WO2008037693A1 - Continuous method for producing citral - Google Patents

Abstract

The invention relates to a method for producing 3,7-dimethyl-octa-2,6-dienal (citral). Said method consists of the following steps and enables the product to be mass produced: a) Initially, 3-methyl-3-butene-1-ol (isoprenol) is produced from isobutane and formaldehyde, b) 3-methyl-2-butenal (prenal) and 3-methyl-3-butenal (isoprenol) are produced from 3-methyl-3-butene-1-ol (isoprenol) by oxidative dehydration by means of an oxygen-containing gas on a silver support catalyst, c) additional 3-methyl-2-butenal (prenal) is produced from a mixture containing 3-methyl-3-butenal by isomerisation d) the 3-methyl-2-butene-1-ol (prenol) is produced from 3-methyl-3-butene-1-ol (isoprenol) by isomerisation, e) the unsaturated acetal 3-methyl-2-butenal-diprenylacetal is produced from 3-methyl-2-butene-1-ol (prenol) and 3-methyl-2-butenal (prenal) using an acidic catalyst and f) the citral is obtained from 3-methyl-2-butenal-diprenylacetal by splitting and subsequently rearranging.

Description

 Continuous process for the production of citral

description

The present invention relates to a process for the preparation of citral (3,7-dimethyl-octa-2,6-dienal) in which the product citral can be obtained in an environmentally friendly manner in high yields starting from cheap products. For many years, various synthetic routes for the production of citral are known. Citral is a valuable intermediate for the production of various odors and fragrances, such as geraniol. In addition, citral has also gained importance as a source material for the production of vitamins, in particular of vitamin A. Citral can also be obtained from lemongrass in a complex process, which, however, is not economical in view of the high demand for citral. Another conventional large-scale process, which is based on ß-pinene as starting material, can not be carried out economically on a large scale due to the environmentally harmful by-products.

It is an object of the present invention to provide a preferably continuously process for the production of citral, in which citral can be obtained in high yield starting from cost-effective starting materials in a complete process comprising a few steps. In this case, the intermediates or the unreacted starting materials formed in the various reaction stages should preferably be recirculated back into the production process in order to reduce the amount of by-products which pollute the environment.

This object is achieved by a process for the preparation of 3,7-dimethyl-octa-2,6-dienal (citral), in which citral is essentially obtained in six process steps:

a) starting from isobutene and formaldehyde, 3-methyl-3-buten-1-ol (isoprenol) is first prepared,

b) starting from 3-methyl-3-buten-1-ol (isoprenol) by oxidative dehydrogenation with an oxygen-containing gas in a continuous process on a supported silver catalyst 3-methyl-2-butenal (Prenal) and 3-methyl 3-butenal (isoprenal) (prepared as a mixture), c) starting from a 3-methyl-3-buten-al-containing mixture is prepared by isomerization in the presence of a catalyst such as sodium acetate, further 3-methyl-2-buten-al (prenal),

d) starting from 3-methyl-3-buten-1-ol (isoprenol) by isomerization in

Presence of hydrogen and a noble metal-containing catalyst, for example in suspension or fixed-bed operation, the 3-methyl-2-buten-1-ol (prenol) is prepared,

e) starting from 3-methyl-2-buten-1-ol (prenol) and 3-methyl-2-butenal

(Prenal) using an acidic catalyst and separating the water formed in the reaction, the unsaturated acetal 3-methyl-2-butenal-diprenylacetal is prepared, and

f) starting from 3-methyl-2-butene-al-diprenylacetal by cleavage in the presence of a catalyst and subsequent Claisen rearrangement of the resulting butadienyl ether, a 2,4,4-trimethyl-3-formyl-1, 5-hexadiene is prepared, from which is then obtained by Cope rearrangement of the 3,7-dimethyl-octa-2,6-dien-al (citral).

In the process according to the invention, starting from isobutene and formaldehyde, the end product citral is obtained in high yield as cost-effective reactants in a process which is also easy to control industrially. The overall process is preferably carried out continuously in the individual stages, in particular as a fully continuous overall process.

Another embodiment of the invention relates to a process for the preparation of citral, wherein the process is carried out fully continuously and the by-products from process steps b) and d) are introduced again into the process.

An embodiment of the invention relates to a process in which the process is carried out fully continuously and the by-products from process steps a), b), d) and e) are introduced again into the process in each case.

A further embodiment of the invention relates to a process for the preparation of citral, wherein in the isomerization of 3-methyl-3-butenal sodium acetate is used as a catalyst and that the entire process is carried out fully continuously and the by-products from the process steps a), b), c), d) and e) are each reinjected into the procedure. The subject matter is also a process for the preparation of citral, wherein the oxidative dehydrogenation of 3-methyl-3-buten-1-ol (isoprenol) with an oxygen-containing gas on a supported silver catalyst (process step b) is carried out such that 3-methyl-3-buten-1-ol is evaporated, then this vapor is added to the oxygen-containing gas, the thus obtained, oxygen-containing alcohol vapor initially at a temperature above the dew point of 3-methyl-3-butene-1 but below the starting temperature of the reaction by an at least 0.5 cm thick layer of supported silver catalyst is passed and then the oxygen-containing vapor of 3-methyl-3-buten-1-ol in a desired capacity corresponding number of arranged parallel and flow around with a fluid heat transfer reaction tubes having an inner diameter of 1 to 2 cm and a length of 35 to 60 cm and filled with the supported silver catalyst t, is reacted at temperatures of 300 ° to 600 ⁰ C to a mixture of isomeric aldehydes.

A further embodiment of the invention relates to a process for the preparation of citral, wherein in the isomerization of 3-methyl-3-buten-1-ol (isoprenol) to 3-methyl-2-buten-1-ol (prenol) (process step d ) in the presence of hydrogen as a noble metal-containing fixed bed catalyst, a catalyst containing palladium and selenium or tellurium or a mixture of selenium and tellurium on a silica support is used.

