WO2009146794A1

WO2009146794A1 - Method and device for producing cyclohexanone

Info

Publication number: WO2009146794A1
Application number: PCT/EP2009/003548
Authority: WO
Inventors: Evin Hizaler Hoffmann; Leslaw Mleczko; Ralph Schellen; Stephan Schubert; Sha PENG
Original assignee: Bayer Technology Services Gmbh
Priority date: 2008-05-29
Filing date: 2009-05-19
Publication date: 2009-12-10
Also published as: DE102008025842A1

Abstract

The present invention relates to a method for the production of cyclohexanone by catalytic hydration of phenol with hydrogen, wherein the reaction is carried out on 5 to 30 serially connected catalyst beds under adiabatic conditions, and a reactor system for carrying out the method.

Description

Process and apparatus for the production of cyclohexanone

The present invention relates to a process for the preparation of cyclohexanone by catalytic hydrogenation of phenol with hydrogen, wherein the reaction is carried out on 5 to 30 successive catalyst beds under adiabatic conditions, and a reactor system for carrying out the method.

Cyclohexanone is generally prepared under the catalytic influence of palladium supported on aluminum with calcium oxide from gaseous phenol and gaseous hydrogen in an exothermic, catalytic equilibrium reaction according to formula (I):

C ₆ H ₆ O (g) + 2 H ₂ (g) →C ₆ Hi ₀ O (g); ΔH = -57.8 kJ / mol (I)

The cyclohexanone prepared by the reaction of formula (I) forms an essential starting material for the production of caprolactam and synthesis of perlon.

The removal and use of the heat of reaction is an important issue in carrying out the cyclohexanone synthesis. An uncontrolled increase in temperature can lead to permanent damage to the catalyst, on the other hand, it comes at high temperatures to an unfavorable shift of the reaction equilibrium in the direction of the starting materials with a corresponding deterioration of the yield. In addition, at high temperatures, there is the potential for side reactions to produce more or less large amounts of, inter alia, cyclohexanol as an impurity. It is therefore advantageous to maintain the temperature of the catalyst bed controlled in the course of the process at a level which allows a rapid conversion with minimization of side reactions and / or catalyst deactivation.

EP 0 602 499 (B1) discloses a process in which phenol is hydrogenated to cyclhexanone in the gas phase in the presence of a palladium catalyst. The disclosed temperatures and pressures for carrying out the process are from 150 ⁰ C to 250 ° C and from 0.8 to 8 bar. It is further disclosed that an excess of 3 to 30 mol / mol of hydrogen based on phenol can be used. The catalysts used can be applied to various oxidic support materials. The process is further characterized by a special treatment of the catalysts characterized. It is always disclosed only a one-step process at a fixed temperature. An adaibate mode of operation is not disclosed.

The process disclosed in EP 0 602 499 (B1) is disadvantageous because it requires a complicated regeneration of the catalysts by special treatment. This in turn results in downtime of the process in which no conversion to cyclohexanone takes place. Furthermore, by operating at fixed temperatures (isothermal operation) in only one reaction zone, it is not possible to introduce the reaction into its equilibrium limitation as a function of the process gas composition. In addition, the released in the reaction energy in the form of heat is not used, which is economically unfavorable. In particular, the heat is not used for advantageous control of the reaction temperature, so that the method is not capable of ensuring optimum conversion with high energy efficiency.

