FI120197B - Process for dimerization of olefins - Google Patents

Description

BACKGROUND OF OLEFHNI DIMERATION

The present invention relates to the selective production of Ce-Cu hydrocarbons useful as components of vehicle fuels. In particular, the present invention relates to a process S for dimerization of lower olefinic hydrocarbons in the presence of a catalyst under conditions which permit selective dimerization.

Light olefin dimers are useful media for the preparation of various products such as alcohols, ketones and carboxylic acids. Multi-branched trimethylolefins and 10-trimethyl paraffins are useful in increasing the octane number of gasoline.

Several methods are known for the use of light olefins in the production of high quality vehicle fuels. The MOGD process (Mobil Olefin to Gasoline and Distillate) converts propylene and butylene into olefinic distillates in high yields. The MOG-15 method (Mobil Olefins to Gasoline) is an extension of the MOGD method. In MOG, reaction conditions allow the formation of aromatic hydrocarbons. The oligomerization of isobutene from C4 olefins with zeolite catalyst has been the subject of several US patents.

An essential feature of selective dimerization of light olefins is to prevent oligomerization, which is a reaction following dimerization. Selective dimerization of isobutene by trimerization and higher oligomerization is known from U.S. Patents 3,325,465, DE 3,542,171 and International Patent Application WO 01/46095. US 3,325,465 discloses the use of nickel and cobalt ions in 13Xzeolite. In DE 3,542,171, the selective dimerization of isobutene to trimethylpentene is carried out by means of a bismuth or lead-lubricated zeolite. WO 01/46095 states that the high porous zeolite Beta is selective for dimerization of isobutene to trimethylpentene in the presence of oxygenates. It is further clarified that dimerization can be carried out in the presence of alcohol.

A few commercial dimerization methods are in use. The Institut Franfais du Petrole (1FP) has developed a dimerization method for light olefins (Dimersol). The Octol method, 2 developed by UOP and Huls AG, produces linear octenes, which are the raw materials for plasticizers. The processes employ homogeneous catalysts.

Several methods are based on the use of ion exchange resins as a dimerization catalyst. Such technical solutions are described, for example, in U.S. Patent Specifications 4,375,576 and 4,100,220. These methods have many good properties, but all have the drawback that they are completely dependent on an oxygenate retardant for improved selectivity. This retarder must be recycled And the dimerized product generally has a significant amount of oxygen-containing by-products. They are highly undesirable since dimerization is used to remove oxygenates from the components. In addition, oxygenation makes it difficult to hydrogenate the dimerized product.

All problems with oxygenates can be resolved if these compounds are completely removed from the process. Therefore, we have described in previous patent application WO 2004/08093 5 a dimerization process in which an olefinic raw material comprising C 3 -C 10 iso-olefins is contacted with a mid-porous zeolite to dimerize the iso-olefins to Q-C 10 dimers. In this process, dimerization is carried out in the presence of a heterogeneous zeolite catalyst. The catalyst may be any zeolite which is active in dimerization reactions. Such zeolites include, for example, natural and synthetic mid-pore zeolites such as ZSM-5, ferrerite, ZXM-22, and ZSM-23. These and similar catalysts have a pore size in the range of 2-8 A (0.2-0.8 nm), preferably 4-6 A (0.4-0.6 nm), are active and selective for trimethylolefins.

It is an object of the present invention to provide an alternative process for the preparation of dimers which are useful, for example, as components of vehicle fuels, particularly gasoline, and as octane indexers.

The invention is based on the idea of dimerizing iso-olefins and optionally linear olefins in C3-C6 * olefinic feed in liquid form in a reaction chain comprising at least one reaction zone and at least one separation region. The reaction is carried out under conditions in which at least part of the C4-C6 iso-olefins are dimerized. The separation area is arranged after the 3 reaction areas, and the circulating flow can be circulated from the separation area back to the dimerization. The process is carried out in principle in the absence of polar compounds.

According to the invention, the dimerization is carried out by contacting the raw iso-olefins with an acidic mesoporous molecular sieve catalyst. In particular, the present invention utilizes an extractor active catalyst material having a mesoporous molecular sieve, for example, implanted in a zeolite structure.

In addition, according to a preferred embodiment, the dimerization is carried out in principle in the absence of butadiene, or more generally in the absence of hydrocarbons containing polyunsaturated hydrocarbons and water in the raw material.

The invention also enables the use of a catalyst material having an acidic mesoporous molecular sieve as an acid catalyst, such as a mesoporous 1S molecular sieve, which is embedded in a zeolite structure to provide a dimerized olefinic feed containing unsaturated hydrocarbons selected from the group consisting of 2-butene, linear C5-olefins and branched C5-olefins

The method according to the present invention is further characterized in the characterizing parts of claims 20 and 20, in particular.

The present invention achieves considerable advantages. In conventional methods, selective dimerization is a more desirable but difficult to achieve objective. Using the process of the present invention, high selectivity can be achieved along with high conversion rates.

A particular advantage of the present invention is that the process can be used for very long periods of time and with extended maintenance intervals, which reduces operating costs and makes production more efficient than those previously employed.

30 4

In conventional techniques, catalyst regeneration often threatens the continuous use of the process. In the present invention, the catalyst can be continuously regenerated during the use of the process. Easy regeneration makes it possible to handle feeds containing nitrogen and sulfur impurities, as well as conventional amounts of polyunsaturated hydrocarbons and water, which is a major advantage. By eliminating unsaturated hydrocarbons containing at least two double bonds or at least one triple bond, the regeneration interval can be significantly extended, even in industrial production, resulting in significant cost savings.

10 When using FCC feeds, which generally contain high levels of linear olefins, with the traditional acid catalyst or ion exchange resins, increasing concentrations of oxygenates are obtained. Hydrogenation of such compounds is difficult, if not impossible. By using catalysts such as the present one, formation of oxygenates can be avoided and the time intervals of the operation can be extended.

15

The invention will now be further explored using the following detailed description with reference to the accompanying drawings.

Figure 1 schematically illustrates the configuration of a basic technical solution of the invention; Figure 2 graphically illustrates the effect of feed composition on catalyst activity based on the results of Example 2 (Experiments 1 and 2); Figure 3 is a graph showing the effect of butadiene on catalyst activity based on the results of Example 3 (Experiments 3 and 4); Figure 4 graphically shows the conversion rates of butylenes as a function of time in Experiments 5 and 6 of Example 25 (a); Figure 5 illustrates the effect of feed dehydration on the amount of catalyst deactivation based on the results of Experiments 5 and 7 of Example 4 (b); and Figure 6 is a graph showing the activity of the catalyst after oxidative regeneration based on Experiments 7 and 9 of Example 5.

