Improved process for preparation of 3, 7-dimethyl-oct-2, 6-dienal

Abstract

Disclosed is a process for the preparation of 3, 7-dimethyl-oct-2, 6-dienal (citral), the process comprising the steps of: a) continuous condensation of pentenol with pentenal in a reaction column in the presence of at least one catalyst wherein the condensed water is distilled off as a pentenal-water azeotrope in vapor form, the vapour is at least partially condensed and the condensate is separated into an aqueous phase and an organic phase, and the organic phase is partially directed as a reflux to the reaction column and is partially discharged as a purge stream while an acetal fraction of dipentenyl acetal comprising pentenal is continuously discharged from the reaction column, wherein the reaction temperature is below 100 DEG C, the catalyst is nitric acid, and the concentration of nitric acid is below 500 ppm; b) continuously subjecting the acetal fraction to cracking conditions in a cracking column in the presence of at least one catalyst, in which pentenol is removed while continuously discharging from the cracking column the acetal fraction containing pentenyl (3-methyl-butadienyl) ether and 2, 4, 4-trimethyl-3-formyl-1, 2, 4-triazole-1, 2, 4-triazole-1, 2, 4-triazole-1, 2, 4-triazole-1, 2, 4-triazole-1, 2, 4-triazole-1, 2-triazole-1 a cleavage fraction containing at least one of 1, 5-hexadiene and optionally citral, in which the conversion of the dipentenyl acetal of the pentenal in step b) is maintained above 90% and below 100% and the unreacted dipentenyl acetal is at least partially contained in the discharged cleavage fraction; c) reacting the cracked fraction in a plug flow reactor to obtain citral; and d) recycling a portion of the pentenol obtained in step b) to step a). The process allows the formation of unwanted by-products to be inhibited during the production of citral.

Description

[0001] The present invention relates to an improved process for preparing citral (3,7-dimethyl-octa-2,6-dienal), by which citral can be obtained in high yield. Citral is a mixture of the isomeric compounds neral and geranal. Citral is a valuable intermediate for the production of various flavors and fragrances, such as geraniol. In addition, citral is also becoming increasingly important as a starting material for the production of vitamins, in particular vitamin A.

[0002] DE 198 46 056 describes a process for preparing citral by thermal cleavage of 3-methyl-2-butene-1-aldehyde-dipentenyl acetal to cis / trans-pentenyl-(3-methyl-butadienyl)-ether, optionally in the presence of an acid catalyst, with cleavage of 3-methyl-2-butene-1-ol (pentenol), Claisen rearrangement of the butadienyl ether to 2,4,4-trimethyl-3-formyl-1,5-hexadiene, and subsequent Cope rearrangement thereof to obtain citral. Pentenol, intermediates and citral are continuously distilled off from the reaction mixture.

[0003] WO 2008 / 037693 discloses a method for producing citral. The method involves the following steps:

[0004] a) Production of 3-methyl-3-butene-1-ol (isopentenol) from isobutene and formaldehyde;

[0005] b) production of 3-methyl-2-butenal (pentenal) and 3-methyl-3-butenal (isopentenal) from 3-methyl-3-buten-1-ol (isopentenol) by oxidative dehydration with the aid of oxygen-containing gases on silver-supported catalysts;

[0006] c) producing further 3-methyl-2-butenal (pentenal) from a mixture containing 3-methyl-3-butenal (isopentenal) by isomerization;

[0007] d) Production of 3-methyl-2-buten-1-ol (pentenol) from 3-methyl-3-buten-1-ol (isopentenol) by isomerization;

[0008] e) producing the unsaturated acetal 3-methyl-2-butenal-dipentenyl acetal from 3-methyl-2-buten-1-ol (pentenol) and 3-methyl-2-butenal (pentenal) using an acidic catalyst; and

[0009] f) Citral is obtained from 3-methyl-2-butenal-dipentenyl acetal by cleavage and subsequent rearrangement.

[0010] This complicated, multi-stage method is prone to the unwanted side reaction that reduces the available citral yield. In other words, the single step shows a selectivity less than 100%, and the amount of the by-product formed may be higher than desired. Such by-products reduce the desired conversion selectivity and must usually be removed from the citral product before subsequent use. A large amount of energy is needed to separate the by-product from citral, and citral can encounter significant losses usually. Such losses may cause the use of other favorable reaction sequences to be unattractive commercially.

[0011] Until now, relatively little is known about the nature of the unwanted byproducts and the mechanisms of their formation.

[0012] The present invention therefore seeks to propose reaction conditions which effectively suppress the formation of unwanted by-products during the production of citral, as well as to remove the by-products which are inevitably formed without impairing the formation of citral and its building blocks.

[0013] This problem is solved by the following method and its preferred embodiments.

[0014] The present invention relates to a method for preparing 3,7-dimethyl-octa-2,6-dienal (citral), comprising the following steps:

[0015] a) continuously condensing pentenols with pentenals in a reaction column in the presence of at least one catalyst, wherein the water of condensation is distilled off in vapor form as a pentenal-water azeotrope, the vapors are at least partially condensed and the condensate is separated into an aqueous phase and an organic phase, and the organic phase is partly conducted as reflux to the reaction column and partly discharged as a purge stream, while an acetal fraction comprising the dipentenyl acetal of pentenals is continuously discharged from the reaction column, wherein the reaction temperature is below 100° C., the catalyst is nitric acid, and the concentration of the nitric acid is below 500 ppm;

[0016] b) subjecting the acetal fraction to cracking conditions continuously in a cracking column in the presence of at least one catalyst, wherein pentenol is removed, while a cracking fraction containing at least one of pentenyl (3-methyl-butadienyl) ether and 2,4,4-trimethyl-3-formyl-1,5-hexadiene and optionally citral is continuously discharged from the cracking column, wherein the conversion of the dipentenyl acetal of pentenal in step b) is maintained above 90% and below 100%, and unreacted dipentenyl acetal is at least partially contained in the discharged cracking fraction;

[0017] c) reacting the cleavage fraction in a plug flow type reactor to obtain citral; and

[0018] d) recycling a portion of the pentenol obtained in step b) to step a).

[0019] The overall reaction sequence is illustrated by the following reaction scheme.

[0020]

[0021] In step a), the unsaturated acetal 3-methyl-2-butenal-dipentenyl acetal (referred to herein as "dipentenyl acetal of pentenal" or "dipentenyl acetal") is formed from pentenol and pentenal using a catalyst. For this purpose, pentenal is reacted together with pentenol in the presence of a catalytic amount of nitric acid, wherein the concentration of nitric acid is below 500 ppm and the reaction temperature is below 100° C., and wherein the water formed during the reaction in the reaction column is separated off.

[0022] In step b), the obtained 3-methyl-2-butenal dipentenyl acetal (dipentenyl acetal) of step a) is cleaved in a cracking tower in the presence of a catalyst, and 3-methyl-2-butene-1-ol (pentenol) is removed to give pentenyl (3-methylbutadienyl) ether. Claisen rearrangement of the obtained pentenyl (3-methylbutadienyl) ether gives 2,4,4-trimethyl-3-formyl-1,5-hexadiene, which then undergoes Cope rearrangement to give 3,7-dimethyl-2,6-octadienal (citral).

[0023] In order to make the method economically feasible, the pentenol obtained in step b) is recycled to step a). However, it has now been found that unwanted by-products can accumulate in the recycle stream. Until a large amount of by-products accumulate in the recycle loop, the problem will become apparent. Therefore, the present invention proposes to discharge a portion of organic entrainer liquid (i.e., the condensed distillate after the water phase is removed) at the top of the reaction tower (wherein pentenol is condensed with pentenal) to remove unwanted by-products. The by-products are concentrated in this organic entrainer liquid, so that the loss of valuable products is minimized.

[0024] The structures of the by-products which may accumulate in the recycle loop from step b) to step a) have been elucidated. The by-products were identified as by-products 1a and 1b as shown below.

[0025] In a preferred embodiment, the rate of the purge stream is such that the total fixed concentration of byproducts 1a and 1b is

[0026]

[0027] In the feed to step a) it is kept below 7 wt.-%, preferably between 2 and 7 wt.-%.

[0028] "Feed to step a)" means the total supply of fresh pentenols and pentenals and pentenols recycled from step b).

[0029] The following by-products 2, 3, 4, 5a and 5b have been identified which are formed under unfavourable reaction conditions in steps b) and c):

[0030]

[0031] It has been unexpectedly found that when the conversion rate of the dipentenyl acetal of pentenal in step b) is driven to complete conversion rate, the concentration of by-products increases sharply. According to the present invention, the conversion rate of the dipentenyl acetal of pentenal in step b) remains on and is higher than 90% and lower than 100%. Preferably, the conversion rate of the dipentenyl acetal of pentenal in step b) remains equal to or lower than 99.5%, preferably equal to or lower than 99%, such as equal to or lower than 98%, or equal to or lower than 97.5%, or equal to or lower than 97%. Preferably, the conversion rate of the dipentenyl acetal of pentenal in step b) remains on and is higher than 91%, such as higher than 92%, or higher than 93%, or higher than 94%, or higher than 95%. In a suitable embodiment, the conversion rate of the dipentenyl acetal of pentenal in step b) is higher than 94% and equal to or lower than 99%, such as higher than 95% and equal to or lower than 98%. Lower conversion rates will cause the method to be economically unprofitable, or will require recovery and recycling of unreacted dipentenyl acetal in addition. However, complete conversion in step b) is undesirable because it causes a decrease in the yield of citral building blocks and increases by-product formation. Conversion rates are controlled by various parameters, including cracking temperature, the nature and concentration of one or more catalysts in step b) and the residence time in step b) (i.e., in a cracking tower).

[0032] Furthermore, the present invention is characterized in that the distillation conditions in the cracking column are effective to at least partially distill off unreacted dipentenyl acetal.

