EP2379677A2

EP2379677A2 - Process for preparing benzene

Info

Publication number: EP2379677A2
Application number: EP09796946A
Authority: EP
Inventors: Ralph Schellen; Evin Hizaler Hoffmann; Leslaw Mleczko; Stephan Schubert; Bharat Marwaha
Original assignee: Bayer Technology Services GmbH
Current assignee: Bayer Intellectual Property GmbH
Priority date: 2008-12-20
Filing date: 2009-12-04
Publication date: 2011-10-26
Also published as: US20110303581A1; WO2010069487A2; WO2010069487A3; DE102008064276A1

Abstract

The present invention relates to a process for the endothermic, catalytic gas-phase reaction of naphtha with hydrogen to form benzene, in which the reaction is carried out under adiabatic conditions in from 5 to 12 reaction zones connected in series.

Description

Process for the preparation of benzene

The present invention relates to a process for the endothermic, catalytic gas phase reaction of naphtha with hydrogen to benzene, wherein the reaction is carried out in 5 to 12 successive reaction zones under adiabatic conditions.

Naphta is an untreated petroleum distillate from the refining of oil or natural gas, from which, among other things, benzene is commonly extracted. Benzene, on the other hand, is an essential starting material for many other petrochemical products.

For example, benzene is used in the chemical industry for the synthesis of many compounds, such as aniline, styrene, nylon, synthetic rubber, plastics, detergents, insecticides, dyes and many more. In addition, many aromatics such as phenol, nitrobenzene, aniline, chlorobenzene, hydroquinone and picric acid are obtained by substitution.

Another product, which has now taken a back seat from the environmental point of view, is the use of benzene as fuel for internal combustion engines after the Otto cycle.

The reactions which are essential in connection with the preparation of naphtha-type benzene are shown in the following formulas (I to IV). The formula (I) refers to the conversion of cyclohexane as the proportion of naphtha to benzene, which is very endothermic and the formula (U) refers to the conversion of hexane to cycloxane, which in turn according to formula (I) are converted to benzene can. The side reactions according to the formulas (in and IV), which can also take place during the production of benzene, are also indicated and are in particular exothermic reactions.

C ₆ H ₁₂ → → C ₆ H ₆ + 3 H ₂ (I)

C ₆ H ₁₄ ^ C ₆ H ₁₂ + H ₂ (H)

C ₆ H ₁₂ + 2 H ₂ → 2 C ₃ H ₈ (YE)

C ₆ H ₁₄ + H ₂ → 2 C ₃ H ₈ (IV)

The reaction according to formula (I) is, as well as those according to formula (II) equilibrium limited.

The benzene obtained from the reaction of formula (I) forms an essential starting material for further reaction, for example, to the aforementioned products. The controlled supply of heat in processes for the production of benzene is important, since the position of the equilibrium of the abovementioned reaction according to the formula (I) is highly dependent on the temperature of the reaction zone and thus the yields and / or selectivities with respect to benzene can be controlled thereby , In particular, the undesirable side reactions according to formulas (IV-IV) can thus be at least partially suppressed.

An uncontrolled temperature drop through the endothermic reaction according to the formula (I) can thus the formation of more or less large amounts of cyclohexane (according to formula I by reverse reaction) and / or propane (according to formula HI) and / or hexane (according to formula II by back reaction ), which is disadvantageous for the further use of the benzene, since such secondary constituents must first be separated off.

It is therefore advantageous to keep the temperature of the reaction zones controlled in the course of the process at a level which allows a rapid conversion while minimizing the side reactions.

Thus, EP 0 601 398 A1 discloses that for the preparation of BTX aromatics (benzene, toluene and xylene) the temperature level and the catalyst used have a significant influence on

Yield and conversion to target product. According to the disclosure of EP 0 601 398 A1 this is

Temperature level at which the reaction should be carried out, essentially determined by the type and composition of the naphtha used, which is usually by his

Boiling point is characterized. This underlines the importance of accurate temperature control in such processes.

EP 0 601 398 A1 also discloses that it is now common practice to carry out catalytic reforming processes in a plurality of series-connected reactors in which catalysts are in the form of a fixed bed.

In EP 0601398 Al isothermal procedure using a salt bath is disclosed, by means of which the disclosed method in the temperatures of about 500 ⁰ C in the

Reaction zone can be adjusted. An adiabte mode of operation is not disclosed. The im

The catalyst used in the disclosure of EP 0 601 398 A1 consists of a

Support material, which is preferably alumina, on which a layer of a

Platinum group metal is located with a promoter metal from group IVB of the Periodic Table of the Elements.

The method according to the disclosure of EP 0 601 398 A1 is disadvantageous since the disclosed isothermal procedure is very complicated and thus very expensive. In particular, in the production of basic chemicals, which include the benzene, but already affect minor procedural disadvantages serious on the economic usability of the overall process, which also discloses the EP 0 601 398 Al.

The possibility of adiabatic operation is disclosed by J. Ancheyta-Juarez et al. in "Modeling and Simulation of four catalytic reactors in series for natal reforming" in Energy & Fuels (2001) 15: 887-893.

