Shell and tube heat exchange reactor for performing catalytic gas phase partial oxidation reaction and method for performing catalytic gas phase partial oxidation reaction

Abstract

Disclosed is a shell and tube heat exchange reactor for performing a catalytic gas phase partial oxidation reaction, 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 tube. The reaction tube includes a reactant preheating zone adjacent to the inlet and a reaction zone downstream of the reactant preheating zone, the reaction zone having a catalytically active wire matrix insert having a catalytically active noble metal at least on a portion of its surface. The reactor requires less frequent maintenance in the form of regeneration and / or replacement of the catalyst. The catalyst may be readily placed into and removed from the reaction tube. Only the portion of the entire reactant stream that travels near the thermal reaction tube wall is heated. Thus, the portion of the reactant stream flowing in the center of the reaction tube is not heated to the reaction temperature, and thus blind reactions of the unstable starting material are reduced or even avoided.

Description

[0001] The present invention relates to a shell-and-tube heat exchange reactor for carrying out a catalytic gas phase partial oxidation reaction, 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 tube, and to a method for carrying out a catalytic gas phase partial oxidation reaction in the shell-and-tube heat exchange reactor. The present invention further relates to a method for preparing pentenol, 3,7-dimethyl-octa-2,6-dienal (citral), menthol and linalool.

[0002] Catalytic gas phase reactions in the chemical industry, such as oxidation, hydrogenation, dehydrogenation, nitration, alkylation, etc., are usually carried out using a shell-and-tube reactor using a solid catalyst arranged in a fixed bed. Such reactions carried out in a shell-and-tube reactor can be endothermic or exothermic. The position of the fixed bed is in the reaction tube of the shell-and-tube reactor. Nowadays, commonly used shell-and-tube reactors can have at least 5,000 and up to 45,000 reaction tubes. The multiple reaction tubes are referred to as reaction tube bundles, which are generally annular and arranged vertically and surrounded by a reactor shell. Both ends of the reaction tube bundle are sealed in a tube sheet. The feed gas stream is usually introduced at the top of the shell-and-tube reactor via a cover and fed to the reaction tube via a tube sheet. The resulting product gas mixture is discharged at the bottom of the shell-and-tube reactor via the opposite tube sheet and cover. Alternatively, the feed gas stream is introduced at the bottom of the shell-and-tube reactor and leaves the shell-and-tube reactor via an upper tube sheet and a cover.

[0003] It is known that catalysts comprising a catalytically active noble metal such as copper or silver on a suitable support can be used to catalyze certain chemical oxidation reactions. Common catalysts used in fixed beds for prior art processes involve, for example, porous solid-state catalysts impregnated with silver, for example for the oxidation of ethylene to ethylene oxide, or shell-type catalysts coated with silver, for example for the oxidation of primary alcohols to aldehydes, such as the oxidation of isopentenol to pentenal.

[0004] Introducing such catalysts consisting of individual catalyst bodies into reaction tubes and settling the catalyst onto a fixed bed is a complex and time-consuming process, since, for example, fill level measurements are required to verify that a sufficient amount of catalyst is present inside the reaction tubes.

[0005] Vapor phase fixed bed tubular catalytic reactors often exhibit an undesirable temperature profile along the length of the reactor tube. Typically, the temperature profile of an exothermic catalytic reaction is low at the inlet, rises to a maximum and then decreases as the reactant stream is starved of reactants.

[0006] For a given reaction, the reactor temperature outside the optimum temperature range causes lower selectivity owing to the formation of undesirable products. This results in that the organic component of the feed gas stream is deposited on the surface of the active catalyst material in the reaction tubes in the form of carbonaceous deposits (for example, in the form of coke) when the temperature is lower than the reaction ignition temperature. As a result of this deposition, a part of the active catalyst material is deactivated and pressure drop is also increased along with deposition increase. Therefore, it is necessary to regenerate and / or even replace the regular maintenance of the catalyst form. For example, when using the above-mentioned shell-type catalyst coated with silver that isoamyl alcohol is oxidized to amyl aldehyde, it is necessary to carry out a maintenance step in the form of regeneration once a week. As a result of this, the annual operating hours significantly reduce, and the existing production capacity can not be fully utilized.

[0007] Furthermore, in the case of catalytic materials with high thermal conductivity, such as copper or silver, "migratory hot spots" may occur, i.e. hot spots that migrate in the direction of the reaction tube inlet. As a result, the residence time of the target product formed is prolonged, which in turn leads to subsequent reactions of the target product and reduces the selectivity for the target product.

[0008] The occurrence of pronounced and undesirable temperature "humps" or "hot spots" is caused by poor heat transfer between the catalyst and the heat transfer medium in the reactor tubes. In order to alleviate this problem, attempts have been made to precede the catalyst-filled reaction zone of the reactor tubes with a non-catalytic heat exchange zone in which the reactant stream is heated to the reaction start temperature before it comes into contact with the catalyst. In addition, in order to avoid possible further undesirable reactions of the products to a certain extent and to recover the products in a form that is as unchanged as possible, it has been proposed to provide a heat exchange zone downstream of the reaction zone, in which the effluent stream is heat-exchanged with a heat transfer medium and thereby cooled.