[0033] The acetal fraction is continuously subjected to cleavage conditions in a cleavage column. "Cleavage conditions" means reaction conditions selected so that the dipentenyl acetal contained in the acetal fraction is cleaved to pentenyl (3-methylbutadienyl) ether (which can subsequently rearrange to 2,4,4-trimethyl-3-formyl-1,5-hexadiene and citral).

[0034] The acetal fraction comprises dipentenyl acetal as main component. The acetal fraction need not necessarily be composed of pure dipentenyl acetal, but can also comprise pentenol, pentenal and citral building blocks. In this respect, it is desirable that the dipentenyl acetal content in the acetal fraction is maintained at a level higher than a certain level, so as to avoid excessive carrying of pentenol and pentenal to the separation of downstream reaction steps and subsequent pentenal. In addition, in this respect, ensuring that the dipentenyl acetal content is suitably high in the acetal fraction also means minimizing the by-products such as the by-product 1 illustrated below that pentenol is derived. Therefore, in a preferred embodiment, the acetal fraction comprises at least 65wt.-% of dipentenyl acetal, preferably at least 75wt.-%, for example 75 to 90wt.-%.

[0035] Increasing the content of dipentenyl acetal in the acetal fraction beyond a certain point reaches a point of diminishing returns. An economic balance must be struck between the improvement due to the increased content and the cost of achieving such a content. The temperature stress caused by the additional separation stage and the increased residence time may even lead to decomposition of the dipentenyl acetal, which may be caused by traces of one or more acetalization catalysts.

[0036] Since cleavage of the dipentenyl acetal may already occur to a certain extent under the conditions of the condensation reaction between the pentenals and pentenols, attempts to drive the condensation to completion may result in premature formation of citral building blocks, which may lead to the undesirable formation of by-products 3 and 5 (i.e., resulting from the reaction of the citral building blocks with the pentenals and pentenols).

[0037] Step b) is carried out in a cracking tower. Suitably, the cracking tower is a distillation tower equipped with an evaporator and a condenser. Suitable internals for the cracking tower are plates, packings, and in particular structured packings made of metal sheets or metal meshes. The theoretical number of plates of the cracking tower may be in the range of 5 to 100.

[0038] In an embodiment, the pyrolysis temperature in step b) is higher than 150°C and lower than 200°C, preferably higher than 155°C and lower than 180°C.

[0039] Step b) is carried out in the presence of a catalyst, preferably an acid catalyst. The catalyst can be a single catalytic substance or a combination of two or more different catalytic substances. Suitable acid catalysts are selected from non-volatile proton acids, such as sulfuric acid, p-toluenesulfonic acid and phosphoric acid. In an embodiment, one or more catalysts in step b) are phosphoric acid. In a preferred embodiment, the concentration of phosphoric acid in the bottoms of the cracking tower remains at a level higher than 100ppm and lower than 1500ppm, preferably higher than 200ppm and lower than 1000ppm. Higher concentrations of (acid) catalysts may result in a reduction in the yield of the citral building block.

[0040] Suitably, the continuous cleavage in the cracking column of step b) can be carried out in the lower part or sump of a distillation column acting as a cracking column. Preferably, the acetal fraction and / or one or more catalysts are introduced into the lower part of the distillation column, into the sump of the distillation column or into the evaporator of the distillation column. If necessary, the volume of the sump of the cracking column can be increased by a container in order to provide a larger reaction volume.

[0041] Typically, the bottoms from the cracking column are a mixture of high boilers including C5-oligomers resulting from the thermal instability of the dipentenyl acetal and citral building blocks.

[0042] If desired, a high boiling point inert compound can be introduced into the sump of the cracking column in order to ensure a minimum filling level of the sump and the evaporator. Suitable high boiling point inert compounds are selected from liquid compounds that are inert under the reaction conditions and have a higher boiling point than citral and dipentenyl acetal. For example, the high boiling point inert compound can be selected from hydrocarbons, such as tetradecane, pentadecane, hexadecane, octadecane, eicosane; or ethers, such as diethylene glycol dibutyl ether; white oil; kerosene; or mixtures thereof.

[0043] A portion of the bottoms from the cracking column is continuously discharged. This serves to avoid accumulation of high boilers. Since the key feature of the present invention is that the cleavage of the dipentenyl acetal in step b) is carried out to a less than complete conversion, the discharged portion of the bottoms from the cracking column may also contain unreacted dipentenyl acetal.

[0044] Suitably, the distillation conditions are selected so that the dipentenyl acetal is mainly retained in the lower part of the distillation column or in the storage tank. During the cleavage reaction, the cleavage fraction is continuously discharged from the cleavage column, and the cleavage fraction contains at least one of pentenyl (3-methyl-butadienyl) ether and 2,4,4-trimethyl-3-formyl-1,5-hexadiene and optionally contains citral. For ease of reference, pentenyl (3-methyl-butadienyl) ether, 2,4,4-trimethyl-3-formyl-1,5-hexadiene and citral are collectively referred to as "citral building blocks". This is because the former is an intermediate on the reaction pathway leading to citral and can be converted into citral in the subsequent step c).

[0045] The conversion of the dipentenyl acetal of pentenals in step b) remains below complete conversion and the unreacted dipentenyl acetal is at least partially contained in the discharged cleavage fraction. The concentration of unreacted dipentenyl acetal in the cleavage fraction may be from 1 to 6 wt.-%, relative to the citral building blocks.

[0046] Additionally, the pentenols formed during the cleavage reaction in step b) are continuously removed from the reaction mixture, typically at the top of the cleavage column.

[0047] The cleavage fraction together with the pentenols formed can be discharged at the top of the distillation column.

[0048] Alternatively and preferably, the cracked fraction can also be discharged in liquid or vapor form at a side draw of the distillation column. The side draw is preferably located in the middle or lower part of the cracking column, in particular 2 to 20 theoretical plates above the addition point of the acetal fraction. Preferably, the cracking column comprises 2 to 80 theoretical plates above the side draw.

[0049] The feed rate of the acetal fraction is adjusted to control the residence time of the dipentenyl acetal in step b). In an embodiment, the residence time in step b) is higher than 5 min and lower than 90 min, preferably higher than 15 min and lower than 80 min. The residence time in this interval represents a good trade-off between the yield of the reaction product and the minimum value of by-product formation.

[0050] In step c), the cleavage fraction is reacted in a plug flow reactor to obtain citral. To this end, the cleavage fraction is led through a plug flow reactor at a suitable temperature of 100° C. to 200° C. for carrying out one or more rearrangement reactions to obtain citral.

[0051] Applicants have found that by using a combination of a highly back-mixed cracking column and a plug flow reactor, the selectivity and yield of the cracking reaction can be increased. In step b), preferably all the catalyst required for the cracking reaction is introduced into the cracking column, and preferably, no catalyst is introduced into the plug flow reactor.

[0052] The acetal fraction subjected to step b) is formed in step a) from pentenols and pentenals. According to the invention, in step a), pentenols are continuously condensed with pentenals in a reaction tower in the presence of at least one catalyst, while the acetal fraction of pentenals is continuously discharged from the reaction tower. Further optimization of the process relates to the conditions of step a).

[0053] In step a), the formation of the following by-products has been observed under unfavorable reaction conditions:

[0054]

[0055] It is known that unsaturated acetal is prepared by reacting olefinic unsaturated aliphatic compounds with allyl alcohol in the presence of distillable acid in a reaction tower. For this purpose, a mixture of at least 2mol of pentenols and 1mol of pentenal can be introduced into the reaction tower, and the water formed during the reaction is distilled out at the top of the tower and removed by a phase separator. Then, dipentenyl acetal can be removed from the bottoms or evaporator of the reaction tower as a crude product. Due to the thermal instability of dipentenyl acetal, a large amount of purification of thick dipentenyl acetal is usually undesirable. However, the concentration (for example, in a short-path evaporator) of thick dipentenyl acetal may be advantageous. Thus, before thick dipentenyl acetal is directed to a cracking tower, unreacted pentenal and pentenol are removed from thick dipentenyl acetal. Suitably, aldehyde is substantially absent in the thick dipentenyl acetal directed to a cracking tower.

[0056] Suitably, the device for preparing unsaturated acetal comprises a distillation tower used as a reaction tower. The steam rising at the top of the reaction tower is condensed in a condenser and passed into a phase separation container, in which water is separated into a lower layer phase. The upper layer is mainly composed of organic compounds such as unreacted aldehyde (i.e. pentenal), unreacted alcohol (i.e. pentenol) and low-boiling secondary compounds (e.g., the formic acid ester of pentenol). Most of the organic phase is recycled to the top of the reaction tower as reflux, and the discharge minor portion is to remove secondary components.

[0057] Suitably, the amount of reflux per 1000 kg of freshly added aldehyde is in the range of 200 kg to 50000 kg, preferably 1000 kg to 20000 kg. Depending on the purity of the feed, the amount of the purge portion per 1000 kg of freshly added aldehyde is in the range of 1 kg to 400 kg, preferably 5 kg to 200 kg.

[0058] The reaction temperature in step a) is below 100° C., preferably in the range of 70 to 80° C. In an embodiment, the reaction pressure in step a) is below 150 mbar, preferably in the range of 90 to 110 mbar. Suitably, the residence time of the reaction mixture in the reactor is in the range of 0.1 s to 10 h, preferably 60 s to 2 h.

[0059] According to the present invention, step a) is carried out in the presence of a catalyst, wherein the catalyst is nitric acid. Relative to the total amount of starting material pentenols and pentenals, the concentration of nitric acid is lower than 500ppm, more preferably in the scope of 100 to 300ppm. Lower amounts of (acid) catalysts may cause the low conversion in the reaction tower. Higher amounts of (acid) catalysts may disadvantageously cause the formation of by-products to increase and cause selectivity to decrease.

[0060] Suitably, the catalyst is added to the reaction tower, preferably to the lower part of the reaction tower. In a preferred embodiment, the catalyst is added to the evaporator. Nitric acid can also be added at different locations, for example at two or more points in the reaction tower.

