Title of Invention	PROCESS FOR THE MANUFACTURE OF OLEFIN OLIGOMERS
Abstract	The present invention relates to a process for producing polyolefms wherein a feedstock comprising n-olefin or n mixture of nolofins is dimorized in the pres ci ce of a solid noidic catalyst by passing the feedstock to a catalytic distillation apparatus comprising either a) a combination of distillation column and a reactor comprising at least one catalyst bed, or b) a distillation column connectcd to one or more sidde reacors comprising at least one catalyst layer, recovering the unre- acted n olefin from the distillation column or the combination of the distillation solumn and this reactor at the upper port thereof as sidc-streain to be coinbined with the feedstock, and the reactor product from the dimerization is hydrozen- aled.

Title of Invention

PROCESS FOR THE MANUFACTURE OF OLEFIN OLIGOMERS

Abstract

The present invention relates to a process for producing polyolefms wherein a feedstock comprising n-olefin or n mixture of nolofins is dimorized in the pres ci ce of a solid noidic catalyst by passing the feedstock to a catalytic distillation apparatus comprising either a) a combination of distillation column and a reactor comprising at least one catalyst bed, or b) a distillation column connectcd to one or more sidde reacors comprising at least one catalyst layer, recovering the unre- acted n olefin from the distillation column or the combination of the distillation solumn and this reactor at the upper port thereof as sidc-streain to be coinbined with the feedstock, and the reactor product from the dimerization is hydrozen- aled.

