DE102008007081A1 - Producing n-butene-oligomers and 1-butene from mixture-I made of hydrocarbons having four carbon atoms, comprises e.g. optionally washing the mixture to produce a mixture-II and removing 1,3-butadiene from mixture-I to produce a mixture-III - Google Patents

Abstract

Production of n-butene-oligomers and 1-butene from a technical mixture-I made of 4 carbon hydrocarbons, comprises: e.g. (a) optionally washing the technical mixture-I to produce a technical mixture-II; (b) removing 1,3-butadiene from technical mixture-I or -II to produce a technical mixture-III; and (c) separating the isobutene containing in the technical mixture-IIIa, by reacting with an excess of methanol in a two-stage methyl-tert-butylether-synthesis and separating the obtained methyl-tert-butylethers and the non-converted methanol under the production of a technical mixture-IV. Production of n-butene-oligomers and 1-butene from technical mixtures-I made of 4C-hydrocarbons, which contain at least 1-butene, isobutene, butane, 2-butene and many unsaturated 4C-hydrocarbons, comprises: (a) optionally washing the technical mixture-I under the production of a technical mixture-II; (b) removing of 1,3-butadiene from the technical mixture-I or -II under the production of a technical mixture-III; (c) separating the isobutene containing in the technical mixture-IIIa, which is identical with the technical mixture-III or is obtained by mixing the mixture-III with a 4C-hydrocarbon mixture, by reacting with an excess of methanol in a two-stage methyl-tert-butylether-synthesis and separating the obtained methyl-tert-butylethers and the non-converted methanol under the production of a technical mixture-IV, where the first step of the methyl-tert-butylether-synthesis is carried out in three solid bed reactors and the second step of the methyl-tert-butylether-synthesis is carried out on a solid acid ion-exchange resin, which exhibits sulfonic acid groups, in a reactive distillation column, which is operated in a pressure range of 0.5-1.5 MPa and at a temperature in the reaction zone of 50-90[deg] C at a reflux ratio of 0.5-1.5, (d) removing of many unsaturated 4C-hydrocarbons from the technical mixture-IV under the production of a technical mixture-V, by selective hydrogenation of many unsaturated 4C-hydrocarbon containing in the technical mixture, with hydrogen in liquid phase on a solid bed catalyst, which exhibits 0.1-2 mass-% of palladium on a carrier, in the presence of 0.05-10 mass-ppm of carbon monoxide, relative to the mass of the technical mixture-IV, at a temperature of 0-7500; (e) distillative separation of the technical mixture-V into a low boiling fraction L1, which contains 1-butene and isobutane and is largely free of n-butane and 2-butene, and into a high boiling fraction S1, which contains 2-butene and n-butane; (f) distillative separation of low boiling fraction L1 in 1-butene and isobutane; (g) optionally treating the high boiling fraction S1 with a molecular sieve with an average pore diameter of greater than 0.4-1.5 nm at a pressure of 0.1-5 MPa and a temperature of 20-160[deg] C under the production of technical mixture-VI; and (h) oligomerizing the n-butene containing in the high boiling fraction S1 or in the technical mixture-VI, under the production of oligomers at 0-200[deg] C and pressure of 0.1-7 MPa on a solid catalyst containing nickel, silicon and aluminum.

Description

The invention relates to a process for the preparation of n-butene oligomers and 1-butene from technical mixtures I of C ₄ -hydrocarbons which contain at least 1-butene, isobutene, butanes, 2-butenes and polyunsaturated C ₄ -hydrocarbons.

1-Butene is obtained, in addition to other C ₄ hydrocarbons such as isobutene and 2-butenes, in large quantities from technical C ₄ cuts, for example the C _{4 cut} from steam crackers or FCC units. These C ₄ cuts consist essentially of butadiene, the monoolefins isobutene, 1-butene and the two 2-butenes and the saturated hydrocarbons isobutane and n-butane. Because of the low boiling point differences of the ingredients and their low separation factors, a distillative work-up is difficult and not economical. The recovery of linear butenes and other products is therefore usually through a combination of chemical reactions and physical separation operations.

The first step, common to all refining variants by default, is the removal of most of the butadiene it contains. This is separated either by means of extractive distillation or selectively hydrogenated to butenes. In both cases, a hydrocarbon mixture remains (so-called raffinate I or hydrogenated crack C ₄ ), which contains the olefins isobutene, 1-butene and 2-butenes (cis and trans) in addition to the saturated hydrocarbons n-butane and isobutane polyunsaturated C ₄ hydrocarbons are typically present in less than 1% proportion.

Because of the closely spaced boiling points of 1-butene and isobutene, no 1-butene can be separated economically from these mixtures by simple distillation. From raffinate I or hydrogenated crack C ₄ isobutene is therefore as far as possible removed by a selective chemical reaction. After removal of the isobutene remains a hydrocarbon mixture (raffinate II) containing the linear butenes and the saturated hydrocarbons isobutane and n-butane. This mixture can be further fractionated by distillation, for example in isobutane and 1-butene and a mixture of the two 2-butenes and n-butane. From the 1-butene-containing fraction 1-butene can be obtained in high purity, which contains only small amounts of isobutene in further distillation steps. This is necessary because 1-butene is widely used as a comonomer in ethylene polymerization where isobutene impurities are undesirable. Typical specifications of 1-butene therefore limit the content of isobutene in 1-butene to below 2000 ppm.

For the selective chemical reaction of isobutene are different Known method. A way to separate from Isobutene is the reaction with alcohols, for example methanol or ethanol, to the corresponding tertiary butyl ethers. The advantage of this reaction is that the isobutene in Presence of linear butenes with high selectivity almost can be fully implemented (without a noticeable Sales of n-butenes occurs). For this purpose, various procedural Variants developed. The technique of reactive distillation has become thereby particularly useful for the delivery of high isobutene sales proved.

The Technically most important process is the conversion of isobutene with methanol to methyl tert-butyl ether (MTBE), mainly as fuel additive finds great use.

Another possibility for the chemical reaction of the isobutene is the reaction with water to give tert-butyl alcohol (TBA). Because of the low solubility of water in C ₄ hydrocarbons, this route is technically more complicated than the ether synthesis.

A Another possibility is to oligomerise the isobutene and to separate the oligomerizate. The disadvantage here is that at complete isobutene removal by oligomerization also a large part of the present linear butenes too Co- or homo-oligomers is implemented. Another disadvantage is the partial isomerization of 1-butene to the 2-butenes.

Another possibility for the chemical reaction of isobutene in the presence of other C ₄ hydrocarbons is the reaction with formaldehyde. The products thus obtained are further processed, for example, to form isoprene.

Many of the known reactions of isobutene, for example to tert-butyl alcohol (TBA) or isobutene oligomers, do not give complete isobutene conversion or, at high conversions in the presence of linear butenes, only poor selectivities. As a solution, for example, a combination of these Processes with a simultaneous ( EP 0 048 893 . DE 29 44 457 ) or subsequent conversion of the remaining isobutene to tert-butyl ethers.

In US 4,797,133 inter alia, a method is described in which in a first reaction zone isobutene z. B. by reaction to tert-butyl alcohol (TBA) separated from the starting hydrocarbon mixture, and the remaining radical is reacted in an etherification step.

In DE 103 02 457 describes a process for the preparation of isobutene oligomers and tert-butyl ethers from isobutene-containing C ₄ streams, in which the isobutene can be completely removed from a substantially butadiene-free C ₄ hydrocarbon stream with only slight losses of linear butenes. In this method, a portion of the isobutene is oligomerized in a first reaction step of acidic catalysts and in a second reaction step, the remaining isobutene is removed by reaction with alcohol to a tert-butyl ether in a reactive distillation column.

In DE 25 21 964 describes a two-stage process for the preparation of alkyl tert-butyl ethers, in which isobutene is reacted with alcohol in a first stage, the resulting ether is separated and the remaining residue is passed to a reaction of the remaining isobutene in a second reaction stage.

In DE 10 2005 062 700 describes a two-step process for the separation of 1-butene from a technical mixture of C ₄ hydrocarbons containing at least 1-butene, isobutene, n-butane and 2-butenes, in which in a first stage isobutene is chemically higher is reacted as a 1-butene boiling product. The resultant 1-butene mixture is separated by distillation into a fraction comprising 1-butene and, if appropriate, isobutene and a fraction comprising 2-butenes and n-butane, and the fraction comprising 1-butene is chemically separated in a second stage to give a higher than 1-butene boiling product is reacted, which is separated by distillation from 1-butene. The possibility of using the separated 2-butenes to produce n-butene oligomers is mentioned.