One embodiment of the invention relates to a process for the preparation of citral, which comprises the starting materials 3-methyl-2-buten-1-ol (prenol ) and 3-methyl-2-butenal (Prenal) are only partially reacted in the reaction column and the resulting unsaturated acetal is concentrated in at least two successive evaporator stages and the recovered starting materials are recycled back to the reaction column.

A further embodiment of the invention relates to a process for the preparation of citral, which comprises the intermediates 3-methyl-2-butene in the thermal cleavage of the 3-methyl-2-butenealdiprenyl acetal and the subsequent rearrangements (process step f)). 1-ol (prenol); Prenyl (3-methyl-butadienyl) - ethers and 2,4,4-trimethyl-3-formyl-1, 5-hexadiene and the final product citral during the reaction continuously removed by distillation from the reaction mixture.

The invention also relates to a process for the preparation of citral, in which the citral obtained by thermal cleavage of the 3-methyl-2-butenal-diprenyl acetal and the subsequent rearrangements is separated off from the crude mixture by rectification, where: a) the crude mixture is introduced laterally into an intake column with a reinforcing part located above the feed point and a driven part located below the feed point,

b) an upper communicating with the upper end of the reinforcing member

Union column with condenser at the upper end of the column and communicates with the lower end of the stripping section communicating lower union column with heaters at the lower end of the column,

c) provides a withdrawal column communicating with the upper union column and the lower union column,

d) removed from the withdrawal column the citral as a side draw and deducting the higher than citral boiling compounds at the top of the upper union column lower boiling than citral compounds and in the bottom of the lower union column.

The individual process steps are described in more detail below.

The synthesis of 3-methyl-3-buten-1-ol (isoprenol) starting from formaldehyde and isobutene (process step a) has been known to the skilled person for decades.

Isomerization of isoprenol to 3-methyl-2-buten-1-ol (prenol) can be carried out in the presence of hydrogen and a catalyst (process step d), wherein a fixed bed catalyst can be used as the catalyst, in particular the palladium and selenium or tellurium or containing a mixture of selenium and tellurium on a silica support. From J. Am. Chem. Soc., 85 (1963), pages 1549 to 1550, is the isomerization of an unsaturated alcohol with a carbonyl compound of a metal of VIII. Group of the Periodic Table of the Elements known as a catalyst. In this process, numerous by-products and secondary products are obtained, for example the corresponding aldehydes.

DE-C-19 01 709 describes a process for the preparation of buten-2-ol-4 compounds in which buten-1-ol-4 compounds are reacted in the presence of palladium or palladium compounds and hydrogen. However, when pure palladium is used in the presence of hydrogen, the double bond of the compounds is significantly hydrogenated and a saturated product is formed.

In addition, low-boiling compounds such as hydrocarbons and aldehydes are formed as by-products, for example by hydrogenolysis and isomerization.

The hydrogenation of the double bond is undesirable because some butenols only low boiling point differences between unreacted starting material and hydrogenation exist. Thus, the boiling point of 3-methyl-3-buten-1-ol, for example, at 1020 mbar 131, 5 ⁰ C, the boiling point of the corresponding hydrogenation product 3-methyl-1-butanol 130.9 ⁰ C. This makes a distillative separation of Hydrogenation product and starting product difficult.

From DE-A 27 51 766 a process for the isomerization of 3-buten-1-ol compounds to the corresponding 2-buten-1-ol compounds is known, wherein the isomerization in the presence of palladium and selenium or tellurium as a catalyst and hydrogen is carried out. The catalyst used is palladium and selenium on activated carbon. Other usable supports include barium sulfate, silica gel, alumina and zeolites. The catalysts can also be used without carriers. Relatively high proportions of low-boiling components, such as isoprene and methylbutene, are formed. The known catalytic isomerizations are carried out batchwise, for example in a stirred tank in suspension mode. Since the double bond isomerization of substituted butenols represents an equilibrium reaction, however, no complete conversions are obtained, but it always remains a part of the starting material, which must be separated for the further use of by-products formed. In order to carry out the isomerization more economically, the reaction should be able to be carried out continuously and lead to a minimal proportion of hydrogenation products or low-boiling components.

EP-A 841 090 discloses the provision of a process for the continuous preparation of 2-buten-1-ol compounds by isomerization of 3-buten-1-ol compounds, the proportion of resulting hydrogenation products and low-boiling components being very low. A fixed-bed catalyst containing palladium and selenium or tellurium or a mixture of selenium and tellurium on a silica support is used, and a BET surface area of 80 to 380 m ² / g and a pore volume of 0.6 to 0.95 cm ³ / g in the pore diameter range of 3 nm to 300 microns, wherein 80 to 95% of the pore volume in the pore diameter range of 10 to 100 nm.

The catalyst contains 0.1 to 2.0 wt .-% palladium and 0.01 to 0.2 wt .-% selenium, tellurium or a mixture of selenium and tellurium, based on the total weight of the catalyst. The BET surface area is, for example, 100 to 150 m ² / g, in particular 110 to 130 m ² / g. The BET surface area is determined by N ₂ adsorption in accordance with DIN 66131. The pore volume in the pore diameter range from 3 nm to 300 μm is preferably 0.8 to 0.9 cm ³ / g, in particular 0.8 to 0.85 cm ³ / g. In this case, 80 to 95%, preferably 85 to 93% of this pore volume in the pore diameter range of 10 to 100 nm. The pore volume is determined by Hg porosimetry. Preferably, the catalyst contains 0.2 to 0.8 wt .-%, in particular from 0.4 to 0.6% by weight of palladium. Preferably, the catalyst contains 0.02 to 0.08, in particular 0.04 to 0.06 wt .-% selenium, tellurium or a mixture of selenium and tellurium, preferably selenium. In addition to the active components mentioned, further metals can be present on the catalyst to a small extent. Preferably only palladium, selenium and / or tellurium, in particular only palladium and selenium, are present on the silica support. The described isomerization of isoprenol to prenol on a fixed bed catalyst is also described in EP-A 841 090, to which reference is expressly made.