30 5 Definitions

For the purpose of the present invention, the '' reaction zone '' means at least one, typically two or three reactors. The reactor may be any continuous reactor on which a solid catalyst may be placed and which may handle liquid reagents. It is useful if the reactor is a simple tubular reactor, a fixed bed reactor or a fluidized bed reactor. The reactor may be a tubular reactor having a plurality of tubes filled with catalyst. Other possibilities include a reactive distillation unit with side reactors. The operating pressure of the reactors will depend on the type of reactor and the composition of the feed, usually attempting to keep the reaction mixture in the liquid phase. In order for the catalyst to be regenerated during operation of the reactor, it is often advantageous to use at least two reactors which can be regenerated alternately. Another useful use is to operate a reactor in which the catalyst can be continuously regenerated.

15 "Separation Range" means a separation system comprising, in one embodiment, a distillation system comprising one or more distillation columns. The feed plate can be selected for each column to be as useful as possible for the overall process. The distillation column may be any column suitable for distillation, such as a packed column or a column fitted with a valve, a sieve or a clock base.

20 '' Isooctene 'and' diisobutene 'are both dimerization products of isobutene. Both of them can thus be used to designate 2,4,4-trimethyl-2-pentene and 2,4,4-trimethyl-2-pentene or mixtures thereof.

The "hydrogenated paraffinic hydrogenated paraffinic compounds" comprising "iso-octane" and "diisobutene" include a dimerization reaction in the reaction area of the desired product. When only C3 olefins, only C1 olefins or only C5 olefins are added to the process, it is clear that the product obtained from the reaction of the olefins produces dimers. When C3, C4 and Cs olefins are present in the feed, in addition to dimerization, reactions may also occur between different olefins.

The word "dimer" is also used in the specification of reaction products for simplicity, but it is to be understood that when various lower olefins are present in the feed, the reaction mixture will generally also contain some Cg olefins.

6

According to a first preferred embodiment of the invention, a hydrocarbon feed containing iso-olefins such as isobutene or linear olefins such as butanes or mixtures thereof is contacted with an acid catalyst in a reaction system substantially free of oxygenates and comprising at least one reaction zone and . The conditions are such that at least part of the iso-olefin is dimerized to iso-octene. The flow from said reaction zone is directed to a separation zone where most of the 10 dimerized reaction product is separated from the unreacted product. It is useful if at least part of the unreacted product is recycled from the separation area to be dimerized.

It is important that the hydrocarbon feed is substantially free of polyunsaturated hydrocarbons such as 1,2-and 1,3-butadiene, propylene (acetylene) and isoprene, i.e. compounds containing at least two optionally linked double bonds or at least one triple bond as opposed to the desired components of the feed, containing a single double bond (iso- or 1-olefins). The term '' substantially free of polyunsaturated hydrocarbons '' refers to levels well below the typical concentration levels of such compounds, for example, in traditional industrial hydrocarbon feedstocks (refinery stream) 20 such as the FCC. These refinery streams contain 0.3 wt% or more but up to 1 wt% butadiene. In the present invention, up to 0.1% by weight or even less is desirable, i.e., less than 500-1000 ppm, particularly less than 100 ppm or even less than 50 ppm. In addition, the dimerization feed should be substantially anhydrous. The term "essentially anhydrous" refers to a concentration of 100 ppm or less. More preferably, the content is less than 30 ppm, and preferably 25 is 10 ppm or less when the raw material is "substantially anhydrous". As discussed in more detail in Examples 2 and 5, both polyunsaturated compounds and water have a strongly detrimental effect on catalyst activity.

In order to remove the polyunsaturated hydrocarbons from the feed, it is preferably subjected to selective hydrogenation which can be carried out in a manner known in the art, for example by contacting the feed with a suitable catalyst, eg metal such as cobalt or aluminum-supported nickel or noble metal (eg palladium or ruthenium) acid. Most of the available technologies can reduce the butadiene content to below 100 pp. Dehydration can be accomplished using activated aluminum and molecular sieves as is known in the art.

According to one embodiment of the invention, the hydrocarbon feed containing olefins is contacted with a particular acid catalyst, namely a catalyst comprising a mesoporous molecular sieve implanted in a zeolite and wherein the catalytic material 10 is thermally stable at a temperature of at least 900 ° C. olefins dimerize. A description of such a structure and the preparation of a particularly suitable catalyst is given in patent application FI20041675, the contents of which are incorporated herein by reference. If the olefin feed comprises C 3 -C 6 olefins, there are also reactions between different olefins to form higher (up to C 12) olefins. The reaction area effluent is fed to a separation area where the majority of the dimerized reaction product is separated to form a first product containing unreacted hydrocarbons and a second product containing dimerized olefins.

The feed of the process of the present invention is a hydrocarbon mixture containing olefins. The feed comprises olefins dimerized at a concentration of at least 10% by weight, preferably at least 20% by weight, based on the total weight of the raw material. As stated above, the olefins are propylene, linear 1-or 2-butene, isobutene, and linear or branched C5-olefins Alternatively, the feed comprises dimerizable components: either C4-olefins, preferably isobutene, for the production of iso-octene, or C5-olefins. to produce substituted C 10 olefins. It is clear that the feed may contain both C4 and C5 olefins to produce a variety of different products. The composition of the product stream will be discussed later. Generally, the feed may contain olefins between C3-C6 olefins, although C4 and C1 olefins, especially C4 and C5 iso-olefins, are of most interest.

According to a first preferred embodiment wherein the C4 hydrocarbons are dimerized, the feed hydrocarbon mixture comprises at least 10% by weight, preferably at least about 15% by weight, of 8 isobutenes. The feed may consist of pure isobutene, but in practice the ready-to-use raw material comprises C4-based hydrocarbon fractions from oil refining. Preferably, the feed comprises a fraction obtained from dehydrogenation of isobutane, whereby the feed consists mainly of isobutene and isobutane and optionally small amounts of C 3 and C 5 hydrocarbons. The feed thus comprises generally 40-60% by weight of isobutane and 60-40% by weight of isobutane, and isobutene is generally 5-20% by weight less than isobutane. Thus, the ratio of isobutene to isobutane is approximately 4: 6-5: 5.5. An example of a dehydrogenation fraction of isobutane is: 45% by weight of isobutene, 50% by weight of isobutane and other unreacted C4 hydrocarbons, and a total of about 5% by weight of C3, Cs and heavier hydrocarbons.