In EP 1 029 839 describes a process for the oligomerization of butenes in the presence of a transition metal catalyst, in which the starting material used is a hydrocarbon stream which is wholly or partly from a hydrocarbon stream ( 1 ), which contains predominantly butenes and which is isolated by separation from a stream of hydrocarbons ( 2 ) having a lower content of butenes, in which the hydrocarbon stream ( 1 ) by extractive distillation of a stream containing saturated and unsaturated C ₄ hydrocarbons ( 2 ) with a polar extractant to give a saturated fraction enriched in top fraction ( 3 ) and a sump fraction enriched in unsaturated hydrocarbons and containing the polar extractant ( 4 ), distillative separation of the bottoms fraction ( 4 ) into a top fraction ( 5 ) which contains butenes as unsaturated hydrocarbons and a bottoms fraction ( 6 ) containing the polar extractant, washing at least part of the top fraction ( 5 ) with water or an aqueous solution and drying the portion of the top fraction treated in step c) ( 5 ) to obtain a substantially anhydrous butene-containing hydrocarbon stream ( 1 ).

All of the processes cited in the prior art do not show how both 1-butene and n-butene oligomers can be obtained in a simple manner from a C ₄ -hydrocarbon technical mixture.

It Therefore, the object of the present invention was an alternative To provide processes for the preparation of n-butene oligomers and 1-butene, which preferably has one or more of the disadvantages of the methods of the prior art does not have.

Surprisingly, it has now been found that n-butene oligomers and 1-butene from technical mixtures I of C ₄ hydrocarbons containing at least 1-butene, isobutene, butanes, 2-butenes and polyunsaturated C ₄ hydrocarbons, in a simple manner by optionally washing the technical mixture I to obtain a technical mixture II, removing butadiene from the technical mixture I or II to obtain a technical mixture III, separating it into a technical mixture IIIa which is identical to the technical mixture III or by mixing the mixture III was obtained with a C ₄ hydrocarbon stream, isobutene by reaction with methanol in a two-stage methyl tert-butyl ether synthesis and separating the resulting methyl tert-butyl ether (MTBE) to obtain a technical mixture IV, removing of polyunsaturated C ₄ hydrocarbons to give a technical mixture V, by selective Hydri tion of polyunsaturated C ₄ hydrocarbons with hydrogen in the liquid phase on a fixed bed catalyst of 0.1 to 2 mass% palladium on a support, in the presence of 0.05 to 10 mass ppm carbon monoxide, based on the masses techni rule mixture V, at a temperature of 0 to 75 ° C, distillative separation of the technical mixture V in a low-boiling fraction L1, which has 1-butene and isobutene and is substantially free of n-butanes and 2-butenes, and in a high-boiling fraction S1 which comprises 2-butenes and n-butane, distillative separation of the low-boiling fraction L1 into 1-butene and isobutane, optionally treating the high-boiling fraction S1 with a molecular sieve having an average pore diameter of greater than 0.4 to 1.5 nm at a pressure of 0 , 1 to 5 MPa and a temperature of 20 to 160 ° C to obtain a technical mixture VI and oligomerizing the n-butenes contained in the high-boiling fraction S1 or the technical mixture VI to obtain oligomers O at temperatures of 0 to 200 ° C and Pressing 0.1 to 7 MPa on a solid nickel, silicon and aluminum catalyst can be produced.

Depending on the embodiment of the method according to the invention, various advantages can be achieved. By an optional wash of the mixture I of C ₄ hydrocarbons at the beginning of the process according to the invention, catalyst poisons which could interfere both in the removal of polyunsaturated hydrocarbons and in the MTBE synthesis and the oligomerization are separated from the process very early. As a result, not only a deactivation of the respective catalysts is avoided but also the formation of by-products which can be formed by reacting the catalyst poisons with one of the starting materials or by one of the catalyst poisons catalyzed reaction of the reactants with each other.

By the heat integration of several process steps can optimizes the energy needed for the overall process be used in the z. B. dissipated heat used elsewhere in the process for heating purposes.

following the process according to the invention becomes exemplary described, without the invention whose scope is made the claims and the entire description, on it should be limited. Also the claims themselves belong to the disclosure of the present invention. are in the following text ranges or preferred ranges indicated, so should also all lying in these areas, theoretically possible Subareas and individual values to the disclosure content of the present The invention includes, without these for reasons have been explicitly mentioned for better clarity.

• a) optionales Waschen der technischen Mischung I unter Erhalt einer technischen Mischung II,
• b) Entfernen von Butadien aus der technischen Mischung I oder II unter Erhalt einer technischen Mischung III,
• c) Abtrennung des in einer technischen Mischung IIIa, die mit der technischen Mischung III identisch ist oder durch Mischen der Mischung III mit einem C₄-Kohlenwasserstoffgemisch erhalten wurde, enthaltenen Isobutens durch Umsetzung mit einem Überschuss an Methanol in einer zweistufigen Methyl-tert.-butylether-Synthese und Abtrennen des erhaltenen Methyl-tert.-butylethers und des nicht umgesetzten Methanols unter Erhalt einer technischen Mischung IV, wobei die erste Stufe der Methyl-ter.-butylether-Synthese in drei Festbettreaktoren und die zweite Stufe der Methyl-tert.-butylether-Synthese an einem festen sauren Ionenaustauscherharz, das Sulfonsäuregruppen aufweist, in einer Reaktivdestillationskolonne, die in einem Druckbereich von 0,5 bis 1,5 MPa und bei einer Temperatur in der Reaktionszone von 50°C bis 90°C bei einem Rücklaufverhältnis zwischen 0,5 und 1,5 betrieben wird, durchgeführt wird,
• d) Entfernen von mehrfach ungesättigten C₄-Kohlenwasserstoffen unter Erhalt einer technischen Mischung V, durch selektive Hydrierung der mehrfach ungesättigten C₄-Kohlenwasserstoffe mit Wasserstoff in flüssiger Phase an einem Festbettkatalysator der 0,1 bis 2 Massen-% Palladium auf einem, vorzugsweise inerten Träger aufweist, in Gegenwart von 0,05 bis 10 Massenppm Kohlenmonoxid, bezogen auf die Massen technischen Gemisches V, bei einer Temperatur von 0 bis 75°C,
• e) Destillative Auftrennung der technischen Mischung V in eine Leichtsiederfraktion L1, die 1-Buten und Isobutan aufweist und weitgehend frei von n-Butanen und 2-Butenen ist, und in eine Schwersiederfraktion S1, die 2-Butene und n-Butan aufweist,
• f) Destillative Auftrennung der Leichtsiederfraktion L1 in 1-Buten und Isobutan,
• g) optionale Behandlung der Schwersiederfraktion S1 mit einem Molekularsieb mit einem mittleren Porendurchmesser von größer 0,4 bis 1,5 nm bei einem Druck von 0,1 bis 5 MPa und einer Temperatur von 20 bis 160°C unter Erhalt einer technischen Mischung VI und
• h) Oligomerisierung der in der Schwersiederfraktion S1 oder der technischen Mischung VI enthaltenen n-Butene unter Erhalt von Oligomeren O bei Temperaturen von 0 bis 200°C und Drücken von 0,1 bis 7 MPa an einem heterogenen Nickel, Silizium und Aluminium aufweisenden Katalysator,

aufweist.The inventive method for the preparation of n-butene oligomers and 1-butene from technical mixtures I of C ₄ hydrocarbons containing at least 1-butene, isobutene, butanes, 2-butenes and polyunsaturated C ₄ hydrocarbons, characterized thereby from that it at least the procedural steps