Starting from 3-methyl-3-buten-1-ol (isoprenol), oxidative dehydrogenation with an oxygen-containing gas yields another intermediate, the reactive 3-methyl-2-butenal (prenal). Since Prenal is a sensitive intermediate, care must be taken to ensure that the reaction conditions for the catalytically accelerated dehydrogenation are strictly adhered to. A particular problem with this reaction (process step b)) is that it is highly exothermic, the reaction rate is highly dependent on the reaction temperature and that the reactants and the reaction products are extremely unstable.

It is known from US 2,042,220 that isoprenol with an excess of oxygen at 360 to 550 ⁰ C in the presence of metal catalysts, eg. As copper and silver catalysts, to 3-methyl-3-butenal (isoprenal) can oxidize. The catalysts may be, for example, alloys, metal compounds or elemental metal. Preference is given to activated catalysts. Activation operations are surface amalgamation of the metal followed by heating of the metal surface. The preparation of copper and silver catalysts in the examples consists in the reduction of copper oxide particles at 300 ° C under hydrogen atmosphere or in the amalgamation and heating of silver wire nets. According to DE 20 41 976 (BASF, 1972), however, considerable amounts of undesired by-products are formed in this process known from US Pat. No. 2,042,220 (Shell, 1936).

In DE 25 17 859 (Teijin, 1976), the dehydrogenation of unsaturated alcohols on a copper catalyst having a specific surface area of 0.01 to 1.5 m ² / g is described essentially in the absence of oxygen at 150 to 300 ^° C. , In the case of α, ß-unsaturated alcohols as starting materials ß, γ-unsaturated aldehydes and saturated aldehydes are formed as by-products; the selectivity for α, ß- unsaturated aldehydes is low. Such mixtures have to be broken down into their components in complex separation operations. DE 20 20 865 (BASF, 1971) and DE 20 41 976 (BASF, 1972) describe the dehydrogenation of β, γ-unsaturated alcohols or α, β-unsaturated alcohols to α, β-unsaturated aldehydes. As dehydrogenation catalysts and mixed catalysts, for. B. called those of copper and silver. The disadvantage, however, is that you must add significant amounts of nucleophilic substances. In the case of the reaction of 3-methyl-3-buten-1-ol good results are obtained only in incomplete conversions, which leads to the teaching of DE 22 43 810 (BASF, 1973) to difficulties in the separation of the unreacted starting material.

If isoprenol is dehydrogenated by the process of DE 25 17 859 without addition of oxygen to metallic copper, considerable quantities of isovaleraldehyde are formed and the activity of the catalysts drops rapidly within a few days, so that frequent regeneration is required.

FR-A-2 231 650 (Givaudan, 1974) describes the preparation of aldehydes and ketones from the corresponding alcohols by air oxidation at 250 to 600 ^° C. in the presence of a gold catalyst. The advantage of the gold catalyst is the higher selectivity compared to copper and silver catalysts, thus reducing by-product formation. Disadvantages of this process are the high catalyst costs, since a solid gold contact is used.

The patents DE 27 15 209 (BASF, 1978) and EP-B 55 354 (BASF, 1982) describe the oxidative dehydrogenation of 3-alkyl-buten-1-ols on catalysts consisting of layers of silver and / or copper crystals exist, with the addition of molecular oxygen described. The amounts of oxygen are in the range of 0.3 to 0.7 mol, based on the starting material. A disadvantage of this method is that high catalyst costs occur through the use of solid silver and good selectivities can only be achieved if defined catalyst grain sizes or catalyst particle size distributions in a layer structure, possibly even certain mixtures of layers of copper and Silver crystals are used. This means both a complex filling of the reactor and a complex catalyst recovery. In addition, occurs at the high reaction temperatures used sintering of the metal crystals, which leads to pressure increase and short service life.

According to EP-A 244 632 (BASF, 1987), the continuous oxidative dehydrogenation even of unsaturated aliphatic alcohols to the corresponding aldehydes quite advantageously, if the reaction is carried out in the gas phase at 300 to 600 ⁰ C on a suitable catalyst, which in between Tube sheets arranged short and thin reaction tubes, which is located in a fluid heat transfer medium flow through the side. The process according to EP-A 244 632 turned out to be very advantageous when the vaporous gas mixture emerging from the tube bundle reactor was mixed shortly after contact with the catalyst with condensed reaction mixture containing water and / or unreacted alcohols at temperatures of -20.degree to +50 ⁰ C, from the condensate obtained, the aldehydes is separated off, as described in EP-A 55 354 and the unreacted alcohols then leads back into the process.

A disadvantage of this very advantageous combination of processes is that in the continuous implementation of the process despite frequent burning off of soot and other incrustations on the catalyst after a few weeks, both the conversion and the selectivity drops sharply, so that the catalyst must be replaced. When opening such a reactor showed that the reaction tubes are clogged for the most part. The plugs in the individual tubes are so hard that the affected tubes must be drilled with a drill, which is technically very complicated and can lead to damage to the pipes. The slow clogging of the individual tubes can be prevented permanently by pre-switching an ordered packing of metal mesh as a pre-filter nor by overlaying the catalyst layer with inert particles, such as glass or porcelain balls, as a filter. Another disadvantage of this method is that the construction of tube bundle reactors with very short reaction tubes for such tube bundle reactors in which high capacity, d. H. to work with a very large number of reaction tubes, prepares more technical difficulties.

It is therefore important to carry out the above-described process step according to EP-A 244 632 for the continuous industrial preparation of unsaturated aliphatic aldehydes by oxygenating dehydrogenation of the corresponding alcohols with an oxygen-containing gas over a supported catalyst consisting of copper, silver and / or gold on an inert support in one Shell and tube reactor, rapid cooling of the reaction gases and separating the aldehydes from the resulting condensate to improve so that the disadvantages of the prior art do not occur, d. h., That you can operate the continuous process for long periods, possibly over several years, without shutting down the system and without costly drilling out of the reaction tubes in a readily manufacturable reactor.