10

The feed required for the preparation of iso-octene may also be the C4 fraction or the C4 fraction of FCC, TCC, DCC and RCC after removal of butadiene, also known as the ethylene unit Raffinate 1. Most preferred are FCC, RCC, TCC and Raffinate 1 , since the hydrocarbon fractions can be used as such, possibly after removal of the heavier (Cg +) fractions. The composition of raffinate 1 is generally as follows: about 50% by weight of isobutene, about 25% by weight of linear butenes and about 25% by weight of paraffins. The composition of the product obtained from the FCC is typically as follows: 10-50% by weight, preferably 10-30% by weight, isobutene, 20-70% by weight of 1- and 2-butene and about 5-40% by weight of butane. As an example of a typical FCC blend, about 17% by weight of isobutene, about 17% by weight of 1-butene, about 33% by weight of 2-butene and about 33% by weight of butane, and others.

It is also possible to use isobutene made from chemicals as feed.

According to another preferred embodiment of the invention, the olefins in the olefinic feed are linear and branched C 5-olefins such as linear pentene, 2-methyl-1-butene, 2-methyl-2-butene, 3-methyl-1-butene and mixtures thereof.

According to one embodiment of the invention, the raw material comprises aromatic hydrocarbons, paraffins and mixtures thereof.

30 9

If the present invention is used for the conversion of linear butenes, the linear butenes are preferably selectively isomerized to 2-butene as completely as possible. In this case, it is preferable to add a separate side reactor cycle to the process configuration. The temperature of this reactor is preferably higher than in the pre-reactor or in the circulating reactor to increase the dimerization conversion.

FCC and similar hydrocarbon streams are suitable for use, for example, in situations where a conventional MTBE unit is used in a mixture of iso-octene and MTBE. manufacture.

10

According to another preferred embodiment of the invention, wherein the C 1 olefins are dimerized, the feed comprises olefins which are linear and branched C 1 olefins or mixtures thereof. Thus, the olefins generally present in the feed include linear pentene, 2-methyl-1-butene, 2-methyl-2-butene, 3-methyl-1-butene. The feed may also contain some amounts of Ce-15 olefins, generally at least 5% by weight.

Typically, in another preferred embodiment, the feed is FCC gasoline, light FCC gasoline, pyrolysis C 1 gasoline, TCC gasoline, RCC gasoline, and coke gasoline, generally the C 1 fraction of FCC gasoline. The feed may also comprise some Ce olefins. It is useful to fractionate the 20 FCC fractions in order to obtain as pure a C5-olefin fraction as possible, with less than 15 wt%, preferably less than 5 wt%, of other C1 hydrocarbons. It is possible to use a fraction which also comprises Ce-olefins. It is typical that the feed then comprises 20-60% by weight, in particular 30-50% by weight, C 1- olefins, 10-30% by weight, in particular 15-25% by weight, Ce-olefins and 15% by weight, or less parafunic hydrocarbon pentanes.

25

According to a third preferred embodiment, the feed comprises both C 4 and C 1 olefins. In this case, the feed is generally the FCC, TCC, DCC and RCC or C4 fraction after removal of butadiene, also called ethylene unit raffinate 1, FCC gasoline, light FCC gasoline, pyrolysis C5 gasoline, TCC gasoline, RCC gasoline and coke gasoline . The ready-to-use available fraction comprises FCC C4 and C1 fractions It is useful to use a fraction comprising at least 10% by weight, preferably at least 15% by weight, C4-olefins and 10 at least 10% by weight, preferably at least 15% by weight -%, Cs olefins. Generally, the amounts of C4-olefins and C5-olefins are approximately equal, although it is also common for C4-olefins to be slightly more numerous.

According to a first preferred embodiment, the C4 olefins are dimerized. Feed compositions are already processed Product compositions are as follows:

When predominantly isobutene dimers are produced, they are generally present in the product stream in an amount of at least 85% by weight, preferably at least 90% by weight. Other components 10 generally present in the product stream include isobutene trimers, 15 wt% or less, preferably 10 wt% or less, isobutene tetramers, less than 0.2 wt%, and other hydrocarbons, less than 1 wt%, preferably less than 0, 1% by weight.

Despite the intended product composition, most of the dimers prepared by the process (65-100% by weight, generally 85-100% by weight, preferably 95-100% by weight) are 2,4,4-trimethylpentenes. When the product stream is hydrogenated, a mixture comprising isooctane is obtained. The fraction of other trimethylpentanes in the mixture (e.g. 2,3,4,4-trimethylpentane) and the fraction of dimethylhexanes in the mixture remain very small. Thus, the fuel component has a high octane number (RON), generally at least 95, preferably from about 20 to about 98.

In another preferred embodiment, dimers of C 1-olefins are produced. The product is generally as follows: 25% by weight, preferably at least 70% by weight, Cs dimers, 5-32% by weight, preferably 5-29% by weight, olefin trimers, less than 1% by weight, preferably less than 0, 5% by weight, olefin tetramers. Since no oxygenates are fed into the process, the amount of oxygenates present in the process and in the final product is very small. When the structure is hydrogenated, a structure that can be used as a fuel component is obtained

According to a third embodiment, dimers of both C4 and C5 olefins are produced. In addition, the C4 and C5 olefins also react to form C5> olefins. The product structure thus comprises at least 65% by weight, preferably at least 70% by weight, of Cs-dimers, C4-dimers and C9-olefins, 5-32% by weight, preferably 5-28.5% by weight, of olefin trimers, less than 1% by weight - 5%, preferably less than 0.5% by weight, of olefin tetramers. When the structure is hydrogenated, a structure is obtained which can be used as a fuel component.

Despite the product structure sought, most (50-100% by weight, generally 60-100% by weight, preferably 90-100% by weight) of the dimers and C5-olefins prepared by the process contain 10 iso-octene, tetramethylpentene and trimethylhexene. When the product stream is hydrogenated, a mixture of corresponding hydrogenated hydrocarbons is obtained. The relative amounts of the individual components will vary depending upon the ratio of the reactive C4 and Cs components in the feed. When the product stream is hydrogenated, a mixture of isooctane, tetramethylpentanes and trimethylhexanes is obtained. Thus, the fuel component has a high octane number (RON), typically at least 95, preferably about 98-100.

The dimer fraction of the reaction product for the feed comprising (among other less reactive compounds) both C4 and C1j iso-olefins (45:55) contains 20-30% by weight of trimethylpentenes, in particular 25-28% by weight, tetramethylpentenes and 20% by weight. trimethylhexenes 20-30% by weight, especially 20-25% by weight, tetramethylhexenes 4-8% by weight, especially 5-6% by weight, and trimethylhexenes 2-5% by weight, especially 3-4% by weight. The remainder of the dimer product is less branched olefins.