• a) optional washing of the technical mixture I to obtain a technical mixture II,
• b) removing butadiene from the technical mixture I or II to obtain a technical mixture III,
• c) separation of the isobutene contained in a technical mixture IIIa which is identical to the technical mixture III or obtained by mixing the mixture III with a C ₄ -hydrocarbon mixture by reaction with an excess of methanol in a two-stage methyl-tert. butyl ether synthesis and separation of the obtained methyl tert-butyl ether and the unreacted methanol to obtain a technical mixture IV, wherein the first stage of the methyl tert-butyl ether synthesis in three fixed bed reactors and the second stage of methyl tert. butyl ether synthesis on a solid acidic ion exchange resin having sulfonic acid groups in a reactive distillation column operating in a pressure range of 0.5 to 1.5 MPa and at a temperature in the reaction zone of 50 ° C to 90 ° C at a reflux ratio between 0.5 and 1.5 is operated,
• d) removing polyunsaturated C ₄ hydrocarbons to obtain a technical mixture V, by selective hydrogenation of the polyunsaturated C ₄ hydrocarbons with hydrogen in the liquid phase over a fixed bed catalyst of 0.1 to 2 mass% palladium on one, preferably inert Carrier, in the presence of 0.05 to 10 ppm by mass carbon monoxide, based on the masses of technical mixture V, at a temperature of 0 to 75 ° C,
• e) Distillative separation of the technical mixture V into a low-boiling fraction L1 which comprises 1-butene and isobutane and is substantially free of n-butanes and 2-butenes, and into a high-boiling fraction S1 which comprises 2-butenes and n-butane,
• f) Distillative separation of the low-boiling fraction L1 into 1-butene and isobutane,
• g) optional treatment of the high-boiling fraction S1 with a molecular sieve having an average pore diameter of greater than 0.4 to 1.5 nm at a pressure of 0.1 to 5 MPa and a temperature of 20 to 160 ° C to obtain a technical mixture VI and
• h) Oligomerization of those contained in the high-boiling fraction S1 or the technical mixture VI n-butenes to give oligomers O at temperatures of 0 to 200 ° C and pressures of 0.1 to 7 MPa on a heterogeneous nickel, silicon and aluminum catalyst,

having.

Process step a): washing the technical mixture

The optional washing of the technical mixture I can with water or an aqueous solution as a washing medium. We prefer demineralized water or drinking water as a washing medium used.

By washing the water, hydrophilic components of the technical mixture I can be completely or partially removed. Such hydrophilic compounds can be about. Nitrogen compounds or oxygen compounds. Examples of nitrogen compounds are, for example, acetonitrile or N-methylpyrrolidone, which can originate from a 1,3-butadiene extractive distillation. An example of an oxygen compound is acetone, which may be from FC crackers. The laundry can be done in one or more stages. Accordingly, the technical design is adapted, for example as an extraction column or as a mixer-settler combination. Preferably, a single-stage wash is carried out by mixing the C ₄ hydrocarbons with the washing medium and subsequent phase separation. The phase separation can be supported by using special filters.

Preferably is the mass ratio of washing medium to technical mixture I 1 to 1 to 1 to 100, preferably 1 to 3 to 1 to 30.

The I technical mixture is after a water wash with watery Wash medium saturated with water. To a two-phase should be avoided in the subsequent process steps in the reactor There, the reaction temperature about 10 ° C over the temperature of the water wash lie.

Process step b): removal of butadiene

The Removing butadiene from the technical mixture I or, if the optional process step a) has been carried out the technical mixture II to obtain a technical mixture III, can be done in different ways.

Thus, the removal of butadiene, in particular by separation of the 1,3-butadiene by extraction, for. With N-methylpyrrolidone as in US 6,337,429 described, or by selective hydrogenation of the butadiene, in particular the 1,3-butadiene. Preferably, the butadiene is removed by selective hydrogenation.

The selective hydrogenation of the butadiene to butenes is preferably carried out in the liquid phase with hydrogen over a solid catalyst. Preferred is the selective hydrogenation as in DE 195 24 971 described, carried out in several series-connected loop reactors. Particularly preferably, the process is carried out in two stages in two reactors, wherein hydrogen is supplied in both stages of the process. The reactors used are preferably adiabatic fixed bed reactors.

Depending on its origin, the technical mixture used I or II z. B. from 30 to 75% by mass, preferably from 30 to 60% by mass of polyunsaturated hydrocarbons. For example, the content of 1,3-butadiene in the C ₄ fraction of a steam cracker is often about 45% by mass. The technical mixture I or II can be used dry or optionally wet.

In order to achieve a high selectivity of the formation of linear butenes, it may be advantageous to limit the concentration of polyunsaturated hydrocarbons, predominantly 1,3-butadiene, in the feed to the first stage of process step b). The feed concentrations of the polyunsaturated C ₄ compounds in the technical mixture I or II is preferably less than 20% by mass, preferably from 3 to 15% by mass and more preferably from 5 to 10% by mass. When technical mixtures I or II with a higher concentration of polyunsaturated C ₄ compounds are used, the first stage reaction product is preferably recycled, that is to say the first reactor (the first stage) is operated in loop mode.

In the second stage, the adiabatic fixed bed reactor is preferably operated with recycling of product. The ratio of recirculation to fresh feed is preferably from 50/1 to 2/1, preferably from 35/1 to 5/1.

The selective hydrogenation according to the embodiment The process of the present invention is preferably incorporated herein by reference a temperature in the reaction zone of 10 to 100 ° C, preferably carried out from 25 to 80 ° C. These medium hydrogenation temperatures may be in use be the same or different of several reactors in these reactors.

Of the used as starting material butadiene-containing stream (technical mixture I or II) represents the liquid phase in the three-phase system Therefore, such hydrogenations are preferably in a Pressure of 0.5 to 5 MPa, preferably from 0.5 to 3 MPa and very particularly preferably from 0.5 to 2 MPa, measured in each case at the inlet of the reactors carried out.

The empty tube velocities of the liquid phase in process step b) can vary within wide limits. In particular, it can go beyond the laminar range significantly. Preferred empty tube velocities with which the method step b) according to the invention can be carried out are from 50 to 700 m ³ / m ² / h, preferably from 100 to 450 m ³ / m ² / h. The empty tube velocities may be the same or different in the presence of multiple reactors in the different reactors.

In the case that more gas (hydrogen) is fed into the reactor than is soluble in the liquid supplied, this gas can be dispersed by means of static mixers. Suitable mixers are for example in the application DE 10 2004 021 128 described.

The The gas used for the hydrogenation can be pure hydrogen or a hydrogen having gas, such as. B. hydrogen that is up to 50% by volume Inert gases (gases that are incompatible with the educts and / or the product still react to change the activity of the catalyst) having. Inert gases may be, for example, nitrogen or Methane, but not carbon monoxide. The content of carbon monoxide in the gas should be because of its reduction often caused by him the catalyst activity in the range of a few ppm by volume preferably less than 10 ppm by volume. In the invention Process is preferably hydrogen with a purity greater 99.5% by volume used as gas.

Palladium-supported catalysts are preferably used for the selective butadiene hydrogenation. The palladium concentration in the supported catalyst is preferably from 0.1 to 2.0% by mass, preferably from 0.2 to 1.0% by mass. As carrier materials, for example MgO, Al ₂ O ₃ , SiO ₂ , TiO ₂ , SiO ₂ / Al ₂ O ₃ , CaCO ₃ or activated carbon or mixtures thereof can be used. Preferred support materials are Al ₂ O ₃ and SiO ₂ . For the hydrogenation in particular shell catalysts can be used, ie catalysts in which the hydrogenation-active substances are in the edge zone of the catalyst body.

The Catalysts are preferably used in a form in which they offer a low flow resistance, z. In Form of granules, pellets or shaped articles such as tablets, Cylinders, balls, extruded extrusions or rings. In each of Reactors or reaction zones may be the same or different Catalysts are used.

Process step c): MTBE synthesis

From the technical mixture III obtained in process step b) or a mixture IIIa obtained by mixing the mixture III with a C ₄ -hydrocarbon mixture, isobutene is separated off in process step c) by means of an MTBE synthesis. The C ₄ -hydrocarbon mixtures used to obtain the technical mixture IIIa preferably have a content of 1,3-butadiene of less than 1 mol%. Typical C ₄ -hydrocarbon mixtures which can be mixed with the technical mixture III to a mixture IIIa are, for. B. Raffinate I and Raffinate II.

Raffinate I is generally understood to mean a C ₄ -hydrocarbon mixture which is obtained by complete or substantial removal of the 1,3-butadiene (maximum residual concentration of 1,3-butadiene of 1% by mass, preferably 0.1% by mass) from a C ₄ hydrocarbon mixture comprising butadiene, olefins such as isobutene, 1-butene and 2-butenes and saturated hydrocarbons such as isobutene and n-butane, such as. B. Crack-C ₄ , was obtained. By raffinate II is generally a raffinate I understood from the isobutene, z. B. by reaction of isobutene with methanol or ethanol to methyl tert-butyl ether or ethyl tert-butyl ether and separation of the ether, was completely or substantially removed.