[0009] In order to optimize the function of the heat exchange areas used on an industrial scale, various packing materials such as balls or rings, which are present as separate elements, have been recommended as flow obstacles in the heat exchange areas. However, these packing materials are disadvantageous because, on the one hand, they lead to a significant loss of pressure and, in addition, to a rapid deposition of combustion residues due to their high specific surface area. In addition, the heat exchange via the reaction tube wall is not efficient because the reactant stream is forced to pass through the gaps in the packing of the separate element.

[0010] US2007 / 274882 relates to a reactor comprising at least: (a) a reaction zone comprising at least one solid catalyst; and (b) a coolable heat exchanger zone comprising at least one shell at least partially accommodating a plug, wherein the reaction zone and the coolable heat exchanger zone are in fluid communication.

[0011] US2012 / 0277473 describes a method for preparing a catalyst by forming a C1-C 10 Oxidative dehydrogenation of alcohols to produce C1-C 10 The shaped catalyst body can be obtained by three-dimensional shaping and / or spatial arrangement of silver-containing fibers and / or wires. The average diameter or average diagonal length of the substantially rectangular or square cross section of these silver-containing fibers and / or wires is in the range of 30 μm to 200 μm.

[0012] It is therefore an object of the present invention to provide a reactor for carrying out catalytic gas phase partial oxidation reactions which requires less frequent maintenance in the form of regeneration and / or replacement of the catalyst.Another object of the present invention is to facilitate the placement of the catalyst into the reaction tubes and the removal of the catalyst therefrom.

[0013] Therefore, the present invention relates to a shell and tube heat exchange reactor for carrying out catalytic gas phase partial oxidation reaction, which comprises

[0014] - a shell-side heat exchange channel for circulating a heat transfer medium and a reaction channel comprising a plurality of reaction tubes;

[0015] - an inlet for introducing a reactant stream into the reaction channel; and

[0016] - an outlet from the reaction channel for recovering an effluent stream from the reaction tube;

[0017] The reaction tube includes

[0018] a reactant preheating zone adjacent to the inlet, and

[0019] A reaction zone downstream of the reactant preheating zone has a catalytically active wire matrix insert having a catalytically active precious metal on at least a portion of its surface.

[0020] As used herein, "downstream" or "upstream" is relative to the direction of flow of the reactant stream.

[0021] The term "reactant preheating zone" means a section of the reaction tubes, i.e., a section inside the reaction tubes, in which substantially no catalytic gas-phase partial oxidation reaction occurs and in which the gaseous stream passing through the reaction tubes exchanges heat with a circulating heat transfer medium via the tube walls. The preheating zone upstream of the reaction zone involves a net heat flow into the reaction tubes 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.

[0022] Upon contact with the catalytic surface, the oxidation reaction begins immediately. Otherwise, coke formation may occur in the case 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 the coke on the catalytic surface.

[0023] 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.

[0024] 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.

[0025] 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 oxidation 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.

[0026] 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.

[0027] In this article, the term "reaction zone" means a region in a reaction tube where a catalytic gas phase partial oxidation reaction occurs. According to the present invention, the reaction zone includes a catalytically active wire matrix plug-in having 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 a single 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 flow. Due to the unique flow characteristics of the reactant flow through 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 flow 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.

[0028] 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.

[0029] 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.

[0030] 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.

[0031] 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.

[0032] 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.

[0033] The wire loop may be formed from one or more than one wound wire, preferably 4 wound wires.

[0034] The wire matrix insert included in the reaction zone has a catalytically active precious metal on at least a portion of its surface. The wires constituting the wire loop can be coarse precious metal wires or wires coated with precious metals. The core wire can be made of a brass alloy or high-grade steel. The precious metal coating superimposed on the core surface has a thickness of, for example, 10 μm. However, in general, coarse precious metal wires have better service life and are preferred. If the wire loop is formed by more than one wound wire, at least one of the wound wires is made of coarse precious metal wires or wires coated with precious metals, and the other wound wires can be made of inert materials.

[0035] In general, the catalytically active noble metal is selected from copper, silver, palladium, platinum, ruthenium, and rhodium, preferably silver. 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 helically 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.

[0036] 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.

[0037] Suitable structures for wire matrix inserts are known per se. Some inserts of this type are disclosed in GB 2 097 910 and 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.

[0038] In an embodiment, the reaction zone of the reaction tube has a void fraction of 0.60 to 0.99, preferably 0.80 to 0.97, more preferably 0.89 to 0.94. "Void fraction" is defined as the ratio of the void volume (i.e., the total volume of the void spaces within the cylinder enclosing the wire matrix plug) to the total volume occupied by the wire matrix plug (i.e., the volume of the cylinder enclosing the wire matrix plug). In other words, the void fraction is defined as the ratio of the void volume to the sum of the void volume plus the volume occupied by the wire loops and elongated cores that make up the wire matrix plug.