The product has a vapor pressure of 10-20 kPA (Reid) and a distillation point (90% by weight, ASTM D86) of 180 to 25 ° C or less.

Part of the first reaction product is generally recycled from the separation area back to the reaction area. It is to be understood that although the following description refers to one side stream, which is a typical configuration, it is also possible to draw two or more side streams and recycle all of them for dimerization.

12

According to a useful embodiment, the reaction zone comprises two parallel reactors. The feed comprising the fresh olefinic feed and the recycled first product may be fed to one reactor and the second reactor may be made to operate simultaneously. The reaction area effluent is diverted to a separation area where most of the dimerized reaction product is separated to form a first product containing unreacted hydrocarbons and a second product containing dimerized olefins.

The selectivity of the dimerization reaction in the process according to our invention is high. According to an embodiment, the selectivity of dimerized olefins, expressed as a ratio of the molar amount of dimeric compounds to the total molar amount of olefins converted, is greater than 0.8, in particular greater than 0.9.

According to the invention an acid catalyst is used. It has been found that prior art problems with zeolite catalysts and mesoporous catalysts can be avoided or at least significantly reduced by the use of new catalyst material, which is a mesoporous molecular sieve implanted in a zeolite and is highly mechanically, thermally and hydrothermally. The new mesoporous molecular sieve immersed in the zeolite is thermally stable at a temperature of at least 900 ° C in the presence of air.

Therefore, the present invention utilizes a number of novel mesoporous molecular sieves that are mechanically, thermally, and hydrothermally stable. These materials are highly reproducible, as can be seen from the examples, and have superior properties in many hydrocarbon conversion reactions, Mesopore refers to materials having pores in the range of 2 to 15 nm and having a regular pore system.

25

The mesoporous molecular sieve in which the zeolite has been implanted comprises a mesoporous molecular sieve selected from the M4 IS group comprising mesoporous materials having an ordered pore system. Preferably, the mesoporous molecular sieve is selected from mesoporous aluminosilicates known as MCM-41.

30 13

The mesoporous molecular sieve is implanted with a zeolite selected from middle pore zeolites, such as MFI, MTT, TON, AEF, MWW, and FER structures, and large porous zeolites, such as 12-membered zeolites, BEA, FAU and MOR structures. Examples of said zeolite groups 5 are ZSM-5, ZSM-23, ZMC-22, SAPO-11, MCM-22, fenierite, beta, Y and X zeolites and mordenite. Preferably, the zeolite is an MFI, MTT, AEF or BEA zeolite.

The catalyst material contains 0.01 to 10% by weight of aluminum (Al).

The catalyst, which is particularly suitable for industrial and commercial use, comprises, in accordance with the invention, a mesoporous molecular sieve loaded with zeolite and a carrier selected from alumina, silica, clay and any other prior art carrier and combinations thereof. Preferably the support comprises alumina or silica. The amount of carrier ranges from 10 to 90% by weight based on the total weight of the catalyst.

The novel group of catalytic materials having a mesoporous molecular sieve implanted with a zeolite structure according to the invention has a high specific surface area (BET) of 1400 to 500 m2 / g, preferably 1200 to 600 m2 / g.

20

The X-ray powder diffraction pattern of the catalytic material of the invention shows the structures of the mesoporous molecular sieve and the zeolite. The dimension of the zeolite unit cabinet varies according to the amount of Al in the catalytic material. The unit cabinet size decreases with aluminum content, from 1.982 nm for catalyst material containing 0.2 wt% Al, to 1.972 nm for catalyst material containing 3.9 wt% AI when the zeolite type was MFI

(material code is MM5). The change in unit cabinet size is the opposite of the changes generally observed with zeolites. Unit cabinet sizes were 1.428 nm and 1.430 nm for the zeolite type BEA (material code MMBE). The distance in the dioome microporous molecular sieve MCM-41 decreases with increasing zeolite content. In MM5 the diode 30 ranges from 4.38 nm to 3.81 nm (MM5). The diode distance ranges from 4.09 nm to 3.97 nm (MMBE). Unit cabinet sizes are the same for pure zeolite and MCM-41 14 phases and their mechanical mixtures. Changes in the distance between the diode and the unit cabinet size are a clear indication of the actual chemical bond between the mesoporous molecular sieve and the implanted zeolite in the catalytic material of the invention.

5

The characteristic features of the catalytic material of the invention, a mesoporous molecular sieve in which a zeolite has been implanted, were measured by X-ray powder diffraction, scanning electron microscopy, transmission electron microscopy, measurement of the specific surface using N

The total number of acidic sites can be measured as the capacity of the catalyst material to bind strongly basic molecules such as ammonia or pyridine. Total acidity was measured by ammonia temperature-programmed desorption (TPD) and Bronsted and Lewis-15 acidity by pyridine infrared spectroscopy (FTIR). The acidity of the catalyst material can be tailored to the amount of aluminum added in the structure and to varying the aluminum (Al) content in the zeolite, MCM-41 and MM phases.

In the absence of international standards for the determination of acidity, the methods used herein are described below.

The acidity assay was performed with NH3-TPD. The total acidity of the catalytic materials was measured by temperature-programmed desorption of ammonium (NH3-TPD) using an Altamira AMI-100 instrument. The sample size was 40 mg. Total acidity was measured by Nik desorption as a function of temperature. The acidity of the samples was calculated from the amount of ΝΗΐΐη adsorbed at 200 ° C and desorbed between 100 ° C and 500 ° C. The NH3 TPD instrument was equipped with a thermal conductivity detector (TCD) manufactured by GowMac. A rise rate of 20 ° C / min was used and the temperature was linearly increased to 500 ° C where it was held for 30 minutes.

Quantitation was performed using pulses of known volume of 10% NH3 in helium.

15

The acidity assay was also determined by pyridine FTIR. The acidity of the samples was measured by infrared spectroscopy (ATI Mattson FTIR.) Using pyridine (99.5% a.r.) as a test molecule for qualitative and quantitative determination of both Brensted and Lewis acid sites. The samples were compressed into thin self-supporting plates (10-12 mg / cm 2). Pyridine 5 was first adsorbed for 30 min at 100 ° C and then desorbed by evacuation at various temperatures (250,300 and 450 ° C) to obtain a distribution of the intensity of the acid sites. All spectra were recorded at 100 ° C with a spectral resolution of 2 cm * 1. Spectral bands at 1545 cm -1 and 1450 cm -1, respectively, were used to identify Bronsted (BAS) and Lewis acid sites (LAS). The amounts of BAS and LAS were calculated from the intensity of the respective spectral bands using molar extinction coefficients.