Full Text	Process for the manufacture of olefin oligomers Technical field The invention relates to the preduction of high-grade base oils and to selective dimerization of n-olefins usinj; a solid and acidic catalyst. Particularly, the inven- tion is directed to a process wherein n-olefins are dimerized in a catalytic distilla- tion apparatus, followed by hy Irogenation to give polyolefins. State of the art Saturated olefinic oligomers are a significant group of high-grade synthetic base oils. Poly-alpha-olefins known as PAOs are typically produced by oligomerization of alpha-olefins in the presence of homogsneous Friedel-Crafts catalysts, such as boron trifluoride (BF3) and a promoter at slightly elevated BF3 pressures, and temperatures below 100 °C. Water or an alcohol normally serves as the promoter. In the PAO process, 1 -decent¦ is typically used as the feedstock, mainly giving trimers and tetramers of the feedstock olefin as the product. Base oils may also be produced by dimerization of n-olefins heavier than decene. Base oils of the PIO (poly(intt rnal olefin)) group are produced by dimerization of internal n-olefins, typically C-15-C16 n-clefins, using BF3 catalysts. Among the products of the PIO process, fimers are particularly suitable feedstocks for base oil production. A catalyst separation step is ..lways necessary in PAO and PIO processes using homogeneous catalysts. As is known, n-olefins refer to linear olsiins or linear olefins with no more than one branch, that is, to slightly oranched olcfins. Oligomerization refers to a reaction where molecules of at least one type react with each other resulting in the increase of the molecular weight, said increase being the added molecular we ight of at least three molecules. Oligomerization may be illustrated with the following equation: where A may be identical with or different from B, x is 0 or an integer, y is 0 or an integer, x+y > 2, Sx = n, Sy = m, A and B are olefmic molecules, and n and m are integers. The term oligomtr refers to a. repetitive combination of monomeric units of at least the one type, this number of said units ranging from 3 to 100. The term dimerization refers to a reaction where molecules of at least one type react with each other resulting in the increase of the molecular weight, said in- crease being the added molecular weight of at least two molecules. Dimerization may be illustrated with the following equation: where A may be identical with or different from B, A and B being olefmic mole- cules. The term dimer refers to a combination of two monomeric units of at least one type. The term polyolefin refers to a combination comprising at least two olefmic monomeric units of at least one type. Several alternative heterogene jus catalysts for the dimerization of heavy olefins are known e.g. from patents LS 4,417,088, US 5,053,569, US 5,453,556 and US 6,703,356. In addition to the dssired dimerization reaction, the use of said hetero- geneous acid catalysts disclose d in said documents resulted in cracking and isom- erization reactions of the feedstock olefins, as well as formation of heavier oli- gomers, mainly trimers and tetamers. Particularly detrimental isomerization reac- tions of the feedstock olefins i lclude reactions yielding products that may not be dimerized, such as naphthenes. Attainable yields of the base oil are lowered by the cracking products formed as well as naphthenes resulting from the isomerization of monomeric olefins. Because relatively valuable olefin are used as the process feedstock, any undesirable side reactions tiereof also have a considerable impact on the feasibility of the process. The document US 5,053,569 discloses the dimerization of a-olefins using an acidic calcium/montmorillonite catalyst. The document US 6,703,356 discloses an oligomerization process of alpha-olefins using alpha-olefins with a carbon number ranging between 10 and 30, or a mixture thereof as the feedstock. The catalyst is specified as a crystal catalyst "laving a constraint index of less than 3. The con- straint index is a measure of the product selectivity of the catalyst. In the exam- ples, 1-hexadecene, l-tetradeene and 1-octadecne are used as feedstocks, whereas MCM-22, MCM-56, USY, Beta, ZSM-12 and WOx/on ZrO2 are used as catalysts. In the examples of the above documents US 5,053,569 and US 6,703,356, highest conversions attained in the dimsrization of heavy olefins using heterogeneous acid catalysts are 92 % and 87 %. A conversion of clearly below 100 % may be attrib- uted to the formation of naphtr enes from feedstock olefins. The product obtained in the examples of the documeit US 5,053,569 contained trimers and heavier oli- gomers in a total amount of 47 %, which i,3 not desirable for the properties of the base oil product. Selectivity fo • dimers is higher (70 %) in the example described in the document US 6,703,3.66, having a high conversion (87 %), yet high amounts of cracked products (4 %) were obtained. The term catalytic distillation refers generally to the combination of a chemical reaction with the product separation. The rsaction and product separation are car- ried out together in an inseparable manner. A catalytic distillation apparatus nor- mally comprises a distillation ;olumn incorporating one or more catalytic zones. In said catalytic zones, streams from a specified level or plate of the distillation column are treated to give desired reaction products. Thereafter, the product stream is fractionated using it distillation means. Said catalytic zone may be placed within or outside the distillation means. In industrial applications, catalytic distillation is used in the prodi ction of ethers, said process being also known e.g. for dehydration of alcohols, and oxydation of paraffins. Use of the catalytic distillation is also known in the oligomerization process. US 4,935,577 discloses an oligomerization process wherein alpha-olefins having from 3 to 12 carbon atoms are passed to a distillation column containing Lewis acid catalyst for the reaction. The temperature of the catalytic distillation apparatus is no more than about 150 °C, ths typical operation temperature range being below 50 °C. A combination catalysi is used in the distillation apparatus, and accord- ingly, a unit for the separation of the Lev/is acid prior to recycling to the distilla- tion apparatus is an essential part of the apparatus US 2,198,937 discloses an apparatus comprising a distillation column and a side reactor for the polymerization of hydrocarbons. The apparatus may be operated under conditions similar to those in catalytic distillation. The document FI 96852 disclc ses a process and an apparatus for oligomerizing olefins. In this process, C3-C20 )lefins or mixtures thereof are passed to a catalytic distillation system where the f :edstock olefins are contacted with a catalyst at a temperature of above 150 °C, hus yielding a product containing oligomers. The catalyst used in this process manly consists of zeolite, the products formed being middle distillates and lubricant'. The catalytic distillation system used in the proc- ess may also be a distillation column for the product separation, connected to at least one side reactor containing the catalys:. Based on above teachings it may be seen that there is an obvious need for a novel improved process for the production of polyolefins from n-olefms, said novel process eliminating or at least s ubstantially reducing problems and deficits associ- ated with the solutions of the state of the art. Objects of the invention An object of the invention is to provide a process for the production of polyolefins from n-olefms. Another object of the inventior is to provide a process for the production of poly- olefins from n-olefms using a s slid acidic catalyst. Still another object of the invention is to provide a process for the production of polyolefins from C8-C30 n-olefms using a solid acidic mesoporous catalyst. Further, an object of the inven:ion is to provide a process for the dimerization of C8-C30 n-olefms using a solid acidic mesoporous catalyst to give polyolefins. An object of the invention is al.ro a process for producing a base oil component. Yet another object of the invention is the use of a catalytic distillation apparatus for the dimerization of n-olefins, particularly C8-C30 n-olefins, An object of the invention is a. so to provide di-n-olcfins consisting of two identi- cal or different n-olefms having carbon chain lengths of C8-C30. Summary of the invention The present invention relates to a process for producing high-grade products use- ful as base oils and base oil components from n-olefins using a solid and acidic catalyst, by dimerizing an n-olrfin or a mixture of n-olefms in a catalytic distilla- tion apparatus, wherein the catalyst is placed into a distillation column or into a side reactor outside the distillation column, followed by hydrogenation of the product. In the process of the invention, the: feedstock comprising an n-olefin or a mixture of n-olefms, is dimerzed in the presence of a solid acidic catalyst by passing the feedstock to a catalytic distillation apparatus either comprising a) a combination of a distillation column and a reactor having at least one catalyst layer, or b) a distillation column connected to one or more side reactors having at least one catalyst layer, recovering the unreacted n-olefin at the upper part of the distillation column or the combination of the distillation column and the reactor as a side-stream to be combined with the feedstock, whereas the reaction product from the dimerization is hydrogenated. Impurities still present in the dimerization product or final product may optionally be removed using an additional distilla- tion step. The invention and some alternative embodiments thereof are illustrated by ap- pended figures 1, 2 and 3, however, without wishing to limit the invention to these presented embodiments. Figures Figure 1 schematically shows an embodiment of the invention for producing a base oil component. Figure 2 schematically shows an embodiment of the invention wherein the dimerization of n-olefins is car'ied out in a catalytic distillation apparatus with the catalyst in separate side reactois. Figure 3 schematically shows an embodiment of the invention wherein the dimerization of n-olefins is performed in a combination of a distillation column and a reactor. Figure 1 is a schematic presentution of the basic solution of the inventive process. The feedstock of the process, 01 stream 1, containing C8-C30 n-olefins or a mixture thereof is obtained from a feed tank (not shown in the figure). Stream 4 recovered at the upper part of a distillatioi column B is combined with stream 1. Streams 1 and 4 together form the stream 2, which is passed to side reactor A containing the dimerization catalyst layer / bed D. Stream 3 is obtained as the product of the side reactor A, said stream 3 containing C8-C30 n-olefm monomers and as reaction products mainly dimeric produsts. Stream 3 is passed to a distillation column B where C8-C30 n-olefm monomers rise to the upper part of said distillation column B, followed by recycling the n-olefm monomer fraction thus obtained as stream 4 to the side reactor A in the reaction stage. Stream 5 is removed at the top of the distillation column B, said stream 5 containing components of non-dimerizable monomer fraction in the upper part of the distillation column. With stream 5 the accumulation of the componei.ts of non-dimerizable monomer fraction such as branched compounds is prevented in the catalytic distillation apparatus. The prod- uct in the form of a product stream 6 mainly consisting of dimers is obtained from the bottom of the distillation column B, said stream 6 being passed to a hydro- genation reactor C. In the hydrogenation reactor C, olefins arc hydrogenated to give the product, suitable as a base oil component, followed by the removal of said product from the hydrogen ition reactor C as the product stream 7. In figure 2, a preferable soluticn of the process of the invention is shown. In this solution, a catalytic distillatioi apparatus having dimerization catalyst layers D and E placed in separate side re actors A and B is used. The feedstock of the proc- ess, or stream 1, containing C8-C30 n-olefins or a mixture thereof is obtained from a feed tank (not shown in the figure). Stream 4 recovered at the upper part of a distillation column C is combined with stream 1. Streams 1 and 4 together form the stream 2 being either passe i to a side rsactor A or side reactor B. Stream 2 is passed to the side reactor containing fresh catalyst or regenerated catalyst. Simul- taneously, the other side reactor contains the exhausted catalyst, or the catalyst is being regenerated. There are two side reactors in the system, and accordingly, it is not necessary to shut down the process due to lowered catalytic activity. Stream 3 is obtained as the product of the side reactor A or B, said stream 3 containing Cs- C30 n-olefin monomers and as i eaction products mainly dimeric products. Stream 3 is passed to a distillation column C where C8-C30 n-olefin monomers rise to the upper part of said distillation column C, followed by recycling the n-olefin monomer fraction thus obtainec as stream 4 to the side reactor A or B in the reac- tion stage. Stream 5 is removed at the upper part of the distillation column C, said stream containing components of non-dimerizable monomer fraction at the upper part of the distillation column C. Accumulation of the components of non- dimerizable monomer fraction such as branched compounds in the catalytic distil- lation apparatus is prevented by stream 5. From the lower part of the distillation column C, a product stream 6 mainly consisting of dinners, subsequently passed to a hydrogenation reactor for hydrogenation (not shown in the figure), and a bottom product 7 are obtained. Figure 3 shows another preferable solution of the inventive process, using a cata- lytic distillation apparatus having the dimerizing catalyst layer B placed inside the distillation column A. The feed stream 1 comprising C8-C30 n-olefms or a mixture thereof is combined with the n-olefin monomer fraction 4 from the top of the dis- tillation column A to give a stieam 2 for passing to the distillation column A, to the upper part of the catalyst la/er B. In the distillation column A, the feed passes to the catalyst layer B, in which the dimerization reaction mainly proceeds. In the catalyst layer B, n-olefin monomer is in the form of a vapour/liquid mixture and the liquid n-olefin monomer ashes the dimers and oligomers formed from the catalyst, n-olefin monomer fration 3 is obtained from the top of the distillation column A, said fraction being plit to give a monomer fraction 4 to be recycled to the reactor, and a monomer steam 5. Accumulation of non-dimerizable compo- nents of the monomer fraction in the catalytic distillation apparatus is prevented by removal of said monomer siream 5. The product in the form of a dimer stream 6 is obtained from