EP 1 251 951 (B1) discloses a device and the possibility of carrying out chemical reactions in the device, wherein the device is characterized by a cascade of reaction zones in contact with one another and heat exchanger devices which are arranged in a composite with one another. The method to be carried out here is thus characterized by the contact of the various reaction zones with a respective heat exchanger device in the form of a cascade. There is no disclosure regarding the utility of the apparatus and method for the synthesis of cyclohexanone from gaseous hydrogen and phenol. It thus remains unclear how, starting from the disclosure of EP 1 251 951 (B1), such a reaction is to be carried out by means of the device and the method carried out therein. Further, for the sake of uniformity, it is to be understood that the method disclosed in EP 1 251 951 (B1) is carried out in a device the same as or similar to the disclosure regarding the device. As a result, by the large-area contact of the heat exchange zones according to the disclosure with the reaction zones, a significant amount of heat takes place by heat conduction between the reaction zones and the adjacent heat exchange zones. The disclosure with regard to the oscillating temperature profile can therefore only be understood as meaning that the temperature peaks ascertained here would be stronger if this contact did not exist. Another indication of this is the exponential increase in the disclosed temperature profiles between the individual Temperature peaks. These indicate that there is some heat sink of appreciable but limited capacity in each reaction zone which can reduce the temperature rise in it. It can never be ruled out that a certain dissipation of heat (eg by radiation) takes place, but a reduction in the possible heat removal from the reaction zone would indicate a linear or degressive temperature gradient, since at higher temperatures the equilibrium of exothermic reactions shifted to the side of the starting materials and thus reduce the reaction rate and with it the generated heat of reaction. Thus, EP 1 251 951 (B1) discloses multi-stage processes in cascades of reaction zones from which heat in an undefined amount is removed by heat conduction. Accordingly, the disclosed method is disadvantageous in that accurate temperature control of the process gases of the reaction is not possible.

It would therefore be advantageous to provide a process which can be carried out in simple reaction devices and which allows accurate, simple temperature control so that it allows high conversions with the highest possible purity of the product. Such simple reaction devices would be easily scaled up and are inexpensive and robust in all sizes.

For the catalytic gas-phase oxidation of phenol with hydrogen to cyclohexanone, as described above, neither suitable reactors have yet been described, nor are suitable processes disclosed which permit this.

It is therefore an object to provide a process for the catalytic hydrogenation of phenol with hydrogen to cyclohexanone, which is feasible under precise temperature control in a simple reaction apparatus and thereby high conversions at high purities of the product allowed, the heat of reaction, for example, to set an advantageous temperature levels can be used for the reaction or steam generation.

It has surprisingly been found that a process for the preparation of cyclohexanone, which is characterized in that it has a conversion of phenol with hydrogen to cyclohexanone according to formula (I)

C ₆ H ₆ O (g) + 2 H ₂ (g) <→ C ₆ H ₁₀ O (g); ΔH = -57.8 kJ / mol (I) in 5 to 30 successive reaction zones under adiabatic conditions in the presence of heterogeneous catalysts, this task can be solved.

Phenol, in the context of the present invention, refers to a process gas which is introduced into the process of the invention and comprises phenol. Usually, the process gas comprises a proportion of between 90 and 100 mol%, preferably between 95 and 100 mol% of phenol.

Hydrogen, in the context of the present invention, denotes a process gas which is introduced into the process according to the invention and comprises hydrogen. Usually, the process gas comprises a proportion of between 90 and 100 mol%, preferably between 95 and 100 mol% of hydrogen.

In addition to the essential components of the process gases hydrogen and phenol, these may also include secondary components. Non-exhaustive examples of minor components which may be included in the process gases include nitrogen, carbon dioxide, water and cyclohexanol.

In general, in connection with the present invention, process gases are understood as gas mixtures which comprise hydrogen and / or phenol and / or cyclohexanone and / or secondary components. Essentially, however, process gases include hydrogen and / or phenol and / or cyclohexanone.

According to the invention, carrying out the process under adiabatic conditions means that the reaction zone does not receive any active heat from the outside, either essentially

Heat is withdrawn. It is well known that full isolation against

Heat supply or removal only by complete evacuation to the exclusion of

Possibility of heat transfer by radiation is possible. Therefore, in the context of the present invention, adiabat means that no heat supply or removal measures are taken.