The acid sites are located on the surface of the catalyst material. Total surface area and pore volume were estimated using N2 adsorption and desorption. The average mesopore area and mesopore diameter were estimated from N2 desorption using the BJH equation (Barrer-Joyner-15 Halenda). The pore diameter has a size-limiting effect on both reactants and products. The size of the micropores depends on the structure of the zeolite. Pores less than 2 nm in diameter are defined by IUPAC as micropores, and pores in the range of 2 to 50 nm are defined as mesopores.

The area and total pore volume decrease when the zeolite is implanted into a mesoporous molecular sieve, as can be seen in Table 1 below, which shows the area, pore volume, and pore diameter values for MM5, MMBE, and for comparison, data for MCM-41, MFI and BE A.

25 30 16

Table 1. Area and porosity

Sample BET area BJH Total BJH

(m2 / g) pore area of meso-pore pore volume volume (m2 / g) (cm3 / g) (nm)

Na-MCM-41-20 949 947 0.829 1 £ ______ _ ii45 0.814 2 ^ 5

Na-MM5-4Z5 820 1009 0.713 2/6 _____ _ __ 0.794 I!

Na-MM5-4Z5-2A125 733 599 0.656 2Ä _____ _ __ 0.351 2Ö *

Na-MMBE-4B 879 884 0.692 2/7

Na-MMBE-4B-2Ä1 844 ~ 859 0/742 2/7 ________ __ _ 0.835 Z6 BEA 585 85 0.254 nothing. * = size of the intra-particle cavity

A process for preparing a mesoporous molecular sieve in which a zeolite has been implanted comprises the steps of: a) preparing zeolite cores from a silicon source and an aluminum source and a structural directing agent (template R), or a silicate or aluminosilicate precursor for the zeolite core; B) preparing a gel mixture of a mesoporous molecular sieve from a source of silicon; i. optionally, an aluminum source and a surfactant (S); ii. introducing the zeolite nuclei or precursor of silicate or aluminosilicate prepared in step a) into the mesoporous molecular sieve gel mixture obtained in step b) and homogenizing and dispersing the zeolite nucleus or precursor of silicate or aluminosilicate in the molecular sieve; c) performing the gel maturation of the mixture of step b); D) performing the hydrothermal synthesis of the mixture of step c) by maintaining the mixture under suitable conditions at a temperature ranging from about 100 ° C to about 200 ° C under static and dynamic mixing until crystals are formed; 5 (f) recovering the crystals; g) washing the solid product; h) drying the solid product; and i) partially or completely removing the surfactant (S) by a stepwise calcination process and optionally removing the template (R), unless removed in step a), to provide a catalyst which is a mesoporous molecular sieve in which zeolite has been inoculated.

In step a), zeolite cores are prepared from a source of silicon and an aluminum source and a structure controlling agent (Temple R). The source of silicon is selected from silica, preferably colloidal silica, solid silica and smoky silica.

The source of aluminum is selected from aluminum sulfates (Al 2 (SO 4) 3. 18H 2 O), hydrogenated aluminum, hydroxides of mines, aluminates, aluminum isoproxide and alumina.

A suitable temple is selected to provide the desired zeolite structure. Examples template is typically used are hydroxides al · kyyliammoniumhalogenidit, alkyyliamiinihydroksidi and alkyvliamiinihalogenidit as tetrapropyyliammoniumbromidi, tetrametyyliammonitmihydroksidi, tetrabutylammonium bromide, tetraethylammonium bromide, tetraethylammonium hydroxide, 25, piperidine, pyrrolidine, octylamine, ethylenediamine, 1,6-diaminohexane and hexamethylene · »imine.

The temperature in step b) is in the range of 40 ° C to 200 ° C and is prepared in static or dynamic mode. Finally, in step a), the temple is optionally removed by a heat treatment process known as a stepwise calcination procedure.

The treatment temperature is in the range of 350 to 900 ° C at removal. Alternatively, the temple 18 may be removed in step i) if not removed in step a), but preferably the temple is removed in step a).

In step b), a mesoporous molecular sieve gel is prepared from sources of silicon, optionally from sources of aluminum and a surfactant. The zeolite cores prepared in step a) or the silicate or aluminosilicate precursor in step c) are added to the molecular sieve gel with stirring. The resulting mixture is homogenized and the zeolite cores or the silicate or aluminosilicate precursor is dispersed. An additional source of aluminum may be added to adjust the acidity of the product. The mixing rate in step c) is in the range of 50-1000 rmp.

10 Processing time is in the range of 10 to 500 minutes.

In the following steps, the resulting gel is agitated, the hydrothermal synthesis is carried out at a temperature of 100-200 ° C, the crystals are recovered, for example by filtration, and the solid product thus isolated is washed extensively with water. The product 15 is then dried and the surfactant removed by heat treatment, e.g. calcination at 350-900 ° C. Further details of the manufacturing process are disclosed in patent application FI20041675.

A catalyst like the above can be used in the dimerization step. The catalyst is placed, for example, in a fixed bed reactor. The temperature of the reaction zone is generally 50 to 200 ° C, preferably 80 to 150 ° C. The upper limit of the temperature range is determined by the desire to avoid certain side reactions. WHSV is about 2 to 200 h'1.

The reaction area effluent is led to a separation area where the components are separated.

25 The structure of the product stream depends on the method parameters and the feed structure. As already discussed, the process of the invention can be used to make a dimerized product from an olefinic feedstock. The feed olefins may be either C 3 olefins, Crolefins, C 1 olefins or mixtures thereof. Thus, it is clear that the structure of the product stream is essentially dependent on the fraction used as the raw material.

30 19

According to a first preferred embodiment, the C4 olefins are dimerized. Feed structures are already processed Product structure is as follows:

The dimer fraction of the reaction product for the feed comprising (among other less reactive compounds) both C4 and C1 olefins (45:55), contains 20-30% by weight of trimethylpentenes, in particular 25-28% by weight, tetramethylpentenes and trimethylhexenes 20- 30% by weight, especially 20-25% by weight, tetramethylhexenes 4-8% by weight, especially 5-6% by weight, and trimethylheptenes 2-5% by weight, especially 3-4% by weight. The remainder of the dimer product is less branched olefins.

10

According to one embodiment, a portion of the first non-recyclable product is subjected to alkylation.

According to one embodiment, the second product is hydrogenated In order to obtain a partially or completely hydrogenated product. However, it must be borne in mind that even a non-hydrogenated product is a valid and even useful fuel component.

Figure 1 shows a simplified configuration of the dimerization process of the present invention.