In process step c) of the process according to the invention that in the mixtures III or IIIa of Isobutene containing C ₄ hydrocarbons by reaction with an excess of methanol in a two-stage methyl tert-butyl ether synthesis (MTBE synthesis) separated. In principle, the MTBE synthesis, as in DE 101 02 082 described, done.

The first stage of MTBE synthesis in the invention Process step c) is carried out in fixed bed reactors, the second stage of the reaction is carried out in a reactive distillation. The first stage of methyl tertiary butyl ether synthesis is preferred carried out in three fixed bed reactors. As reactors, in which methanol with the isobutene close to the thermodynamic Balance can be implemented, conventional Fixed bed reactors (tube bundle reactors, adiabatic fixed bed reactors, Circulation reactors) can be used. You can with or without partial recycle, where appropriate, the recycle stream can be cooled, operated.

In In the first stage, the turnover of the isobutene takes place up to the setting the thermodynamic equilibrium of MTBE, methanol and isobutene, preferably adjusted to an isobutene conversion of 94 and 96% or is reached. As catalysts, they are preferably the same Catalysts used in the second stage. The first stage reactors are preferably at a temperature from 20 to 110 ° C, preferably 25 to 70 ° C and a Pressure of 0.5 to 5 MPa, preferably 0.7 to 2 MPa operated.

There the thermodynamic equilibrium between methanol / isobutene and Ether at low temperature predominantly on the side of Ethers lies, it is preferred, the first of the reactors at higher Temperature for the purpose of high reaction rate than the following, in which the equilibrium position is exploited to operate. Preferably becomes the first of the reactors at a temperature of 35 to 70 ° C operated and the subsequent reactors are at a temperature operated from 25 to 50 ° C.

The molar ratio of methanol to isobutene in the feed to first reactor of the first stage of process step c) is located preferably in the range of 10 to 1 to 1 to 1, more preferably from 5 to 1 to 1.1 to 1, and most preferably in the range from 1.8 to 1 to 1.2 to 1.

The second stage of the methyl tertiary butyl ether synthesis is carried out in a reactive distillation column operating in a pressure range of 0.5 to 1.5 MPa, preferably 0.75 to 1.0 MPa, at a temperature in the reaction zone of 50 ° C to 90 ° C, preferably 55 to 70 ° C at a reflux ratio between 0.5 and 1.5, preferably between 0.7 and 0.9 is operated, carried out on an acidic ion exchange resin. The reflux ratio is defined as the ratio of the reflux flow in the column to the discharged distillate stream. The relatively low reflux ratio can lead to a significant steam saving, whereby the inventive method has a low energy consumption. Surprisingly, despite the relatively low reflux ratios and temperatures in the catalyst packings, a relatively high isobutene conversion is achieved. This is particularly surprising since in the literature, for example from Lawrence A. Smith, D. Hearn, in Catalytic Distillation, Proc. Intersoc. Energy Conversions. Conf. (1984) 19th, (Vol2), p998-1002 and in further in DE 10102082 quoted writings the opposite is described. The optimum reflux ratio can be dependent on the throughput, composition of the column feed and the column pressure. However, it is preferably always in the above ranges.

By means of the two-stage MTBE synthesis, in particular isobutene residual concentrations in the technical mixture IV of less than 1000 ppm by mass (wppm), preferably 800 ppm by weight and particularly preferably less than 500 ppm by weight (based on the C ₄ mixture in the distillate) can be obtained.

The Temperature of the column inlet is independent of this Composition, the reaction pressure in the column and the throughput preferably between 50 ° C and 90 ° C, preferably between 60 ° C and 75 ° C.

Of the Feed to the reactive distillation column can be above or below, preferably carried out below the catalyst zone. The feed to the reactive distillation column is preferably carried out below the reactive packing, preferably 3 to 13, more preferably 4 to 10 theoretical plates below the reactive packing.

In the column feed the second stage may be more methanol than is needed for the complete implementation of the remaining isobutene. However, the methanol excess should be limited in such a way that on the one hand a sufficient amount of methanol for the forming azeotrope of metha nol and C ₄ hydrocarbons is present, on the other hand, not so much that methanol geriete in the bottom product, so that preferably a MTBE is obtained with a methanol content below 5000 wppm. The residual content of isobutene in the feed to the second stage is preferably less than 2% by mass, preferably less than 1% by mass.

optional additional methanol can be fed to the second stage become. This can be done together with the feed from the first stage or but also at one point or at several points of the reactive distillation column done, z. B. at the top of the column and / or on, between and / or under the catalyst bed.

The reactive distillation column preferably contains the catalyst in the amplifier column and below and above the catalyst packing are preferably separators or distillation packs. The catalyst can either be integrated in a package (such as in KataMax® ^® packings EP 0 428 265 ), KataPak ^® packages ( EP 0 396 650 ) or MultiPak ^® packages ( German utility model no. 298 7 007.3 ) as described in the cited documents, or be polymerized onto moldings ( US 5,244,929 ). KataMax® ^® packings are preferably used.

Prefers indicates the reactive distillation column above the catalyst packing a range of purely distillative separation. Preferably, the Zone above the catalyst packing 5 to 20, in particular 10 to 15 stages on. The separation zone below the catalyst comprises 12 to 36, in particular 20 to 30 separation stages. The catalyst zone can with a distillative effect of 1 to 5 theoretical plates per Meter pack height be estimated. The height the catalyst zone / reactive zone can be dependent of the desired isobutene conversion by simple preliminary tests determine. The amount of catalyst is preferably chosen so large that an isobutene conversion of 75 to 99%, preferably 85 to 98% and more preferably from 95 to 97% based on the isobutene content is reached in the feed to the reactive distillation.

The reactive distillation column in the second stage of process step c) of the process according to the invention is preferably operated with a hydraulic load of the catalytic packing of 10% to 110%, preferably 20% to 70% of its flood point load. Hydraulic loading of a distillation column is understood to be the uniform flow stress of the column cross-section due to the ascending vapor mass flow and the returning liquid mass flow. The upper load limit indicates the maximum load of steam and return fluid, above which the separation effect due to entrainment or jamming of the return fluid by the rising vapor flow decreases. The lower load indicates the minimum load below which the separation effect decreases or collapses due to irregular flow or emptying of the column -. B. the floors. ( Vauck / Müller, "Basic Operations of Chemical Process Engineering", p. 626, VEB Deutscher Verlag für Grundstoffindustrie ).

At the point of flooding, the shear stresses transferred from the gas to the liquid become so great that the entire amount of liquid is entrained in the form of drops with the gas or that phase inversion occurs in the column ( J. Mackowiak, "Fluid Dynamics of Columns with Modern Packages and Packings for Gas / Liquid Systems", Otto Salle Verlag 1991 ).

When Catalysts become solid acidic ion exchange resins in MTBE synthesis, have the sulfonic acid groups used. suitable Ion exchange resins are, for example, those obtained by sulfonation of phenol / aldehyde condensates or of cooligomers of aromatic Vinyl compounds are produced. Examples of aromatic Vinyl compounds for the preparation of the cooligomers are: styrene, vinyltoluene, Vinylnaphthalene, vinylethylbenzene, methylstyrene, vinylchlorobenzene, Vinylxylol and divinylbenzene. In particular, the cooligomers, which are formed by reaction of styrene with divinylbenzene, as Preliminary stage for the production of. ion exchange resins used with sulfonic acid groups. The resins can gel-shaped, macroporous or sponge-shaped getting produced. The properties of these resins, especially specific ones Surface, porosity, stability, swelling or shrinkage and exchange capacity, can by the manufacturing process can be varied.

In the process according to the invention, the ion exchange resins can be used in their H form. Strong acid resins of the styrene-divinylbenzene type are sold, inter alia, under the following trade names: Duolite C20, Duolite C26, Amberlyst 15, Amberlyst 35, Amberlite IR-120, Amberlite 200, Dowex 50, Lewatit SPC 118, Lewatit SPC 108, K2611, K2621 , OC 1501 are preferably used (each Rohm & Haas) or Lewatit ^® K2621 (Lanxess) as ion exchange resins, the types Amberlyst ^® 15, Amberlyst ^® 35th

The Pore volume of the ion exchange resins used as catalysts is preferably from 0.3 to 0.9 ml / g, in particular from 0.5 to 0.9 ml / g. The grain size of the resins is preferably from 0.3 mm to 1.5 mm, in particular of 0.5 mm to 1.0 mm. The particle size distribution can be elected narrower or wider. So, for example Ion exchange resins with very uniform grain size (Monodisperse resins) are used. The capacity of the Ion exchanger is, based on the delivery form, preferred from 0.7 to 2.0 eq / l, in particular from 1.1 to 2.0 eq / l or preferably from 0.5 to 5.5 mol / kg, in particular from 0.8 to 5.5 mol / kg. The Capacities in mol / kg refer to each to constant weight in a warm stream of nitrogen at z. 105 ° C dried ion exchange resin.