In an alternative embodiment of the method according to the invention, however, a heat transfer can be reduced, for example, by insulation by means of generally known insulation means, such as polystyrene insulating materials, or by sufficiently large distances to heat sinks or heat sources, the insulation means being air. An advantage of the adiabatic driving method according to the invention of the 5 to 30 reaction zones connected in series with respect to a non-adiabatic mode of operation is that no means for heat removal must be provided in the reaction zones, which entails a considerable simplification of the construction. This results in particular simplifications in the manufacture of the reactor and in the scalability of the process and an increase in reaction conversions. In addition, the heat generated in the course of the exothermic reaction progress can be utilized in the individual reaction zone to increase the conversion in a controlled manner by allowing the process gases and the reaction zone to increase in temperature to near equilibrium limitation during the passage of the reaction zone.

Another advantage of the method according to the invention is the possibility of very accurate temperature control by the close staggering of adiabatic reaction zones. It can thus be adjusted in that reaction zone advantageous in the reaction progress temperature.

The catalysts used in the process according to the invention are usually catalysts which consist of a material which, in addition to its catalytic activity for the reaction of the formula (I), is characterized by sufficient thermal stability and by a high specific surface area. Catalyst materials characterized by such thermal stability and high surface area are, for example, noble metal catalysts supported on oxidic materials.

Preferred is palladium on an alumina carrier which also comprises calcium oxide.

Specific surface area in the context of the present invention refers to the area of the catalyst material that can be reached by the process gases, based on the mass of catalyst material used.

A high specific surface area is a specific surface area of at least 100 m / g, preferably of at least 200 m ² / g.

The catalysts of the invention are each in the reaction zones and can be used in all known forms, e.g. Fixed bed, fluidized bed, fluidized bed, are present.

Preferred are the forms of fixed bed and fluidized bed. The fixed bed arrangement comprises a catalyst bed in the strict sense, ie loose, supported or unsupported catalyst in any form and in the form of suitable packings. The term catalyst bed as used herein also includes contiguous areas of suitable packages on a carrier material or structured catalyst supports. These would be, for example, to be coated ceramic honeycomb carrier with comparatively high geometric surfaces or corrugated layers of metal wire mesh on which, for example, catalyst granules is immobilized. As a special form of packing, in the context of the present invention, the presence of the catalyst in monolithic form is considered.

If a fixed-bed arrangement of the catalyst is used, the catalyst is preferably present in beds of particles with average particle sizes of 1 to 10 mm, preferably 1, 5 to 8 mm, particularly preferably 2 to 5 mm.

Also preferably, the catalyst is in a fixed bed arrangement in monolithic form. In the case of a fixed bed arrangement, particular preference is given to a monolithic catalyst which consists of the support of aluminum oxide and is coated with palladium and calcium oxide.

If a fluidized bed arrangement of the catalyst is used, the catalyst is preferably present in loose beds of particles, as have also previously been described for the fixed bed arrangement.

Beds of such particles are advantageous because the particles have a high outer surface of the catalyst material compared to the process gases hydrogen and phenol due to their size and thus a high conversion rate can be achieved. Thus, the mass transport limitation of the reaction by diffusion can be kept low. At the same time, however, the particles are not yet so small that disproportionately high pressure losses occur when the fixed bed flows through. The ranges of the particle sizes given in the preferred embodiment of the process, comprising a reaction in a fixed bed, are thus an optimum between the achievable conversion from the reaction according to formula (I) and the pressure drop produced when carrying out the process. Pressure loss is directly coupled with the necessary energy in the form of compressor performance, so that a disproportionate increase in the same would result in an inefficient operation of the process. In a preferred embodiment of the process according to the invention, the conversion is carried out at 6 to 25, more preferably 7 to 15 reaction zones connected in series.

A preferred further embodiment of the method is characterized in that the process gas emerging from at least one reaction zone is subsequently passed through at least one heat exchange zone downstream of said reaction zone.

In a particularly preferred further embodiment of the process, after each reaction zone is at least one, preferably exactly one heat exchange zone, through which the process gas leaving the reaction zone is passed.

The reaction zones can either be arranged in a reactor or arranged divided into several reactors. The arrangement of the reaction zones in a reactor leads to a reduction in the number of apparatuses used.