the bottom of the distillation column A, said stream 6 being passed to a hydrogenation read or (not shown in the figure) for hydrogenation. The catalytic distillation apparatus comprises either a) a combination of a distilla- tion column and a reactor having at least one catalyst layer, or b) a distillation column connected to one or more side reactors having at least one catalyst layer. Detailed description of the in mention It was surprisingly found that the use of a catalytic distillation apparatus with a solid acidic catalyst allows for the selective dimerization of n-olefins to give base oil useful as a lubricant, or base oil component, with an excellent yield that may even be above 95 %. In the process of the invention, the dimerization of n-olefins or n-olefin mixtures to yield d. mers is perlbrmed in a catalytic distillation appara- tus comprising either a) a combination of a distillation column and a reactor com- prising at least one catalyst layer, or b) a distillation column connected to one or more side reactors comprising at least one catalyst layer, said catalyst being in the side reactor outside the distillation column. In this way, the temperature of the dimerization reaction may be maintained low, typically below 150 CC, thus pre- venting the formation of naplithenes and cracking of products. Dimerization is followed by hydrogenation to produce the base oil component. Optionally after the dimerization and/or hydroj enation, the product may, if necessary, be passed to a distillation apparatus where my monomer residues are removed from the dimer- ized product, or the dimers are separated from heavier trimers and tetramers. The quality of the base oil and'or base oil component obtained as the product use- ful as a lubricant, after hydrogenation is excellent due to, among other things, low number of undesirable reactions. In the process of the inventio 1, n-olefin or a mixture of n-olefins is passed to a catalytic distillation apparatus where the single conversion of n-olefins may be suitably controlled. The term single conversion refers to the conversion of an n- olefm to give another compound during the reaction stage. Single conversion of n- olefin monomers is defined as follows: Single conversion of n-olefin monomers (%) = 100 x (proportion of n-olefins in the feedstock prior to the reaction stage proportion of n-olcfins in the product after the reaction stage) / propo/tion of n-olofins in the feedstock prior to the reac- tion stage. In the process of the invention, it is preferable to use recycling of the monomeric olefins particularly in embodiments where the catalyst is placed in a side reactor. In this case, total conversion and dimerizE.tion selectivity are improved by effi- cient recycling of the monomeric olefins. A requirement for the efficient recycling of the monomeric olefins is the adjustment of the reaction conditions to values preventing undesired side reactions such as cracking and isomerization of the monomer resulting in naphthcnes. Moreover, relatively low single conversion between 5 and 50 % is anotier requirement for high dimerization selectivity meaning selectivity ranging approximately between 80 and 100 %. In the process of the invention, high total con/ersion, ranging between 95 and 100 %, is attained with efficient recycling, in spile of the relatively low single conversion. Efficient recycling means that the exten of recycling is such that between 50 and 95 % of the original monomer stream contacts the catalyst layer more than once. Extent of recycling refers to the level o1 the stream passed from the separation step to the catalyst layer. Dimensioning of the components of the apparatus, and thus the investment costs depend on tie level of said stream, and therefore, reasonable extent of recycling is preferab' e for said investment costs. Also energy consump- tion of the process is influence! by the extent of recycling. Single conversion is adjusted :o a suitable value on the basis of the carbon num- ber. Suitable single conversion means that the extent of side reactions such as formation of heavier oligomer! and cracking is adjusted to be low, that is less than 10 % of the total conversion of the monomeric olefms is due to side reactions. Reduced formation of heavy c ligomers has a positive impact on the properties of the base oil product, and retarc s the deactivation rate of the catalyst. The process of the invention is now described in more detail. In the process, the feedstock is passed to a catalyst layer in a catalytic distillation apparatus, either above the catalyst layer in a distillation column where the feedstock migrates to the catalyst layer, or to a catalyst layer in a side reactor, followed by passing the reaction mixture to a distillation column or to a lower part of the distillation col- umn where the monomer fract on migrates to a higher part of the distillation col- umn, and any impurities present in the monomer lighter than the monomers in the distillation range leave the di; filiation column at the top thereof or may be re- moved in a subsequent separat on unit downstream of the distillation column, and any impurities present in the n.onomer heavier than the monomers in the distilla- tion range are removed from the column at the bottom or may be removed from the bottom product in a separat; separation unit. The dimerization reaction and some simultaneous oligomerization proceed in the catalyst bed. A monomeric friction such as n-hexadecene with a boiling point between 280 and 290 °C / 1 atm, rises in the column as the lighter component upwards from the feeding plate and is removed from the distillation column, typi- cally at the top thereof as a s :de stream to be recycled and combined with the feedstock. The dimer produced as the raw product, e.g. C32 olefin, migrates downward in the distillation column, followed by separation of the dimer stream at a lower part thereof. Oligomerization product such as C48+ olefin is obtained as the bottom product stream, In case unreacted olefins and side reactions such as formation of naphthenes corresponding to monomers with respect to their carbon number may not be totally avoided, accumulation thereof in the process streams may, however, be prevented removing part of the side stream to be recycled at the upper part of the distillation column. A sqlvent or solvent mixture may be added to the feedstock, thus retarding deacti- vation of the dimerization catalyst. The solvent is selected among hydrocarbons such as n-paraffins, isoparaffns and aromatic solvents. The solvent may be re- moved from the distillation column as a respective side stream via cooling to a separation unit where the solvent is separated from the reaction product. The sol- vent may also be recycled in the catalytic distillation apparatus without a separate separation unit. The feedstock may be optional y dried for removal of any water and other impuri- ties present therein. Drying may be carried out with known drying means such as with commercially available Secular sieves e.g. zeolite 3A molecular sieve used for drying hydrocarbons, or using other suitable known methods. Process feedstock comprises at least one n-olefin selected from the group consist- ing of C8-C30n-olefins. Examples of suitable n-olefins include 1-decene, mixture of decenes, 1-dodecene, mixture of dodecenes, 1-hexadecene, mixture of hexade- cenes, 1-octadecene, mixture of octadecemss, and C20-C22 1-olefins, preferably 1- hexadecene, or a mixture of h;xadecenes. n-Olefins may be synthetic olefins or olefins of biological origin, pioduced from biological starting materials such as vegetable oil and animal fats. In case a catalytic distillation a jparatus where the catalyst is placed in a side reac- tor outside the distillation coluinn is used in the dimerization step of the process of the invention, the feedstock such as 1-hexadecene is directly introduced into the catalyst layer in the side reacto:1 where a dimer is produced. The product stream of said side reactor is passed to 1 distillation column where the monomer fraction such as n-hexadecene having a boiling point between 280 and 290 °C / 1 atm will rise upward as the lighter component in the distillation column from the feeding plate and is removed, typically at the top of the column as a side stream and recy- cled to the side reactor. The dimer formed as the reaction product such as C32 ole- fin migrates downward in the distillation column, the dimer stream being sepa- rated at a lower part thereof. an oligomerization product such as C48+ olefin is obtained as the bottom product stream. The diiner product obtained above and the bottom product produced as the by- product are hydrogenated in a hydrogenation step operated continuously or batch wise in the presence of hydrogsn. Known, hydrogenation catalysts containing met- als from the groups VIII and/cr VIA of the Periodic System of the Elements may be used. Preferable hydrogen ition catalysts include supported Pd, Pt, Ni, Cu, CuCr, NiMo or CoMo catalys" s, the support being preferably alumina and/or sil- ica. The hydrogenation step is carried out at a pressure of 5 to 100 bar, preferably at 10 to 70 bar, and at a temper iture of 100 to 400 °C, preferably at 150 to 250 °C. In the dimerization step of the process of the invention, the temperature in the distillation column or in the catalyst layer of the side reactor ranges between 25 and 200 °C, preferably between 50 and 150 °C. Pressure range varies according to the carbon number of the feedi tock and reaction temperature, and further, the lo- cation of the catalyst either in the catalytic distillation apparatus or in a side reac- tor has an impact on the pressure range. The pressure range thus varies from re- duced pressures to elevated pressures. The pressure may vary between 0.001 mbar and 50 bar, preferably 0.5 bar und 30 bar, particularly preferably between 1 mbar and 20 bar. WHSV is properly adjusted on the basis of the single conversion and deactivation rate of the catalyst. WHSV typically ranges between 0.1 and 50 h"1 with respect to the feedstock, pi eferably between 0.5 and 20 h'1. Selection of the pressure of the dimerization reaction as a function of the tempera- ture is illustrated in the table 1 below. Table 1 shows the vapour pressure of 1- decene as a function of the ten.perature. In case the desired reaction temperature in the catalyst layer of the cataytic distillation apparatus is 50 °C, the pressure is selected to be 11 mbar, or in case the desired temperature in the catalyst layer of the catalytic distillation apparatus is 150 CC, the pressure is selected to be 580 mbar. If a catalytic distillation apparatus where the catalyst is placed in a side reactor is used, the pressure of the reactor is selected to give the product mixture in liquid phase. High-grade base oil or a base oil component may be produced by the process of the invention from n-olefins us .ng a solid aid acidic catalyst not requiring parts of the catalyst to be subsequently separated and recycled to the distillation apparatus. Moreover, solid acidic catalysis always result in isomerization of double bonds. Particularly suitable solid acidic catalysts include catalysts having a mesopore surface area of above 100 m2/g as measured by nitrogen adsorption and calculated with the following BJH equation: Suitable solid acidic catalyst materials include materials having a mesopore sur- face area based on the BJH eq lation of more than 100 m2/g, preferably more than 300 m2/g. Such materials comprise amorphous aluminium silicates, preferably acidic amorphous aluminium silicates, and particularly preferably acidic amor- phous aluminium silicates with Bronsted acidic sites; zeolites, preferably dealu- minated Y zeolites; and mesoporous materials with a regular porous structure, containing silicon and aluminium, among which mesoporous molecular sieves with inserted zeolite are preferable. The zeolite is preferably ZSM-5, beta-zeolite or MCM-22, whereas the mes jporous molecular sieve is MCM-41 with a regular molecular structure. The acidic solid catalyst material is particularly preferably a mesoporous molecular sieve wiih imbedded MCM-22 zeolitic structures. The expression mesoporous malecular sieve imbedded with zeolite refers to a catalyst having mesoporous molecular sieve structure and zeolitic structure in the same material and where said mesoporous molecular sieve structure and zeolitic structure are bonded together tt rough a chemical bond. The mesoporous molecu- lar sieve with imbedded zeolite, and the production thereof is disclosed in the pat- ent application FI 20041675. The aluminium content of the catalyst material varies between 0.2 and 30 % by weight as determined by alum inium content assays, while the level of the acid sites varies between 50 and 50) umol/g as determined by the analysis according to the NH3-TPD method. Suitable catalyst materials have BrBnsted acidity, the level thereof being more than 1) umol/g as measured by the proton NMR method. The mesoporous surface area c f the catalyst material as calculated with the BJH equation is more than 100 m2/g preferably more than 300 m2/g. The catalyst also comprises a support material to provide a readily mouldable catalyst with mechanical resists nee. Said support material is typically an inorganic oxide such as alumina or silica. In the process of the invention the olefins and isomers of the olefins in the feed- stock are preferably recycled, ind the dimsrized and oligomerized fractions thus produced are obtained as products from the bottom or as a side stream. In this manner, the total conversion h the process is high, preferably more than 95 %, and more than 99 % at best. An apparatus comprising 1) an optional drying means for the feedstock, 2) a cata- lytic distillation apparatus comprising either a) a combination of a distillation col- umn and a reactor having at bast one catalyst layer, or b) a distillation column connected to one or more side reactors having at least one catalyst layer, 3) hy- drogenation reactor and optimally a distillation column upstream and/or down- stream of the hydrogenation reactor, is suitable for the process of the invention, and particularly for dimerizaticn of Cs-Cao n-olefins. Regeneration of the deactivat sd dimerization catalyst may be performed in the same reactor as the dimerizat on reaction. The regeneration is carried out at an elevated temperature using a gaseous mixture that may contain oxygen. During the regeneration, the temperature is equal or higher than that of the dimerization reaction. Thermal resistance of the catalyst and the feasibility of the process are influenced by the maximum legeneration temperature, lower regeneration tem- peratures being preferable with respect to energy consumption. However, the re- generation temperature should be sufficiently high for the removal of coal formed and any impurities adsorbed in the catalyst. The process of the invention I; endowed with several advantages. The process is continuous, and thus continuoos production of the base oil