Out the reaction mixture obtained in the second step of the MTBE synthesis are the obtained methyl tert-butyl ether and the unreacted Separated methanol to obtain a technical mixture IV. The separation of the MTBE from the reaction mixture already takes place in: the reactive distillation column, where the MTBE as the bottom product accrues. As the top product of the reactive distillation column a methanol-containing stream is obtained, the 1-butene, 2-butenes and contains butane.

The resulting in the reactive distillation column as bottom product MTBE can be used for various purposes. Since it contains only extremely small amounts of methyl sec-butyl ether (MSBE), it is suitable for the production of high-purity isobutene by its cleavage, since virtually no linear butenes can be formed by cleavage of the methyl sec-butyl ether. The MTBE cleavage can z. B. according to DE 100 20 943.2 be performed. Due to the low content of by-products (MSBE and C ₈ olefins), the MTBE thus obtained can be used as a solvent in analytics or in organic syntheses after separation from the remaining alcohols. Furthermore, its use as a component for gasoline fuels is possible.

Preferably, the process step c) of the inventive method is carried out so that in the second stage, a top product containing methanol and a C ₄ hydrocarbon mixture having an isobutene content of less than 1000 wppm, based on the C ₄ hydrocarbon mixture, and a bottom product containing MTBE is obtained, and that the overhead product is separated into a technical mixture IV and methanol. Further purification of the MTBE is no longer required when used as a component for gasoline fuels. From the top product, the methanol can be separated, for example, by adsorption on molecular sieves, membrane processes, entrainer distillation or extraction, preferably by extraction to obtain a methanol-lean or -free technical mixture IV.

The Separation of the methanol from the top product can in particular by Extraction with water or an aqueous solution take place as a washing medium. Preference is given to an aqueous Solution with a pH greater than or equal 8, preferably used from 8 to 12. The adjustment of the pH can z. B. by adding sodium hydroxide solution and / or sulfuric acid respectively. This extraction according to the known standard methods The technique can, for example, in an extraction column or in a cascade of mixers and separation tanks. Compared to the other methods, it has different Benefits, such as low investment and low operating costs.

Preferably, the separation of the methanol from the top product of the reactive distillation column is carried out with water or an aqueous solution as a washing medium in an extraction column. The residual content of alcohol is preferably lowered to below 0.2% by mass, more preferably below 500 ppm by weight, very particularly preferably below 50 ppm by weight. The extraction column preferably has from 2 to 25, more preferably from 5 to 15 theoretical plates and is preferably operated at temperatures of 10 to 90 ° C and pressures of at least 0.1 MPa above the vapor pressure of the C ₄ hydrocarbons. The mass ratio of washing medium to top product fed (technical mixture IV plus methanol) is preferably from 1 to 5 to 1 to 40.

The Alcohol-laden wash water from the extraction is preferred worked up in a separate unit and at least partially returned to the extraction. The workup can be done for example by a distillation in which a virtually alcohol-free water fraction in the bottom and methanol as top product to be obtained. The methanol can be returned to the MTBE synthesis be driven.

Prefers the overhead product of the reactive distillation column is transferred to an extraction column, in in countercurrent an extractant such. As water, over a feed located at the top is fed. The extractant can be removed via the drain at the bottom of the column. At the top of the extraction column can as a product of the extraction of a technical mixture IV are obtained.

Process step d): Selective hydrogenation

From the technical mixture IV obtained in process step c) polyunsaturated C ₄ -hydrocarbons are removed to obtain a technical mixture V, by selective hydrogenation of the polyunsaturated C ₄ hydrocarbons with hydrogen. The hydrogenation is preferably carried out analogously to that in DE 31 43 647 described method.

The Hydrogenation takes place in the liquid phase on a palladium-containing Fixed bed catalyst with hydrogen with the addition of carbon monoxide as a moderator.

hydrogen and carbon monoxide are complete in the hydrocarbon mixture solved. At least the amount of hydrogen added is the stoichiometric for the hydrogenation of the multiply unsaturated compounds to the monoenes is necessary. It can be made from the composition of the to be hydrogenated calculate technical mixture IV.

The amount of CO 2 to be added to the mass of technical mixture IV is at least 0.05 ppm. By default, amounts above 20 ppm do not lead to any further substantial improvement in the hydrogenation results, so that amounts of 0.05 to 10 ppm are preferred. The optimum amount of CO to be metered in the respective process can easily be experimentally determined, as in DE 31 43 647 described, be determined.

Of the Catalyst has 0.1 to 2 mass% of palladium on one, preferably inert carrier. Belong to such carriers For example, alumina, silica gel, aluminosilicate and activated carbon. Per liter of catalyst used is preferably an amount of hydrocarbon in the range of 5 to 300 liters enforced.

The Temperature at which the hydrogenation is carried out is from 0 to 75 ° C. So that no free water is generated, the hydrogenation is appropriate at higher Temperature operated as the extraction in step c).

Of the Process pressure must be sufficiently large to allow the liquid phase upright at the chosen temperature and around a sufficient amount of hydrogen and carbon monoxide in solution bring to. The reaction pressure is preferably smaller 6 MPa, preferably less than 2 MPa. A typical reaction pressure is 1.5 MPa.

Prefers the hydrogenation is multi-stage, particularly preferably two-stage, carried out. The supply of hydrogen takes place before each of the reactors, preferably of carbon monoxide in the first of the reactors. The reactors can be under recirculation be operated by product.

Process step e): distillation

Of the Process step e) comprises a distillative separation of the technical Mixture V into a top fraction L1 which is essentially 1-butene, Iso-butane and optionally low-boiling and largely is free from n-butanes and 2-butenes, and a bottoms fraction S1, which has at least 2-butenes and n-butane.

Out the technical mixture V obtained in process step d), in The specialist literature often referred to as Raffinate II, can Isobutane and 1-butene by distillation, in whole or in part, be separated. The remaining mixture S1, too referred to as raffinate III, usually contains mainly 2-butenes, n-butane and optionally one Proportion of 1-butene. The mixture S1 is feed mixture for the process step g) of the invention Process.

The Distillative separation can be used for separation of such Hydrocarbon mixtures usually used equipment, z. B. distillation or Fraktionierungskolonnen done.

In a preferred embodiment, the distillative separation is carried out in a Superfraktionierkolonne. The feed to this column is preferably in the lower half, preferably in the lower third of the column. Because of the close boiling position of the mixture to be separated, process step e) is preferably carried out in a column which has more than 100, preferably more than 125, more preferably more than 150 theoretical plates and very preferably from 150 to 200 theoretical plates. The column can be designed as a packed or tray column. Preferably, the Ausfüh tion designed as a tray column.

The Return ratio (return flow to Distillate) in the column, depending on the realized number of stages and the operating pressure, preferably less than or equal to 20, preferably less than 14, more preferably less 11. The condensation can be carried out against cooling water or air become. The distillate container is preferably deposited as liquid-liquid executed. As a result, if appropriate, in the feed stream contained water as a second phase in the distillate tank be separated and it may be a technically anhydrous bottom product S1 are obtained.

The separation according to process step e) is preferably carried out _absolutely at a pressure of 0.4 to 1.0 MPa _absolute , preferably at a pressure of 0.5 to 0.7 MPa. The temperature at which the separation is carried out is preferably from 35 to 80 ° C, preferably 40 to 65 ° C.

For heating the evaporator of the column used in step e), a conventional heat transfer medium, such as. As steam or hot water and preferably waste heat from other processes can be used. In the latter case, it may be advantageous to equip the column with more than one evaporator. The column is preferably equipped as a simple column with at least one evaporator and at least one condenser. Because of the high energy demand and the small temperature difference between the bottom and top of the column energy saving circuits are particularly preferred embodiments. As an example, reference is made to the method of vapor compression. Another particularly preferred circuit is the double-effect distillation in integration with a second column. The second column may preferably be a parallel-connected column with the same or different separation task. In this case, one of the columns is driven at such high pressure that their condensation temperature is sufficient to heat the other column. In the thermal connection of columns with different separation tasks, in principle any suitable column from the process according to the invention but also a column which is present outside the process according to the invention at the plant location, with the column of process step e) are connected. The second column is particularly preferably the C ₄ separating column from process step f).