[0039] Due to the higher void fraction of the wire matrix included in the reaction zone compared to a packing of individual elements such as a fixed bed of individual catalyst particles, a larger proportion of the reaction heat is rejected by radiation to the reaction tube walls and does not have to be dissipated by the reactant stream.

[0040] In an embodiment, the catalytically active wire matrix plug is adapted to enable radial mixing of the laminar boundary layer of the reactant stream into the bulk reactant stream passing through the reaction tube. Convective heat transfer generally involves the exchange of thermal energy between the surface and the moving fluid. The deployed wire matrix plug destroys the flow boundary layer near the wall, and the reactant stream forms a relatively weak vortex near the wall, thereby reducing the thermal resistance of the wall fluid. The spiral offset of the wire matrix plug causes the fluid to rotate, so that the fluid flows from the center of the tube to the wall, and again impacts the vortex generated by the wire ring near the wall and mixes with it, thereby enhancing heat transfer.

[0041] Preferably, the catalytically active wire matrix plug-in is adapted to enable the reaction stream flow characterized by Reynolds number being 12000 or less, preferably 8000 or less, more preferably 2300 or less. In this kind of low Reynolds flow state, the pressure drop through the tube caused by the wire matrix plug-in is not a significant problem, because the low speed and relatively low level of turbulence in the flow. Due to these low Reynolds numbers, compared with the common fixed bed filler of a single element (such as a ball or annulus), the heat transfer between the catalytically active wire matrix plug-in and the reaction tube inner wall is better by 3 to 5 times. Under such low Reynolds numbers, the heat transfer via conduction and radiation in the direction to the reaction tube wall is better than convection. Both conduction and radiation are improved by open wire matrix plug-in. Suitably, the region such as reactant preheating zone with a substantially free cross section or with catalytic activity being zero or limited wire matrix plug-in is adapted to enable the reaction stream flow characterized by Reynolds number being 12000 or less, preferably 8000 or less.

[0042] For example, both the active catalyst surface and the mass of catalyst per volume unit can be manipulated via the thickness of the bonded wire. In an embodiment, the ratio of the inner diameter of the reaction tube to the diameter of the wire is in the range of about 10 to 100, preferably about 10 to 50, more preferably about 20 to 40. The wire preferably has a wire diameter in the range of 50 μm to 5000 μm, more preferably 200 μm to 2000 μm.

[0043] Furthermore, mass transfer in the boundary layer around the catalytically active wires can be increased. The mass transfer rate in the boundary layer of such thin wires is higher compared to typically used rings or spheres due to the lower characteristic diameter of the wires used in the wire matrix insert.

[0044] In an embodiment, the ratio of the length of the reaction zone to the length of the reactant preheating zone is in the range of 0.01 to 100, preferably 0.05 to 5, more preferably 0.1 to 1. The ratio allows the reactant preheating zone in the reaction tube to have a suitable length.

[0045] In an embodiment, the reactor tubes include an effluent cooling zone downstream of the reaction zone. The term "effluent cooling zone" refers to a section of the reactor tubes where substantially no catalytic gas phase partial oxidation reaction occurs and where the gaseous stream passing through the reactor tubes exchanges heat with a circulating heat transfer medium via the tube walls. This involves a net heat flow out of the reactor tubes. The hot effluent stream is cooled and the heat transfer medium external to the reactor tubes absorbs the heat dissipated by the effluent stream.

[0046] The effluent cooling zone preferably has a substantially free cross section or has a wire matrix insert with zero or limited catalytic activity. Preferably, the effluent cooling zone comprises a wire matrix insert. A cooling zone downstream of the reaction zone has the benefit of avoiding continuous reactions like overoxidation to carbon oxides.

[0047] The characteristic design of the shell and tube heat exchange reactor is known per se to the technician. The shell and tube heat exchange reactor is limited by a "reactor shell" which is appropriately cylindrical, and is limited by an "upper cover" and a "lower cover" on the upper and lower ends of the reactor shell, wherein the upper cover and the lower cover are connected to the reactor shell in a gas-tight manner. Inside the shell and tube heat exchange reactor, a plurality of vertically arranged "reaction tubes" exist in such a way that the reactor shell surrounds a plurality of reaction tubes. Generally, the term "multiple" reaction tubes refers to a large number of reaction tubes present inside the shell and tube heat exchange reactor, for example at least 5,000 and up to 45,000 reaction tubes. In a preferred embodiment, the reaction tube has an inner diameter preferably in the range of 0.10 cm to 5.0 cm; more preferably in the range of 0.50 cm to 3.0 cm. Preferably, the reaction tube has a length of at least 5 cm, preferably in the range of 10 to 100 cm, especially in the range of 25 to 60 cm. The upper end of the reaction tube is connected to the "upper tube sheet", and the lower end of the reaction tube is connected to the "lower tube sheet", each connected in a gas-tight manner. In other words, both ends of the reaction tube are sealed in the tube sheet. Thus, a gas-tight region is formed by the space between the upper cover and the upper tube sheet, the interior of the reaction tube, and the space between the lower tube sheet and the lower cover. In said region, the feed gas mixture is introduced into the shell and tube heat exchange reactor, is subjected to the chemical reaction envisaged in the shell and tube heat exchange reactor, such as catalytic gas phase partial oxidation reaction, inside the reaction tube, and is then removed from the shell and tube heat exchange reactor.