component is possible without interruptions typical for batch reactors. In the process of the inventior, the reactor feedstock may be adjusted as desired since the composition may be controlled very precisely at the side exit of the dis- tillation column. It is thus poss ible to provide the reactor with a stream containing only a monomer fraction without any lighter or heavier fractions due to possible unselective reactions, said lighter or heavier fractions yielding products with un- desirable carbon numbers durir g the reaction stage. Desired compositions are obtai led as the reactor products since the outlet sites for the products may be selected. This is particularly useful in cases where a mixture of n-olefins having different carbon numbers serves as the feedstock. It was surprisingly found that selective dimerization of olefins is provided at a reaction temperature below If 0 °C. Selective dimerization means here a dimeriza- tion selectivity of more than 80 %. Dimer yields are remarkably high since any unreacted monomer present in the reaction product may be recycled to the feed- stock for the reactor by adjusting the recycle ratio. Reactive components may betimmediately separated in the distillation column, and thus side reactions detrir ental to the process may be quickly stopped. Other separation units e.g. two disti lation columns are not necessary in the process for concentration of the feedstock and the reaction product since the concentration of the product and fractionation if the feedstock may be carried out in the same col- umn, which accordingly resultss in a simplers apparatus and lower investment costs. In addition, the regeneration of the dimerization catalyst may be performed in the same reactor as the reaction. The invention is now illustraisd with the following examples without wishing to limit the scope thereof. Examples Example 1 (comparative example) Dimerization of olefins in a bitch reactor Dimerization process was pjrformed in a batch reactor using 1-hexadecene (Neodene 16®). The process temperature was 200 °C, the reactor pressure being 20 bar. The amount of the fe sdstock was 50 g, and that of the catalyst 2 g. The total reaction time was 24 hours. Commercially available Y zeolite (TOSOH Co.), beta-zeolites (TOSOH Co,) aid a mesoporous material MCM-41 (produced ac- cording to a method disclosed in Catalysis Letters 36 (1996) 103) were used as catalysts. Total Cu conversion (= conversion of C16 hydrocarbons to give prod- Dimerization of 1-decene was also performed in a batch reactor. The process tem- perature was 120 °C, the realtor pressure being 20 bar. Commercially available aluminium silicate catalyst Nikki Chemical Co. Ltd) was used as the catalyst, in an amount of 4 % by weight of the feed. The results are shown in the figure 4. The results indicate that the dimsrization selectivity was clearly lower in the batch reactor in comparison to the selectivity of the process of the invention. Example 2 Dimerization in a system corrssponding to a catalytic distillation apparatus, com- prising dimerization in a flow reactor and distillation a) Dimerization in a flow reac or 1-Hexadecene (Neodene 16®] was introduced at a rate of 10 g/h into a flow reac- tor (1 bar (a)). The flow reactcr was packed with 5 g of aluminium silicate catalyst (Nildci Chemical Co. Ltd) wth an aluminium content of 13 % by weight, the number of acid sites of 120 xmol/g, and mesoporous surface area > 300 m2/g, diluted with silicon carbide (\ (catalyst) : V (SiC) = 1:3), followed by the dimeri- zation reaction at a temperaturs of 120 °C in the reactor. The dimerization product was collected into a product container. Results of the dimerization, that is, total The conversion of the C16 hydrocarbons declined the longer the catalyst remained in the hydrocarbon stream. Accumulation of heavy hydrocarbon oligomers in the catalyst (coke formation) was reason to the reduced conversion. b) Distillation of the products The experiment of a) was refeated, and the products obtained from the experi- ments (1766 g) were pooled. The unreactsd monomer fraction (1103 g), middle fraction containing C20-C30 hy irocarbons (9 g), dimer fraction (452 g) and heavy bottom product (199 g) were separated from said pooled products by distillation. c) Recycling of the monomers ' dimerization The monomer fraction obtained above was used as the feedstock in dimerization. The distilled monomer fractior mainly consisted of internal C16 olefins. The com- position of the distilled mono ner fractior. is shown in the table 4 below in the form of surface area percentages as measured by GC-MS-analysis. The dimerized product of example 2, separated by distillation, and the bottom product of the distillation wers hydrogenated as separate batches in a batch reactor using a heterogeneous nickel catalyst. In the hydrogenation, the reaction time was 2 hours, the temperature was 200 °C, and the pressure was 50 bar. The properties of the hydrogenated dimer and the bottom product are presented in table 6 below. Example 4 Regeneration of the catalyst In the dimerization according to example 2 a), the hydrocarbon stream was stopped after 96 hours. There ifter, the catalyst was purged with nitrogen stream (30 1/h) for 1.5 hours at 200 °C At the beginning of the regeneration, the nitrogen stream was replaced with a stieam of synthetic air (8 1/h). The reactor was heated from 200 °C to 500 °C with the temperature elevation rate of 1.5 °C/min. The regeneration was continued fcr 2 hours at 500 °C. Then, the temperature was re- duced again to 200 °C, and the air stream was replaced with nitrogen stream (30 1/h) for 1 hour. Dimerization was carried out using the regenerated catalyst in a flow reactor (120 °C) as descr.bed in example 2 a). The results are shown in table 7 below. Example 5 Dimerization in a flow reactor using a mesoporous catalyst A flow reactor was packed wi ;h a fresh mesoporous H-MM-4MW22-2A1 catalyst (mesoporous molecular sieve embedded with zeolite, the production of which is described in the patent application FI 20041675) (5 g) having an aluminium con- tent of 2.2 % by weight, number of acid sites of 180 urnol/g, and mesoporous sur- face area of > 700 m2, diluted with silicon carbide (support) (V (catalyst) : V (SiC) = 1:3). Dimerization of 1-hexadecene was performed using the mesoporous H-MM-4MW22-2A1 catalyst n the flow reactor (120 °C) as described in the ex- ample 2 a). The results are shewn in the table 8 below. Example 6 Regeneration of the H-MM-4.V1W22-2A1 catalyst The H-MM-4MW22-2A1 catalyst used in example 5 was regenerated using the regenerating treatment described in example 4. Dimerization of 1-hexadecene was performed using this mesopor jus H-MM-4MW22-2A1 catalyst in the flow reactor (120 °C) as described in the e? ample 2 a). The results are shown in table 9. Example 7 Hydrogenation of the dimerizs tion produces, and properties of the base oil product Products from examples 6 and 7 were pooled (1448 g). From the pooled product, unreacted Ci6 fraction (980 g£ was separated by distillation, and from the bottom product containing dimers (4d2 g) middle fraction (6 g) was separated. The bot- tom product containing dimeis was hydrogenated according to example 3. Table 10 shows the composition and properties of the hydrogenated bottom product. Example 8 Dimerization of 1-decene in a catalytic distillation reactor Dimerization of 1-decene was carried out in a catalytic distillation reactor where the catalyst was placed inside a distillation column, the amount of the catalyst being 4 % by weight based on the amount of 1-decene. An aluminium silicate catalyst (Nikki Chemical Co. Ltd.) was used as the catalyst. Figure 4 graphically shows the comparison of the reactions either carried out in a batch reactor at a temperature of 120 °C, or in a catalytic distillation reactor at the same tempera- ture, respectively using the same catalyst / feedstock level. As can be seen from figure 4, with the same conversion, a dimer selectivity of more than 90 % may be attained using a catalytic distillation apparatus, while selectivities in the batch experiment were below 80 % at identical temperatures.. Example 9 Effect of pressure on the dimer yield in a catalytic distillation apparatus Dimerization of 1-dccene was carried out in a catalytic distillation apparatus where the catalyst was placed inside a distillation column, the catalyst being amorphous aluminium silicate (Nikki Chemical Co. Ltd.), the amount thereof be- ing 4 % by weight based on the amount of 1-decene. The conversion of 1-decene was raised to the desired level by adjusting the pressure of the catalytic distillation apparatus. During the experiment, the pressure was adjusted to values ranging between 0.17 and 1 bar, the temperature at the bottom of the distillation column elevating from 80 to 300 °C. Dimer selectivity for all conversion levels was >80 %. Base oil yield was 92 % for the decene conversion of 94 %. Example 10 Dimerization of 1-hexadecene in a catalytic distillation reactor Dimerization of 1-hexadecene was performed in a catalytic distillation apparatus where the catalyst was placed in a distillation column, the amount of the catalyst being 6.5 % by weight of 1-hcxadecene. The catalyst was amorphous aluminium silicate (Nilcki Chemical Co. Ltd.). The pressure during the experiment was 0.002 bar, the temperature at the bo torn being elevated from 130 °C to 235 °C as the dimerization reaction proceeded. Dimer selectivity for all conversion levels was >80 %. Base oil and dimer yie ds were 99 % and 82 %, respectively, for the hexa- decene conversion of 99.3 %. Claims /PCI7FI2007/05035'V Amended 14.4.2008 1. Process for pro iucing polyolefms, characterized in that a feedstock comprising at lea:;t one n-olefin selected from the group consisting of C8-C30 n-olefins or a mixture of the n-olefins is dimerized at a tempera- ture ranging between 25 and 200 °C and at a pressure ranging between 0.001 mbar and 50 bar in the presence of a solid acidic catalyst by pass- ing the feedstock to a catalytic distillation apparatus comprising either a) a combination of a distillation column and a reactor comprising at least one catalyst bed, or b) a distillation column connected to one or more side reactors comprising at least one catalyst layer comprising solid acidic catalyst material with a mesoporous surface area of more than 100 m2/g, an aluminium content between 0.2 and 30 % by weight, the amount of the acid sites of the material ranging between 50 and 500 umol/g, and the naterial being selected from the group consisting of amorphous alumiriumsilica*es, zeolites, and mesoporous materials con- taining silicon anc aluminium, recovering the unreacted n-olefin at the upper part of the distillation column or at the upper part of the combi- nation of the disti lation column and the reactor as a side stream to be combined with tht feedstock, and the reaction product from the dimeri- zation is hydrogenated. 2. Process accord ng to claim 1, characterized in that the feedstock comprises at least one n-olefin selected from the group consisting of 1- decene, mixtures of decenes, 1-dodecene, mixtures of dodecenes, 1- hexadecene, mixtures of he:xadecenes, 1-tetradecene, 1-octadecene, mixtures of octadesenes, and C20-C22 1-olefins and mixtures thereof. 3. Process accordiig to claim 1 or 2, characterized in that the catalyst layer comprises solid acidic catalyst material selected from the group consisting of acidic amorphous aluminium silicates, dealuminated Y zeolites, mesoporous molecular sieves imbedded with zeolite, said zeo- lite being preferably ZSM-5, beta-zeolite or MCM-22, whereas the mesoporous mole;ular sieve is MCM-41 with a regular molecular structure. 4. Process according to any one of claims 1-3, characterized in that the mesoporous sirface area of the catalyst material is more than 300 m2/g. 5. Process accordi lg to any one of claims 1-4, characterized in that catalyst material i.; amorphous aluminium silicate with a mesoporous surface area of more than 300 m2/g. 6. Process according to any one of claims 1-5, characterized in that part of the side stieam is removed from the distillation column or the combination of the distillation column and the reactor at the upper part thereof. 7. Process according to any one of claims 1 - 6, characterized in that a solvent or a solvent mixture selected from the group of hydrocarbons is added to the feedstock. 8. Process according to any one of claims 1 - 7, characterized in that the feedstock is subjected to drying. 9. Process according to any one of claims 1 - 8, characterized in that the dimerization is performed at a temperature ranging between 50 and 150 °C, and at a prjssure ranging between 0.5 bar and 30 bar. 10. Process accord ng to any one of claims 1 - 9, characterized in that the hydrogenation is performed at a pressure ranging between 5 and 100 bar, preferably between 10 and 70 bar, and at a temperature rang- ing between 100 md 400 °C, preferably between 150 and 250 °C, in the presence of a 1' ydrogenation catalyst. 11. Process according to any one of claims 1 - 10, characterized in that the feedstock comprises I -decene. 12. Process according to any one of claims 1 - 10, characterized in that the feedstock comprises -hexadecenc. 13. Process according to any one of claims 1 - 12, characterized in that the feedstock comprises olefins of natural origin produced from biological starting materials. 14. Process according to any one of claims 1 - 13, characterized in that distillation is carried out after the dimerizalion or the hydrogena- tion or after both. 15. Process according to any one of claims 1 - 14, characterized in that base oil or bass oil component is produced with said process. 16. Use of an apparatus comprising 1) an optional drying means for the feedstock, 2) a catalytic distillation apparatus comprising either a) a combination of a distillation column and a reactor comprising at least one catalyst bed of a solid and acidic catalyst, or b) a distillation col- umn connected to one or more side reactors comprising at least one catalyst bed of a solid and acidic catalyst, 3) hydrogenation reactor and optionally a distillation column downstream of the catalytic distillation apparatus and/or hydrogenation reactor, in a process according to any of the claims 1 - 16 The present invention relates to a process for producing polyolefms wherein a feedstock comprising n-olefin or n mixture of nolofins is dimorized in the pres ci ce of a solid noidic catalyst by passing the feedstock to a catalytic distillation apparatus comprising either a) a combination of distillation column and a reactor comprising at least one catalyst bed, or b) a distillation column connectcd to one or more sidde reacors comprising at least one catalyst layer, recovering the unre- acted n olefin from the distillation column or the combination of the distillation solumn and this reactor at the upper port thereof as sidc-streain to be coinbined with the feedstock, and the reactor product from the dimerization is hydrozen- aled.