[0061] The addition point of fresh aldehyde and / or fresh alcohol is not critical. Aldehyde and alcohol can be added separately at different points in the reaction column. Preferably, aldehyde and alcohol are combined with the effluent from the condenser. The amount of freshly added alcohol is controlled so that the ratio of alcohol to aldehyde is in the range of 1 to 3, preferably 1.5 to 2.5.

[0062] In step d), the pentenol removed in step b) is recycled to step a). This allows increased yields to be achieved in the process according to the invention.

[0063] The applicant has found that it is important to control the concentration of contaminants contained in the pentenols recycled to step a). In particular, the concentration of 2,4,4-trimethyl-3-formyl-1,5-hexadiene in the pentenols recycled from step b) to step a) is preferably controlled so that the concentration of 2,4,4-trimethyl-3-formyl-1,5-hexadiene in step a) is lower than 1 wt.-%, preferably lower than 0.5 wt.-%, relative to the total weight of pentenols and pentenals. Similarly, the concentration of citral in the pentenols recycled from step b) to step a) is preferably controlled so that the concentration of citral in step a) is lower than 1 wt.-%, preferably lower than 0.1 wt.-%, relative to the total weight of pentenols and pentenals.

[0064] The inventors have found that it is advantageous to control the amount of 2,4,4-trimethyl-3-formyl-1,5-hexadiene and / or citral introduced into step a) together with the recycled pentenol from step b). As mentioned above, higher concentrations of 2,4,4-trimethyl-3-formyl-1,5-hexadiene and / or citral lead to the formation of by-products.

[0065] Prenol useful as a starting material for the present invention can be obtained by reacting at least one formaldehyde source and isobutylene in a reactor, usually at elevated temperature and pressure, to obtain 3-methylbut-3-en-1-ol (isopentenol), and subjecting the obtained isopentenol to isomerization.

[0066] In one embodiment, the prenol is obtained by mixing at least one formaldehyde source and isobutylene and injecting them into a reactor through at least one nozzle and reacting the at least one formaldehyde source and isobutylene under supercritical conditions. To achieve supercritical conditions, formaldehyde and isobutylene are preferably reacted at a temperature of at least 220° C., for example in the range of 220° C. to 290° C., and an absolute pressure of at least 200 bar absolute. The reaction of isobutylene and formaldehyde can be carried out in the presence of at least one catalyst such as an amine base, for example hexamethylenetetramine (urotropine).

[0067] In one embodiment, isopentanol is obtained by mixing at least one formaldehyde source and isobutylene and injecting them into a reactor through multiple nozzles operated in parallel and reacting the formaldehyde source and isobutylene under supercritical conditions; wherein the reactor includes a vertically arranged container, a side wall, an upper part and a lower part; and wherein the formaldehyde source and isobutylene are injected into a mixing chamber arranged in the upper part of the reactor, and a fluid containing formaldehyde and / or isobutylene and / or isopentanol enters a post-reaction chamber arranged in the lower part from the mixing chamber; and a guide tube is provided below each nozzle arranged substantially concentrically in the mixing chamber, and these guide tubes provide a descending conduit inside the guide tube and an ascending conduit outside the guide tube, so that the formaldehyde source and isobutylene injected through the nozzle travel generally downward in the descending conduit, and then the fluid containing formaldehyde and / or isobutylene and / or isopentanol is turned in a generally upward direction in the ascending conduit, and the fluid is backmixed with the injected formaldehyde source and isobutylene.

[0068] The formaldehyde may be provided as a liquid, for example as a solution of paraformaldehyde.Preferably, the at least one formaldehyde source comprises or is an aqueous formaldehyde solution.

[0069] Although initial rapid and intense mixing of the reactants is desirable, it may be advantageous to continue and complete the reaction under limited backmixing conditions. Therefore, the reaction mixture may be passed into a post-reaction chamber disposed behind the reactor or in the lower portion of the reactor. In the post-reaction chamber, backmixing is limited.

[0070] In one embodiment, the reactor comprises an upper part and a lower part. The injection and mixing of the reactants occur in a mixing chamber disposed in the upper part of the reactor, and the fluid containing formaldehyde and / or isobutylene and / or isopentenol enters a post-reaction chamber disposed in the lower part from the mixing chamber.

[0071] Further details about reacting at least one formaldehyde source and isobutylene to obtain prenol can be found in WO 2020 / 049111 A1.

[0072] In one embodiment, reacting at least one formaldehyde source and isobutylene comprises mixing at least one formaldehyde source and isobutylene and injecting through at least one nozzle into one or more first conduits into an internal loop reactor comprising:

[0073] - a vertically arranged cylindrical container comprising a side wall;

[0074] - at least one flow guide cylinder having a cylinder inlet end and a cylinder outlet end arranged vertically within the container, the one or more flow guide cylinders being arranged concentrically with the one or more nozzles and having an inner surface and an outer surface, wherein the one or more flow guide cylinders provide one or more first conduits within the one or more flow guide cylinders, and second conduits outside the one or more flow guide cylinders and within the sidewalls, the one or more first conduits being in fluid communication with the second conduits;

[0075] - reactor fluid outlet means;

[0076] wherein the inner surface of the one or more flow guide sleeves is convexly curved so that the cross-section of the one or more first conduits exhibits an annular constriction between the sleeve inlet end and the sleeve outlet end; wherein the constriction is located closer to the sleeve inlet end; wherein the convex curvature of the inner surface of the one or more flow guide sleeves extends over at least 70%, preferably at least 80%, and most preferably at least 90% of the length of the flow guide sleeve; and

[0077] wherein the outer surface of the one or more flow guide sleeves is convexly curved so that the one or more flow guide sleeves exhibit a circumferential bulge between the sleeve inlet end and the sleeve outlet end, and the position of the circumferential bulge is preferably closer to the sleeve outlet end; wherein the convex curvature of the outer surface of the one or more flow guide sleeves extends over at least 70%, preferably at least 80%, and most preferably at least 90% of the length of the flow guide sleeve; and

[0078] The edges of one or more draft tubes are rounded so that at least one formaldehyde source and isobutylene injected through the nozzles travel generally downward in one or more first conduits to obtain a reaction fluid, and then the reaction fluid turns in the opposite direction to travel through a second conduit and subsequently backmix with the injected fluid.

[0079] This configuration of one or more flow tubes allows for control of the boundary layer flowing over the edge of the flow tube. When the angle of attack of the flow relative to the solid reaches a certain limit, the reverse pressure gradient becomes so large that the flow cannot successfully pass over the solid. At this point, the flow separates from the upper surface of the solid, resulting in a condition commonly referred to as stall. This configuration allows for reduced flow separation or correspondingly delayed flow separation. Reducing flow separation allows for reduced liquid friction and therefore results in a lower pressure drop along the streamlines of the recirculating flow, which in turn makes the circulation ratio of this configuration higher. The curved shape of the inner surface of the flow tube wall guides the fluid through the flow tube in an optimized manner comparable to the fluid flow on the airfoil.

[0080] The inner surface of the guide tube is curved in the longitudinal direction of the guide tube, or in other words, the inner surface of the guide tube has a convex shape, so that the first duct exhibits a minimum cross section between the tube inlet end and the tube outlet end. This means that the cross section of the first duct decreases from the cross section at the tube inlet end to the minimum cross section, and increases from the minimum cross section to the cross section at the tube outlet end.

[0081] The guide tube has a curved, approximately conical section between the inlet end of the tube and the contraction, which is wider at the inlet end and narrower at the contraction. At least a portion of the fluid flowing downstream through the guide tube is deflected to flow along the inner surface of the guide tube until the end of the guide tube. The flow passing through the guide tube remains mostly connected, and the resulting pressure loss is small. Near the contraction, the fluid flowing downstream through the guide tube is accelerated. Between the contraction and the outlet end of the tube, the cross-sectional area of ​​the guide tube widens again. Therefore, the change in area combined with the conservation of mass will force the speed through a larger area to be slower than the speed through a smaller area, accompanied by the conversion of dynamic pressure to static pressure. The acceleration of the fluid flowing downstream through the guide tube to the vicinity of the contraction adds a radial velocity component to the flow, thereby increasing the mixing between the circulating flow and the injection flow. By avoiding flow separation in this case, no significant pressure loss will occur.

[0082] In a preferred embodiment, the nozzle is a two-component nozzle. It is particularly preferred that the two-component nozzle is designed to provide an annular jet of isobutylene surrounding a central jet of at least one formaldehyde source, and the injection speeds of the two jets are different. In this embodiment, the jet of isobutylene has a large shear surface directed toward both the central jet of at least one formaldehyde source and the reaction mixture in the reactor, thereby allowing advantageous rapid mixing of the reactants.

[0083] In a preferred embodiment, the loop reactor comprises deflector means arranged between the nozzle and the draft tube, the deflector means being adapted to deflect the fluid travelling in the second conduit in an opposite direction.

[0084] The deflector device suitably comprises a surface which is recessed relative to the end of the limiting tube inlet end of the guide tube. In a preferred embodiment, the deflector device has a partial annular surface. It is particularly preferred that the deflector device is arranged in the shape of an upper part of a torus which is divided into two in a plane parallel to the annular direction. This shape can achieve particularly efficient deflection of the fluid traveling in the second conduit. The deflector device can achieve stabilization of the injected fluid flow. This is particularly important in the following situation: the flow velocity of the fluid traveling in the second conduit is not uniform over the cross section of the reactor, which may cause the injected fluid flow to be eccentric. If unattended, this eccentricity may cause the circulation ratio to drop.

[0085] When the first conduit is a descending conduit and the second conduit is an ascending conduit, it is preferred that the shape of the deflector device constitutes an upper part of a torus bisected in a plane parallel to the annular direction, wherein the torus bisected at least 50% of its height, such as at least 55% or 65% of its height. Thus, the upper part of the torus is the same size as the lower part of the torus or the upper part of the torus is smaller than the lower part of the torus. In another preferred embodiment, the shape of the deflector device constitutes an upper part of a torus bisected in a plane parallel to the annular direction, wherein the torus bisected at most 85% of its height, such as at 80% of its height. Within these ranges, the inlet of the deflector device is angled, which is particularly suitable for fluid deflection.