The object of this reaction step is a process for the continuous preparation of the unsaturated aliphatic aldehyde isoprenal

CH, ^■ CK CHO (A)

by oxygenating dehydrogenation of one of the two or a mixture of both 3-methyl-1-butene-1-ols of the formula (Ba) or (Bb)

CH. CH,

CH. C-CH ₀ -CH ₉ OH CH. CH - CH ₂ OH

(Ba) (Bb)

(Isoprenol and / or Prenol) with an oxygen-containing gas over a supported catalyst consisting of copper, silver and / or gold on an inert support in a tube reactor, rapidly cooling the reaction gases and separating the aldehydes from the resulting condensate,

a) vaporizing the vapor of one or both of the 3-methyl-buten-1-ols,

b) adding an oxygen-containing gas to the alcohol vapor,

c) the obtained oxygen-containing alcohol vapor initially at temperatures above the dew point of the alcohol but below the starting temperature of the

Reaction through an at least 0.5 cm thick layer, preferably a 0.6 to 5 cm thick layer, one of the abovementioned supported catalysts, which preferably occupies the cross section of the entire reactor, conducts, and only thereafter

d) the oxygen-containing alcohol vapor in a number corresponding to the desired capacity of parallel arranged and surrounded by a fluid heat carrier and filled with one of said support catalyst reaction tubes having an inner diameter D of about 0.5 to 3 cm, preferably 1 to 2 cm , and a length of at least 5 cm, preferably 35 to 60 cm, at temperatures of 300 to 600 ⁰ C to the corresponding aldehyde react. With regard to the more precise procedure of the dehydrogenation, reference is made to DE-A 197 22 567 (BASF, 1998).

In process step e) of the above-mentioned production process of citral, the unsaturated acetal 3-methyl-2-butenal-diprenylacetal is prepared starting from prenol and prenal using a catalyst. For this purpose, prenal is reacted together with prenol in the presence of catalytic amounts of acid and with removal of the water formed during the reaction.

The components are reacted only partially in a reaction column, the acetal obtained is concentrated in at least two successive evaporator stages and the recovered components are returned to the reaction column.

The preparation of unsaturated acetals by reacting olefinically unsaturated aliphatic compounds with allyl alcohols in a reaction column in the presence of a distillable acid is known per se from DE 26 25 074 (BASF, 1977). There it is described that a mixture of at least 2 mol of the alcohol and one mole of the aldehyde is introduced into the reaction column and the water formed in the reaction is distilled overhead and withdrawn by means of a phase separator. From the evaporator of the column then the acetal should be removed as a crude product. The reaction column should be operated so that the aldehyde is no longer contained in the bottom of the column, so the aldehyde is completely implemented in the reaction column. In the examples reaction reactions for the aldehyde of more than 94.5% are called.

The method described represents an advance for the production of acetals, but unfortunately has some disadvantages. Since, according to the known state of the art, high conversions of over 90% are desired, the production plant is difficult to control. Small fluctuations in the amount of inflows or in the purities of the starting materials cause the desired conversion in the reaction column can not be met. This fact is particularly noticeable when using feedstocks with impurities. Such impurities are e.g. Derivative products of the aldehyde and the alcohol, e.g. Formates of the alcohol, ethers formed from the alcohol or condensation products of alcohol and aldehyde linked via C-C bonds. As is known, such by-products are formed in greater quantities, in particular when the acetal is subjected to a Claisen and Cope cleavage and rearrangement reaction, as described, for example, in US Pat. in the further manufacturing steps of citral is the case.

Difficult is also the addition of the right amount of acid. Here, even a slight underdosing of the amount of acid to break down the Um- record in the reaction column. But if more than the right amount of acid is added, this leads to a large increase in the amount of high-boiling secondary components and ethers. The control of the amount of acid is made more difficult by the fact that the acid accumulates in the reaction column and as a result an overdose or an underdosing is usually not noticed until after a considerable time delay.

The abovementioned disadvantages have hitherto hindered the economical production of the acetals in a reaction column on an industrial scale.

The apparatus for producing the unsaturated acetals preferably consists of a distillation column which is used as a reaction column. The rising at the top of the column vapors are condensed in the condenser and passed into a phase separation vessel in which the water separates as the lower phase. The upper phase consists essentially of organic compounds (aldehyde, alcohol and low boiling side compounds such as the formates of the alcohol used). Most of the organic phase is recycled as reflux to the top of the column, a smaller portion discharged to remove the minor components.

The amount of reflux is per 1000 kg freshly added aldehyde 200 kg to 50,000 kg, in particular 1000 kg to 20,000 kg. This small amount of reflux and thus the low energy requirement represent a particular advantage of the process step. The amount of discharge is - depending on the purity of the starting materials used - per 1000 kg freshly added aldehyde between 1 kg and 400 kg, in particular between 5 kg and 200 kg. The bottom of the column passes into the evaporator, which is the first stage of the two-stage concentration of the acetal.

The vapors recovered in the evaporator consist of 10 wt .-% to 80 wt .-% alcohol, up to 10 wt .-% of the acetal and 10 wt .-% to 40 wt .-% of the aldehyde. The conversion of aldehyde in the reaction column is therefore below 90%. It is possible to dispense with the condensation of the vapors and to return the vapors in gaseous form to the column. Preferably, however, the vapors are condensed in the condenser. The amount of vapors generated in the vaporizer stage, that is, the amount of reactants recycled in the first vaporizer stage is between twice and thirty times, more preferably between three and twenty times, the amount of freshly added aldehyde.

Too small an amount of the reactants recycled in the first evaporator stage will result in a loss of selectivity. Too high an amount in the first evaporator Although the stage of recycled reactants is favorable for the selectivity of acetal synthesis, it consumes unnecessarily large amounts of energy. A great advantage of the process according to the invention is its easy controllability and stable operation, since the amount of reactants to be recycled in the first evaporator stage can easily be determined by observing the temperatures in the column and the evaporators.