The individual reaction zones and heat exchange zones can also be arranged together in a reactor or in any combination of reaction zones with heat exchange zones in several reactors.

If reaction zones and heat exchange zones are present in a reactor, then in an alternative embodiment of the invention there is a heat insulation zone between them, in order to be able to support the adiabatic operation of the reaction zone.

In addition, each of the series-connected reaction zones can be replaced or supplemented independently of one another by one or more reaction zones connected in parallel. The use of reaction zones connected in parallel allows in particular their replacement or supplementation during ongoing continuous operation of the process.

Parallel and successive reaction zones may in particular also be combined with one another. However, the process according to the invention particularly preferably has exclusively reaction zones connected in series.

The reactors which are preferably used in the process according to the invention can consist of simple containers with one or more reaction zones, as described, for example, in Ulimann's Encyclopedia of Industrial Chemistry (Fifth, Completely Revised Edition, Vol B4, pages 95-104, page 210-216), wherein in each case between the individual reaction zones and / or heat exchange zones heat insulation zones can be additionally provided.

In an alternative embodiment of the method, there is thus at least one heat-insulating zone between a reaction zone and a heat exchange zone. Preferably, there is a heat isolation zone around each reaction zone.

The catalysts or the fixed beds thereof are mounted in a manner known per se on or between gas-permeable walls comprising the reaction zone of the reactor. Particularly in the case of thin fixed beds, technical devices for uniform gas distribution can be provided in the flow direction in front of the catalyst beds. These can be perforated plates or other internals that cause a uniform entry of the process gas into the fixed bed by generating a small but uniform pressure loss.

In a particular embodiment of the process according to the invention, preference is given to using a molar excess of between 0 and 500% of hydrogen, based on the molar flow of phenol, before entering the reaction zone. By a

Increasing the ratio of hydrogen per phenol can on the one hand accelerate the reaction and thus the space-time yield (amount of cyclohexanone produced per mass

On the other hand, the equilibrium of the reaction is shifted positively in the direction of the product (cyclohexanone). As a result, the conversion of the educt phenol can be increased.

In a further particularly preferred embodiment of the method, the inlet temperature of the process gas entering the respective reaction zone is from 10 to 250.degree. C., preferably from 50 to 220.degree. C., more preferably from 140 to 200.degree.

In yet another particularly preferred embodiment of the process, the absolute pressure at the inlet of the first reaction zone is between 0.8 and 6 bar, preferably between 1 and 4 bar, more preferably between 1 and 3 bar.

In a further particularly preferred embodiment of the process, the residence time of the process gas in the overall process is between 0.1 and 15 s, preferably between 0.2 and 5 s, particularly preferably between 0.5 and 2 s. The phenol and the hydrogen are preferably fed only before the first reaction zone. This has the advantage that the entire process gas can be used for the absorption and removal of the heat of reaction in all reaction zones. In addition, by such a procedure, the space-time yield can be increased, or the necessary catalyst mass can be reduced. However, it is also possible to meter in phenol and / or hydrogen in the process gas as needed before one or more of the reaction zones following the first reaction zone. In addition, the temperature profile in the reaction zone can be controlled via the supply of gas between the reaction zones.

In a preferred further development of the process, the process gas leaving the last reaction zone is at least partially reused by being introduced into one of the reaction zones. More preferably, only the portion of the hydrogen is reused by introducing it into the first reaction zone.

This further development is advantageous because the process gas at the outset of the last reaction zone essentially comprises only hydrogen and cyclohexanone and thus a further conversion of the hydrogen with phenol to cyclohexanone can be achieved via reuse.

In a particularly preferred embodiment of the method according to the invention, the process gas is cooled after at least one of the reaction zones used, more preferably after each of the catalyst beds used. For this purpose, the process gas is passed after exiting a reaction zone through one or more of the above-mentioned heat exchange zones, which are located behind the respective reaction zones. These may be used as heat exchange zones in the form of heat exchangers known to those skilled in the art, e.g. Tube bundle, plate, Ringnut-, spiral, finned tube, microstructured heat exchanger be executed. Preference is given to microstructured heat exchangers.