Process step f): 1-butene separation

Of the Process step f) involves the distillative separation of Low-boiling fraction L1 from process step e) in 1-butene and Isobutane and optionally low boilers.

Out the top fraction L1 obtained from process step e), which in Essentially comprises 1-butene, isobutane and optionally low-boiling components, is separated by distillation 1-butene. The separation of the 1-butene can by distillation of fraction L1 in one or more distillation columns respectively.

In a preferred embodiment, the separation of the 1-butene is carried out in a distillation column in which very pure 1-butene is obtained as the bottom product. The top product obtained is an isobutane-rich fraction which optionally also contains low boilers (for example C ₃ hydrocarbons).

Prefers the separation is carried out in a Superfraktionierkolonne. The feed to this column is preferably in the upper half, preferably in the lower half of the upper half the column. Because of the tight Siedelage of the mixture to be separated the column is preferably greater than 100, preferably more than 125, more preferably more than 150 and most preferably run from 150 to 200 theoretical plates. The Column is preferably carried out as a packed or tray column, preferably as a tray column. The return ratio (Return amount to distillate decrease) is dependent of the realized number of stages and the operating pressure, preferably less than or equal to 100, preferably less than 70, particularly preferably less 60. Most preferably, the reflux ratio from 30 to 60. The condensation can be against cooling water or Air to be done. The distillate container is preferably as a liquid-liquid separator executed. As a result, if appropriate, in the feed stream contained water as a second phase in the distillate tank be separated and it may be a technically anhydrous bottom product to be obtained.

For heating the evaporator of the column, a conventional heat transfer medium, such as. As steam or hot water, and preferably waste heat from other processes can be used. In the latter case, it may be advantageous to equip the column with more than one evaporator. The column is preferably equipped as a simple column with at least one evaporator and at least one condenser. Because of the high energy demand and the small temperature difference between the bottom and top of the column are Energy saving circuits particularly preferred embodiments. As an example, reference is made to the method of vapor compression. Another particularly preferred circuit is the double-effect distillation in integration with a second column. The second column may preferably be a parallel-connected column with the same or different separation task. In this case, one of the columns is driven at such high pressure that their condensation temperature is sufficient to heat the other column. In the thermal engineering connection of columns with different separation tasks, in principle any suitable column from the process according to the invention but also a column which is present outside the process according to the invention at the plant location, can be connected to the column according to the invention of process step f). The second column is particularly preferably the C ₄ separating column from process step e). In this case, one of the columns is driven at such high pressure that their condensation temperature is sufficient to heat the other column.

Next this preferred embodiment of the method step according to the invention f) it is also possible from the top fraction L1 in one first distillation column initially optionally present Separate low boilers as top product while in the bottom of the column a mixture is obtained, mainly 1-butene and isobutane. In a second column, like carried out the above-described Superfraktionierkolonne may be, this sump mixture in 1-butene, which is obtained as bottoms, and an isobutane-rich fraction (overhead product) are separated.

With produced by the process according to the invention, Pure 1-butene preferably contains less than 5000 wppm (Mass ppm), preferably less than 2000 wppm, and more preferably less than 1500 wppm isobutene and is a sought after intermediate. It can be used, for example, as a comonomer in the production of polyethylene (LLDPE or HDPE) and used by ethylene-propylene copolymers become. It is still used as alkylating agent and is Starting material for the production of butan-2-ol, butene oxide, Valeraldehyde. Another use of the invention produced almost isobutene-free 1-butene is the preparation of n-butene oligomers, especially after the octol process.

In addition to the 1-butene fall, depending on the starting composition of C ₄ hydrocarbons, in the working up by distillation in process step f) isobutane-rich fractions. These can be purified further, preferably to pure isobutane. The isobutane obtained in the work-up preferably has a purity of at least 90% by mass of isobutane, particularly preferably 95% by mass of isobutane and preferably contains less than 1000 ppm by weight, more preferably less than 200 ppm by weight olefins. Purification to pure isobutane can be carried out, for example, by complete hydrogenation of the alkenes still present to give alkanes and subsequent distillation.

Process step g): molecular sieve treatment

The high-boiling fraction S1 obtained in process step e) can be treated with an adsorber in an optional process step g) in order to remove trace components which in process step h) could poison the catalyst or reduce its activity. Suitable adsorbers are for example in DE 198 45 857 or DE 38 25 169 or DE 39 14 817 described. The product of the treatment according to process step g), the technical mixture VI is obtained.

Prefers Molecular sieves are used as adsorbers, which have a medium Pore diameter of 0.4 to 1.5 nm. At the molecular sieves can be both crystalline, natural aluminum silicates, such as As layered silicates, as well as synthetically produced Act molecular sieves. These include, for example, commercial available zeolites of the A, X and Y type. The molecular sieves In addition, other metals, such as copper, silver, Zinc or mixtures of these metals. The content of these Metals are preferably less than 2 mass%.

Of the Process step g) according to the invention is preferably at a pressure of 0.1 to 5 MPa and a temperature of 20 to 160 ° C in the liquid phase. Typical trace components that come with the cleaning the adsorber are removed, for. As sulfur, nitrogen, Oxygen and / or halogen compounds.

Process step h): Oligomerization

The technical mixture VI obtained in process step g) or the n-butenes present in fraction S1 contained in process step e) are oligomerized to give oligomers O on a catalyst comprising heterogeneous nickel, silicon and aluminum. The this process step too Underlying procedures have found entrance into the literature as OCTOL process, which in Hydrocarbon Process., Int. Ed. (1986) 65 (2nd section 1), pages 31 to 33 as well as in the scriptures DE 39 14 817 . EP1 029 839 and DE 10 2004 018753 is described.

The oligomerization takes place in the presence of heterogeneous nickel-containing supported catalysts. As support materials, the catalysts may comprise, for example, silica and alumina, aluminosilicates or zeolites. The carriers may optionally be sulfated. In addition, the catalyst may contain sulfur. Such catalysts are known in the literature and, for example, in DE 43 39 713 or WO 01/37989 described.

For the preparation of the used nickel-supported catalysts There are different ways. For example, such Catalysts by co-precipitation of nickel compounds and carrier material, such as. B. aluminum and silicon compounds, Filtration and subsequent annealing are made. A Another possibility is to nickel compounds to apply one of the abovementioned carrier materials, for example, by impregnation or spraying, and then calcining the catalyst precursor. To prepare the catalysts, nickel compounds, such as nickel nitrate, nickel chloride or amine complexes, be used. Suitable carrier materials are commercially available available carriers such. B. aluminosilicates of Company Grace with the type designation Grace DAVICAT or pyrogenic Aluminosilicates or amorphous zeolites (MCM4) from Mobil Oil.

Especially Titanium-free carriers or supported catalysts are preferred The formally used essentially consists of nickel oxide, aluminum oxide and silica. These catalysts preferably contain from 5 to 50% by mass of nickel, in particular from 10 to 30% by mass. The contents of aluminum are in the range of 5 to 30% by mass, in particular in the range of 7 to 20% by mass. The proportions of silicon are in the range of 10 to 40% by mass, in particular in the range of 20 to 30 mass% (in each case based on the total metal content). When other components, these catalysts can be 0.1 to 2% by mass of alkali metal oxide, alkaline earth oxide, lanthanum oxide or oxides of rare earths and possibly forming aids.

The Nickel catalysts are useful in one Form used in which they have a low flow resistance offer, for. In the form of granules, pellets or shaped bodies, such as tablets, cylinders, spheres, extruded extrudates or rings.

The Oligomerization in process step h) is at (reaction) temperatures from 0 to 200 ° C, preferably 50 to 130 ° C, and pressures from 0.1 to 70 MPa, preferably from 0.1 to 10 MPa, and more preferably from 0.5 to 3 MPa.

The oligomerization according to process step h) is preferably operated in multiple stages. Each of the stages consists of a reaction part which can be formed by one or more reactors or reaction zones in which the oligomers are formed and a distillation which separates the oligomers from the C ₄ hydrocarbons. The separated C ₄ hydrocarbons are preferably recycled partly as a recycle stream in the reaction part, the rest is driven to the next reaction stage. The number of stages is preferably from 2 to 6, preferably 2 to 4.