20

According to the embodiment shown in the figure, the process comprises a reaction zone 1 and a separation zone 2. In addition, upstream of the reaction zone 1 is a selective hydrogenation zone 3 and a dehydration zone 4, preferably connected in series, as shown in Figure 1. The effluent of the dehydration zone 4 is led to a stripper 5 where light components such as hydrogen 25 are evaporated in the gas phase while the remaining liquid effluent is introduced into the reaction zone 1.

The dimerization product is formed in reaction zone 1, and the dimers are separated from the unreacted components in separation zone 2, as shown in the figure. The unreacted components 30 are recycled to stream 6. The unreacted components leave the process in stream 7, 20

The reaction zone comprises one or more reactors. Many continuous reactors capable of depositing a solid catalyst and a liquid reagent are suitable for the present invention. According to an embodiment of the invention, the reactor must allow regeneration of the catalyst. Regeneration can be performed during continuous production. Alternatively, 5 or more parallel reactors can be used. One reactor can be regenerated while another is used.

A typical dimerization system consists of one or more reaction areas followed by product separation, and the inert processing of non-reactive materials. Multiple reaction and product 10 separation steps can be combined with Saija if the conversion requirement is high.

The reaction zone comprises any reactor type suitable for use in a liquid phase, and in which a solid catalyst can be used. Such reactor types include a fixed bed reactor, a moving bed reactor, a stirred tank reactor, a fluidized bed reactor, or a jet bed reactor or a combination of these reactors.

In order to meet the requirements of continuous production, the dimerization catalyst must be regenerated regularly. Often the reactor system must also include space for catalyst regeneration. If continuous use is not necessary, it is of course possible to suspend production to regenerate the catalyst. However, in industrial production, it is preferable to use multiple reactors which can be regenerated one by one while others are used. An example of such a sequester according to the present invention is the combination of two or more fixed bed reactors so that each can be separated from production for catalyst replacement or regeneration.

25

Another popular alternative for dimerization production is the use of a reactor from which the catalyst can be continuously extracted and recycled through a regeneration site. In one embodiment of the invention, a fluidized bed reactor is used or the catalyst from the jet bed reactor run can be continuously removed and recycled through regeneration.

21

According to a preferred embodiment of the invention, the separation area comprises a distillation column. The reaction zone product reference comprises light hydrocarbons remaining in the hydrocarbon feed and oligomers formed in the reactor having a boiling point substantially higher than the boiling point of the feed. This makes separation by distillation easy.

5

According to one embodiment of our invention, the separation region is preferably a distillation region. The reactants are monomers and the product is a mixture of oligomers, so they have significantly different boiling points, which makes separation by distillation easy. For ease of separation, the expansion evaporation separator tank, evaporator, stripper or fractionator and other distillation devices known in the art can be used.

In another preferred embodiment of the invention, the reaction zone comprises two parallel reactors which are operated alternately. This means that when one reactor is regenerated, the second reactor is used for dimerization. The separation area comprises a distillation column.

1S Part of the feed is removed from the process and another part is directed back to the separation area to increase the yield of the reaction area.

The reactor may be a fluidized bed reactor and the catalyst is continuously regenerated in a regeneration unit.

20

As discussed above, the step of bringing the C4-C6 iso-olefins into contact with the catalyst material under conditions leading to dimerization is preferably carried out in such a manner that the raw material is substantially free of butadiene and water. Thus, the raw material should contain less than 1000 ppm butadiene and less than 100 ppm water, whereby the butadiene and water contents are calculated based on the weight of the 25 raw materials. It is particularly preferred that the raw material contains less than 500 ppm butadiene by weight of the raw material. Likewise, it is preferred that the raw material is substantially free of compounds containing two or more double carbon-carbon bonds or at least one triple carbon-carbon bond. Such compounds include, for example, isoprene, acetylene, and other compounds having fused or conjugated double bonds, or compounds having at least one triple bond (carbon-carbon).

22

It is also important that the raw material is absolutely free of polar compounds. For this purpose, according to one embodiment, the raw material is selectively hydrogenated prior to dimerization to remove unsaturated hydrocarbon compounds containing triple bonds or more than two double bonds. Preferably, the raw material S is dehydrated before dimerization.

Based on the above, as an example of a particularly preferred embodiment, the process for producing dimers useful as vehicle fuels or components after possible hydrogenation comprises the following steps: providing an industrial refinery feedstock containing C 1 -C 6; -isoleolefins, -the raw material is selectively hydrogenated to reduce the content of unsaturated hydrocarbon compounds containing two or more unsaturated double (carbon-carbon) 15 or at least one triple bonds to less than 1000 ppm to provide hydrogenated crude feeding the hydrated and dehydrated raw material to dehydrate to reduce the water content to less than 10 ppm, based on the weight of the raw material to provide the hydrated and dehydrated raw material, and .

•

Selective hydrogenation, meaning hydrogenation of particularly unsaturated compounds except the desired starting material for the dimerization process, i.e. iso-olefins or linear olefins, in particular isobutylene, is carried out in the presence of excess hydrogen and the free hydrogen is removed after the hydrogenation step in the stripper. The free hydrogen is removed along with any water recovered during the dehydration step in the stripper. The liquid effluent is recovered from the stripper and dimerized.

30 23

The following non-limiting examples illustrate the invention:

Example 1

Synthesis of mesoporous, MIFI-implanted molecular sieve 5 Na-MM5-96h-4ZS-2Al-35 using an aluminum source

Example 1a

Synthesis of Na-MM5-96h-4ZS-2Al-35

Synthesis of 10 Na-MM5-96h-4ZS-2Al-35 was performed by preparing solutions A, B, and C. Solution A

was prepared by mixing 4.5 g of fumed silica with 51.7 g of distilled water under continuous stirring (r.m.p. 196) for 20 minutes. Solution B was prepared by adding 18.1 g of tetramethylammonium silicate to 11.7 g of sodium silicate with continuous stirring (r.m.p.

180) and the resulting mixture was stirred for 20 minutes. Solution C was prepared by dissolving 26.3 g of tetradecyltrimethylammonium bromide in 174.3 g of distilled water with vigorous stirring (r.m.p. 336) for 20 minutes. Solution B was added slowly (15 minutes) to Solution A; with vigorous stirring (r.m.p. 320) and after addition of solution B, the resulting mixture was further stirred for 20 min. Solution C was added slowly (20 min) to the mixture (A + B) with vigorous stirring (r.m.p. 336) and after addition of solution C, the mixture was stirred for another 20 min. Thereafter, 4.2 g of MFI cores prepared according to Example 6 of FI Application 2004 1664 were dispersed in gel solution (A + B + C) with vigorous stirring (r.m.p. 340) for 20 minutes. The homogenization of the dispersed MFI was performed by vigorously mixing the gel (r.m.p.