[0048] During the "operation mode" of the shell and tube heat exchange reactor, the reaction envisaged in the shell and tube heat exchange reactor, for example a catalytic gas phase partial oxidation reaction, is carried out. The reaction is carried out at a reaction temperature, which is usually an elevated temperature.

[0049] Typically, in order to control the temperature of the shell and tube heat exchange reactor and / or provide an elevated temperature, a heat exchange medium is circulated in the liquid-tight region between the upper tube sheet, the lower tube sheet, the interior of the reactor shell and the exterior of the reaction tube. The heat exchange medium may be, for example, a low melting point metal such as sodium or mercury or an alloy of different metals, or a liquefied salt melt of a eutectic mixture containing a nitrate portion, such as a mixture of at least two salts selected from alkali metal nitrates, alkali metal nitrites and alkali metal carbonates, preferably a mixture of two or three of the salts potassium nitrate, sodium nitrate and sodium nitrite.

[0050] It is very important to provide stable reaction conditions, i.e., to react at a constant temperature. In order to provide such stable reaction conditions, a pump is used to circulate the heat transfer medium appropriately through the liquid-tight area in the overall longitudinal direction at the expected temperature. Doing so enables the reaction tubes to be cooled or heated, and then the heat transfer medium is cooled or heated in a heat exchanger (e.g., an externally arranged heat exchanger). In general, for this purpose, the reaction tubes are arranged in a shell and tube heat exchange reactor so that they are equidistant from each other.

[0051] Rapid cooling of the effluent stream emerging from the outlet from the reaction channel is highly desirable. This can be achieved by quenching the effluent stream with an aqueous phase.

[0052] The shell and tube heat exchange reactor of the present invention is particularly, but not exclusively, suitable for carrying out catalytic gas phase partial oxidation reactions, such as the oxidation of alcohols to aldehydes. Therefore, the present invention further relates to the use of the shell and tube heat exchange reactor of the present invention for carrying out catalytic gas phase partial oxidation reactions, such as the production of aldehydes.

[0053] The present invention further relates to a method for carrying out a catalytic gas phase partial oxidation reaction, the method comprising introducing a reactant stream into the inlet of a shell and tube heat exchange reactor as described above. In the method, the reactant stream comprises a partially oxidizable organic substrate and molecular oxygen (e.g. in the form of air).

[0054] In an embodiment, the noble metal is silver and the partially oxidizable organic substrate is ethylene which is catalytically oxidized to ethylene oxide.

[0055] In an embodiment, the noble metal is silver and the partially oxidizable organic substrate is methanol which is catalytically oxidized to formaldehyde.

[0056] In an embodiment, the noble metal is silver and the partially oxidizable organic substrate is an alcohol that is oxidized to an aldehyde.

[0057] Various ethylenically unsaturated aldehydes can be produced by the present invention.In one embodiment, ethylenically unsaturated alcohol is 3-methylbut-2-ene-1-ol (pentenol), and the ethylenically unsaturated aldehyde produced is pentenal (3-methyl-2-butene-1-aldehyde), or ethylenically unsaturated alcohol is 3-methylbut-3-ene-1-ol (isopentenol), and the ethylenically unsaturated aldehyde produced is 3-methyl-3-butene-1-aldehyde (isopentenal).Carrying out isopentenol dehydration to isopentenal may be followed by partial isomerization to become pentenal.

[0058] Typically, the method comprises the following steps:

[0059] a) vaporizing (iso)prenol;

[0060] b) mixing the (iso)prenol vapor with an oxygen-containing gas; and

[0061] c) introducing the reactant stream into the inlet of a shell and tube heat exchange reactor as described above and reacting the gaseous mixture to form (iso)amylaldehyde.

[0062] For example, in the case of the catalytic gas phase partial oxidation reaction of (iso)prenol to (iso)prenaldehyde, the reaction temperature is in the range of 300°C to 500°C, preferably 350°C to 400°C.

[0063] Prenol useful as a starting material for the present invention can be obtained by reacting a formaldehyde source and isobutylene in a reactor generally at elevated temperature and pressure to obtain 3-methylbut-3-en-1-ol (isopentenol), and subjecting the obtained isopentenol to isomerization.

[0064] In one embodiment, the prenol is obtained by mixing a formaldehyde source and isobutylene and injecting them into a reactor through at least one nozzle and reacting the formaldehyde source and isobutylene under supercritical conditions. To achieve supercritical conditions, the 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 a catalyst such as an amine base, for example hexamethylenetetramine (urotropine).