[0086] In the reaction of formaldehyde and isobutylene, high temperature is needed to obtain high prenol yield. Effectively removing heat is critical for product quality and process safety. The heat removed from prenol is used to raise its temperature before isobutylene enters the reactor. The hot prenol stream contains sensible heat from the chemical reaction. Sensible heat is a potential recyclable energy that can be reused.

[0087] Advantageously, reacting at least one formaldehyde source and isobutylene preferably comprises heat exchanging a hot isopentanol stream discharged from the reactor with an isobutylene stream directed to the reactor; wherein the heat exchange is carried out in at least two shell and tube heat exchangers; each of these heat exchangers comprises a plurality of tubes and shell-side heat exchange channels; wherein the hot isopentanol is directed through the tubes of the heat exchangers; and the isobutylene is directed through the shell-side channels, and at least two of the heat exchangers are connected in series with respect to both the shell-side stream and the tube-side stream.

[0088] Such a configuration allows for extended operating intervals between maintenance disruptions in such a process. The term "maintenance disruption" is intended to mean the periodic process shutdowns required to enable clearing of the tubes of the heat exchangers clogged with fouling. Typically, a maintenance disruption is indicated when the isobutylene leaving the last heat exchanger is insufficiently preheated and even subsequent heaters are unable to add additional external heat to the isobutylene to bring it to the desired temperature before it enters the reactor. In the configuration of the present invention, the preheating of the isobutylene stream can be maintained for a longer period of time at a sufficiently high level so that the desired temperature of the isobutylene can be readily achieved before it enters the reactor.

[0089] One particular area that is susceptible to fouling in conventional shell and tube heat exchangers is the area of ​​the tubes near the tube sheet, close to the inlet, where the tube side fluid leaves the individual tubes. Excessive fouling in this area can lead to plugging of individual tubes and fluid stagnation along the entire length of these tubes. Fluid stagnation generally results in reduced heat transfer performance.

[0090] As another consequence of the reduced heat transfer performance caused by fouling, the energy required in the heater to adjust the temperature of the preheated isobutylene stream to the desired reaction temperature increases. As a result, more additional external heat becomes necessary (which is disadvantageous in terms of energy requirements and process economics) and often has a negative impact on the carbon dioxide footprint of the product.

[0091] By using two or more heat exchangers, the effect of fouling in individual tubes on the overall heat exchange capacity is reduced compared to an arrangement using only a single heat exchanger. As a result, the heat transfer rate is maintained at a desired level for a longer period of time, thereby extending the operating interval between maintenance interruptions, and less additional external heat is required for preheating of the isobutylene stream compared to a plant with a single heat exchanger in a severely fouled state.

[0092] It is not easy to separate prenol from unreacted formaldehyde. This difficulty is caused by the fact that both monomeric formaldehyde (as well as polymeric formaldehyde) forms hydrates with water and hemiformals with prenol. Hydrates and hemiformals with different degrees of formaldehyde polymerization have mixing boiling points.

[0093] However, it has been found that formaldehyde can be almost completely separated from prenol by distillation at a temperature at which hemiformal is cleaved into formaldehyde and prenol, so that formaldehyde can be easily separated from prenol.

[0094] Therefore, crude isopentenol can be purified by subjecting a crude isopentenol stream containing isopentenol, water and formaldehyde or an isopentenol containing fractions thereof to distillation in a low-boiler separation column operated at a pressure of 2 bar absolute or more, preferably 2.5 bar absolute or more, to obtain a distillate stream containing aqueous formaldehyde and a bottom stream containing isopentenol substantially free of formaldehyde.

[0095] In particular, it has been found that formaldehyde can be separated almost completely from prenol and concentrated aqueous formaldehyde suitable for recycling to the prenol synthesis can be obtained in a distillation train comprising a first distillation carried out at a temperature at which the equilibrium shifts towards the hemiformal of formaldehyde and prenol, so that substantially all the formaldehyde remains in the bottoms of the distillation, and a second distillation carried out at a temperature at which the hemiformal is cracked into formaldehyde and prenol, so that the formaldehyde can be easily separated from the prenol.

[0096] In order to allow the first distillation to be carried out at a temperature lower than the prenol-formaldehyde dissociation temperature and the second distillation to be carried out at a temperature higher than the prenol-formaldehyde dissociation temperature, two low-boiling-point separation towers operating at different pressures are envisaged. Therefore, under the relatively low pressure existing in the first low-boiling-point separation tower, a first distillate containing water and low-boiling-points substantially free of formaldehyde is obtained. Under the relatively high pressure existing in the second low-boiling-point separation tower, almost all formaldehyde is separated from the prenol. Therefore, the method allows the prenol substantially free of formaldehyde to be obtained.

[0097] Therefore, in a more preferred embodiment, the purification method comprises

[0098] (i) directing the crude isopentenol stream to a first low boiler separation column operated at a pressure of 1.5 bar absolute or less to obtain a first bottom stream containing isopentenol and formaldehyde and a first distillate stream containing water and low boilers;

[0099] (ii) directing the first bottom stream to a second low boiler separation column operated at a pressure of 2 bar absolute or higher to obtain a second distillate stream containing aqueous formaldehyde and a second bottom stream containing isopentenol; and

[0100] (iii) the second bottom stream is directed to a refining column to obtain pure isopentanol as a distillate stream, and a bottom stream containing high boiling substances.

[0101] The second distillate stream constitutes concentrated aqueous formaldehyde suitable for recycling to the prenol synthesis.

[0102] Suitably, the second low boiler separation column is operated at a pressure of 2.5 bar absolute or higher, preferably 2.8 bar absolute or higher, most preferably 2.9 bar absolute or higher. The bottoms temperature of the second low boiler separation column is preferably in the range of 160° C. to 200° C., more preferably 170° C. to 185° C., most preferably 175° C. to 180° C. The top temperature of the second low boiler separation column is preferably in the range of 115° C. to 160° C., more preferably 125° C. to 145° C.

[0103] In a particularly preferred embodiment, the second low boiler separation column is operated at a pressure in the range of 2.9 to 3.5 bar absolute, a bottoms temperature in the range of 175 to 180°C and a top temperature in the range of 130 to 140°C.

[0104] Further information on methods for recovering essentially formaldehyde-free prenyl alcohol can be found in WO 2022 / 189652 A1.

[0105] The prenol obtained may be subjected to catalytic isomerization by contacting a reactant stream comprising prenol with at least one heterogeneous isomerization catalyst in order to obtain prenol.

[0106] The isomerization of isopentenol to 3-methyl-2-butene-1-ol (pentenol) can be carried out on a supported noble metal, preferably in the presence of hydrogen. Preferred catalysts are fixed bed catalysts containing palladium and selenium or tellurium or a mixture of selenium and tellurium supported on silicon dioxide. Isomerization is carried out at a temperature of 50 ℃ to 150 ℃ to produce a reaction mixture of prenol and isopentenol. Isopentenol can be recycled. Additional details are provided in WO 2008 / 037693.

[0107] When subjecting isopentenol to catalytic isomerization, it may be advantageous to maintain a weight ratio of formaldehyde to isopentenol in the reactant stream of less than 0.04, preferably less than 0.03, particularly less than 0.02, or less than 0.01. In still more preferred embodiments, the weight ratio of formaldehyde to isopentenol is maintained at less than 0.002, or less than 0.001.

[0108] It has been found that the presence of formaldehyde in the reactant stream is detrimental to the activity and selectivity of the process and may accelerate catalyst deactivation and / or poisoning.

[0109] Hydrogen is required for the isomerization of prenol to prenol. The poisoning mechanism resulting from the presence of formaldehyde is believed to involve the dehydrogenation of formaldehyde to carbon monoxide, which chemisorbs more strongly on the catalyst than hydrogen.

[0110] Another cause of catalyst deactivation, which may occur in combination with the previously mentioned causes of catalyst poisoning, is the formation of paraformaldehyde or trioxane, which can deposit on the catalyst in solid form and sequester the prenol being processed from the catalytically active surface. This leads to a gradual deactivation of the catalyst.

[0111] Additional by-products, such as prenaldehyde or isoamyl alcohol, which are typically contained in the isopentenol due to the manufacturing process are expected to have little or no influence on the performance of the isomerization reaction.

[0112] The weight ratio of formaldehyde to prenol in the reactant stream can be maintained at a specific level or less. However, the weight ratio of formaldehyde to prenol in the reactant stream is reduced to a point above a certain point, and a point of rapidly diminishing returns will be reached. The removal of formaldehyde involves additional equipment and operating costs. An economic balance must be achieved between the improvement brought by reducing the ratio and the cost of maintaining such a ratio. Therefore, the weight ratio of formaldehyde to prenol is preferably not less than 0.0005, or is not less than 0.005 in some cases.

[0113] The presence of formaldehyde in the reaction stream is due to two main sources. Formaldehyde can be contained in the fresh feed stream sent to the reactor, i.e. as an impurity originating from the prenol manufacturing step. All formaldehyde that cannot be separated in the purification step after the prenol synthesis ends up in the reaction stream.

[0114] In addition, formaldehyde is also generated in situ. Part of the prenol decomposes back to isobutylene and formaldehyde.

[0115] Since the double bond isomerization of prenol is an equilibrium reaction, conversion is necessarily incomplete. For the economic operation of the method, unconverted prenol must be removed and recycled. If the step of purifying the stream containing unreacted prenol is not taken, the recycling of prenol may unintentionally (re)introduce formaldehyde into the isomerization step.

[0116] Reducing the weight ratio of formaldehyde to isoprenol in the reactant stream can be achieved in several different ways. In an embodiment, formaldehyde is removed from the unreacted isoprenol stream before combining the unreacted isoprenol stream with the fresh feed stream.