Full Text

Process for the manufacture of olefin oligomers
Technical field
The invention relates to the preduction of high-grade base oils and to selective
dimerization of n-olefins usinj; a solid and acidic catalyst. Particularly, the inven-
tion is directed to a process wherein n-olefins are dimerized in a catalytic distilla-
tion apparatus, followed by hy Irogenation to give polyolefins.
State of the art
Saturated olefinic oligomers are a significant group of high-grade synthetic base
oils. Poly-alpha-olefins known as PAOs are typically produced by oligomerization
of alpha-olefins in the presence of homogsneous Friedel-Crafts catalysts, such as
boron trifluoride (BF3) and a promoter at slightly elevated BF3 pressures, and
temperatures below 100 °C. Water or an alcohol normally serves as the promoter.
In the PAO process, 1 -decent¦ is typically used as the feedstock, mainly giving
trimers and tetramers of the feedstock olefin as the product.
Base oils may also be produced by dimerization of n-olefins heavier than decene.
Base oils of the PIO (poly(intt rnal olefin)) group are produced by dimerization of
internal n-olefins, typically C-15-C16 n-clefins, using BF3 catalysts. Among the
products of the PIO process, fimers are particularly suitable feedstocks for base
oil production.
A catalyst separation step is ..lways necessary in PAO and PIO processes using
homogeneous catalysts.
As is known, n-olefins refer to linear olsiins or linear olefins with no more than
one branch, that is, to slightly oranched olcfins.
Oligomerization refers to a reaction where molecules of at least one type react
with each other resulting in the increase of the molecular weight, said increase
being the added molecular we ight of at least three molecules. Oligomerization
may be illustrated with the following equation:

where A may be identical with or different from B, x is 0 or an integer, y is 0 or
an integer, x+y > 2, Sx = n, Sy = m, A and B are olefmic molecules, and n and m
are integers. The term oligomtr refers to a. repetitive combination of monomeric
units of at least the one type, this number of said units ranging from 3 to 100.
The term dimerization refers to a reaction where molecules of at least one type
react with each other resulting in the increase of the molecular weight, said in-
crease being the added molecular weight of at least two molecules. Dimerization
may be illustrated with the following equation:

where A may be identical with or different from B, A and B being olefmic mole-
cules. The term dimer refers to a combination of two monomeric units of at least
one type.
The term polyolefin refers to a combination comprising at least two olefmic
monomeric units of at least one type.
Several alternative heterogene jus catalysts for the dimerization of heavy olefins
are known e.g. from patents LS 4,417,088, US 5,053,569, US 5,453,556 and US
6,703,356. In addition to the dssired dimerization reaction, the use of said hetero-
geneous acid catalysts disclose d in said documents resulted in cracking and isom-
erization reactions of the feedstock olefins, as well as formation of heavier oli-
gomers, mainly trimers and tetamers. Particularly detrimental isomerization reac-
tions of the feedstock olefins i lclude reactions yielding products that may not be
dimerized, such as naphthenes. Attainable yields of the base oil are lowered by the
cracking products formed as well as naphthenes resulting from the isomerization
of monomeric olefins. Because relatively valuable olefin are used as the process
feedstock, any undesirable side reactions tiereof also have a considerable impact
on the feasibility of the process.
The document US 5,053,569 discloses the dimerization of a-olefins using an
acidic calcium/montmorillonite catalyst. The document US 6,703,356 discloses an
oligomerization process of alpha-olefins using alpha-olefins with a carbon number
ranging between 10 and 30, or a mixture thereof as the feedstock. The catalyst is
specified as a crystal catalyst "laving a constraint index of less than 3. The con-
straint index is a measure of the product selectivity of the catalyst. In the exam-
ples, 1-hexadecene, l-tetradeene and 1-octadecne are used as feedstocks,
whereas MCM-22, MCM-56, USY, Beta, ZSM-12 and WOx/on ZrO2 are used as
catalysts.
In the examples of the above documents US 5,053,569 and US 6,703,356, highest
conversions attained in the dimsrization of heavy olefins using heterogeneous acid
catalysts are 92 % and 87 %. A conversion of clearly below 100 % may be attrib-
uted to the formation of naphtr enes from feedstock olefins. The product obtained
in the examples of the documeit US 5,053,569 contained trimers and heavier oli-
gomers in a total amount of 47 %, which i,3 not desirable for the properties of the
base oil product. Selectivity fo • dimers is higher (70 %) in the example described
in the document US 6,703,3.66, having a high conversion (87 %), yet high
amounts of cracked products (4 %) were obtained.
The term catalytic distillation refers generally to the combination of a chemical
reaction with the product separation. The rsaction and product separation are car-
ried out together in an inseparable manner. A catalytic distillation apparatus nor-
mally comprises a distillation ;olumn incorporating one or more catalytic zones.
In said catalytic zones, streams from a specified level or plate of the distillation
column are treated to give desired reaction products. Thereafter, the product
stream is fractionated using it distillation means. Said catalytic zone may be
placed within or outside the distillation means. In industrial applications, catalytic
distillation is used in the prodi ction of ethers, said process being also known e.g.
for dehydration of alcohols, and oxydation of paraffins.
Use of the catalytic distillation is also known in the oligomerization process. US
4,935,577 discloses an oligomerization process wherein alpha-olefins having from
3 to 12 carbon atoms are passed to a distillation column containing Lewis acid
catalyst for the reaction. The temperature of the catalytic distillation apparatus is
no more than about 150 °C, ths typical operation temperature range being below
50 °C. A combination catalysi is used in the distillation apparatus, and accord-
ingly, a unit for the separation of the Lev/is acid prior to recycling to the distilla-
tion apparatus is an essential part of the apparatus
US 2,198,937 discloses an apparatus comprising a distillation column and a side
reactor for the polymerization of hydrocarbons. The apparatus may be operated
under conditions similar to those in catalytic distillation.
The document FI 96852 disclc ses a process and an apparatus for oligomerizing
olefins. In this process, C3-C20 )lefins or mixtures thereof are passed to a catalytic
distillation system where the f :edstock olefins are contacted with a catalyst at a
temperature of above 150 °C, hus yielding a product containing oligomers. The
catalyst used in this process manly consists of zeolite, the products formed being
middle distillates and lubricant'. The catalytic distillation system used in the proc-
ess may also be a distillation column for the product separation, connected to at
least one side reactor containing the catalys:.
Based on above teachings it may be seen that there is an obvious need for a novel
improved process for the production of polyolefins from n-olefms, said novel
process eliminating or at least s ubstantially reducing problems and deficits associ-
ated with the solutions of the state of the art.
Objects of the invention
An object of the invention is to provide a process for the production of polyolefins
from n-olefms.
Another object of the inventior is to provide a process for the production of poly-
olefins from n-olefms using a s slid acidic catalyst.
Still another object of the invention is to provide a process for the production of
polyolefins from C8-C30 n-olefms using a solid acidic mesoporous catalyst.
Further, an object of the inven:ion is to provide a process for the dimerization of
C8-C30 n-olefms using a solid acidic mesoporous catalyst to give polyolefins.
An object of the invention is al.ro a process for producing a base oil component.
Yet another object of the invention is the use of a catalytic distillation apparatus
for the dimerization of n-olefins, particularly C8-C30 n-olefins,
An object of the invention is a. so to provide di-n-olcfins consisting of two identi-
cal or different n-olefms having carbon chain lengths of C8-C30.
Summary of the invention
The present invention relates to a process for producing high-grade products use-
ful as base oils and base oil components from n-olefins using a solid and acidic
catalyst, by dimerizing an n-olrfin or a mixture of n-olefms in a catalytic distilla-
tion apparatus, wherein the catalyst is placed into a distillation column or into a
side reactor outside the distillation column, followed by hydrogenation of the
product. In the process of the invention, the: feedstock comprising an n-olefin or a
mixture of n-olefms, is dimerzed in the presence of a solid acidic catalyst by
passing the feedstock to a catalytic distillation apparatus either comprising a) a
combination of a distillation column and a reactor having at least one catalyst
layer, or b) a distillation column connected to one or more side reactors having at
least one catalyst layer, recovering the unreacted n-olefin at the upper part of the
distillation column or the combination of the distillation column and the reactor as
a side-stream to be combined with the feedstock, whereas the reaction product
from the dimerization is hydrogenated. Impurities still present in the dimerization
product or final product may optionally be removed using an additional distilla-
tion step.
The invention and some alternative embodiments thereof are illustrated by ap-
pended figures 1, 2 and 3, however, without wishing to limit the invention to these
presented embodiments.
Figures
Figure 1 schematically shows an embodiment of the invention for producing a
base oil component.
Figure 2 schematically shows an embodiment of the invention wherein the
dimerization of n-olefins is car'ied out in a catalytic distillation apparatus with the
catalyst in separate side reactois.
Figure 3 schematically shows an embodiment of the invention wherein the
dimerization of n-olefins is performed in a combination of a distillation column
and a reactor.
Figure 1 is a schematic presentution of the basic solution of the inventive process.
The feedstock of the process, 01 stream 1, containing C8-C30 n-olefins or a mixture
thereof is obtained from a feed tank (not shown in the figure). Stream 4 recovered
at the upper part of a distillatioi column B is combined with stream 1. Streams 1
and 4 together form the stream 2, which is passed to side reactor A containing the
dimerization catalyst layer / bed D. Stream 3 is obtained as the product of the side
reactor A, said stream 3 containing C8-C30 n-olefm monomers and as reaction
products mainly dimeric produsts. Stream 3 is passed to a distillation column B
where C8-C30 n-olefm monomers rise to the upper part of said distillation column
B, followed by recycling the n-olefm monomer fraction thus obtained as stream 4
to the side reactor A in the reaction stage. Stream 5 is removed at the top of the
distillation column B, said stream 5 containing components of non-dimerizable
monomer fraction in the upper part of the distillation column. With stream 5 the
accumulation of the componei.ts of non-dimerizable monomer fraction such as
branched compounds is prevented in the catalytic distillation apparatus. The prod-
uct in the form of a product stream 6 mainly consisting of dimers is obtained from
the bottom of the distillation column B, said stream 6 being passed to a hydro-
genation reactor C. In the hydrogenation reactor C, olefins arc hydrogenated to
give the product, suitable as a base oil component, followed by the removal of
said product from the hydrogen ition reactor C as the product stream 7.
In figure 2, a preferable soluticn of the process of the invention is shown. In this
solution, a catalytic distillatioi apparatus having dimerization catalyst layers D
and E placed in separate side re actors A and B is used. The feedstock of the proc-
ess, or stream 1, containing C8-C30 n-olefins or a mixture thereof is obtained from
a feed tank (not shown in the figure). Stream 4 recovered at the upper part of a
distillation column C is combined with stream 1. Streams 1 and 4 together form
the stream 2 being either passe i to a side rsactor A or side reactor B. Stream 2 is
passed to the side reactor containing fresh catalyst or regenerated catalyst. Simul-
taneously, the other side reactor contains the exhausted catalyst, or the catalyst is
being regenerated. There are two side reactors in the system, and accordingly, it is
not necessary to shut down the process due to lowered catalytic activity. Stream 3
is obtained as the product of the side reactor A or B, said stream 3 containing Cs-
C30 n-olefin monomers and as i eaction products mainly dimeric products. Stream
3 is passed to a distillation column C where C8-C30 n-olefin monomers rise to the
upper part of said distillation column C, followed by recycling the n-olefin
monomer fraction thus obtainec as stream 4 to the side reactor A or B in the reac-
tion stage. Stream 5 is removed at the upper part of the distillation column C, said
stream containing components of non-dimerizable monomer fraction at the upper
part of the distillation column C. Accumulation of the components of non-
dimerizable monomer fraction such as branched compounds in the catalytic distil-
lation apparatus is prevented by stream 5. From the lower part of the distillation
column C, a product stream 6 mainly consisting of dinners, subsequently passed to
a hydrogenation reactor for hydrogenation (not shown in the figure), and a bottom
product 7 are obtained.
Figure 3 shows another preferable solution of the inventive process, using a cata-
lytic distillation apparatus having the dimerizing catalyst layer B placed inside the
distillation column A. The feed stream 1 comprising C8-C30 n-olefms or a mixture
thereof is combined with the n-olefin monomer fraction 4 from the top of the dis-
tillation column A to give a stieam 2 for passing to the distillation column A, to
the upper part of the catalyst la/er B. In the distillation column A, the feed passes
to the catalyst layer B, in which the dimerization reaction mainly proceeds. In the
catalyst layer B, n-olefin monomer is in the form of a vapour/liquid mixture and
the liquid n-olefin monomer ashes the dimers and oligomers formed from the
catalyst, n-olefin monomer fration 3 is obtained from the top of the distillation
column A, said fraction being plit to give a monomer fraction 4 to be recycled to
the reactor, and a monomer steam 5. Accumulation of non-dimerizable compo-
nents of the monomer fraction in the catalytic distillation apparatus is prevented
by removal of said monomer siream 5. The product in the form of a dimer stream
6 is obtained from the bottom of the distillation column A, said stream 6 being
passed to a hydrogenation read or (not shown in the figure) for hydrogenation.
The catalytic distillation apparatus comprises either a) a combination of a distilla-
tion column and a reactor having at least one catalyst layer, or b) a distillation
column connected to one or more side reactors having at least one catalyst layer.
Detailed description of the in mention
It was surprisingly found that the use of a catalytic distillation apparatus with a
solid acidic catalyst allows for the selective dimerization of n-olefins to give base
oil useful as a lubricant, or base oil component, with an excellent yield that may
even be above 95 %. In the process of the invention, the dimerization of n-olefins
or n-olefin mixtures to yield d. mers is perlbrmed in a catalytic distillation appara-
tus comprising either a) a combination of a distillation column and a reactor com-
prising at least one catalyst layer, or b) a distillation column connected to one or
more side reactors comprising at least one catalyst layer, said catalyst being in the
side reactor outside the distillation column. In this way, the temperature of the
dimerization reaction may be maintained low, typically below 150 CC, thus pre-
venting the formation of naplithenes and cracking of products. Dimerization is
followed by hydrogenation to produce the base oil component. Optionally after
the dimerization and/or hydroj enation, the product may, if necessary, be passed to
a distillation apparatus where my monomer residues are removed from the dimer-
ized product, or the dimers are separated from heavier trimers and tetramers.
The quality of the base oil and'or base oil component obtained as the product use-
ful as a lubricant, after hydrogenation is excellent due to, among other things, low
number of undesirable reactions.
In the process of the inventio 1, n-olefin or a mixture of n-olefins is passed to a
catalytic distillation apparatus where the single conversion of n-olefins may be
suitably controlled. The term single conversion refers to the conversion of an n-
olefm to give another compound during the reaction stage. Single conversion of n-
olefin monomers is defined as follows:
Single conversion of n-olefin monomers (%) = 100 x (proportion of n-olefins in
the feedstock prior to the reaction stage proportion of n-olcfins in the product
after the reaction stage) / propo/tion of n-olofins in the feedstock prior to the reac-
tion stage.
In the process of the invention, it is preferable to use recycling of the monomeric
olefins particularly in embodiments where the catalyst is placed in a side reactor.
In this case, total conversion and dimerizE.tion selectivity are improved by effi-
cient recycling of the monomeric olefins. A requirement for the efficient recycling
of the monomeric olefins is the adjustment of the reaction conditions to values
preventing undesired side reactions such as cracking and isomerization of the
monomer resulting in naphthcnes. Moreover, relatively low single conversion
between 5 and 50 % is anotier requirement for high dimerization selectivity
meaning selectivity ranging approximately between 80 and 100 %. In the process
of the invention, high total con/ersion, ranging between 95 and 100 %, is attained
with efficient recycling, in spile of the relatively low single conversion. Efficient
recycling means that the exten of recycling is such that between 50 and 95 % of
the original monomer stream contacts the catalyst layer more than once. Extent of
recycling refers to the level o1 the stream passed from the separation step to the
catalyst layer. Dimensioning of the components of the apparatus, and thus the
investment costs depend on tie level of said stream, and therefore, reasonable
extent of recycling is preferab' e for said investment costs. Also energy consump-
tion of the process is influence! by the extent of recycling.
Single conversion is adjusted :o a suitable value on the basis of the carbon num-
ber. Suitable single conversion means that the extent of side reactions such as
formation of heavier oligomer! and cracking is adjusted to be low, that is less than
10 % of the total conversion of the monomeric olefms is due to side reactions.
Reduced formation of heavy c ligomers has a positive impact on the properties of
the base oil product, and retarc s the deactivation rate of the catalyst.
The process of the invention is now described in more detail. In the process, the
feedstock is passed to a catalyst layer in a catalytic distillation apparatus, either
above the catalyst layer in a distillation column where the feedstock migrates to
the catalyst layer, or to a catalyst layer in a side reactor, followed by passing the
reaction mixture to a distillation column or to a lower part of the distillation col-
umn where the monomer fract on migrates to a higher part of the distillation col-
umn, and any impurities present in the monomer lighter than the monomers in the
distillation range leave the di; filiation column at the top thereof or may be re-
moved in a subsequent separat on unit downstream of the distillation column, and
any impurities present in the n.onomer heavier than the monomers in the distilla-
tion range are removed from the column at the bottom or may be removed from
the bottom product in a separat; separation unit.
The dimerization reaction and some simultaneous oligomerization proceed in the
catalyst bed. A monomeric friction such as n-hexadecene with a boiling point
between 280 and 290 °C / 1 atm, rises in the column as the lighter component
upwards from the feeding plate and is removed from the distillation column, typi-
cally at the top thereof as a s :de stream to be recycled and combined with the
feedstock. The dimer produced as the raw product, e.g. C32 olefin, migrates
downward in the distillation column, followed by separation of the dimer stream
at a lower part thereof. Oligomerization product such as C48+ olefin is obtained as
the bottom product stream, In case unreacted olefins and side reactions such as
formation of naphthenes corresponding to monomers with respect to their carbon
number may not be totally avoided, accumulation thereof in the process streams
may, however, be prevented removing part of the side stream to be recycled at
the upper part of the distillation column.
A sqlvent or solvent mixture may be added to the feedstock, thus retarding deacti-
vation of the dimerization catalyst. The solvent is selected among hydrocarbons
such as n-paraffins, isoparaffns and aromatic solvents. The solvent may be re-
moved from the distillation column as a respective side stream via cooling to a
separation unit where the solvent is separated from the reaction product. The sol-
vent may also be recycled in the catalytic distillation apparatus without a separate
separation unit.
The feedstock may be optional y dried for removal of any water and other impuri-
ties present therein. Drying may be carried out with known drying means such as
with commercially available Secular sieves e.g. zeolite 3A molecular sieve used
for drying hydrocarbons, or using other suitable known methods.
Process feedstock comprises at least one n-olefin selected from the group consist-
ing of C8-C30n-olefins. Examples of suitable n-olefins include 1-decene, mixture
of decenes, 1-dodecene, mixture of dodecenes, 1-hexadecene, mixture of hexade-
cenes, 1-octadecene, mixture of octadecemss, and C20-C22 1-olefins, preferably 1-
hexadecene, or a mixture of h;xadecenes. n-Olefins may be synthetic olefins or
olefins of biological origin, pioduced from biological starting materials such as
vegetable oil and animal fats.
In case a catalytic distillation a jparatus where the catalyst is placed in a side reac-
tor outside the distillation coluinn is used in the dimerization step of the process of
the invention, the feedstock such as 1-hexadecene is directly introduced into the
catalyst layer in the side reacto:1 where a dimer is produced. The product stream of
said side reactor is passed to 1 distillation column where the monomer fraction
such as n-hexadecene having a boiling point between 280 and 290 °C / 1 atm will
rise upward as the lighter component in the distillation column from the feeding
plate and is removed, typically at the top of the column as a side stream and recy-
cled to the side reactor. The dimer formed as the reaction product such as C32 ole-
fin migrates downward in the distillation column, the dimer stream being sepa-
rated at a lower part thereof. an oligomerization product such as C48+ olefin is
obtained as the bottom product stream.
The diiner product obtained above and the bottom product produced as the by-
product are hydrogenated in a hydrogenation step operated continuously or batch
wise in the presence of hydrogsn. Known, hydrogenation catalysts containing met-
als from the groups VIII and/cr VIA of the Periodic System of the Elements may
be used. Preferable hydrogen ition catalysts include supported Pd, Pt, Ni, Cu,
CuCr, NiMo or CoMo catalys" s, the support being preferably alumina and/or sil-
ica. The hydrogenation step is carried out at a pressure of 5 to 100 bar, preferably
at 10 to 70 bar, and at a temper iture of 100 to 400 °C, preferably at 150 to 250 °C.
In the dimerization step of the process of the invention, the temperature in the
distillation column or in the catalyst layer of the side reactor ranges between 25
and 200 °C, preferably between 50 and 150 °C. Pressure range varies according to
the carbon number of the feedi tock and reaction temperature, and further, the lo-
cation of the catalyst either in the catalytic distillation apparatus or in a side reac-
tor has an impact on the pressure range. The pressure range thus varies from re-
duced pressures to elevated pressures. The pressure may vary between 0.001 mbar
and 50 bar, preferably 0.5 bar und 30 bar, particularly preferably between 1 mbar
and 20 bar. WHSV is properly adjusted on the basis of the single conversion and
deactivation rate of the catalyst. WHSV typically ranges between 0.1 and 50 h"1
with respect to the feedstock, pi eferably between 0.5 and 20 h'1.
Selection of the pressure of the dimerization reaction as a function of the tempera-
ture is illustrated in the table 1 below. Table 1 shows the vapour pressure of 1-
decene as a function of the ten.perature. In case the desired reaction temperature
in the catalyst layer of the cataytic distillation apparatus is 50 °C, the pressure is
selected to be 11 mbar, or in case the desired temperature in the catalyst layer of
the catalytic distillation apparatus is 150 CC, the pressure is selected to be 580
mbar.
If a catalytic distillation apparatus where the catalyst is placed in a side reactor is
used, the pressure of the reactor is selected to give the product mixture in liquid
phase.
High-grade base oil or a base oil component may be produced by the process of
the invention from n-olefins us .ng a solid aid acidic catalyst not requiring parts of
the catalyst to be subsequently separated and recycled to the distillation apparatus.
Moreover, solid acidic catalysis always result in isomerization of double bonds.
Particularly suitable solid acidic catalysts include catalysts having a mesopore
surface area of above 100 m2/g as measured by nitrogen adsorption and calculated
with the following BJH equation:

Suitable solid acidic catalyst materials include materials having a mesopore sur-
face area based on the BJH eq lation of more than 100 m2/g, preferably more than
300 m2/g. Such materials comprise amorphous aluminium silicates, preferably
acidic amorphous aluminium silicates, and particularly preferably acidic amor-
phous aluminium silicates with Bronsted acidic sites; zeolites, preferably dealu-
minated Y zeolites; and mesoporous materials with a regular porous structure,
containing silicon and aluminium, among which mesoporous molecular sieves
with inserted zeolite are preferable. The zeolite is preferably ZSM-5, beta-zeolite
or MCM-22, whereas the mes jporous molecular sieve is MCM-41 with a regular
molecular structure. The acidic solid catalyst material is particularly preferably a
mesoporous molecular sieve wiih imbedded MCM-22 zeolitic structures.
The expression mesoporous malecular sieve imbedded with zeolite refers to a
catalyst having mesoporous molecular sieve structure and zeolitic structure in the
same material and where said mesoporous molecular sieve structure and zeolitic
structure are bonded together tt rough a chemical bond. The mesoporous molecu-
lar sieve with imbedded zeolite, and the production thereof is disclosed in the pat-
ent application FI 20041675.
The aluminium content of the catalyst material varies between 0.2 and 30 % by
weight as determined by alum inium content assays, while the level of the acid
sites varies between 50 and 50) umol/g as determined by the analysis according
to the NH3-TPD method. Suitable catalyst materials have BrBnsted acidity, the
level thereof being more than 1) umol/g as measured by the proton NMR method.
The mesoporous surface area c f the catalyst material as calculated with the BJH
equation is more than 100 m2/g preferably more than 300 m2/g.
The catalyst also comprises a support material to provide a readily mouldable
catalyst with mechanical resists nee. Said support material is typically an inorganic
oxide such as alumina or silica.
In the process of the invention the olefins and isomers of the olefins in the feed-
stock are preferably recycled, ind the dimsrized and oligomerized fractions thus
produced are obtained as products from the bottom or as a side stream. In this
manner, the total conversion h the process is high, preferably more than 95 %,
and more than 99 % at best.
An apparatus comprising 1) an optional drying means for the feedstock, 2) a cata-
lytic distillation apparatus comprising either a) a combination of a distillation col-
umn and a reactor having at bast one catalyst layer, or b) a distillation column
connected to one or more side reactors having at least one catalyst layer, 3) hy-
drogenation reactor and optimally a distillation column upstream and/or down-
stream of the hydrogenation reactor, is suitable for the process of the invention,
and particularly for dimerizaticn of Cs-Cao n-olefins.
Regeneration of the deactivat sd dimerization catalyst may be performed in the
same reactor as the dimerizat on reaction. The regeneration is carried out at an
elevated temperature using a gaseous mixture that may contain oxygen. During
the regeneration, the temperature is equal or higher than that of the dimerization
reaction. Thermal resistance of the catalyst and the feasibility of the process are
influenced by the maximum legeneration temperature, lower regeneration tem-
peratures being preferable with respect to energy consumption. However, the re-
generation temperature should be sufficiently high for the removal of coal formed
and any impurities adsorbed in the catalyst.
The process of the invention I; endowed with several advantages. The process is
continuous, and thus continuoos production of the base oil component is possible
without interruptions typical for batch reactors.
In the process of the inventior, the reactor feedstock may be adjusted as desired
since the composition may be controlled very precisely at the side exit of the dis-
tillation column. It is thus poss ible to provide the reactor with a stream containing
only a monomer fraction without any lighter or heavier fractions due to possible
unselective reactions, said lighter or heavier fractions yielding products with un-
desirable carbon numbers durir g the reaction stage.
Desired compositions are obtai led as the reactor products since the outlet sites for
the products may be selected. This is particularly useful in cases where a mixture
of n-olefins having different carbon numbers serves as the feedstock.
It was surprisingly found that selective dimerization of olefins is provided at a
reaction temperature below If 0 °C. Selective dimerization means here a dimeriza-
tion selectivity of more than 80 %. Dimer yields are remarkably high since any
unreacted monomer present in the reaction product may be recycled to the feed-
stock for the reactor by adjusting the recycle ratio.
Reactive components may betimmediately separated in the distillation column,
and thus side reactions detrir ental to the process may be quickly stopped. Other
separation units e.g. two disti lation columns are not necessary in the process for
concentration of the feedstock and the reaction product since the concentration of
the product and fractionation if the feedstock may be carried out in the same col-
umn, which accordingly resultss in a simplers apparatus and lower investment costs.
In addition, the regeneration of the dimerization catalyst may be performed in the
same reactor as the reaction.
The invention is now illustraisd with the following examples without wishing to
limit the scope thereof.
Examples
Example 1 (comparative example)
Dimerization of olefins in a bitch reactor
Dimerization process was pjrformed in a batch reactor using 1-hexadecene
(Neodene 16®). The process temperature was 200 °C, the reactor pressure being
20 bar. The amount of the fe sdstock was 50 g, and that of the catalyst 2 g. The
total reaction time was 24 hours. Commercially available Y zeolite (TOSOH Co.),
beta-zeolites (TOSOH Co,) aid a mesoporous material MCM-41 (produced ac-
cording to a method disclosed in Catalysis Letters 36 (1996) 103) were used as
catalysts. Total Cu conversion (= conversion of C16 hydrocarbons to give prod-
Dimerization of 1-decene was also performed in a batch reactor. The process tem-
perature was 120 °C, the realtor pressure being 20 bar. Commercially available
aluminium silicate catalyst Nikki Chemical Co. Ltd) was used as the catalyst, in
an amount of 4 % by weight of the feed. The results are shown in the figure 4. The
results indicate that the dimsrization selectivity was clearly lower in the batch
reactor in comparison to the selectivity of the process of the invention.
Example 2
Dimerization in a system corrssponding to a catalytic distillation apparatus, com-
prising dimerization in a flow reactor and distillation
a) Dimerization in a flow reac or
1-Hexadecene (Neodene 16®] was introduced at a rate of 10 g/h into a flow reac-
tor (1 bar (a)). The flow reactcr was packed with 5 g of aluminium silicate catalyst
(Nildci Chemical Co. Ltd) wth an aluminium content of 13 % by weight, the
number of acid sites of 120 xmol/g, and mesoporous surface area > 300 m2/g,
diluted with silicon carbide (\ (catalyst) : V (SiC) = 1:3), followed by the dimeri-
zation reaction at a temperaturs of 120 °C in the reactor. The dimerization product
was collected into a product container. Results of the dimerization, that is, total
The conversion of the C16 hydrocarbons declined the longer the catalyst remained
in the hydrocarbon stream. Accumulation of heavy hydrocarbon oligomers in the
catalyst (coke formation) was reason to the reduced conversion.
b) Distillation of the products
The experiment of a) was refeated, and the products obtained from the experi-
ments (1766 g) were pooled. The unreactsd monomer fraction (1103 g), middle
fraction containing C20-C30 hy irocarbons (9 g), dimer fraction (452 g) and heavy
bottom product (199 g) were separated from said pooled products by distillation.
c) Recycling of the monomers ' dimerization
The monomer fraction obtained above was used as the feedstock in dimerization.
The distilled monomer fractior mainly consisted of internal C16 olefins. The com-
position of the distilled mono ner fractior. is shown in the table 4 below in the
form of surface area percentages as measured by GC-MS-analysis.
The dimerized product of example 2, separated by distillation, and the bottom
product of the distillation wers hydrogenated as separate batches in a batch reactor
using a heterogeneous nickel catalyst. In the hydrogenation, the reaction time was
2 hours, the temperature was 200 °C, and the pressure was 50 bar. The properties
of the hydrogenated dimer and the bottom product are presented in table 6 below.