Technically, the oligomerization reaction is preferably carried out in one or more shell and tube reactors. The C ₄ -hydrocarbons (mixture VI or fraction S1), optionally together with a C ₄ -Recyclestrom, passed through the tubes equipped with catalyst. The resulting heat of reaction can be removed via a cooling liquid on the shell side, so that a flat temperature profile is maintained over the catalyst in the tubes. One or more reactors can be used, which can be connected in parallel or in series. For example, in-line operation may be advantageous to limit the reactor length or to take advantage of the ability to operate the reactors at different temperatures. In the latter option, heating or cooling of the reaction mixture can additionally take place between the rectors.

The reaction mixture obtained from the reactors is separated in each stage in a distillation column into C ₄ hydrocarbons and a fraction having oligomers O fractionated. The distillation is preferably carried out at a pressure of 0.1 to 2 MPa, preferably at 0.2 to 0.5 MPa.

In a preferred embodiment of process step h), the oligomers O obtained in the bottom of the distillation columns are transferred from the front stages into the distillation column of the last stage led and there distilled together with the resulting in the last stage reaction. Such a preferred embodiment of process step h) is in 1 shown. This embodiment of the process step h) according to the invention has the advantage that in all other stages, the separation of C ₄ hydrocarbons and oligomers must be made less accurate, so still residues of C ₄ hydrocarbons in the oligomers and / or residues of oligomers in the C ₄ hydrocarbon remain. As a result, the necessary amount of steam for the operation of the columns is lowered. Preferably, however, not more than 1% by mass, more preferably 5000 ppm, of oligomers O are left in the C ₄ hydrocarbons and not more than 5000 ppm of C ₄ hydrocarbons in the oligomers.

Oligomerization of the n-butenes contained in S1 or VI gives in particular olefins having eight, twelve, sixteen, twenty or more carbon atoms as oligomers O. These olefins can z. B. for the production of plasticizer alcohols (C ₉ - or C ₁₃ alcohols) or of alcohols (C ₁₃ -, C ₁₇ - or C ₂₁ alcohols) are used for the preparation of detergent raw materials. Preferably, they are worked up before further processing by distillation into one or more fractions, wherein preferably the separation into a fraction with dibutene (mainly C ₈ olefins), a fraction with tributes (C ₁₂ olefins) and a fraction with higher oligomers (C ₁₆₊ olefins). Hydroformylation, hydrogenation and distillation of the dibutene give isononyl alcohols, which are widely used as plasticizer alcohols. From the tributene can be obtained by analogous reactions Isotridecylalkohole. From the C ₁₆₊ fraction, hybrids to the paraffins give access to mixtures of highly pure paraffins.

In a preferred embodiment, the distillative separation of the oligomers is carried out in each case a dibutene, tributene and C ₁₆₊ fraction in two operated under vacuum columns. In the first column, the dibutene is obtained as the top product _completely at a pressure of 100 to 800 hPa. The bottom product is separated in a second column at a pressure of 10 to 300 hPa _absolute in tributes (top product) and the C ₁₆₊ fraction (bottom product).

If desired, the fractions can be further separated, for example the C ₁₆₊ fraction in tetrabutene and C ₂₀₊ olefins. Also, the dibutene can be further separated for certain applications, such a separation is, for example, in EP 1 029 839 described.

In addition to the oligomers, a technical mixture VII is obtained in process step h), which has the unreacted in process step h) C ₄ hydrocarbons such as butane and optionally n-butenes.

Process step i): Optional further work-up of the residual C ₄ -hydrocarbons VII

There a complete conversion of butenes in the oligomerization (Process step h) economically usually does not make sense, falls as mixture VII a mixture of n-butenes and butane, the purely distillative not economical can be separated. This mixture VII can be used directly be, for example, as a raw material in a hydrocracker. Prefers However, this mixture VII is worked up either to n-butane or by extractive distillation in a butene-enriched and Butene-depleted fraction separated.

The Work-up to n-butane can be achieved by complete hydrogenation the still contained butenes take place. Technically, the hydrogenation according to the prior art, preferably on fixed-bed catalysts performed. Preference is given to the use of palladium-supported catalysts.

The n-butane obtained from the hydrogenation can be further purified, for example by distillative separation of isobutane or C ₅ hydrocarbons.

Further purification of the obtained, n-butane-containing stream, z. B. be done by this z. B. is washed in a wash column with water or an aqueous solution. The water-containing n-butane-containing stream which is obtained in this wash can then be worked up in a further distillation column into an almost anhydrous n-butane-containing stream. For this purpose, this distillation column can have a decanter in which the condensed overhead product is separated into an organic phase, which is recycled to the column at 50 to 100 mass%, and an aqueous phase which is discharged. The stream not recycled to the column, which contains low boilers such as isobutane, can be reused elsewhere. The n-butane-containing bottom product obtained in this distillation column has a water content of less than 50 ppm by weight (ppm by mass), more preferably less than 5 ppm by weight. This distillation column is before preferably with a head temperature of 40 to 60 ° C, more preferably operated with a head temperature of 45 to 55 ° C. By washing and / or subsequent separation of the water from the stream having n-butanes, polar impurities, which in particular influence the odor of the n-butanes, can be separated as far as possible.

Around It may be advantageous to remove traces of other impurities In addition, the current containing n-butanes in addition to one Subsequent purification, in which a treatment with one or more adsorbers, for. B. in adsorbent beds takes place. Common adsorbents are z. As activated carbon and molecular sieves.

The separation of butenes and n-butane by extractive distillation is also known in the art and is described, for example, in US Pat DE 102 42 923 and DE 10 2005 023 549 whose contents are expressly referred to.

feedstocks

In the process according to the invention, it is possible to use all customary technical C ₄ -hydrocarbon mixtures. Suitable isobutene-containing C ₄ streams are, for example, light petroleum fractions from refineries, C ₄ fractions from crackers (for example steam crackers, hydrocrackers, crackers), mixtures of Fischer-Tropsch syntheses, mixtures from the dehydrogenation of butanes, mixtures from skeletal isomerization of linear butenes and Mixtures formed by metathesis of olefins. These techniques are described in the specialist literature ( K. Weissermel, HJ Arpe, Industrial Organic Chemistry, Wiley-VCH, 5th Ed., 1998, pages 23 to 24; 65 to 99; 122 to 124 ).

C ₄ fractions are preferably used from steam crackers which are operated primarily for the production of ethene and propene and in which, for example, refinery gases, naphtha, gas oil, LPG (liquefied petroleum gas) and NGL (natural gas liquid) are used as raw materials, or C. ₄ fractions from cat crackers. Depending on the cracking process, the by-produced C ₄ cuts contain different amounts of isobutene, 1,3-butadiene, 1-butene, c-2-butene, t-2-butene, n-butane and i-butane.

The hydrocarbon mixtures with isobutene and linear butenes used in the process according to the invention preferably have the following compositions. Table 1: Typical compositions of technical hydrocarbon mixtures which can be used in the process according to the invention. steam cracker catcracker component CC ₄ (S) CC ₄ (K) Isobutane [% by mass] 0.6-6 37 n-butane [% by mass] 0.5-8 13 1-butene [mass%] 9-25 12 Isobutene [% by mass] 10-35 15 2-butenes [mass%] 4-20 23 1,3-butadiene [% by mass] 25-70 <1

explanation

• - CC ₄ (S): typical of a C ₄ mixture obtained from the crack C _{4 of} a steam cracker (High Severity).
• - CC ₄ (K): typical composition of a Crack-C ₄ obtained from a Katcracker.

Description of the figures

On the hand of the figure 1 the method according to the invention is explained in more detail below, without the method being restricted to the embodiments exemplarily illustrated there.