[0065]

[0026] In this document, it is to be understood that any reference to "a catalyst" includes the possibility that the catalyst is a single catalytic species or a combination of two or more different catalytic species.

[0066] The formaldehyde may be provided as a liquid, for example as a solution of paraformaldehyde. Preferably, the formaldehyde source is an aqueous formaldehyde solution. The formaldehyde source may be a single formaldehyde source or a combination of two or more different formaldehyde sources.

[0067] 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.

[0068] 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.

[0069] Further details about reacting a formaldehyde source and isobutylene to obtain prenol can be found in WO 2020 / 049111 A1.

[0070] In one embodiment, reacting the formaldehyde source and the isobutylene comprises mixing the formaldehyde source and the isobutylene and injecting them through at least one nozzle into the one or more first conduits into an internal loop reactor comprising:

[0071] - a vertically arranged cylindrical container comprising a side wall;

[0072] - at least one flow guide cylinder having a cylinder inlet end and a cylinder outlet end arranged vertically in 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 in the one or more flow guide cylinders, and second conduits outside the one or more flow guide cylinders and in the sidewalls, the one or more first conduits being in fluid communication with the second conduits;

[0073] - reactor fluid outlet means;

[0074] 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

[0075] wherein the outer surface of one or more flow guide sleeves is convexly curved, so that 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 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

[0076] The edges of one or more draft tubes are rounded so that the 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 is turned in the opposite direction to travel through a second conduit and subsequently backmixed with the injected fluid.

[0077] 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.

[0078] 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.

[0079] 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.

[0080] 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 the 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 the formaldehyde source and the reaction mixture in the reactor, thereby allowing advantageous rapid mixing of the reactants.

[0081] 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.

[0082] 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.

[0083] 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.

[0084] 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.

[0085] Advantageously, reacting the formaldehyde source and the isobutylene preferably comprises heat exchanging a hot isopentanol stream withdrawn 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 the 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 exchanger; 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.

[0086] 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 a heat exchanger 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. One aspect of the present invention is that the preheating of the isobutylene stream can be maintained at a sufficiently high level for an extended period of time so that the desired temperature of the isobutylene can be readily achieved before it enters the reactor.

[0087] 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.

[0088] 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.

[0089] 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.

[0090] 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.

[0091] 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.

[0092] 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.

[0093] 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.

[0094] 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.

[0095] Therefore, in a more preferred embodiment, the purification method comprises

[0096] (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;

[0097] (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

[0098] (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.

[0099] The second distillate stream constitutes concentrated aqueous formaldehyde suitable for recycling to the prenol synthesis.

[0100] 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.

[0101] 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.

[0102] Further information on methods for recovering essentially formaldehyde-free prenyl alcohol can be found in WO 2022 / 189652 A1.

[0103] The prenol obtained can be subjected to catalytic isomerization in order to obtain prenol.

[0104] 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.

[0105] The (iso)prenol obtained as described above can be subjected to the catalytic gas phase partial oxidation reaction of the present invention. Before contacting with the catalytically active wire matrix insert, the (iso)prenol can 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 can be depleted of organically bound nitrogen by this method.

[0106] 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.

[0107] 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.

[0108] 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.

[0109] 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.

[0110] 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.

[0111] 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.

[0112] 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.

[0113] 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.

[0114] 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.

[0115] 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.

[0116] 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.

[0117] 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.

[0118] When subjecting isopentenol to the catalytic gas-phase partial oxidation reaction of the present invention, 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.

[0119] 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.

[0120] 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.

[0121] 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.

[0122] 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 pollution in the reactant 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 reactant 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.

[0123] According to the present invention, reducing the weight ratio of formaldehyde to prenol in the reactant stream can be achieved in several different ways. In an embodiment, formaldehyde is removed from the unreacted prenol stream before it is combined with the fresh feed stream.

[0124] In an example, an unreacted isopentenol stream is combined with a fresh feed stream and formaldehyde is removed from the combined stream.

[0125] 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.

[0126] In an embodiment, the fresh feed stream comprising prenol is derived from a process for reacting isobutylene and formaldehyde and is purified to a weight ratio of formaldehyde to prenol of less than 0.04, preferably less than 0.03, in particular less than 0.02, or less than 0.01. In a still more preferred embodiment, the fresh feed stream is purified to a weight ratio of formaldehyde to prenol of less than 0.002, or less than 0.001.

[0127] 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.

[0128] The pentenals produced by the present invention are useful intermediates in the preparation of citral. Citral is a mixture of the isomeric compounds neral and geranal.

[0129] 3,7-Dimethyl-octa-2,6-dienal (citral) can be prepared by obtaining pentenal by the method as described above, which further includes the steps of: condensing pentenal with pentenol to obtain the diprenyl acetal of pentenal; and subjecting the diprenyl acetal of pentenal to cleavage conditions to obtain citral via pentenyl (3-methyl-butadienyl) ether and 2,4,4-trimethyl-3-formyl-1,5-hexadiene.