In the context of the present invention, microstructured denotes that the

Heat exchanger for the purpose of heat transfer comprises fluid-carrying channels, which are characterized in that they have a hydraulic diameter between 50 microns and 5 mm. The hydraulic diameter is calculated four times the flow cross-sectional area of the fluid-conducting channel divided by the circumference of the channel.

In a particular embodiment of the method, steam is generated during cooling of the process gas in the heat exchange zones by the heat exchanger.

Within this further embodiment, it is preferable in the heat exchangers, which include the heat exchange zones, to carry out evaporation on the side of the cooling medium, preferably partial evaporation.

Partial evaporation referred to in the context of the present invention, an evaporation in which a gas / liquid mixture of a substance is used as a cooling medium and in which there is still a gas / liquid mixture of a substance after heat transfer in the heat exchanger.

The carrying out of evaporation is particularly advantageous because in this way the achievable heat transfer coefficient from / to process gases on / from the cooling / heating medium becomes particularly high and thus efficient cooling can be achieved.

Performing a partial evaporation is particularly advantageous because the absorption / release of heat by the cooling medium thereby no longer results in a temperature change of the cooling medium, but only the gas / liquid ratio is shifted. This has the consequence that over the entire heat exchange zone, the process gas is cooled to a constant temperature. This in turn reliably prevents the occurrence of temperature profiles in the flow of the process gases, which improves the control over the reaction temperatures in the reaction zones and in particular prevents the formation of local overheating by temperature profiles.

In an alternative embodiment, instead of evaporation / partial evaporation, a mixing zone can also be provided upstream of the inlet of a reaction zone in order to standardize the temperature profiles in the flow of process gases which may arise during cooling by mixing transversely to the main flow direction.

In a further preferred embodiment of the process, the successively connected reaction zones are increased in the reaction zone to the reaction zone or operated decreasing average temperature. This means that within a sequence of reaction zones, the temperature can be both increased and decreased from reaction zone to reaction zone. Thus, it may be particularly advantageous to allow the average temperature to first increase from reaction zone to reaction zone to increase the conversion and then to allow the average temperature in the following last reaction zone to decrease again to shift the equilibrium. This can be adjusted, for example, via the control of the heat exchange zones connected between the reaction zone. Further options for setting the average temperature are described below.

The thickness of the flow-through reaction zones can be chosen to be the same or different and results according to laws generally known in the art from the residence time described above and the process gas quantities enforced in the process. The erf Massenndungsgemäß enforceable by the method mass flows of product gas (cyclohexanone), which also results in the amounts of process gas to be used, are usually between 0.01 and 25 t / h, preferably between 0.1 and 20 t / h, more preferably between 1 and 15 t / h.

The maximum outlet temperature of the process gas from the reaction zones is usually in a range from 170 ° C to 300 ° C, preferably from 180 ° C to 250 ° C, more preferably from 190 ⁰ C to 220 ° C. The control of the temperature in the reaction zones is preferably carried out by at least one of the following measures: dimensioning of the adiabatic reaction zone, control of heat dissipation between the reaction zones, addition of gas between the reaction zones, molar ratio of the reactants / excess of hydrogen used, addition of inert gases, in particular Nitrogen, carbon dioxide, before and / or between the reaction zones.

The composition of the catalysts in the reaction zones according to the invention may be identical or different. In a preferred embodiment, the same catalysts are used in each reaction zone. However, it is also advantageous to use different catalysts in the individual reaction zones. Thus, especially in the first reaction zone, when the concentration of the reaction educts is still high, a less active catalyst can be used and in the further reaction zones the activity of the catalyst can be increased from reaction zone to reaction zone. The control of catalyst activity can also be achieved by dilution with Inert materials or carrier material take place. Also advantageous is the use of a catalyst in the first and / or second reaction zone, which is particularly stable against deactivation at the temperatures of the process in these reaction zones.