In 1 schematically a preferred embodiment of the method step h) is shown. In 1 is a two-stage execution of the process step h) shown. In each stage there is a reaction part or reactor R1 or R2 and one distillation D1 or D2 each. The technical mixture S1 or VI 1 first passes into the reaction part R1, in which n-butenes are oligomerized. The reaction mixture obtained from the reaction part R1 2 is passed into the distillation of the first stage D1 and there in a top stream 3 containing unreacted C4 hydrocarbons and a bottoms stream 4 containing the oligomers. A partial flow 3a the head current 3 can be returned before the reactor R1. The rest of the electricity 3b the head current 3 is fed to the second stage reaction section R2. The reaction mixture obtained from the reaction part R2 5 is fed to the distillation of the second stage D2. The bottom stream obtained in distillation D1 4 is also fed to the distillation D2. In the distillation D2 becomes a top stream 6 obtained containing unreacted C ₄ hydrocarbons and partially as a partial stream 6a can be returned before the reaction part R2. The rest part 6b the head current 6 can be used for further use. As a bottoms product 7 are obtained in the distillation D2, the oligomers O. On the presentation of procedural usual streams such. As cooling water streams and / or conventional equipment such. B. heat exchangers or separators was omitted in favor of a better overview.

QUOTES INCLUDE IN THE DESCRIPTION

This list The documents listed by the applicant have been automated generated and is solely for better information recorded by the reader. The list is not part of the German Patent or utility model application. The DPMA takes over no liability for any errors or omissions.

Cited patent literature

• EP 0048893 [0010]
• DE 2944457 [0010]
• US 4797133 [0011]
• - DE 10302457 [0012]
• - DE 2521964 [0013]
• - DE 102005062700 [0014]
• - EP 1029839 [0015, 0093, 0105]
• US 6337429 [0028]
• DE 19524971 [0029]
• - DE 102004021128 [0036]
• - DE 10102082 [0042, 0047]
• - EP 0428265 [0053]
• - EP 0396650 [0053]
• - DE 2987007 U [0053]
• - US 5244929 [0053]
• - DE 10020943 [0061]
• DE 3143647 [0067, 0070]
• - DE 19845857 [0090]
• - DE 3825169 [0090]
• - DE 3914817 [0090, 0093]
• - DE 102004018753 [0093]
• - DE 4339713 [0094]
• WO 01/37989 [0094]
• - DE 10242923 [0112]
• - DE 102005023549 [0112]

Cited non-patent literature

• Lawrence A. Smith, D. Hearn, in Catalytic Distillation, Proc. Intersoc. Energy Conversions. Conf. (1984) 19th, (Vol2), p998-1002 [0047]
• - Vauck / Müller, "Basic Operations of Chemical Process Engineering", p. 626, VEB Deutscher Verlag für Grundstoffindustrie [0055]
• - J. Mackowiak, "Fluid Dynamics of Columns with Modern Packages and Packings for Gas / Liquid Systems", Otto Salle Verlag 1991 [0056]
• - Hydrocarbon Process., Int. Ed. (1986) 65 (2nd section 1), pages 31 to 33 [0093]
• K. Weissermel, HJ Arpe, Industrial Organic Chemistry, Wiley-VCH, 5th Ed., 1998, pages 23 to 24; 65 to 99; 122 to 124 [0113]

Claims (20)

Process for the preparation of n-butene oligomers and 1-butene from technical mixtures I of C ₄ -hydrocarbons, which contain at least 1-butene, isobutene, butanes, 2-butenes and polyunsaturated C ₄ -hydrocarbons, characterized by the process steps a ) optional washing of the technical mixture I to obtain a technical mixture II, b) removal of 1,3-butadiene from the technical mixture I or II to obtain a technical mixture III, c) separation of the mixture in a technical mixture IIIa, with the technical mixture III is identical or obtained by mixing the mixture III with a C ₄ -hydrocarbon mixture, isobutene contained by reaction with an excess of methanol in a two-stage methyl tert-butyl ether synthesis and separating the resulting methyl tert-butyl ether and the unreacted methanol to give a technical mixture IV, wherein the first stage of the methyl tert-butyl ether synthesis in three fixed bed reactors and the second stage of methyl tertiary butyl ether synthesis on a solid acidic ion exchange resin having sulfonic acid groups in a reactive distillation column operating in a pressure range of 0.5 to 1.5 MPa and at a temperature in the reaction zone d) removal of polyunsaturated C ₄ hydrocarbons from the technical mixture IV to obtain a technical mixture V, by selective hydrogenation is carried out from 50 ° C to 90 ° C at a reflux ratio between 0.5 and 1.5 the polyunsaturated C ₄ hydrocarbons contained in the technical mixture with hydrogen in the liquid phase on a fixed bed catalyst of 0.1 to 2 mass% palladium on a support, in the presence of 0.05 to 10 ppm by mass carbon monoxide, based on the masses of the technical mixture IV, at a temperature from 0 to 7500 e) Distillative separation of the technical mixture V into a Leic low-boiling fraction L1, which has 1-butene and isobutane and which is substantially free from n-butanes and 2-butenes, and a high-boiling fraction S1 which comprises 2-butenes and n-butane, f) distillative separation of the low-boiling fraction L1 in 1- G. Optional treatment of the high-boiling fraction S1 with a molecular sieve having an average pore diameter of greater than 0.4 to 1.5 nm at a pressure of 0.1 to 5 MPa and a temperature of 20 to 160 ° C to obtain a technical mixture VI and h) oligomerization of the n-butenes contained in the high-boiling fraction S1 or in the technical mixture VI to obtain oligomers O at temperatures of 0 to 200 ° C and pressures of 0.1 to 7 MPa on a solid nickel, silicon and aluminum containing catalyst. Method according to claim 1, characterized in that that in process step a) as the washing medium, water or an aqueous Solution is used. Method according to claim 1 or 2, characterized that in method step a) a mass ratio of Washing medium to technical mixture I of 1 to 1 to 1 used to 100 becomes. Method according to at least one of claims 1 to 3, characterized in that the 1,3-butadiene is obtained in process step b) by a selective hydrogenation from the technical mixture I or II, wherein the selective hydrogenation at a pressure of 0.5 to 2 MPa and a temperature of 25 to 80 ° C in the presence of hydrogen over a palladium-supported catalyst having as a carrier Al ₂ O ₃ and SiO ₂ and 0.2 to 1 mass% of palladium based on the carrier. Method according to claim 4, characterized in that that the selective hydrogenation is carried out in two stages with the first stage reactor operating in loop mode becomes. Method according to at least one of the claims 1 to 5, characterized in that the first of the fixed bed reactors in the first stage of process step c) at a temperature from 35 to 70 ° C and the subsequent fixed bed reactors be operated at a temperature of 25 to 50 ° C. Method according to at least one of the claims 1 to 6, characterized in that the column feed to the reactive distillation column in the second stage of process step c) a temperature of 50 to 90 ° C. Method according to at least one of claims 1 to 7, characterized in that in step c) from the top product of the reactive distillation column of the second stage, the methanol is separated by extraction, wherein the mass ratio of washing medium to top product 1 to 5 to 1 to 40 be wearing. Method according to at least one of the claims 1 to 8, characterized in that the reactive distillation column in process step c) operated at a pressure of 0.7 to 0.9 MPa becomes. Method according to at least one of the claims 1 to 9, characterized in that the reactive distillation column in process step c) with a hydraulic load of the catalytic Pack of 10% to 110% is operated. Method according to at least one of the claims 1 to 10, characterized in that the pressure in process step d) is less than 20 MPa. Method according to at least one of the claims 1 to 11, characterized in that the method step d) is carried out in two stages. Method according to at least one of the claims 1 to 12, characterized in that the method step e) is carried out in a column that is from 150 to 200 having theoretical plates. Method according to at least one of the claims 1 to 13, characterized in that in process step e) the Distillation column with a reflux ratio is operated by less than 11. Method according to at least one of the claims 1 to 14, characterized in that the method step f) is carried out in a column that is from 150 to 200 having theoretical plates. Method according to at least one of the claims 1 to 15, characterized in that in method step f) the Distillation column with a reflux ratio from 30 to 60 is operated. Method according to at least one of the claims 1 to 16, characterized in that in process step g) zeolites of the A, X or Y type can be used as a molecular sieve. Method according to at least one of the claims 1 to 17, characterized in that in process step h) a Catalyst containing from 5 to 50% by mass of nickel, from 5 to 30% by mass of aluminum and from 10 to 40% by mass of silicon based on the total metal content. Method according to at least one of the claims 1 to 18, characterized in that method step h) in several stages operated with a number of 2 to 4 stages. Method according to claim 19, characterized that in each case one stage in process step h) a reaction part and having a distillation and that in the bottom of the distillation column / n resulting oligomers O from the front stages / the front stage led to the distillation column of the last stage and there together with the incurred in the reaction part of the last stage Be distilled off reaction.