[0117] In an example, an unreacted isopentenol stream is combined with a fresh feed stream and formaldehyde is removed from the combined stream.

[0118] Alternatively, it is feasible to mix the unreacted prenol stream with an amount of a sufficiently purified fresh feed stream in order to obtain the desired weight ratio of formaldehyde to prenol in the combined stream.

[0119] Formaldehyde may be removed from the isopentenol stream by conventional separation methods such as distillation, selective adsorption and or selective reaction, in particular by a purification process involving pressure swing distillation as described above.

[0120] Isopentenol, in particular the isopentenol obtained as described above can be converted into pentenal, involving isomerization and oxidative dehydrogenation in any order.Therefore, isopentenol can be first isomerized to pentenol, and then pentenol is oxidized to pentenal; Or, isopentenol oxidative dehydrogenation can be first to isopentenal, and optionally at least a portion of isopentenal isomerized to pentenal.

[0121] Pentenols, in particular pentenols obtained as described above, can be oxidized by contacting a reactant stream comprising pentenols with at least one oxidizing agent and at least one oxidation catalyst, preferably in the presence of a liquid phase, in order to obtain pentenals.

[0122] Suitable oxidizing agents include hydrogen peroxide and oxygen, especially oxygen.

[0123] The oxidation is preferably carried out in the presence of a liquid phase and with oxygen as the oxidant. The liquid phase preferably comprises at least 25 wt.-% of water, more preferably at least 50 wt.-% of water or at least 70 wt.-% of water, based on the total weight of the liquid phase, determined at a temperature of 20° C. and a pressure of 1 bar. It has been found that these conditions allow a simple and efficient process for preparing pentenals from pentenols.

[0124] The oxidation is typically carried out in the presence of at least one oxidation catalyst selected from the group consisting of platinum, palladium and gold. Preferably, the at least one oxidation catalyst comprises platinum. In a preferred embodiment, the at least one oxidation catalyst is a supported catalyst.

[0125] The oxidation is suitably carried out at a temperature of 20° C. to 100° C., preferably at a temperature of 20° C. to 70° C. The oxidation is suitably carried out at an oxygen partial pressure of between 0.2 and 8 bar.

[0126] The oxidative dehydrogenation of prenol typically comprises contacting a reactant stream, particularly a gaseous reactant stream, comprising prenol with at least one heterogeneous oxidative dehydrogenation catalyst, particularly at least one silver-containing heterogeneous oxidative dehydrogenation catalyst, in the presence of molecular oxygen. The at least one heterogeneous catalyst may be composed of an inert support having a smooth surface containing a silver active layer. Alternatively, a solid (all-metal) silver body may be used.

[0127] In an embodiment, the oxidative dehydrogenation is carried out by passing the isopentenol through a plurality of reaction tubes of a shell and tube heat exchange reactor comprising

[0128] - a shell-side heat exchange channel for circulating a heat transfer medium and a reaction channel comprising a plurality of reaction tubes;

[0129] - an inlet for introducing a reactant stream into the reaction channel; and

[0130] - an outlet from the reaction channel for recovering an effluent stream from the reaction tubes;

[0131] These reaction tubes include

[0132] a reactant preheating zone adjacent to the inlet, and

[0133] In a reaction zone downstream of the reactant preheating zone, the reaction zone has a catalytically active wire matrix insert having silver on at least a portion of its surface.

[0134] The term "reactant preheating zone" means a section of the reaction tube, i.e., a section inside the reaction tube, in which substantially no catalytic oxidative dehydrogenation reaction occurs and in which the gaseous stream passing through the reaction tube exchanges heat with a circulating heat transfer medium via the tube wall. The preheating zone upstream of the reaction zone involves a net heat flow into the reaction tube and ensures that the reactant stream is sufficiently heated to a temperature close to or at the reaction temperature when it reaches the reaction zone.

[0135] Upon contact with the catalytic surface, the oxidative dehydrogenation reaction begins immediately. Otherwise, coke formation may occur when a "cold" reactant stream reaches the catalytic surface so that the reaction start temperature of the reaction is not reached. Less coke formation advantageously leads to extended reactor operation without having to burn off coke on the catalytic surface.

[0136] Preferably, the reactant preheating zone is adapted to allow laminar flow of reactants inside the reactant preheating zone. This means that the reactant preheating zone does not have any obstacles to the reactant flow that triggers the laminar to turbulent flow transition. Therefore, the reactant preheating zone preferably has a substantially free cross section, i.e. the preheating zone is empty.

[0137] In the case of a "substantially free cross section", the reactant preheating zone can be empty. Alternatively, the reactant preheating zone can accommodate fixtures made of materials with zero or limited catalytic activity, which have negligible cross sections in a plane perpendicular to the longitudinal axis of the reaction tube. The fixtures can be attached to a catalytically active wire matrix present in the reaction zone and allow the wire matrix plug to be easily placed in or removed from the reaction zone. For example, the negligible fixtures can be stainless steel wires or rods.

[0138] This arrangement allows heating only the portion of the entire reactant stream that travels near the hot reaction tube wall. Therefore, the portion of the reactant stream that flows in the center of the reaction tube is not heated to the reaction temperature, and therefore reduces or even avoids the blind reaction of unstable starting materials. "Blind reaction" is a non-selective oxidation reaction that occurs in the absence of a catalyst. Once the reactant stream arrives at the reaction zone, an oxidative dehydrogenation reaction is initiated. Due to the exothermic nature of this reaction, energy is released and the remainder of the reactant stream is quickly heated to the reaction start temperature, and the reaction proceeds. This rapid heating of the main part of the reaction mixture reduces undesirable side reactions and therefore produces increased selectivity.

[0139] Alternatively, the reactant preheating zone can have a wire matrix plug-in with zero or limited catalytic activity. The wire matrix plug-in can reduce or eliminate the temperature gradient without any obstacle to the flow that will promote turbulent flow characteristics. If the wire matrix plug-in does not catalyze the gas phase partial oxidation reaction discussed to a significant extent, and the chemical composition of the flow passing through the wire matrix plug-in does not change significantly, then the wire matrix plug-in is considered to have zero catalytic activity (or in other words, is considered to be "inert"). Similarly, if the catalytic activity of the matrix plug-in is less than the activity of the reaction zone, the matrix plug-in is considered to have limited catalytic activity. In an embodiment, the wire matrix plug-in with zero or limited catalytic activity is made of inert material, preferably stainless steel.

[0140] In this article, the term "reaction zone" means the area in the reaction tube where the catalytic gas phase partial oxidation reaction occurs. The reaction zone includes a catalytically active wire matrix plug-in, which has a catalytically active precious metal at least on a portion of its surface. Since the wire matrix included in the reaction zone is more open in structure than the filler of the individual element, a larger proportion of the reaction heat is discharged to the reaction tube wall by radiation and does not have to be dissipated by the reactant stream. Due to the unique flow characteristics of the reactant stream of the reaction tube with the wire matrix plug-in in the appropriate position, the heat transfer via the tube wall is improved. Significant hot spots can be avoided. This in turn avoids the deposition of the organic components of the reactant stream on the surface of the active catalyst material, and is accompanied by a pressure drop. In general, less regular maintenance in the form of regeneration and / or replacement of the catalyst is required. The annual operating hours can be increased, and the existing production capacity can be fully utilized, thereby reducing operating costs and increasing profits.

[0141] The wire matrix insert may be formed continuously or integrally with the catalyst body which is present separately. Therefore, it is very advantageous to place and remove the wire matrix insert in the catalyst receiving area of ​​the reaction tube.

[0142] A "reaction zone" may consist of a single continuous reaction zone. Alternatively, the reaction zone may comprise a series of alternating regions with catalytically active wire matrix inserts and regions with substantially free cross sections or with wire matrix inserts of zero or limited catalytic activity.

[0143] A "wire matrix insert" is understood to be a self-supporting skeleton-like structure made of coiled, bent or curled metal wires, which is adapted to be inserted into a reaction tube of a shell and tube reactor. A wire matrix insert has a structure with a greater volume than the longitudinal wires.

[0144] Fixing devices such as stainless steel wires or rods may be attached to the wire matrix insert, which allows the wire matrix insert to be easily placed into or removed from the reaction zone.

[0145] In an embodiment, the catalytically active wire matrix insert comprises an elongated core having a plurality of wire loops extending from the elongated core, wherein the wire loops are longitudinally arranged and spirally displaced, i.e., adjacent wire loops have an angular offset. The loops can be formed by spirally bending the wire over the length of the wire matrix insert. In view of the ease of manufacture, the elongated core preferably comprises at least two longitudinal core wire members, which are twisted around each other to form a core wire winding, and the wire loops are accommodated in the core wire winding.

[0146] The wire loop may be formed from one or more than one wound wire, preferably 4 wound wires.

[0147] The wire matrix insert included in the reaction zone has catalytically active precious metal silver on at least a portion of its surface. The wire constituting the wire loop can be a solid silver wire or a wire coated with silver. The core wire can be made of a brass alloy or high-grade steel. The silver coating superimposed on the core surface has a thickness of, for example, 10 μm. However, in general, solid silver wire has a better service life and is preferred. If the wire loop is formed by more than one wound wire, at least one of the wound wires is made of a solid silver wire or a wire coated with silver, while the other wound wires can be made of an inert material.

[0148] Silver wire having the same composition throughout its cross section and containing at least 92.5 wt.-% Ag can be suitably used. The silver wire is spirally bent to form a wire loop and is combined with at least two longitudinal core wire members, which are twisted around each other to form a core wire winding, and the wire loop is accommodated in the core wire winding. The longitudinal core wire member can also be a silver wire or an inert metal wire.

[0149] In a preferred embodiment, the catalytically active wire matrix insert comprises an elongated core having a plurality of wire loops extending therefrom, wherein the wire loops are longitudinally arranged and helically displaced, and wherein the wire loops comprise solid silver wires.