Thus, J. Ancheyta-Juarez et al. Disclose that it may be advantageous to carry out the reaction of naphtha to (among others) benzene in three to four, especially four reaction zones connected in series, wherein intermediate cooling may be provided between the aforementioned reaction zones ,

By means of the method disclosed in the disclosure of J. Ancheyta-Juarez et al. achievable yields of benzene (A ₆ ) are very low with a share of only about 4 mol% of the reaction product, which makes the process disadvantageous.

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 as to the usability of the apparatus and method of producing benzene.

It therefore remains unclear how, starting from the disclosure of EP 1 251 951 (B1), such a reaction should be carried out by means of the device and the method carried out therein. In particular, no method comprising endothermic reactions is disclosed.

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 regarding 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 a certain heat sink with appreciable but limited capacity in each reaction zone is present, which can reduce the temperature rise in the same. It can never be ruled out that a certain dissipation of heat (eg by radiation) takes place; However, a reduction of the possible heat removal from the reaction zone would indicate a linear or degressive temperature gradient, since no further addition of reactants is provided and thus, after an exothermic reaction, the reaction becomes slower and thus the generated heat of reaction would decrease.

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. Thus, the disclosed process is not adiabatic and disadvantageous in that accurate temperature control of the reaction is not possible. This is especially true for the undisclosed possibility of an endothermic reaction in the reaction zones.

Starting from the prior art, it would therefore be advantageous to provide a process for the preparation of benzene, which can be carried out in simple reaction devices and which allows a precise, simple temperature control of the endothermic process, so that it high conversions at the highest possible purity of the product while maintaining desired yields and / or selectivities allowed. Such simple reaction devices would be easily scaled up and are inexpensive and robust in all sizes.

For the endothermic, catalytic gas-phase reaction of naphtha to benzene, as has just been shown, no suitable processes have yet been demonstrated which permit this.

It is therefore an object to provide a process for the endothermic, catalytic gas phase reaction of naphtha to benzene, which can be carried out under precise temperature control in simple reaction devices and thereby high conversions at high purities of the product allowed.

It has surprisingly been found that a process for the preparation of benzene from naphtha in the presence of hydrogen in an endothermic, heterogeneous catalytic gas phase reaction, characterized in that it comprises 5 to 12 successive reaction zones with adiabatic conditions, can solve this task.

Benzene in the context of the present invention refers to a process gas which essentially comprises benzene. The benzene may also include levels of hydrogen and other hydrocarbons.

Other hydrocarbons in the context of the present invention refer to substances present as process gas consisting of carbon, hydrogen and optionally oxygen. In essence, however, such hydrocarbons consist of carbon and hydrogen. - -

Such hydrocarbons are usually those which are introduced either as further constituents of naphtha in the process according to the invention, or those which are formed by side reactions in the process according to the invention, for example after the reactions according to formulas (HI and IV).

Non-conclusive examples of hydrocarbons which are introduced as further constituents of the naphtha in the inventive method are, for example, naphthalene, isopentane and toluene.

Non-conclusive examples of hydrocarbons formed by side reactions in the course of the process according to the invention, for example after the reactions according to the formulas (DI and FV), are for example hexane, cyclohexane and propane.

Naphta refers to a mixture of hydrocarbons as a process gas, as is well known to those skilled in the art. In connection with the process according to the invention, naphtha is preferably a mixture of hydrocarbons which essentially comprises cyclohexane.

Hydrogen, in the context of the present invention, denotes a process gas which essentially comprises hydrogen. This hydrogen can be formed, for example, by the reactions according to the formulas (I and H) or else fed to the process as process gas.

A feeding of hydrogen as a process gas in the erfϊndungsgemäße method is preferred. Preheated hydrogen is particularly preferably supplied as process gas to the process.

Such feeding of, in particular, hydrogen is advantageous because in this way the hydrogen can be used as the heat transfer medium in the method for controlling the temperature. Furthermore, the hydrogen prevents deposits of carbon products on the catalyst surfaces of the catalysts located in the reaction zones (coking).

The term essentially refers in the context of the present invention, a mass fraction and / or a molar fraction of at least 80%.

The naphtha used in the process according to the invention, its constituents, the hydrogen, the benzene and the products of the process according to the invention are also referred to collectively below as process gases.

It follows that the entire erfϊndungsgemäße method is carried out in the gas phase.

If substances used in the process, such as the hydrocarbons at room temperature (23 ⁰ C) and ambient pressure (1013 hPa) are not gaseous, it is assumed in the following that such substances before or during their use in the erfϊndungsgemäßen Process can be überfuhrt by increasing the temperature and / or reducing the pressure in the gas phase.

In addition to the essential components of the process gases, these can also include secondary components. Non-exhaustive examples of minor components that may be included in the process gases include argon, nitrogen and / or carbon dioxide.

According to the invention, the implementation of the process under adiabatic conditions means that the reaction zone from the outside is essentially neither actively supplied with heat nor heat withdrawn. It is well known that complete isolation against heat input or discharge is possible only by complete evacuation, excluding the possibility of heat transfer by radiation. 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, e.g. by isolation by means of well-known isolation means, e.g. Polystyrene insulating materials, or by sufficiently large distances to heat sinks or heat sources, wherein the insulating agent is air, a heat transfer can be reduced.