Example 4
Regeneration of the catalyst
In the dimerization according to example 2 a), the hydrocarbon stream was
stopped after 96 hours. There ifter, the catalyst was purged with nitrogen stream
(30 1/h) for 1.5 hours at 200 °C At the beginning of the regeneration, the nitrogen
stream was replaced with a stieam of synthetic air (8 1/h). The reactor was heated
from 200 °C to 500 °C with the temperature elevation rate of 1.5 °C/min. The
regeneration was continued fcr 2 hours at 500 °C. Then, the temperature was re-
duced again to 200 °C, and the air stream was replaced with nitrogen stream (30
1/h) for 1 hour. Dimerization was carried out using the regenerated catalyst in a
flow reactor (120 °C) as descr.bed in example 2 a). The results are shown in table
7 below.

Example 5
Dimerization in a flow reactor using a mesoporous catalyst
A flow reactor was packed wi ;h a fresh mesoporous H-MM-4MW22-2A1 catalyst
(mesoporous molecular sieve embedded with zeolite, the production of which is
described in the patent application FI 20041675) (5 g) having an aluminium con-
tent of 2.2 % by weight, number of acid sites of 180 urnol/g, and mesoporous sur-
face area of > 700 m2, diluted with silicon carbide (support) (V (catalyst) : V
(SiC) = 1:3). Dimerization of 1-hexadecene was performed using the mesoporous
H-MM-4MW22-2A1 catalyst n the flow reactor (120 °C) as described in the ex-
ample 2 a). The results are shewn in the table 8 below.

Example 6
Regeneration of the H-MM-4.V1W22-2A1 catalyst
The H-MM-4MW22-2A1 catalyst used in example 5 was regenerated using the
regenerating treatment described in example 4. Dimerization of 1-hexadecene was
performed using this mesopor jus H-MM-4MW22-2A1 catalyst in the flow reactor
(120 °C) as described in the e? ample 2 a). The results are shown in table 9.

Example 7
Hydrogenation of the dimerizs tion produces, and properties of the base oil product
Products from examples 6 and 7 were pooled (1448 g). From the pooled product,
unreacted Ci6 fraction (980 g£ was separated by distillation, and from the bottom
product containing dimers (4d2 g) middle fraction (6 g) was separated. The bot-
tom product containing dimeis was hydrogenated according to example 3. Table
10 shows the composition and properties of the hydrogenated bottom product.
Example 8
Dimerization of 1-decene in a catalytic distillation reactor
Dimerization of 1-decene was carried out in a catalytic distillation reactor where
the catalyst was placed inside a distillation column, the amount of the catalyst
being 4 % by weight based on the amount of 1-decene. An aluminium silicate
catalyst (Nikki Chemical Co. Ltd.) was used as the catalyst. Figure 4 graphically
shows the comparison of the reactions either carried out in a batch reactor at a
temperature of 120 °C, or in a catalytic distillation reactor at the same tempera-
ture, respectively using the same catalyst / feedstock level. As can be seen from
figure 4, with the same conversion, a dimer selectivity of more than 90 % may be
attained using a catalytic distillation apparatus, while selectivities in the batch
experiment were below 80 % at identical temperatures..
Example 9
Effect of pressure on the dimer yield in a catalytic distillation apparatus
Dimerization of 1-dccene was carried out in a catalytic distillation apparatus
where the catalyst was placed inside a distillation column, the catalyst being
amorphous aluminium silicate (Nikki Chemical Co. Ltd.), the amount thereof be-
ing 4 % by weight based on the amount of 1-decene. The conversion of 1-decene
was raised to the desired level by adjusting the pressure of the catalytic distillation
apparatus. During the experiment, the pressure was adjusted to values ranging
between 0.17 and 1 bar, the temperature at the bottom of the distillation column
elevating from 80 to 300 °C. Dimer selectivity for all conversion levels was >80
%. Base oil yield was 92 % for the decene conversion of 94 %.
Example 10
Dimerization of 1-hexadecene in a catalytic distillation reactor
Dimerization of 1-hexadecene was performed in a catalytic distillation apparatus
where the catalyst was placed in a distillation column, the amount of the catalyst
being 6.5 % by weight of 1-hcxadecene. The catalyst was amorphous aluminium
silicate (Nilcki Chemical Co. Ltd.). The pressure during the experiment was 0.002
bar, the temperature at the bo torn being elevated from 130 °C to 235 °C as the
dimerization reaction proceeded. Dimer selectivity for all conversion levels was
>80 %. Base oil and dimer yie ds were 99 % and 82 %, respectively, for the hexa-
decene conversion of 99.3 %.
Claims /PCI7FI2007/05035'V Amended 14.4.2008
1. Process for pro iucing polyolefms, characterized in that a feedstock
comprising at lea:;t one n-olefin selected from the group consisting of
C8-C30 n-olefins or a mixture of the n-olefins is dimerized at a tempera-
ture ranging between 25 and 200 °C and at a pressure ranging between
0.001 mbar and 50 bar in the presence of a solid acidic catalyst by pass-
ing the feedstock to a catalytic distillation apparatus comprising either
a) a combination of a distillation column and a reactor comprising at
least one catalyst bed, or b) a distillation column connected to one or
more side reactors comprising at least one catalyst layer comprising
solid acidic catalyst material with a mesoporous surface area of more
than 100 m2/g, an aluminium content between 0.2 and 30 % by weight,
the amount of the acid sites of the material ranging between 50 and 500
umol/g, and the naterial being selected from the group consisting of
amorphous alumiriumsilica*es, zeolites, and mesoporous materials con-
taining silicon anc aluminium, recovering the unreacted n-olefin at the
upper part of the distillation column or at the upper part of the combi-
nation of the disti lation column and the reactor as a side stream to be
combined with tht feedstock, and the reaction product from the dimeri-
zation is hydrogenated.
2. Process accord ng to claim 1, characterized in that the feedstock
comprises at least one n-olefin selected from the group consisting of 1-
decene, mixtures of decenes, 1-dodecene, mixtures of dodecenes, 1-
hexadecene, mixtures of he:xadecenes, 1-tetradecene, 1-octadecene,
mixtures of octadesenes, and C20-C22 1-olefins and mixtures thereof.
3. Process accordiig to claim 1 or 2, characterized in that the catalyst
layer comprises solid acidic catalyst material selected from the group
consisting of acidic amorphous aluminium silicates, dealuminated Y
zeolites, mesoporous molecular sieves imbedded with zeolite, said zeo-
lite being preferably ZSM-5, beta-zeolite or MCM-22, whereas the
mesoporous mole;ular sieve is MCM-41 with a regular molecular
structure.
4. Process according to any one of claims 1-3, characterized in that
the mesoporous sirface area of the catalyst material is more than 300
m2/g.
5. Process accordi lg to any one of claims 1-4, characterized in that
catalyst material i.; amorphous aluminium silicate with a mesoporous
surface area of more than 300 m2/g.
6. Process according to any one of claims 1-5, characterized in that
part of the side stieam is removed from the distillation column or the
combination of the distillation column and the reactor at the upper part
thereof.
7. Process according to any one of claims 1 - 6, characterized in that a
solvent or a solvent mixture selected from the group of hydrocarbons is
added to the feedstock.
8. Process according to any one of claims 1 - 7, characterized in that
the feedstock is subjected to drying.
9. Process according to any one of claims 1 - 8, characterized in that
the dimerization is performed at a temperature ranging between 50 and
150 °C, and at a prjssure ranging between 0.5 bar and 30 bar.
10. Process accord ng to any one of claims 1 - 9, characterized in that
the hydrogenation is performed at a pressure ranging between 5 and
100 bar, preferably between 10 and 70 bar, and at a temperature rang-
ing between 100 md 400 °C, preferably between 150 and 250 °C, in
the presence of a 1' ydrogenation catalyst.
11. Process according to any one of claims 1 - 10, characterized in
that the feedstock comprises I -decene.
12. Process according to any one of claims 1 - 10, characterized in
that the feedstock comprises -hexadecenc.
13. Process according to any one of claims 1 - 12, characterized in
that the feedstock comprises olefins of natural origin produced from
biological starting materials.
14. Process according to any one of claims 1 - 13, characterized in
that distillation is carried out after the dimerizalion or the hydrogena-
tion or after both.
15. Process according to any one of claims 1 - 14, characterized in
that base oil or bass oil component is produced with said process.
16. Use of an apparatus comprising 1) an optional drying means for the
feedstock, 2) a catalytic distillation apparatus comprising either a) a
combination of a distillation column and a reactor comprising at least
one catalyst bed of a solid and acidic catalyst, or b) a distillation col-
umn connected to one or more side reactors comprising at least one
catalyst bed of a solid and acidic catalyst, 3) hydrogenation reactor and
optionally a distillation column downstream of the catalytic distillation
apparatus and/or hydrogenation reactor, in a process according to any
of the claims 1 - 16

The present invention relates to a process for producing polyolefms wherein a feedstock comprising n-olefin or
n mixture of nolofins is dimorized in the pres ci ce of a solid noidic catalyst by passing the feedstock to a catalytic distillation
apparatus comprising either a) a combination of distillation column and a reactor comprising at least one catalyst bed, or b) a
distillation column connectcd to one or more sidde reacors comprising at least one catalyst layer, recovering the unre- acted n olefin
from the distillation column or the combination of the distillation solumn and this reactor at the upper port thereof as sidc-streain
to be coinbined with the feedstock, and the reactor product from the dimerization is hydrozen- aled.

PROCESS FOR THE MANUFACTURE OF OLEFIN OLIGOMERS

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