Optionally, a reactor can be used, in which the condensate obtained in the condenser and the starting materials alcohol and aldehyde are added. The reactor is used to adjust the thermodynamic equilibrium between alcohol and aldehyde on the one hand and water and acetal on the other. For the reactor, backmixed reactors (such as stirred tanks) but advantageously non-backmixed tubular reactor reactors (such as packed columns) can be used. The residence time of the reaction mixture in the reactor is between 0.1 sec and 10 hours. Since the thermodynamic equilibrium often sets very fast, may be sufficient for a very short residence time. Then, if necessary, can be dispensed with a special apparatus, and serve the already existing pipeline between the mixing point of alcohol, aldehyde and the condensate obtained in the condenser as a reactor. Optionally, to mix the feedstocks with the condensate recovered in the condenser, the usual mixing means such. static mixers or stirred tanks are used.

The acid can be added to the reactor or the evaporator. Preferably, however, the acid is added to the reaction column and added there again advantageously in the lower part of the column or in the evaporator, if it is a volatile and in particular nitric acid. It is also possible to add the acid in different places, e.g. at two or more points in the column or partly in the column and partly in the reactor. The liquid emerging from the reactor should preferably be added in the upper part of the column. It is also possible to add the liquid, optionally mixed with the reflux, directly as reflux to the top of the column.

The addition point of fresh aldehyde and fresh alcohol is not critical. The two starting materials can also be added separately at different points of the column and / or the reactor. The starting materials are preferably combined with the effluent from the condenser. The added amount of freshly added alcohol is controlled so that the ratio of alcohol to aldehyde is between 1 and 3, in particular between 1, 5 and 2.5. A particular advantage of the method step according to the invention is that even the contaminated starting materials described above can be used without problems. The liquid leaving the evaporator usually contains the acetal in a concentration of between 10% by weight and 70% by weight. The remainder consists essentially of the aldehyde and the alcohol. This liquid is placed in a downstream evaporator where the acetal is recovered in a concentration between 30 wt .-% and 99.9 wt .-% as a liquid, in particular from 50 wt .-% to 95 wt .-%.

The vapors rising from the evaporator are preferably returned to the bottom of the column and thus used to heat the column. With regard to the exact reaction conditions for the preparation of the unsaturated acetal, reference is also made to EP-A 1 186 589 (BASF, 2002).

Starting from the diprenyl acetal prepared in process step e), the intermediate 2,4,4-trimethyl-3-formyl-1,5-hexadiene can be obtained by thermal cleavage in the presence of a catalyst and subsequent Claisen rearrangement of the resulting butadiene ether (process step f). , which can then be converted to citral by a Cope rearrangement.

Subject of the method section f) is accordingly an optimized process for the preparation of 3,7-dimethyl-2,6-octadienal (citral) of the formula I.

by thermal cleavage, if appropriate in the presence of an acidic catalyst of 3-methyl-2-butenal-diprenyl acetal from process section e) of the formula II

with cleavage of 3-methyl-2-buten-1-ol (prenol) of the formula to the cis / trans-prenyl- (3-methyl-butadienyl) ether of the formula IV

Claisen rearrangement of the resulting butadienyl ether of formula IV gives 2,4,4-trimethyl-3-formyl-1,5-hexadiene of formula V.

which can then be isomerized by a subsequent Cope rearrangement to the citral of the formula I. In this case, both the prenol formed and the intermediately formed intermediates of the formulas IV and V and the citral already during the reaction can be removed by distillation from the reaction mixture continuously and thus achieve improved yields.

With particular advantage, the process step succeeds in carrying out the thermal cleavage of the acetal of the formula II in the lower part or in the bottom of a distillation column which functions as a cleavage column and has from 5 to 100 theoretical plates. It has proved to be very useful when introducing the acetal of formula II and optionally the acid catalyst in the lower part of the distillation column, in the bottom of the distillation column or in the evaporator of the distillation column.

One possibility for carrying out the process step according to the invention is characterized in that one carries out the thermal cleavage of the acetal of the formula II in the lower part or in the bottom of a distillation column with 5 to 100 theoretical plates, wherein the acetal of the formula II by suitable selection of the distillation conditions in holding lower part or in the bottom of the column, the citral of the formula I formed, the intermediately formed intermediates of formulas IV and V and the split prenol of the formula III collectively together at the top of the column as overhead and then the mixture by distillation, for. In a further distillation column. With particular advantage, the inventive method, if one carries out the thermal cleavage of the acetal of formula II in the lower part or in the bottom of a distillation column with 5 to 100 theoretical plates, wherein the acetal of the formula II by suitable selection of the distillation conditions in the lower part or holds in the bottom of the column, the citral of formula I formed and the intermediately formed intermediates of formulas IV and V together liquid or vapor withdraws on a arranged in the middle or lower part of the column side draw and the cleaved prenol of the formula III at the top of Separate column with the overhead stream.

The mixture of citral withdrawn at the side draw or from the top stream of a column without side draw and the intermediately formed intermediates of the formulas IV and V is then suitably passed through a heated residence time tube in which the intermediates of formulas IV and V at temperatures of 100 be transferred to 200 ⁰ C to citral. The acetal required as the starting compound can, as described above, be prepared by reacting 3-methyl-2-buten-1-ol (prenol) with 3-methyl-2-butenal (Prenal). Depending on the conditions of preparation of the acetal, this may contain between 0.1 and 30% by weight of unreacted prenal and between 0.1 and 60% by weight of unreacted prenol. It is advantageous for the process if the concentration of the acetal of the formula II used in the starting material is above 30% by weight, preferably above 70% by weight.

The cleavage of the acetal of the formula II to prenol, citral and the citral precursors of the formulas IV and V takes place in a distillation column, the so-called split column, which is equipped with an evaporator and a condenser. Suitable internals for the column are trays, fillers and in particular structured packings of sheet metal or metal mesh. The number of theoretical plates of the column should be between 5 and 100. In a preferred embodiment, there is a side draw from which the citral and citral precursors are in vapor or liquid

Form be taken from the column. The side draw is located in the middle or lower part of the column, wherein the side draw is expediently between 2 and 20 theoretical plates above the addition point of the acetal. Above the

Side take-off are still between 2 and 80 theoretical plates. The mixture of citral and the citral precursors is passed through a heated residence time tube in which the citral precursors of formulas IV and V are thermally rearranged to citral. The recovered citral can then be removed at the sampling point.