340) for another 35 minutes. Then 2.3 g of aluminum isopropoxide was added and stirred for 20 min. The resulting gel was then allowed to mature for 3 hours with stirring (r.m.p. 180). The pH of the gel was adjusted and the gel was placed in a Teflon cup, which was then placed in an autoclave. The synthesis was carried out for 96 h at 100 ° C. After completion of the synthesis, the reactor was cooled and the resulting mesoporous material was filtered and washed thoroughly with distilled water. The resulting Na-MM5-96h-4ZS-2Al-35 was dried and calcined using a step calcination procedure in a muffle furnace at 450 ° C.

30 ί 24

Example 1b Preparation of H-MM5-96 h-4ZS-2Al-35

A mixture of 10 g of Na-MM5-96h-4ZS-2Al-35 (prepared above) was ion-exchanged with 1M aqueous ammonium nitrate solution for 24 h at room temperature, after ion exchange the granular porous material was thoroughly washed, dried and calcined for four hours using a step calciner. in a muffle furnace at 4 ° C to 0 ° C.

The following examples illustrate the effect of butadiene and water on the activity of a mesoporous catalyst (H-10 MM5-96h-4ZS-2Al-35) according to Finnish patent application 20041675.

Examples include dimerization experiments performed with isobutylene in a microreactor at 20 bar. The catalyst was packed into a reactor in powder form and mixed with silicon carbide (SiC). The content of isobutylene in the dried (anhydrous) feed was about 20% by weight in Examples 1.3, and 4. In Examples 2, 5, 6.7 and 9, the content of isobutylene in industrial feed was about 17% by weight and the concentration of 1-butylene was about 15% by weight.

Example 2

Effect of feed structure on catalyst activity

The effect of the feed structure on the activity of the catalyst was investigated by Experiments 1 and 2, and their results are shown in Table 2 and Figure 2. The experiments were performed at 80 ° C with a WHSV of 80 h -1 and a dry feed of Experiment 1. In Experiment 1, the content of isobutylene feed was about 20% by weight and the concentration of unreacted i-pentane was about 80%. In Experiment 2 25, an industrial C4 hydrocarbon refinery feed with about 17% by weight of isobutylene was used. The industrial feed contained isobutylene and, in addition to alkanes, other alkenes such as 1- and 2-butylene, and small amounts of butadiene and water.

The results clearly show the effect of feed structure on activity. Conversion of isobutylene 30 decreases much faster when using industrial refinery feed Based on results! We were surprised to find that water and butadiene were apparently strongly deactivated.

Table 2. Conversion of isobutylene X (i = C4) in Experiments 1 and 2 and conversion of 1 + i-butylene 5 in X (l + i = C4) in Experiment 2.

Test 1 Test 2

Time (h) X (i = C4) Time (h) X (l + i = C4) X (i = C4) ~ 24 993 23 53Ό ~ 43 944 46 ~ 33.6 563 ~ 6Ä 923 6? 7 ~ 223 ~ 83 9L5 83 144 105 893 1Ö9 83 1Ϊ6 873 Ϊ3Ϊ 63 17.0 823 Ϊ73 43 21.0 773 2Ϊ3 Z9 ~ 253 743 253 I 43

The results are also shown in Figure 2. As shown, the catalyst deactivated much faster in Experiment 2, in which the feed did not contain butadiene and water.

Example 3

Deactivating effect of butadiene on the catalyst

Table 3 and Figure 3 show the conversion of isobutylene in tests 1.3 and 4. The experiments were performed at 80 ° C and had a WHSV value of 80 h -1 and no feed dehydrated. In both Trials 3 and 4, similar feeds were used that differed from those used in Experiment 1 only in terms of butadiene content. The concentration of butadiene in Experiment 1 was about 0.0% by weight, and the feed used in Experiments 3 and 4 contained about 0.5% by weight of butadiene. The feeds for Experiments 3 and 4 were prepared from the feed for Experiment 1 by adding butadiene. Based on the results of the experiments, we conclude that the butadiene in the feed is very deactivating the catalyst.

According to our experiments, the effect of butadiene on catalyst activity was very rapid, about 2.5 hours after the start of the experiment, the deactivation of the catalyst was no longer essentially dependent on the amount of butadiene in the feed.

26

Table 3. Effect of butadiene on catalyst activity.

Koel Test 3 Test 4

Time (h) X (i = C4) Time (h) X (i = C4) Time (h) X (i = C4) ~ 2Ä 993 "433 Z6 ~~ 44Ö ~ 45 944 46 37 47 523 ~ 6Ä 922 6 / 7 33J 63 533 _ ____ __ _ _ 10.5 893 1Ö9 3Ö2 TTÖ 453 12.6 873 13Ό 333 Ϊ33 4Z9 17.0 823 173 333 172 393 TTÖ 773 241 32Λ) "242 353 _ _ __ __ ___ _ _ I _I _____ — ___ 5

Example 4 (1)

Water Deactivation of Catalyst (1)

Experiments 5 and 6 were performed at 100 ° C and had a WHSV value of 40 h -1 (Figure 4, Table 4).

The feed was introduced into the reactor via a dehydration unit containing approximately 40 ml of a 3 A molecular sieve. In Experiment 5, dehydration was ongoing throughout the experiment. In Experiment 6, the dehydration medium was replaced with another similar regenerated dehydration medium 8.9 hours after the fourth GC sample was taken. After 13.1 hours the dehydration unit was completely decommissioned. The results show that increasing the dehydration power significantly slowed down the deactivation, whereas stopping the feed dehydration during the experiment caused a collapse of the catalyst activity. Therefore, even a small amount of water in the feed has a very strong deactivating effect on the activity of the catalyst.

20 27

Table 4. Conversion of 1 + i-butylene in Experiments 5 and 6.

Experiment 5 ™ Experiment 6 —— ~ Ato (hj X (l + i = C4) AibT (h) Tx (l + i = C4) 33 ~ 653 23 663 'i 33 564 4/7 593 33 5Z8 6.8 543 "53 ~ 47.4 83 ~~ 503 11.9 403 ~ ΓΓ3 48/7 14 342 Ϊ3Τ 464 16Λ 3Ϊ1 17.1 13Λ 3Ö3 254 2Ϊ3 93 ”242 22/7 2 ^ 2 84

Figure 3 shows that catalyst 6 activity collapsed after feed dehydration was stopped during experiments.

·> 5

Example 4 (2)

Water Deactivation of Catalyst (2) Conversion of 3 1 + i-Butylene in Experiments 5 and 7 (100 ° C, 40 h -1) is shown in Table 5 and Figure 5.