[0150] In general, the catalytically active wire matrix insert has a cylindrical cladding surface whose diameter matches the inner diameter of the reaction tube. This includes situations where the diameter of the cylindrical cladding surface of the undeployed wire matrix insert is slightly larger than the inner diameter of the reaction tube. Due to the elastic or resilient nature of the wire matrix insert, it can be inserted into the reaction tube with a slight back pressure so that the wire loop fits tightly against the inner wall of the reaction tube.

[0151] Suitable structures for wire matrix inserts are known per se, see for example GB 2 097 910. Some inserts of this type are disclosed in GB Patent 1 570 530. Other inserts and methods of producing them are disclosed in GB 2 097 910 A. Matrix inserts are commercially available from Cal Gavin Ltd. in the United Kingdom and are marketed under the trade name Sale.

[0152] When subjecting isopentenol to oxidative dehydrogenation, it may be advantageous to maintain a weight ratio of formaldehyde to isopentenol in the reactant stream of less than 0.04, preferably less than 0.03, particularly less than 0.02, or less than 0.01. In still more preferred embodiments, the weight ratio of formaldehyde to isopentenol is maintained at less than 0.002, or less than 0.001.

[0153] The weight ratio of formaldehyde to prenol in the reactant stream can be maintained at a specific level or less. However, the weight ratio of formaldehyde to prenol in the reactant stream is reduced to a point above a certain point, and a point of rapidly diminishing returns will be reached. The removal of formaldehyde involves additional equipment and operating costs. An economic balance must be achieved between the improvement brought by reducing the ratio and the cost of maintaining such a ratio. Therefore, the weight ratio of formaldehyde to prenol is preferably not less than 0.0005, or is not less than 0.005 in some cases.

[0154] It has been found that reactor blockage and pressure drop increase are significantly affected by the presence of formaldehyde in the reactant stream. The catalyst fouling reaction of condensation and polymerization is believed to be the main reaction involved in the formation of carbon or coke on the catalyst. It is believed that this carbon formation involves the thermal condensation of formaldehyde or formaldehyde with the olefinic hydrocarbons isopentenol and (iso) amylene aldehyde. In the presence of a catalyst, the main condensation product tends to undergo dehydrogenation and polymerization type reactions and settles on the catalyst and undergoes further dehydrogenation and decomposition until a carbonaceous deposit is formed.

[0155] The presence of formaldehyde in the reaction stream is due to two main sources. Formaldehyde can be contained in the fresh feed stream sent to the reactor, i.e. as an impurity originating from the prenol manufacturing step. All formaldehyde that cannot be separated in the purification step after the prenol synthesis ends up in the reaction stream.

[0156] In addition, formaldehyde also can be generated in situ. Part of prenol decomposes back to isobutylene and formaldehyde. Because most of continuous industrial processes are operated at a single-pass conversion level of 50% to 60% and when unconverted prenol is recycled, if the step of purifying the stream containing unreacted prenol is not taken, formaldehyde may be present in the recycle stream of unconverted prenol. It has been found that the recycle stream of unconverted prenol constitutes the largest source of formaldehyde pollutants in the reaction stream. The method is usually carried out with a partial conversion, for example, with a conversion of 30% to 70%, preferably 50% to 60%. The unreacted prenol stream is separated from the product stream. The unreacted prenol stream is recycled, that is, merged with the fresh feed stream containing prenol to provide a reaction stream. The unreacted isopentenol stream contains isopentenol as the main component, but may also contain prenaldehyde, isopentenaldehyde, isopentanol, isovaleraldehyde, isovaleric acid, prenol, formaldehyde. It may also contain trace amounts of other C3 and C2 aldehydes and acids.

[0157] Reducing the weight ratio of formaldehyde to isoprenol in the reactant stream can be achieved in several different ways. In an embodiment, formaldehyde is removed from the unreacted isoprenol stream before combining the unreacted isoprenol stream with the fresh feed stream.

[0158] In an example, an unreacted isopentenol stream is combined with a fresh feed stream and formaldehyde is removed from the combined stream.

[0159] Alternatively, it is feasible to mix the unreacted prenol stream with an amount of a sufficiently purified fresh feed stream in order to obtain the desired weight ratio of formaldehyde to prenol in the combined stream.

[0160] Formaldehyde may be removed from the isopentenol stream by conventional separation methods such as distillation, selective adsorption and or selective reaction, in particular by a purification process involving pressure swing distillation as described above.

[0161] Before contacting with at least one oxidative dehydrogenation catalyst or with at least one oxidation catalyst, respectively, the (iso)prenol may advantageously be treated to remove organically bound nitrogen from the (iso)prenol by contacting the (iso)prenol with a weakly acidic solid adsorbent. In other words, the (iso)prenol may be depleted of organically bound nitrogen by this method.

[0162] The term "organically bound nitrogen" is intended to mean any compound containing at least one nitrogen atom directly bound to one or more carbon atoms. For example, such compounds containing at least one nitrogen atom may be selected from amines such as ethylamine, trimethylamine, aniline, pyridine or piperidine. An amine of particular importance in practice is hexamethylenetetramine (urotropine). (Iso)prenol may contain about 5 to 30 ppm of organically bound nitrogen.

[0163] It has been found that weakly acidic solid adsorbents are able to adsorb organically bound nitrogen in the presence of large amounts of (iso)prenol without disturbing reactive carbon-carbon double bonds.

[0164] The weakly acidic adsorbent may comprise an adsorbent material having sufficient acidity to adsorb organically bound nitrogen from (iso)prenol. In an embodiment, the solid adsorbent is a cross-linked resin with phosphonic acid functional groups. Preferably, the resin polymer is a vinyl aromatic copolymer, preferably a cross-linked polystyrene and more preferably a polystyrene divinylbenzene copolymer. Other polymers with phosphonic acid functional groups may also be used. Preferably, the cross-linked resin with phosphonic acid functional groups is of the macroporous type. The preferred solid adsorbent is Purolite S956.

[0165] Resin is typically used in bead form and is loaded into the post. Make (iso) prenol pass through the post, contact resin beads. During contact, the organically bound nitrogen in the (iso) prenol is reacted with functional groups and exchanged, wherein proton is transferred to nitrogen and ionic bond is formed to the anionic site of the resin. Maintain contact until reaching threshold level, i.e. breakthrough concentration. At this breakthrough point, the process reaches equilibrium, and wherein the extra organically bound nitrogen can not be effectively removed. Make flow stop and the post is backwashed with water, preferably deionized water or softened water. By reverse flow, the resin is fluidized, and the solid captured by the beads is loosened and removed.

[0166] In another embodiment, the solid adsorbent is a silica-alumina hydrate.Many silica-alumina catalyst compositions and methods for their preparation are described in the patent literature, see for example US 4,499,197.

[0167] Preferably, the alumina content of the silica-alumina hydrate is from about 10 to about 90 wt.-% Al2O3. A preferred range of alumina content is from about 30 to about 70 wt.-% Al2O3.

[0168] The introduction of silica into alumina results in the introduction of acidic centers. The number of acidic centers can be controlled by the amount of silica introduced. The number of acidic centers increases with the amount of silica introduced up to a maximum number of acidic centers and decreases again with a further increase in the amount of silica after the maximum number of acidic centers has been reached.

[0169] An example of a commercially available silica-alumina hydrate is available from Sasol Germany Gmbh, Hamburg, Germany. Based on orthorhombic aluminum oxide hydroxide (boehmite; AlOOH) and doped with SiO2. Available are various Grades: Siral 1 (Al2O3 / SiO2=99 / 1), Siral 5 (Al2O3 / SiO2=95 / 5), Siral 10 (Al2O3 / SiO2=90 / 10), Siral 20 (Al2O3 / SiO2=80 / 20), Siral 28M (Al2O3 / SiO2=72 / 28), Siral 30 (Al2O3 / SiO2=70 / 30), Siral 40 (Al2O3 / SiO2=60 / 40). Siral 40 is particularly preferred.

[0170] In an embodiment, the (iso)prenol is passed through a bed of a weakly acidic solid adsorbent. Suitably, the step of "passing ... through a bed" means providing a layer ("bed") of a weakly acidic solid adsorbent in a conventional reaction vessel known to the skilled person (e.g. in a stirred tank reactor) which may preferably be equipped with a stirring device. The (iso)prenol is then introduced into the reaction vessel and directed through the reaction vessel in such a way that it comes into contact with the weakly acidic solid adsorbent.

[0171] Alternatively, the weakly acidic solid adsorbent may be provided in reaction tubes, for example reaction tubes of a tubular reactor, and the (iso)prenol is then continuously flowed through one or more of the reaction tubes while contacting the weakly acidic solid adsorbent.

[0172] In an embodiment, after contacting the alcohol stream with the weakly acidic solid adsorbent, the (iso)prenol contains less than 2 ppm of organically bound nitrogen. In this context, "ppm" means the wt.-ppm of compounds incorporating organically bound nitrogen relative to the total weight of the (iso)prenol.

[0173] Suitably, the content of organically bound nitrogen in the (iso)prenol can be determined by Kjeldahlanalysis. Alternatively, an oxidative combustion method according to DIN 51444 with a chemiluminescence detector can be used.

[0174] Citral is a useful intermediate for, for example, menthol or linalool.

[0175] Menthol can be prepared from citral via a process comprising the following steps

[0176] - catalytic hydrogenation of citral to obtain citronellal;

[0177] - cyclizing citronellal in the presence of at least one acidic catalyst to obtain isopulegol; and

[0178] - Catalytic hydrogenation of isopulegol to obtain menthol.

[0179] The overall reaction sequence is illustrated by the following reaction scheme.

[0180]

[0181] The hydrogenation of citral to obtain citronellal can be achieved by carrying out the hydrogenation in the presence of a rhodium-phosphine catalyst.