The recovered from the top stream in the condenser Prenol and optionally 3-methyl-2-butenal can be removed at the sampling point and again for the preparation of the

Acetals of formula II are recycled in the process. The in the process accumulating relatively small amounts of sump drain can be removed at the sampling point.

The acetal is added in the lower part of the cleavage column, but preferably in the bottom of the cleavage column or in the evaporator. Optionally, the bottom of the cleavage column can be enlarged by a container to have a larger reaction volume available. The inflow to the acetal of the formula II is expediently such that the residence time relative to the inflow of acetal is between one minute and 6 hours, preferably between 5 minutes and 2 hours. This relatively short residence time and the resulting smaller cost-effective equipment represent an advantage of the process step. This process section can in principle be carried out batchwise, semicontinuously or continuously. For the semi-continuous process, a certain amount of acetal is initially charged at the beginning of the reaction and then optionally further acetal is added. With particular advantage, the process is carried out continuously.

If desired, a high-boiling inert compound can be introduced to start the apparatus in the bottom of the column to ensure a minimum level of the sump and the evaporator. Suitable are all inert under the reaction conditions liquid substances which have a higher boiling point than citral and in particular as the acetal of formula II, such as the hydrocarbons tetradecane, pentadecane, hexadecane, octadecane, eicosane, or ethers such as diethylene glycol dibutyl ether, white oils, paraffin oils or mixtures of said compounds. The reaction can be carried out without catalyst, ie only by heating. With particular advantage, however, it is carried out in the presence of an acidic catalyst. Non-volatile protic acids, such as sulfuric acid, p-toluenesulfonic acid and, above all, phosphoric acid are particularly suitable as acidic catalysts. The catalyst is advantageously added to the lower part of the cleavage column, preferably into the bottom of the cleavage column or into the evaporator.

With regard to the exact execution of the thermal cleavage, the Claisen rearrangement and the subsequent Cope rearrangement, reference is made to DE-A 198 46 056 (BASF, 2000).

The invention is further illustrated by the following examples, wherein the individual stages shown in the examples are carried out continuously in the process according to the invention: Example 1: Preparation of 3-methyl-3-buten-1-ol (MBE) in a 3000 ml autoclave

Raw material:

2053 g (36.7 mol) isobutene (from a steel bottle)

200 g (3.3 mol) of formaldehyde (aqueous solution, 50%)

1, 4 g urotropin solution (aqueous 15%)

1a) Implementation:

The formaldehyde solution and isobutene from a steel bottle are placed in an autoclave and the autoclave is closed. The autoclave is heated to 260 ⁰ C, whereby an autogenous pressure of about 100 bar is formed. Then the autoclave is pressed with nitrogen to about 250 bar. The mixture is stirred at about 260 ⁰ C and about 250 bar for one hour. After one hour, allow to cool to room temperature and relax. Released isobutene is collected (about 1800 g) and can be used for other approaches. The liquid reaction is balanced and analyzed. The result is a discharge of:

Upper phase: 370 g and lower phase: 29 g,

where an analytical evaluation gives the following:

Upper phase: 270.1 g (73%) MBE 3.7 g (1%) formaldehyde

74.0 g (20%) of water

22.2 g (6%) by-products

Subphase: 2.3 g (8%) MBE 26.7 g (92%) water

Conversion of formaldehyde: 96.3%; Selectivity to MBE: 98.6%.

1 b) Purification by distillation of MBE:

The distillation is carried out in a 2 liter double-jacket four-necked flask with thermostat, column (with 1 m Sulzer EX packing), reflux divider, vacuum pump, distillation templates, thermometer and cold trap with dry ice. Raw material:

1 110 g discharge upper phase from autoclave experiment 1a) in the bottom flask submit (for the distillation insert 1 b) 3 Autoklavenausträge 1 a) are used.).

Execution:

The removal of water is carried out under normal pressure (1013 mbar) at a pot temperature up to 135 ° C and a transition temperature to 101 ⁰ C. The reflux ratio in this case carries loading about 5: 1.

Fraction 1: 430 g (2-phase);

Analysis: 215 g (50%) of water

193.5 g (45%) MBE

The MBE purifying distillation is then carried out at a pressure of about 100 mbar at a bottom temperature of about 155 ° C and a transition temperature to about 73 ° C. The reflux ratio is approximately 5: 1.

This gives two fractions:

Fraction 2: 50 g

Analysis: 25g (50%) formaldehyde

25g (50%) MBE

Fraction 3: 580 g;

Analytics: 577, i g (99, 5%) MBE

The weight of sump discharge after the distillation end is 40 g. The fractions 1 and optionally 2 can be used for further distillations.

Example 2: Preparation of 3-methyl-2-butenal (MBA) in the short tube reactor

For the oxidative dehydrogenation of MBE to MBA on a silver catalyst, a short tube reactor (length: 100 mm, diameter 12 mm) with a sand bath heater is used. Raw material: 1000 g (11, 6 mol) MBE approx. 12 g of 6% silver-shell catalyst (corresponding to a bulk height of 100 mm)

Implementation:

The reactor is heated to a temperature of 130 ⁰ C. The coolers are operated with a brine with a temperature of 3 ° C. The sand bath is heated to a temperature of about 360 ⁰ C, in this case, a nitrogen stream of 50 liters / hour manufacturer is provided. Once the temperatures have been reached, the MBE inlet is started at approx. 110 g / h via the pump and at the same time 51 l / h of air are fed in (the nitrogen is switched off here). The reaction has started as soon as the hot spot measurement is well above 360 ^0C. After MBE inlet end, the system is rendered inert with nitrogen and turned off.