[0130] In particular, 3,7-dimethyl-octa-2,6-dienal (citral) can be prepared by a method comprising the following steps:

[0131] a) condensing pentenals with pentenols in a reaction column in the presence of at least one catalyst, while simultaneously removing an acetal fraction comprising the dipentenyl acetal of pentenals from the reaction column;

[0132] b) subjecting the acetal fraction to cracking conditions in a cracking column in the presence of at least one catalyst, while withdrawing from the cracking column a cracked fraction containing at least one of pentenyl (3-methyl-butadienyl) ether and 2,4,4-trimethyl-3-formyl-1,5-hexadiene and optionally citral; and

[0133] c) reacting the cleavage fraction in a plug flow type reactor to obtain citral.

[0134] The overall reaction sequence is illustrated by the following reaction scheme.

[0135]

[0136] In step a), 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 acid, and the water formed during the reaction is separated in a reaction column. In step b), the resulting 3-methyl-2-butenal dipentenyl acetal (dipentenyl acetal) of step a) is cracked in a cracking column 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 pentenyl (3-methylbutadienyl) ether obtained gives 2,4,4-trimethyl-3-formyl-1,5-hexadiene, which subsequently undergoes a Cope rearrangement to give 3,7-dimethyl-2,6-octadienal (citral).

[0137] Step a) is carried out in the presence of a catalyst, preferably an acid. In an embodiment, the catalyst in step a) is nitric acid.

[0138] Preferably, in step b), the acetal fraction is continuously subjected to cleavage conditions in a cleavage column. "Cleavage conditions" means reaction conditions selected such 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).

[0139] The acetal fraction comprises dipentenyl acetal as main component.The acetal fraction does not necessarily need to consist of pure dipentenyl acetal but may also comprise pentenol, pentenal and citral building blocks.

[0140] Step b) is carried out in the presence of a catalyst, preferably an acid catalyst. Suitable acid catalysts are selected from non-volatile proton acids such as sulfuric acid, p-toluenesulfonic acid and phosphoric acid.

[0141] 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.

[0142] 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.

[0143] 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).

[0144] 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.

[0145] The cleavage fraction together with the pentenols formed can be discharged at the top of the distillation column.

[0146] Alternatively and preferably, the cracked fraction can also be discharged in liquid or vaporous form at a side draw of the distillation column.

[0147] In step c), the cracked fraction is reacted in a plug flow reactor to obtain citral. To this end, the cracked fraction is directed through a plug flow reactor at a suitable temperature for performing one or more rearrangement reactions to obtain citral. By adopting a combination of a highly back-mixed cracking tower and a plug flow reactor, the selectivity and yield of the cracking reaction can be increased. In step b), preferably all catalysts required for the cracking reaction are introduced into the cracking tower, and preferably, no catalyst is introduced into the plug flow reactor.

[0148] In an embodiment, the pentenol removed in step b) is recycled to step a). This allows increased yields to be achieved in the process of the invention.

[0149] Therefore, in one aspect, the present invention relates to an improved process for preparing citral (3,7-dimethyl-octa-2,6-dienal), comprising the following steps

[0150] A) reacting a formaldehyde source with isobutylene to obtain 3-methylbut-3-en-1-ol (isopentenol), and subjecting at least a portion of the isopentenol obtained to isomerization to obtain prenol;

[0151] B) preparing pentenal (3-methylbut-2-en-1-al) from 3-methylbut-2-en-1-ol (pentenol) using the method described above, and / or preparing pentenal via isopentenal (3-methylbut-3-en-1-al) prepared from isopentenol using the method described above; and

[0152] C) condensing pentenal with pentenol to obtain dipentenyl acetal of pentenal; and subjecting the dipentenyl acetal of pentenal to cleavage conditions to obtain citral via pentenyl (3-methyl-butadienyl) ether and 2,4,4-trimethyl-3-formyl-1,5-hexadiene.

[0153] Step A can be carried out as described above or by other methods known in the art, preferably via distillation at a temperature at which hemiformal is cracked into formaldehyde and prenol so that formaldehyde can be easily separated from prenol, and more preferably by subjecting a crude prenol stream containing prenol, water and formaldehyde or a prenol containing fractions thereof to distillation in a low-boiler separation column operated at a pressure of 2 bar absolute or higher, preferably 2.5 bar absolute or higher, to obtain a distillate stream containing aqueous formaldehyde and a bottom stream containing prenol substantially free of formaldehyde.

[0154] Step B comprises the oxidative dehydrogenation of prenol and / or isopentenol.Isopentenol is converted with a catalytically active metal catalyst to form a reaction mixture of 3-methylbut-3-ene-1-aldehyde and 3-methylbut-2-ene-1-aldehyde.Then, the former isomer can be isomerized under base catalysis to give the desired 3-methylbut-2-ene-1-aldehyde.

[0155] Step C can be carried out as described above, for example via steps a) to c).

[0156] The citral thus obtained is a useful intermediate for example for menthol or linalool.