[0182] The cyclization of citronellal to isopulegol can be achieved by carrying out the cyclization in the presence of at least one Lewis acidic aluminum-containing catalyst, such as a bis(diarylphenoxy)aluminum compound, which catalyst can be used in the presence of an auxiliary such as a carboxylic anhydride. Isopulegol can be recovered from the catalyst-containing reaction product by distillative separation to give an isopulegol-rich overhead product and an isopulegol-depleted bottom product. At least one catalyst can be regenerated from the bottom product. The isopulegol obtainable in this way by the cyclization of citronellal can be further purified by suitable separation and / or purification methods, in particular by crystallization, and is at least largely free of undesirable impurities or by-products.

[0183] The hydrogenation of isopulegol can be achieved by carrying out the hydrogenation in the presence of at least one heterogeneous nickel-containing catalyst, preferably at least one heterogeneous nickel- and copper-containing catalyst.

[0184] Further details on the reaction sequence from citral to menthol can be found in US 2013 / 46118 A1 (which is incorporated herein by reference).

[0185] Thus, in one aspect, the invention relates to an improved process for preparing menthol by producing citral using the above process and then producing menthol from the citral.Menthol may be prepared as described herein or by other processes known in the art.

[0186] Linalool can be prepared from citral via a process comprising catalytic hydrogenation of citral to obtain nerol and / or geraniol and isomerization thereof.

[0187] The hydrogenation of citral to obtain nerol and / or geraniol can be achieved by carrying out the hydrogenation in the presence of at least one supported ruthenium, rhodium, osmium, iridium or platinum catalyst, preferably at least one ruthenium catalyst supported on carbon black.

[0188] Isomerization of nerol and / or geraniol to obtain linalool can be achieved by carrying out the isomerization in the presence of at least one tungsten catalyst, in particular a tungsten dioxide (VI) complex. Further details on the isomerization of nerol and / or geraniol can be found in US 7,126,033 B2.

[0189] Thus, in one aspect, the present invention relates to an improved method for preparing linalool by producing citral using the above method and then producing linalool from citral. Linalool may be prepared as described herein or by other methods known in the art.

[0190] The present invention is further illustrated by the following examples.

[0191] Examples

[0192] In these examples, the term "by-product" is abbreviated as "BP." With respect to GC analysis, all "%" are reported as "area-%".

[0193] Example 1

[0194] This example illustrates the conversion and byproduct formation in a model reaction that simulates the cleavage reaction of dipentenyl acetal. The conversion and byproducts are reported as a function of time at various combinations of cleavage temperature and catalyst concentration.

[0195] Dipentenyl acetal (25 g, purity 73.8%, containing 1.6% 2,4,4-trimethyl-3-formyl-1,5-hexadiene, 5.4% pentenyl (3-methylbutadienyl) ether, 5.8% BP 1a / b) was mixed with H3PO4 (85%) in an amount of 250ppm, 530ppm, or 750ppm, respectively, and divided equally into 6 microwave containers (volume 10mL). The microwave containers were sealed with lids and heated to a temperature of 150°C, 160°C, 170°C, or 180°C, respectively. Each microwave container was heated for an operating time of 0min, 5min, 10min, 15min, 30min, and 60min. After subsequently cooling to room temperature, solid K2CO3 (100-200mg) was added to each reaction mixture, and the reaction mixture was filtered using a syringe filter. Afterwards, the neutralized reaction mixture was analyzed with GC. The results are reported in Tables 1 to 7.

[0196]

[0197]

[0198]

[0199]

[0200] For each tested temperature and catalyst concentration, the evolution of by-product 5 versus the residual concentration of dipentenyl acetal (see Tables 1 to 7) is plotted in Figure 1 In. From Figure 1It can be seen that for all temperatures and catalyst concentrations, the formation of by-product 5 increases sharply as the content of dipentenyl acetal approaches 0%, ie, the conversion approaches 100%.

[0201] Figure 2 The concentration of dipentenyl acetal is described over time for each temperature and catalyst concentration tested (see Tables 1 to 7). It can be seen that the higher the temperature and the higher the catalyst concentration, the faster the reaction proceeds.

[0202] Figure 3 The concentration of Σcitral building blocks (pentenyl (3-methyl-butadienyl) ether, 2,4,4-trimethyl-3-formyl-1,5-hexadiene and citral) is plotted over time for each tested temperature and catalyst concentration (see Tables 1 to 7). Figure 3 An initial increase in the concentration of Σcitral building blocks is shown. For runs at temperatures of 160° C. or higher, after reaching a maximum, the content of Σcitral building blocks decreases.

[0203] Figure 4 The concentration of byproduct 5 is plotted over time for each tested temperature and catalyst concentration (see Tables 1 to 7). Figure 4 It can be seen that the formation of byproduct 5 is slow at first, followed by a sharp increase in the concentration of byproduct 5. Figure 4 It can also be seen that the formation of byproduct 5 starts earlier and the formation rate is faster as the cleavage temperature and phosphoric acid concentration increase.

[0204] Example 2

[0205] Example 2 illustrates the continuous reactive distillation of dipentenyl acetal, wherein the cleavage fractions (pentenyl (3-methyl-butadienyl) ether, 2,4,4-trimethyl-3-formyl-1,5-hexadiene and citral) are continuously discharged.

[0206] A distillation column with structured packing (Montz A3-750) having a height of 800 mm was used, equipped with an external evaporator and a circulation pump with a hold up of 240 mL as well as a top condenser and a reflux distributor.

[0207] The tower is operated at a top pressure of 50 mbar. The feed stream of dipentenyl acetal (content 75%) is mixed with the storage tank stream and metered into the evaporator. The H3PO4-catalyst is diluted in pentenol and metered into the storage tank circulation stream as a 1%-solution. For each experiment, a sufficient lead time of 5-6h is selected until steady-state conditions are obtained. In operations 8-1 to 8-4, the distillate stream is taken out at the top of the tower. In operations 8-5 to 8-6, side fractions and top streams are taken out. Samples of the storage tank stream and the distillate stream are taken out and analyzed with GC. The precise H3PO4 concentration in the storage tank is determined via elemental P-analysis. The reaction conditions (residence time, H3PO4 amount and temperature) are shown in Tables 8a and 8b.

[0208]

[0209]

[0210] Tables 8a and 8b show higher yields of citral building blocks and less by-product formation compared to Example 1. Comparison of runs 8-2, 8-3, 8-4, and 8-5 shows that a balance between H3PO4 concentration and reaction temperature must be maintained. Too high a concentration of H3PO4 at high temperatures may be detrimental in terms of the yield of citral building blocks.

[0211] Example 3

[0212] Example 3 illustrates that if the intermediate 2,4,4-trimethyl-3-formyl-1,5-hexadiene is carried to step a) and subjected to the reaction conditions of step a), by-products BP 1a / b, BP 2, BP 3, BP 4 and BP 5a / b are formed. This example also illustrates the fate of 2,4,4-trimethyl-3-formyl-1,5-hexadiene in step a) and the products formed therefrom in step b).

[0213] In a 500mL reactor equipped with a 15cm Vigreux distillation tower, a condenser, a distillation template and a vacuum pump, 2,4,4-trimethyl-3-formyl-1,5-hexadiene (176g, 92% purity, containing 6% citral) was dissolved in pentenol (208g, 2.4mol) and cyclohexane (300mL). After adding HNO3 (500ppm), the reaction mixture was refluxed (80°C-95°C, 200 mbar) and the water formed during the reaction was removed via azeotropic distillation using a Dean-Stark apparatus. After reflux for 3h, 13g of water was removed from the reaction mixture. The reaction mixture was analyzed via GC. The results are shown in Table 9.

[0214] Table 9.

[0215]

[0216] [1] Ratio 1:1

[0217] [2] 2,4,4-Trimethyl-3-formyl-1,5-hexadiene

[0218] [3] nd = not determined

[0219] The vacuum was gradually reduced to 1 mbar and low boiling point compounds (HNO , cyclohexane, pentenol, pentenal, BP1a / b, citral) were distilled from the reaction mixture. 103 g of light yellow storage tank residue was obtained. The storage tank was analyzed via GC. The results are shown in Table 10.

[0220] Table 10.

[0221]

[0222] [1] 2,4,4-Trimethyl-3-formyl-1,5-hexadiene

[0223] [2] nd = not determined

[0224] Catalyst H3PO4 (85%, 0.03 g) was added to the residue and the temperature was gradually raised to 140°C-145°C under a vacuum of 10 mbar. During the temperature rise, the formed product was distilled from the reaction mixture. After 2 h, 55 g of distillate were collected. The distillate was analyzed via GC. The results are shown in Table 11.

[0225] Table 11.

[0226]

[0227] [1] 2,4,4-Trimethyl-3-formyl-1,5-hexadiene

[0228] [2] nd = not determined

[0229] Example 3 confirms that under the reaction conditions of step a), 2,4,4-trimethyl-3-formyl-1,5-hexadiene is mostly converted into BP 2 and BP 3 (wherein pentenol is as a byproduct), wherein only a small amount of citral. The remaining tank fraction after removing low-boiling compounds is mainly composed of BP 2 and BP 3 (see Table 10). It is believed that the tank fraction simulates the pollutants carried to step b) together with the crude dipentenyl acetal. Under the reaction conditions of step b), most of BP 3 is formed by BP 2, wherein pentenol is as a byproduct. Under the reaction conditions of step b), no additional citral is obtained by the tank fraction. Therefore, BP 3 cannot be converted into citral and can be considered as the end point in the reaction route.

[0230] Example 4

[0231] Example 4 illustrates the effect of reaction temperature and nitric acid concentration in step a).

[0232] In the 250mL three-necked flask equipped with oil bath, magnetic stirring apparatus, 15cm Vigreux distillation tower, condenser, distillation template and vacuum pump, add starting material pentenol (120.4g, 1.4mol) and pentenal (58.8g, 0.7mol) and stir at room temperature.Add HNO3 (65%) with the amount of 100ppm, 250ppm or 500ppm respectively, evacuate the reaction mixture to 100 millibars and increase the temperature to 80 ℃-82 ℃ or 90 ℃-91 ℃ respectively, until the reaction mixture begins to reflux.Remove the formed water via azeotropic distillation, and vacuum is gradually reduced to 70 millibars.After 1h, stop the reaction.Analyze the organic phase in storage tank gained and water separator via GC.The results are shown in Table 12.