It results at the following reaction conditions:

Temperature quartz sand outside: approx. 360 ° C Temperature reactor at the hotspot: approx. 450 ° C

Catalyst load: approx. 0.9 t / m2 ^* h

Molar ratio OH / O: about 1, 0: 0.75

Oxygen content in the exhaust gas: approx. 2 Vol%

Isobutene content in the exhaust gas: approx. 5% by volume CO content in the exhaust gas: approx. 4% by volume

CO ₂ content in the exhaust gas: approx. 4% by volume

Sales: 62.8%

Selectivity: 75.6%

Ratio of MBA to IMBA: 1: 4

IMBA is 3-methyl-3-butenal

The result is a discharge of: 920 g upper phase and 80 g lower phase, where an analytical evaluation gives the following:

Upper phase: 360.0 g (40%) MBE

460 g (50%) MBA + IMBA 73.6 g (8%) water; Lower phase: 4 g (5%) MBE

4 g (5%) MBA H H LMBA

72 g (90%) of water.

Example 3: Isomerization of IMBA (3-methyl-3-butenal) to MBA

Raw material:

920 g discharge upper phase from example 2 20 mg sodium acetate (purity: p.A.)

3a) isomerization reaction:

The reaction is carried out in a 2 liter four-necked flask equipped with stirrer, thermometer, oil bath and condenser. The upper phase effluent from the reaction obtained in Example 2 is heated to about 150 ⁰ C together with sodium acetate as the catalyst for about 2 hours. Subsequent analytical examination gives the following products:

368.0 g (40%) MBE

460.0 g (50%) MBA

73.6 g (8%) of water.

Instead of the sodium acetate, which is preferably used as a catalyst, other isomerization catalysts can be used.

3b) Pure distillation of MBA:

It is assumed that 920 g Isomerisierungsaustrag from Example 3a).

In a 2 liter double-jacket four-necked flask with thermostat, distillation column (with 1 m Sulzer EX-packing), reflux divider, vacuum pump, distillation templates, thermometer and cold trap, the cleaning takes place. The purifying distillation takes place at a vacuum of about 250 mbar. Here, the low boilers and the water are distilled off first to a top temperature of about 87 ⁰ C. The reflux ratio here is approximately 10: 1.

Fraction 1 (LS and water): 95 g (2-phase);

Analysis: 69.4 g (73%) water 20.9 g (22%) of low-boiling secondary components

4.8 g (5%) MBA. The MBA pure distillation takes place at a bottom temperature of up to about 165 ⁰ C, a vacuum of 250 mbar and a top temperature to about 94 ⁰ C. The reflux ratio is about 5: 1.

Fraction 2: 465 g;

Analytics: 455.7 g (98%) MBA.

The weight of bottoms discharge (MBE) after distillation is 340 g. The recovered MBE can be used again in the reaction.

Example 4: Preparation of 3-methyl-2-buten-ol (prenol) from MBE by catalytic isomerization

4a) isomerization reaction

For the isomerization of MBE to prenol on a Pd / Se catalyst, a tube reactor is used.

Starting product: 100 g of MBE.

60 ml of catalyst (Pd / Se on ceramic support) are introduced into the tubular reactor. The apparatus is rendered inert with nitrogen. The tube reactor is heated to about 80 ⁰ C. Then about 100 g / h of MBE and about 5 l / h of hydrogen are passed from below through the reactor. After a reaction time of 10 to 100 minutes, the following results:

Sales: approx. 55%

Selectivity: approx. 91%.

The output is: 1000 g; Analysis: 470.0 g (47%) MBE

500.0 g (50%) Prenol 25.0 g (2.5%) iso-amyl alcohol

4 b) Purification by distillation of prenol

Starting material: 1000 g isomerization product from Example 4a)

In a 2 liter double-jacket four-necked flask with thermostat, 1 m column (with Sulzer EX packing), reflux divider, vacuum pump, distillation templates, thermometers and Cold trap is the pure distillation at a vacuum of about 500 mbar. Here, the low boilers and the unreacted MBE are first distilled off up to a top temperature of about 107 ⁰ C. The reflux ratio here is approximately 10: 1.

Fraction 1 (lighter boilers and MBE): 501 g;

Analysis: 35.1 g (7%) of light boilers, of which about 25 g iso-amyl alcohol,

460.9 g (92%) MBE. 5.0 g (1%) Prenol.

The thus recovered MBE is used together with the low-boiling components (especially isoamyl alcohol) preferably in the synthesis of MBA.

The Prenol pure distillation is then carried out at a bottom temperature of up to 165 ⁰ C, a vacuum of 500 mbar and a top temperature to 122 ⁰ C. The Rücklaufver- ratio is in this case 5: 1.

Fraction 2 (prenol): 470 g;

Analysis: 2.4 g (0.5%) MBE

467.7g (99.5%) prenol.

The bottoms discharge is 9 g.

Example 5: Preparation of Citral from MBA and Prenol

5a) Carrying out the acetalization

In a 2 liter double-jacket four-necked flask with thermostat, column (with 0.5m Sulzer EX-packing), water separator, vacuum pump, thermometer and cold trap are used:

450 g (5.244 mol) MBA 950 g (10.98 mol) prenol 0.6 g nitric acid 5% aqueous 50 g MBA Pure product from main fraction into the overflow of the water separator.

The starting materials are charged and heated at a vacuum of 80 mbar up to a bottom temperature of 110 ⁰ C, wherein the resulting water is removed from the system. The trial period is about 8 hours for the acetalization. The result is the following:

Weight of acetalation water: 60 g (3.333 mol) Weight at sump: 1320 g; Swamp analysis: 765.6 g (58%) of MBA-di-prenyl-acetal

158.4 g (12%) MBA

343.2 g (26%) prenol.

The turnover of MBA is: 64.1%

The selectivity to the acetal is: 95.7%.

Sales of prenol are: 63.7%

Selectivity to acetal: 92.0%.

5b) Carrying out the acetal cleavage

In a 2 liter double-jacket four-necked flask with thermostat, column (with 0.5m Sulzer EX-packing), reflux divider, vacuum pump, distillation template, thermometer and cold trap are used:

1320 g of acetalization bottoms (Example 5a) 0.6 g of phosphoric acid (5% aqueous).