[0157] Menthol can be prepared from citral via a process comprising the following steps

[0158] - catalytic hydrogenation of citral to obtain citronellal;

[0159] - cyclizing citronellal in the presence of an acidic catalyst to obtain isopulegol; and

[0160] - Catalytic hydrogenation of isopulegol to obtain menthol.

[0161] The overall reaction sequence is illustrated by the following reaction scheme.

[0162]

[0163] The hydrogenation of citral to obtain citronellal can be achieved by carrying out the hydrogenation in the presence of a rhodium-phosphine catalyst.

[0164] The cyclization of citronellal to isopulegol can be achieved by carrying out the cyclization in the presence of a 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 reaction product containing the catalyst by distillation separation to give an isopulegol-rich overhead product and an isopulegol-depleted bottom product. The 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.

[0165] The hydrogenation of isopulegol can be achieved by carrying out the hydrogenation in the presence of a heterogeneous nickel-containing catalyst, preferably a heterogeneous nickel- and copper-containing catalyst.

[0166] Further details on the reaction sequence from citral to menthol can be found in US 2013 / 46118 A1 (which is incorporated herein by reference).

[0167] 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.

[0168] Linalool can be prepared from citral via a process comprising catalytic hydrogenation of citral to obtain nerol and / or geraniol and isomerization thereof.

[0169] The hydrogenation of citral to obtain nerol and / or geraniol can be achieved by hydrogenation in the presence of a supported ruthenium, rhodium, osmium, iridium or platinum catalyst, preferably a ruthenium catalyst supported on carbon black.

[0170] Isomerization of nerol and / or geraniol to obtain linalool can be achieved by carrying out the isomerization in the presence of a tungsten catalyst, in particular a tungsten(VI) dioxide complex. Further details on the isomerization of nerol and / or geraniol can be found in US 7,126,033 B2.

[0171] 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.

[0172] The present invention is further illustrated by the following examples and drawings. However, it will be understood that these examples and drawings are not intended to limit the scope of the present invention in any way.

[0173] Figure 1 Plotted are the pressure drop versus time for three different catalytically active structures.

[0174] Examples

[0175] Pentenols are continuously vaporized in a double-pipe vaporizer. Pentenol vapor is introduced into the reaction tube at the bottom of the reaction tube at a flow rate of 300g / h, a temperature of 365°C and a pressure of 1atm. Together with the pentenol vapor, air is introduced into the reaction tube at the bottom of the reaction tube at a flow rate of 100Nl / h. The reaction tube has an internal diameter of 12mm and a length of 500mm. The effluent stream is reclaimed at the top of the reaction tube and analyzed by gas chromatography.

[0176] Testing different catalytically active structures in the reaction tube:

[0177] - a wire matrix insert made of thick silver wire (available from Calgavin) having a void fraction of 90.6% and a length of 300 mm (Insert 1),

[0178] - a wire matrix insert made of thick silver wire having a void fraction of 90.6% and a length of 150 mm (available from Calgarvin Corporation) (Insert 2), or

[0179] - A packing of silver-coated steatite catalyst with a diameter of 2 mm.

[0180] The wire matrix insert was placed into the reaction tube so that one end of the wire matrix insert was located at the outlet of the reaction tube, creating a reactant preheating zone of 200 mm (in the case of insert 1) or 350 mm (insert 2) in length. The silver-coated steatite catalyst was placed at the outlet of the reaction tube so that it filled the 300 mm reaction tube, creating a preheating zone of 200 mm in length.

[0181] Three runs were performed for each of Insert 1 and Insert 2 (inventive examples) and for the silver-coated steatite catalyst (comparative example). The results are shown in Tables 1 to 3 and Figure 1 middle.

[0182] The normalized selectivities (selectivity of the silver-coated steatite catalyst = 100%) are shown in Table 1.

[0183] Table 1.

[0184]

[0185] *Comparative Examples

[0186] The conversion of pentenols and the selectivity of pentenals are shown in Table 2.

[0187] Table 2.

[0188]

[0189] *Comparative Examples

[0190] The amounts of gaseous products formed and the average selectivity to pentenals are shown in Table 3.

[0191] Table 3.

[0192] catalyst <![CDATA[Liquid mass loss [g / h] [1] > Silver coated steatite catalyst* 53 Plugin 1 60 Plugin 2 48

[0193] [1] Calculated as the flow rate of liquid feed in g / h minus the flow rate of liquid product in g / h

[0194] *Comparative Examples

[0195] Liquid mass loss is a measure of side reactions that produce gaseous products such as peroxidation to CO and / or CO2.

[0196] Figure 1 Pressure drop versus reaction time is plotted for three different catalytically active structures: Insert 1 and Insert 2 (inventive examples) and a silver-coated steatite catalyst (comparative example). Figure 1 Also included in the run at 350° C. with insert 2. At 350° C., no reaction occurs and, therefore, this run illustrates the lowest possible pressure drop, with no deposits or coke formation at all.