The experiments differed only in the volume-controlled amount of the dehydration technique used (molecular sieve 3Λ). In Experiment 7, the dehydration unit contained 395 mL of molecular sieve of 3 A. As shown, increasing the dehydration of the feed significantly slowed the rate of catalyst deactivation. Examples 4 (1) and 4 (2) clearly show that the catalyst life can be significantly extended by dehydration and water separation. In test 5 h, the average water content of the reactor feed was 100 ppm and in test 7 the concentration was 10 ppm.

20 3 28

Table 5.1 + Conversion of i-butylene in Experiments 5 and 7.

Test 5 Test 7

Time (h) X (l + i = C4) Time (h) X (l + i = C4) _ __ __ _ __ 222 "4ΖΪ ___ __ _ 40.2 __ _ 11.9 4ÖÖ 1ÄÖ 342 __ __ ~ 2Ö2 254" 242 227

Claims (27)

A process for dimerizing olefins in the presence of a catalyst, wherein a hydrocarbon feedstock containing C4-C6 isoolefins is dimerized, characterized in that said process comprises the following steps: - The C4-C6 isoolefins are brought into contact with dimerization conditions a catalytic material comprising an acidic mesoporous molecular sieve, wherein the catalytic material is thermally stable at a temperature of at least 900 ° C, and - the raw material used contains less than 1000 ppm butadiene and less than 100 ppm water, the concentrations of butadiene and water are calculated on the basis of the weight of the raw material. 2. Process according to claim 1, characterized in that the raw material contains less than 500 ppm butadiene and less than 30 ppm water calculated on the basis of the weight of the raw material. 3. A method according to claim 1 or 2, characterized by the catalytic material comprising a mesoporous molecular sieve embedded in a zeolite. 20 4. Process according to any of the preceding claims, characterized in that the raw material does not in principle comprise any compounds which contain two or more double bonds (carbon-carbon) or at least one triple bond (carbon-carbon). 25 5. A process according to any of the preceding claims, characterized in that the raw material does not in principle comprise any polar compounds. Process according to any of the preceding claims, characterized in that the raw material is selectively hydrogenated prior to dimerization to remove unsaturated hydrocarbon compounds comprising triple bonds or more than two double bonds. Process according to any of the preceding claims, characterized in that the raw material is dehydrated before dimerization. 8. A process according to any of the preceding claims, further comprising: - obtaining an industrial refined crude material containing C4-C6 isoolefins; - the raw material is selectively hydrogenated to reduce the concentration of unsaturated hydrocarbon compounds containing several or two unsaturated hydrocarbon compounds containing (carbon-carbon) or at least one triple bond (carbon-carbon) to less than 1000 ppm, to obtain a hydrogenated feedstock. - the hydrogenated raw material is fed to be dehydrated to reduce the water content to less than 100 ppm, calculated on the basis of the weight of the raw material, to obtain a hydrated and dehydrated raw material, and - the hydrogenated and dehydrated raw material is led to dimerization. 15 Method according to claim 8, characterized in that the selective hydrogenation is carried out in the presence of excess hydrogen and that free hydrogen is removed after the hydrogenation step in a stripper. 20 Process according to Claim 9, characterized in that the free hydrogen is removed together with any water recovered in the dehydration step in the stripper. 11. A method according to claim 10, characterized in that the liquid discharge of the stripper is utilized and dimerized. 25 Process according to any of the preceding claims, characterized in that dimerization is carried out at a temperature of 50 - 200 ° C, preferably 80 - 150 ° C. Process according to any of the preceding claims, characterized in that the pressure of the reaction zone is maintained higher than the meadow pressure of the mixture in the reactors. Process according to any of the preceding claims, characterized in that the pressure of the reaction zone is maintained in the range 10-50 bar, preferably about the range 15-25 bar. A process according to any of the preceding claims, characterized in that the specific surface of the catalytic material is in the range 1400 - 500 m2 / g, preferably in the range 1200 - 600 m2 / g. MFI, MTT, TON, AEF and FER zeolites, and a large porous zeolite containing BEA, FAU and MOR zeolites, preferably the zeolite is selected from the group consisting of MFI, MTT, AEF, BEA yes MOR zeolites. Method according to any of the preceding claims, characterized in that the catalytic material comprises a mesoporous molecular sieve selected from the M41S group, preferably the mesoporous molecular sieve is MCM-41 or MCM-48. A process according to any of the preceding claims, characterized in that the catalytic material comprises a medium porous zeolite selected from the group consisting of Method according to any of claims 15-17, characterized in that the mesoporous molecular sieve is MCM-41 or MCM-48 and the zeolite is an MFI or BEA zeolite. Process according to any of claims 15-18, characterized in that the catalyst is in proton form, in cation form or is modified with a metal. 25 Process according to any of claims 15-19, characterized in that the catalyst comprises 90-10% by weight of a catalytic material according to any of claims 9-13 and 10-90% by weight of a support. 30 21. A method according to claim 1, characterized in that the separation zone comprises a distillation zone. Process according to any of the preceding claims, characterized in that the first product of the process comprises the distillation zone distillate and the second product comprises the distillation zone distillation zone. 5 23. A method according to any of the preceding claims, characterized in that the distillation zone is used in such a way that a product is obtained whose steam pressure rises to 10 -20 kPa (Reid) and whose distillation point (90% by volume, ASTM D86) is equal to 180 ° C or less. 10 Process according to any of the preceding claims, characterized in that the dimerization process is carried out in a reactive distillation system containing at least one reaction zone and at least one distillation zone, wherein said at least one reaction zone contains at least one at least one reactor and said at least one reactor and said at least one reactor. a distillation column. 25 Cs olefins; the said liquid phase olefin hydrocarbon material is contacted with an acid catalyst in the reaction zone, the acid catalyst comprising a mesoporous molecular sieve embedded in a zeolite and the catalytic material thermally stable at a temperature of at least 900 ° C; 30. Before said reaction zone, a product containing dimerized olefin is recovered. 5 25. A method according to any of the preceding claims, characterized in that the isoctane is produced from the feed which comprises isobutene, and the isoct is optionally hydrogenated to produce isoctane. 20 26. A process according to any of claims 1 to 25 for dimerization of olefins, lower hydrocarbons, characterized in that: - in a reaction zone is fed an olefinic hydrocarbon material containing unsaturated hydrocarbons, which is isobutene, butene, linear CVolcfincr and branched Process according to any of the preceding claims, characterized in that the step of contacting the olefinic hydrocarbons of said raw material with an acid catalyst is carried out mainly in the absence of oxygenate and polar compounds.