The starting materials are submitted and heated under a vacuum of 50 mbar up to a sump temperature of 160 ⁰ C. Overhead, the resulting prenol and the low boilers are distilled off. The reflux ratio should be about 5: 1. The test duration is about 6 hours for the acetal cleavage. The result is the following:

Distillate weight: 780 g; Analytics: 156 g (20%) MBA

585 g (75%) prenol

Bottom weight: 520 g; Analysis: 426.4 g (82%) citral

5.2 g (1%) of MBA-di-prenyl-acetal.

The conversion of acetal is: 99.3% The selectivity is 87.8%.

Depending on the intended use, the citral thus obtained can be subjected to further purification or used directly as starting material for further syntheses.

Claims

claims

A process for the preparation of 3,7-dimethyl-octa-2,6-dienal (citral) in substantially six process steps, wherein:

a) initially producing 3-methyl-3-buten-1-ol (isoprenol) from isobutene and formaldehyde,

b) starting from 3-methyl-3-buten-1-ol (isoprenol) by oxidative dehydrogenation with an oxygen-containing gas in a continuous process on a supported silver catalyst 3-methyl-2-butenal (Prenal) and 3 Methyl-3-butenal (isoprenal) is produced,

c) starting from a 3-methyl-3-butenal-containing mixture by I-somerisierung in the presence of a catalyst further 3-methyl-2-butenal

(Prenal) is produced,

d) starting from 3-methyl-3-buten-1-ol (isoprenol) by isomerization in the presence of hydrogen and a noble metal-containing fixed-bed catalyst, the 3-methyl-2-buten-1-ol (prenol) is prepared,

e) starting from 3-methyl-2-buten-1-ol (prenol) and 3-methyl-2-butenal (Prenal) using an acid catalyst and separating the water formed in the reaction, the unsaturated acetal 3-methyl 2-butenal-diprenyl acetal is produced, and

f) starting from 3-methyl-2-butenal-diprenyl acetal by cleavage in the presence of a catalyst and subsequent Claisen rearrangement of the resulting butadienyl ether, a 2,4,4-trimethyl-3-formyl-1, 5-hexadiene is produced, then from the Cope rearrangement the 3,7-dimethyl-octa

2,6-dienal (citral) is obtained.

2. A process for the preparation of citral according to claim 1, characterized in that the process is carried out fully continuously and the by-products from the process steps b) and d) are introduced again into the process.

3. A process for the preparation of citral according to one of the claims 1 or 2, characterized in that the process is carried out fully continuously and the by-products from the process steps a), b), d) and e) are introduced in each case again in the process.

4. A process for the preparation of citral according to one of the claims 1 to 3, characterized in that in the isomerization of 3-methyl-3-butenal sodium acetate is used as a catalyst and that the entire process is carried out fully continuously and the by-products from the process steps a ), b), c), d) and e) in each case be reintroduced into the process.

5. Process for the preparation of citral according to one of claims 1 to 3, characterized in that the oxidative dehydrogenation of 3-methyl-3-buten-1-ol (isoprenol) with an oxygen-containing gas over a supported silver catalyst ( Process step b) is carried out so that initially 3-methyl-3-buten-1-ol is evaporated, then this vapor is added to the oxygen-containing gas, the thus obtained oxygen-containing alcohol vapor initially at a temperature above the dew point of -Methyl-

But 3-butene-1-ol is passed below the starting temperature of the reaction through an at least 0.5 cm thick layer of supported silver catalyst and then the oxygen-containing vapor of 3-methyl-3-buten-1-ol in one of desired capacity corresponding number of parallel arranged and flow around with a fluid heat transfer reaction tubes having an inner diameter of 1 to 2 cm and a length of 35 to 60 cm and are filled with the supported silver catalyst, at temperatures of 300 ° to 600 ⁰ C. is converted to a mixture of isomeric aldehydes.

6. A process for the preparation of citral according to one of the claims 1 to 5, characterized in that in the isomerization of 3-methyl-3-buten-1-ol (isoprenol) to 3-methyl-2-buten-1-ol ( Prenol) (process step d) in the presence of hydrogen as a noble metal-containing fixed bed catalyst, a catalyst containing palladium and selenium or tellurium or a mixture of selenium and tellurium on a silica support is used.

7. A process for the preparation of citral according to one of the claims 1 to 6, characterized in that in the preparation of 3-methyl-2-butenal diprenyl acetal (process step e)) the starting materials 3-methyl-2-buten-1-ol (Prenol) and 3-methyl-2-butenal (Prenal) in the reaction column only partially reacted and concentrated the obtained unsaturated acetal in at least two successive evaporator stages and the recovered reactants back into the reaction column.

8. A process for the preparation of citral according to one of the claims 1 to 7, characterized in that in the thermal cleavage of the 3-methyl-2-butenal-diprenyl acetal and the subsequent rearrangements (process step f)), both the intermediates 3-methyl-2 -butene-1-ol (prenol); Prenyl (3-methyl-butadienyl) ether and 2,4,4-trimethyl-3-formyl-1, 5-hexadiene and the

End product citral are removed during the reaction continuously by distillation from the reaction mixture.

9. A process for the preparation of citral according to one of the claims 1 to 8, characterized in that by thermal cleavage of the 3

Citral obtained from the crude mixture separated by methyl-2-butenal-diprenylacetals and the subsequent rearrangements by rectification, where:

a) the crude mixture is introduced laterally into an intake column with a reinforcing part located above the feed point and a driven part located below the feed point,

b) providing an upper merging column having a condenser at the upper end of the column communicating with the upper end of the reinforcing member and a lower merging column communicating with the lower end of the driven member and having heaters at the lower end of the column;

c) provides a withdrawal column communicating with the upper union column and the lower union column,

e) removed from the withdrawal column the citral as a side draw and at the top of the upper union column, the lower than citral boiling compounds and subtracting the higher than citral boiling compounds in the bottom of the lower union column.