[0197] from Figure 1 It can be seen that the pressure drop for the examples according to the invention is lower (almost 2 times lower) than the pressure drop when using the silver coated steatite catalyst (comparative example). This advantageously results in longer operating time and less downtime per year.

Claims

1. A shell and tube heat exchange reactor for catalytic gas phase partial oxidation reaction, 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 a catalytically active precious metal on at least a portion of its surface.

2. The shell and tube heat exchange reactor according to claim 1, wherein: The reactant preheating zone has a substantially free cross section or has a wire matrix insert having zero or limited catalytic activity.

3. The shell and tube heat exchange reactor according to claim 1 or 2, wherein: The ratio of the length of the reaction zone to the length of the reactant preheating zone is in the range of 0.01 to 100, preferably 0.05 to 5, more preferably 0.1 to 1.

4. A shell and tube heat exchange reactor as claimed in any one of the preceding claims, wherein: The reaction zone comprises a series of alternating regions having catalytically active wire matrix inserts and regions having essentially free cross sections or having wire matrix inserts having zero or limited catalytic activity.

5. A shell and tube heat exchange reactor as claimed in any one of the preceding claims, wherein: The reaction tubes comprise an effluent cooling zone downstream of the reaction zone, wherein the effluent cooling zone has a substantially free cross section or has a wire matrix insert with zero or limited catalytic activity.

6. A shell and tube heat exchange reactor as claimed in any one of the preceding claims, wherein: The catalytically active noble metal is selected from the group consisting of copper, silver, palladium, platinum, ruthenium, and rhodium, preferably silver.

7. The shell and tube heat exchange reactor according to any one of claims 2 to 6, wherein: The wire matrix insert having zero or limited catalytic activity is made of an inert material, preferably stainless steel.

8. A shell and tube heat exchange reactor as claimed in any one of the preceding claims, wherein: 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 coarse precious metal wires, or wires coated with precious metals.

9. The shell and tube heat exchange reactor according to claim 8, wherein: The elongated core includes at least two longitudinal core wire members twisted around each other to form core wire windings, and the wire loops are received in the core wire windings.

10. The shell and tube heat exchange reactor according to claim 8 or 9, wherein: The ratio of the inner diameter of the reaction tube to the diameter of the crude noble metal wire or the noble metal-coated wire is in the range of about 10-100, preferably about 10-50, more preferably about 20-40.

11. A shell and tube heat exchange reactor as claimed in any one of the preceding claims, wherein: The reaction zone has a void fraction of 0.60 to 0.99, preferably 0.80 to 0.97, more preferably 0.89 to 0.

94.

12. A shell and tube heat exchange reactor as claimed in any one of the preceding claims, wherein: The catalytically active wire matrix insert is adapted to enable radial mixing of the laminar boundary layer of the reactant flow into the bulk reactant flow passing through the reaction tubes.

13. A process for carrying out a catalytic gas phase partial oxidation reaction, the process comprising introducing a reactant stream comprising a partially oxidizable organic substrate and molecular oxygen into the inlet of a shell and tube heat exchange reactor as claimed in any one of claims 1 to 12.

14. The method of claim 13, wherein: The flow of the reactant stream inside the preheat zone is substantially laminar.

15. The method according to claim 13 or 14, wherein: The flow of the reactant stream inside the reaction zone including the catalytically active wire matrix insert is characterized by a Reynolds number of 12,000 or less, preferably 8,000 or less, more preferably 2,300 or less.

16. The method according to any one of claims 13 to 15, for producing an aldehyde, wherein: The noble metal is silver and the partially oxidizable organic substrate is an alcohol, preferably 3-methylbut-2-en-1-ol (pentenol) or 3-methylbut-3-en-1-ol (isopentenol); The prenol is optionally obtained by isomerization of prenol.

17. The method of claim 16, wherein: The prenol is obtained by reacting 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 α and β: a) mixing the at least one formaldehyde source and isobutylene and injecting the at least one formaldehyde source and isobutylene into one or more first conduits through at least one nozzle into an 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; β) 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.

18. The method according to claim 16 or 17, wherein: The partially oxidizable organic substrate is prenol, wherein the process further comprises at least one of αα, ββ, and γγ: 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; γγ) treating the (iso)prenol by contacting the (iso)prenol with a weakly acidic solid adsorbent to remove organically bound nitrogen from the (iso)prenol prior to contacting with the catalytically active wire matrix insert.

19. A method for preparing 3,7-dimethyl-octa-2,6-dienal (citral), the method comprising obtaining pentenal by the method according to any one of claims 16 to 18, further comprising the following steps: condensing the pentenal with pentenol to obtain a dipentenyl acetal of pentenal; and subjecting the dipentenyl acetal of pentenal to cleavage conditions to obtain citral via pentenyl (3-methyl-butadienyl) ether and 2,4,4-trimethyl-3-formyl-1,5-hexadiene.

20. A method for preparing citral-derived chemicals, the method comprising preparing citral by the method of claim 19, 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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