[0233]

[0234] As can be seen from Table 12, an increase in the reaction temperature and nitric acid concentration in step a) leads to an increase in the formation of by-product 1 and high-boiling point by-products (BP 2 to BP 5), and a decrease in the selectivity of pentenols and pentenals.

Claims

1. A method for preparing 3,7-dimethyl-octa-2,6-dienal (citral), the method comprising the following steps: a) continuously condensing pentenols with pentenals in a reaction column in the presence of at least one catalyst, wherein the water of condensation is distilled off in vapor form as a pentenal-water azeotrope, at least partly condensing the vapor and separating the condensate into an aqueous phase and an organic phase, and conducting the organic phase partly as reflux to the reaction column and discharging the organic phase partly as a purge stream, At the same time, an acetal fraction comprising dipentenyl acetal of pentenal is continuously discharged from the reaction column, wherein the reaction temperature is below 100°C, the catalyst is nitric acid and the concentration of the nitric acid is below 500 ppm; b) continuously subjecting the acetal fraction to cleavage conditions in a cleavage column in the presence of at least one catalyst, wherein pentenol is removed, while continuously withdrawing from the cleavage column a cleavage fraction containing at least one of pentenyl (3-methyl-butadienyl) ether and 2,4,4-trimethyl-3-formyl-1,5-hexadiene and optionally citral, wherein the conversion of the dipentenyl acetal of pentenals in step b) is maintained at more than 90% and less than 100%, and unreacted dipentenyl acetal is at least partially contained in the discharged cleavage fraction; c) reacting the cleavage fraction in a plug flow type reactor to obtain citral; as well as d) recycling a portion of the pentenol obtained in step b) to step a).

2. The method according to claim 1, wherein: The pyrolysis temperature in step b) is above 150°C and below 200°C.

3. The method according to claim 1 or 2, wherein: The catalyst in step b) is phosphoric acid.

4. The method according to claim 3, wherein: The concentration of phosphoric acid in the bottoms of the cracking tower is higher than 100 ppm and lower than 1500 ppm.

5. A method according to any one of the preceding claims, wherein: The residence time in step b) is higher than 5 min and lower than 90 min.

6. A method according to any one of the preceding claims, wherein: The cleavage fraction is discharged from the cleavage column as a side draw, and the removed pentenol is discharged via the top of the cleavage column.

7. A method according to any one of the preceding claims, wherein: The rate of the purge flow is such that the total fixed concentration of byproducts 1a and 1b In the feed to step a) it is kept below 7 wt.-%, preferably between 2 and 7 wt.-%.

8. A method according to any one of the preceding claims, wherein: The reaction pressure in step a) is below 150 mbar.

9. A method according to any one of the preceding claims, wherein: The concentration of 2,4,4-trimethyl-3-formyl-1,5-hexadiene recycled from step b) to the pentenol in step a) is controlled such that the concentration of 2,4,4-trimethyl-3-formyl-1,5-hexadiene in step a) is below 1 wt.-% relative to the total weight of pentenol and pentenal.

10. A method according to any one of the preceding claims, wherein: The concentration of citral in the pentenols recycled from step b) to step a) is controlled such that the concentration of citral in step a) is below 1 wt.-%, relative to the total weight of pentenols and pentenals.

11. A method according to any one of the preceding claims, wherein: Prenol is obtained by reacting at least one formaldehyde source and isobutylene to obtain 3-methylbut-3-en-1-ol (isopentenol), and subjecting at least a portion of the isopentenol obtained to isomerization by contacting a reactant stream comprising isopentenol with at least one heterogeneous isomerization catalyst, preferably in the presence of hydrogen.

12. The method according to claim 11, wherein: the prenol being obtained by reacting the at least one formaldehyde source and isobutylene in a reactor to obtain prenol, Wherein reacting the at least one formaldehyde source and isobutylene preferably comprises at least one of alpha, beta and gamma: a) mixing the at least one formaldehyde source and isobutylene and injecting them into a reactor through a plurality of nozzles operated in parallel, and reacting the formaldehyde source and isobutylene under supercritical conditions; wherein the reactor comprises a vertically arranged container, a side wall, an upper portion and a lower portion; and wherein the formaldehyde source and isobutylene are injected into a mixing chamber of the reactor arranged in the upper portion, and a fluid containing formaldehyde and / or isobutylene and / or isopentanol enters from the mixing chamber into a post-reaction chamber arranged in the lower portion; as well as Providing draft tubes disposed substantially concentrically below each nozzle in the mixing chamber, the draft tubes providing downcomers within the draft tubes and upcomers outside the draft tubes such that the formaldehyde source and isobutylene injected through the nozzles travel generally downward in the downcomers and then the fluid containing formaldehyde and / or isobutylene and / or isopentanol turns in a generally upward direction in the upcomers and backmixes the fluid with the injected formaldehyde source and isobutylene; β) mixing the at least one formaldehyde source and isobutylene and injecting the at least one formaldehyde source and isobutylene into an internal loop reactor through at least one nozzle into one or more first conduits, the internal loop reactor comprising: - a vertically arranged cylindrical container comprising a side wall; - at least one flow guide cylinder having a cylinder inlet end and a cylinder outlet end arranged vertically within the container, the one or more flow guide cylinders being arranged concentrically with the one or more nozzles and having an inner surface and an outer surface, wherein the one or more flow guide cylinders provide the one or more first conduits within the one or more flow guide cylinders and a second conduit outside the one or more flow guide cylinders and within the sidewall, the one or more first conduits being in fluid communication with the second conduit; - reactor fluid outlet means; wherein the inner surface of the one or more flow guide sleeves is convexly curved so that the cross-section of the one or more first conduits exhibits an annular constriction between the sleeve inlet end and the sleeve outlet end; wherein the constriction is located closer to the sleeve inlet end; wherein the convex curvature of the inner surface of the one or more flow guide sleeves extends over at least 70%, preferably at least 80%, and most preferably at least 90% of the length of the flow guide sleeve; and wherein the outer surface of the one or more flow guide sleeves is convexly curved so that the one or more flow guide sleeves exhibit a circumferential bulge between the sleeve inlet end and the sleeve outlet end, and the position of the circumferential bulge is preferably closer to the sleeve outlet end; wherein the convex curvature of the outer surface of the one or more flow guide sleeves extends over at least 70%, preferably at least 80%, and most preferably at least 90% of the length of the flow guide sleeve; and wherein the edges of the one or more draft tubes are rounded so that the at least one formaldehyde source and isobutylene injected through the nozzles travel generally downward in the one or more first conduits to obtain a reaction fluid, which then turns in the opposite direction to travel through the second conduit and subsequently backmix with the injected fluid; γ) exchanging heat between a hot isopentanol stream withdrawn from the reactor and an isobutylene stream directed to the reactor; wherein the heat exchange is performed in at least two shell and tube heat exchangers; each of these heat exchangers comprises a plurality of tubes and shell side heat exchange channels; wherein the hot isopentenol is conducted through the tubes of the heat exchangers; and the isobutylene is conducted through the shell-side channel, and at least two of the heat exchangers are connected in series with respect to both the shell-side flow and the tube-side flow.

13. The method according to claim 11 or 12, further comprising at least one of the following: αα and ββ: αα) purifying the prenol by subjecting a crude prenol stream containing prenol, water and formaldehyde or a prenol containing fraction thereof to distillation in a low boiler separation column operated at a pressure of 2 bar absolute or more, preferably 2.5 bar absolute or more, to obtain a distillate stream containing aqueous formaldehyde and a bottom stream containing prenol substantially free of formaldehyde; ββ) maintaining a weight ratio of formaldehyde to isopentenol in the reactant stream of less than 0.

04.

14. A method according to any one of the preceding claims, wherein: The pentenals are provided by at least one of γ-i) and γ-ii): γ-i) subjecting the prenol to an oxidative dehydrogenation by contacting a reactant stream comprising the prenol with at least one heterogeneous oxidative dehydrogenation catalyst in the presence of molecular oxygen to obtain pentenals and / or isopentenals, and optionally isomerizing at least a portion of the isopentenals to prenals; wherein step γ-i) is optionally characterized by maintaining a weight ratio of formaldehyde to isopentenol in the reaction stream of less than 0.04; γ-ii) oxidizing pentenols to obtain pentenals by contacting a reactant stream comprising pentenols with at least one oxidizing agent and at least one oxidation catalyst, preferably in the presence of a liquid phase.

15. The method according to claim 14, wherein: The oxidative dehydrogenation of step γ-i) is carried out by passing the isopentenyl alcohol through a plurality of reaction tubes of a shell-and-tube heat exchange reactor, the shell-and-tube heat exchange reactor comprising - a shell-side heat exchange channel for circulating a heat transfer medium and a reaction channel comprising a plurality of reaction tubes; - an inlet for introducing a reactant stream into the reaction channel; and - an outlet from the reaction channel for recovering an effluent stream from the reaction tubes; These reaction tubes include a reactant preheating zone adjacent to the inlet, and In a reaction zone downstream of the reactant preheating zone, the reaction zone has a catalytically active wire matrix insert having silver on at least a portion of its surface.

16. The method according to claim 15, wherein: The catalytically active wire matrix insert includes an elongated core having a plurality of wire loops extending therefrom, wherein the wire loops are longitudinally arranged and helically displaced, and wherein the wire loops contain solid silver wires.

17. A process for preparing a citral-derived chemical, the process comprising preparing citral by a process according to any one of the preceding claims, and at least one of ααα, βββ or (βββ plus γγγ): ααα) converting the citral to obtain menthol; βββ) converting the citral into geraniol and / or nerol; γγγ) converting the geraniol and / or nerol to obtain linalool.

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