims (16)

1. Process for the production of polymethylbenzenes containing at least three methyl groups, from alkylaromatic hydrocarbons of general formula: where R 1 is an alkyl radical containing 2-6 carbon atoms, R 2 is an alkyl radical containing 1-3 carbon atoms, R is a methyl radical, n is a factor from 1 - 3 and m is 0 or 1, characterized in that a) they alkylaromatic hydrocarbons are hydrogenated to the corresponding alkylcycloalkanes, the percentage of unreacted hydrocarbons not being greater than 0.001, b) the alkylcycloalkanes are isomerized in the presence of AlCl^ at a temperature not higher than 150°C for a time sufficient to form a reaction product containing mainly gem- structured polymethylcyclohexanes with the same number of carbon atoms as in the starting material, c) the reaction product is brought into contact with an acidic, difunctional dehydrogenation-isomerization catalyst in the presence of hydrogen in a reaction zone maintained at a temperature within the range of 200° - 600°C and a pressure not greater than 50 kg/cm<2>, and d) the reaction product containing polymethylbenzenes with the same number of carbon atoms as in the starting material recovered. 2. Method according to claim 1, characterized in that alkylaromatic hydrocarbons are used where R<1> is an alkyl radical with 2-3 carbon atoms and R 2 is an alkyl radical with 2-3 carbon atoms. 3- Method according to claims 1-2, characterized in that diethylbenzene, cymene or triethylbenzene is used as alkylaromatic hydrocarbon. 4. Method according to claims 1-3, characterized in that the hydrogenation is carried out in the presence of a catalyst containing a metal from group VIII in the periodic table, which metal is deposited on a support selected from silicon dioxide, aluminum oxide, silicon dioxide-aluminium oxide, silicon dioxide- or aluminum oxide-containing zeolitic compounds and mixtures thereof. 5. Method according to claim k, characterized in that the hydrogenation is carried out in the presence of a catalyst containing a metal from group VIII, chosen from nickel, platinum, palladium, ruthenium, osmium, rhodium, mixtures thereof and alloys and mixtures with rhenium. 6. Method according to claims 1-5, characterized in that the hydrogenation reaction is carried out at a temperature in the range 150 - 350°C. 7. Method according to claims 1-6, characterized in that the hydrogen used in the hydrogenation reaction is used in a molar ratio between hydrogen and hydrocarbons of 2 - 30:1. 8. Method according to claims 1-7" characterized in that the hydrogenation reaction is carried out at a pressure in the range of 20 - 70 kg per cm 2. 9. Method according to claims 1 - 8> characterized in that the hydrogenation reaction is carried out at a space velocity for the liquid per hour of 0.5 - 50. 10. Method according to claim 1, characterized in that the isomerization reaction is carried out in the presence of a promoter selected from gaseous HC1, olefins, chlorinated hydrocarbons and mixtures thereof. 11. Method according to claim 1, characterized in that the dehydrogenation-isomerisation reaction uses a catalyst containing a metal selected from platinum and alloys of platinum with rhenium deposited on a carrier selected from silicon dioxide, aluminum oxide, silicon dioxide-aluminium oxide, silicon dioxide- or aluminum oxide-containing zeolitic compounds and mixtures thereof. 12. Method according to claim 11, characterized in that a catalyst containing 0.1 - 3% by weight platinum on a support is used. 13' Method according to claims 1 and 11 - 12, characterized in that the space velocity for the liquid per hour in the combined dehydrogenation-isomerization reaction is in the range 0.5 - 30. l4. Method according to claims 1 and 11 - 13, characterized in that aromatic compounds are added to the reaction zone, and the molar ratio between added material and aromatics is maintained in the range 0.1:1 to 4:1. 15- Method according to claims 1 and 11 - 14, characterized in that the molar ratio between hydrogen and hydrocarbons is maintained at a value in the reaction zone within the range 3:1 to 40:1. 16. Method according to claim 15, characterized in that the molar ratio is maintained between 6:1 and 25:1.