0.1 kg / h to 20 kg / h, preferably 0.5 kg / h to 15 kg / h, particularly preferably 2 kg / h to 10 kg / h of cyclohexanone can be prepared by the process according to the invention per 1 kg of catalyst.

The inventive method is thus characterized by high space-time yields, combined with a reduction of the apparatus sizes and a simplification of the apparatus or reactors. This surprisingly high space-time yield is made possible by the interaction of the novel and preferred embodiments of the novel process. In particular, the interplay of staggered, adiabatic reaction zones with interposed heat exchange zones and the defined residence times allows precise control of the process and the resulting high space-time yields, as well as a reduction in the by-products formed such as cyclohexanol.

Another object of the invention is a reactor system for the reaction of phenol and hydrogen to cyclohexanone, characterized in that it comprises feed lines (Z) for a process gas comprising phenol and hydrogen or for at least two process gases, of which at least one phenol and at least one hydrogen, and 5 to 30 successive reaction zones (R) in the form of fixed beds of a heterogeneous catalyst, between the reaction zones heat insulation zones (I) in the form of insulating material and between these heat exchange zones (W) are in the form of plate heat exchangers with the reaction zones over - And discharges for the process gases are connected and include the supply and discharge lines for a cooling medium.

The reactor system may also comprise 6 to 25 or 7 to 15 reaction zones in the form of fixed beds. The insulating material of the heat insulating zones is preferably a material having a

W

Wärmeleitkoeffizient λ less than or equal to 0.08. Particularly preferred are about

_m - K Polystyrene, polyurethanes, glass wool or air.

The plate heat exchangers are preferably microstructured plate heat exchangers.

The present invention will be elucidated with reference to the drawings without, however, being limited thereto.

1 shows a schematic representation of an embodiment of the reactor system according to the invention, the following reference numbers being used in the figures: Z: supply line (s)

R: reaction zone (s)

I: thermal insulation zone (s)

W: heat exchange zone (s)

2 shows temperature (T) and conversion of phenol (U) over a number of 10 reaction zones (S) with downstream heat exchange zones (according to Example 1).

The present invention will be further illustrated by the following Example 1, without being limited thereto.

Examples

Example 1:

In this example, the process gas flows over a total of 10 fixed catalyst beds from a catalyst comprising palladium on a support of aluminum oxide with calcium oxide, ie through 10 reaction zones. Each after a reaction zone is a heat exchange zone in which the process gas is cooled before it enters the next reaction zone. The process gases used at the beginning of the first reaction zone are pure phenol and pure hydrogen, wherein the volume flow of the phenol is adjusted so that a total of 20 mol% of phenol and 80 mol% are fed into the first reaction zone. The absolute inlet pressure of the process gases directly in front of the first reaction zone is 1 bar. The length of the fixed catalyst beds, ie the reaction zones, is always 0.1 m. The activity of the catalysts is adjusted in the individual reaction zones by dilution with pellets of quartz glass. In the last reaction zone is an undiluted catalyst bed. The activity of the last reaction zone is therefore normalized to 100%. The individual activities of the reaction zones set by dilution are shown in Table 1. There is no subsequent metering of gas before the individual catalyst stages. The residence time in all reaction zones of the plant as a whole is 0.95 seconds.

Table 1: Length of the reaction zones according to Example 1

The results are shown in FIG. Here, the individual reaction zones are listed on the x-axis, so that a spatial course of developments in the process is visible. The temperature of the process gas is indicated on the left y-axis. The temperature profile across the individual reaction zones is shown as a thick, solid line. On the right y-axis the total conversion of phenol is indicated. Of the The course of the conversion over the individual reaction zones is shown as a thick dashed line.

It can be seen that the inlet temperature of the process gas before the first reaction zone is about 190 ° C. Due to the exothermic reaction to cyclohexanone under adiabatic conditions, the temperature in the first reaction zone rises to about 220 ° C, before the process gas is cooled in the downstream heat exchange zone again. The inlet temperature before the next reaction zone is again about 190 ° C. By exothermic adiabatic reaction, it rises again to about 220 ° C. The sequence of heating and cooling essentially continues until the sixth reaction zone. From the seventh reaction zone, the inlet temperatures of the process gas sink in front of the individual reaction zones. This is advantageous since, in the course of the reaction later reaction zones, the amount of reactants capable of the reaction is lower and, accordingly, a more advantageous position of the reaction equilibrium is sought after the conversion. Consequently, the temperature of the process gas can be lowered. Furthermore, this investment costs can be saved, since due to the larger temperature increase per reaction zone from catalyst stage 7, the required number of heat exchangers decreases.

It is a conversion of 100% of the beginning of the first reaction zone used phenol, calculated from the remaining mass output of the last reaction zone obtained. The achieved space-time yield based on the mass used catalyst

5.37 kg cyclohexanone / kg cathath.

Claims

claims

1. Process for the preparation of cyclohexanone, characterized in that it comprises a conversion of phenol with hydrogen to cyclohexanone according to formula (I)

C ₆ H ₆ O (g) + 2 H ₂ (g) →C ₆ H ₁₀ O (g); ΔH = -57.8 kJ / mol (I)

in 5 to 30 successive reaction zones under adiabatic conditions in the presence of heterogeneous catalysts.

2. The method according to claim 1, characterized in that the conversion takes place in 6 to 25, preferably 7 to 15 successive reaction zones.

3. The method according to claim 1 or claim 2, characterized in that the inlet temperature of the entering into the respective reaction zone process gas from 10 to 250 ° C, preferably from 50 to 220 ° C, more preferably from 140 to 200 ° C.

4. The method according to claim 1 to 3, characterized in that the absolute pressure at the entrance of the first reaction zone between 0.8 and 6 bar, preferably between 1 and 4 bar, more preferably between 1 and 3 bar.

5. The method according to any one of the preceding claims, characterized in that the residence time of the process gas in all reaction zones together between 0.1 and 15 s, preferably between 0.2 and 5 s, more preferably between 0.5 and 2 s.

6. The method according to any one of the preceding claims, characterized in that the

Catalysts consist of a material which is characterized in addition to its catalytic activity for the reaction according to formula (I) by sufficient thermal stability and by a high specific surface area.

7. The method according to claim 6, characterized in that the catalysts comprise palladium on an alumina support, which also comprises calcium oxide.

8. The method according to any one of the preceding claims, characterized in that the catalysts are present in a fixed bed arrangement.

9. The method according to claim 8, characterized in that the catalysts are present as monoliths.

10. The method according to any one of claims 1 to 7, characterized in that the catalysts are in fluidized bed arrangement.

11. The method according to claims 8 or 10, characterized in that the

Catalysts in beds of particles having average particle sizes of 1 to 10 mm, preferably 1.5 to 8 mm, more preferably 2 to 5 mm are present.

12. The method according to any one of the preceding claims, characterized in that after at least one reaction zone is at least one heat exchange zone through which the process gas is passed.

13. The method according to claim 12, characterized in that after each reaction zone at least one, preferably a heat exchange zone is located, through which the process gas is passed.

14. The method according to any one of the preceding claims, characterized in that between a reaction zone and a heat exchange zone at least one

Heat insulation zone is located.

15. The method according to claim 14, characterized in that there is a heat insulation zone around each reaction zone.

16. Reactor system for carrying out a method according to one of the preceding claims, characterized in that it comprises feed lines (Z) for a process gas comprising phenol and hydrogen or for at least two process gases, of which at least one phenol and at least one hydrogen and 5 to 30 in a row connected reaction zones (R) in the form of fixed beds of a heterogeneous catalyst, wherein between the reaction zones heat insulation zones (T) in the form of insulating material and between them

Heat exchange zones (W) are in the form of plate heat exchangers, which are connected to the reaction zones via supply and discharge lines for the process gases and the supply and discharge lines for a cooling medium.