An advantage of the adiabatic driving method according to the invention of the 5 to 12 reaction zones connected in series with respect to a non-adiabatic mode of operation is that no means for supplying heat have to 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.

Another advantage of the method according to the invention is the possibility of very accurate temperature control, due to the close staggering of adiabatic reaction zones. It can thus be set and controlled in each 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 according to formula (I), is characterized by sufficient chemical resistance under the conditions of the process and by a high specific surface area.

Catalyst materials characterized by such chemical resistance under the conditions of the process include, for example, catalysts comprising platinum and / or rhenium. Preferred catalyst materials consist of equal proportions by weight of rhenium and platinum.

These catalysts can be applied to support materials. Such support materials usually include alumina and / or titania. Preference is given to support materials of alumina.

Particularly preferred catalysts consist of rhenium and platinum, which are applied to the same weight fraction on an alumina support. Methods for producing such catalysts are generally known to the person skilled in the art, for example from EP 0 601 398 A1.

Specific surface referred to in the context of the present invention, the area of the catalyst material, which can be achieved by the process gas, based on the mass of catalyst used.

A high specific surface area is a specific surface area of at least 1 m ² / g, preferably of at least 10 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, moving bed, present.

Preferably, the appearance is fixed bed.

The fixed bed arrangement comprises a catalyst bed in the true sense, d. H. loose, supported or unsupported catalyst in any form and in the form of suitable packings. The term catalyst bed as used herein also encompasses contiguous areas of suitable packages on a support material or structured catalyst supports. These would be e.g. 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 having mean particle sizes of 1 to 10 mm, preferably 2 to 8 mm, particularly preferably 3 to 7 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 contains the abovementioned metals rhenium and platinum in equal proportions by weight on an aluminum oxide support. Also particularly preferred is a fixed bed arrangement with beds of particles having average particle sizes of 1 to 10 mm, preferably 2 to 8 mm, particularly preferably 3 to 7 mm, wherein the particles are alumina particles to which the aforementioned metals rhenium and platinum in equal proportions by weight are applied.

When a catalyst in monolithic form is used in the reaction zones, in a preferred further development of the invention the monolithic catalyst is provided with channels through which the process gases flow. Usually, the channels have a diameter of 0.1 to 3 mm, preferably a diameter of 0.2 to 2 mm, more preferably from 0.5 to 1.5 mm.

If a fluidized bed arrangement of the catalyst is used, the catalyst is preferably present in loose beds of particles, as have already been described in connection with the fixed bed arrangement.

Beds of such particles are advantageous because the size of the particles have a high specific surface of the catalyst material compared to the process gases 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 reactions according to the formulas (I and II) and the pressure loss produced when carrying out the process. Pressure loss is coupled in a direct manner 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 method.

In a preferred embodiment of the process according to the invention, the conversion takes place in 6 to 10, more preferably 6 to 8 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 obtain 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 preferably used in the process according to the invention can consist of simple containers with one or more reaction zones, as e.g. in Ullmann's Encyclopedia of Industrial Chemistry (Fifth, Completely Revised Edition, VoI B4, page 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, therefore, there is 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, bubble-cap trays, valve trays or other internals which cause a uniform entry of the process gas into the fixed bed by producing a small but uniform pressure loss. In a preferred embodiment of the process, the inlet temperature of the process gas entering a reaction zone is from 740 to 790 K, preferably from 750 to 780 K, particularly preferably from 755 to 775 K.

In a further preferred embodiment of the method, the absolute pressure at the inlet of the first reaction zone is between 10 and 40 bar, preferably between 15 and 35 bar, more preferably between 20 and 30 bar.

In yet another preferred embodiment of the process, the residence time of the process gas in all reaction zones together is between 0.5 and 30 s, preferably between 1 and 20 s, particularly preferably between 5 and 15 s.

The naphtha and optionally the hydrogen are preferably fed only before the first reaction zone. This has the advantage that the entire process gas is available for absorbing 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, before one or more of the reaction zones following the first reaction zone, to meter in naphtha and, if appropriate, hydrogen into the process gas as required. In addition, the temperature of the conversion can be controlled by the supply of these process gases between the reaction zones, if they are preheated.

In preferred embodiments of the process according to the invention, the molar ratio of hydrogen to hydrocarbons contained in the naphtha is set in the range from 3 to 9, preferably from 4 to 8, particularly preferably from 5 to 7, moles of hydrogen per mole of hydrocarbon in the naphtha.

The advantages of such a supply of hydrogen have already been set out above. These apply in particular in connection with the addition of a surplus.

Persons skilled in the art will be aware of suitable means for determining the molar amounts of hydrocarbons in a process gas, such as naphtha. A non-exhaustive example is the quantitative analysis by gas chromatography. If the molar composition of the process gas naphtha is known, the adjustment of the molar ratio of hydrogen can be done by simply adjusting the volume flow ratio of the process gases naphtha and hydrogen.

In a further preferred embodiment of the method according to the invention, the process gas is heated after at least one of the reaction zones used, particularly preferably after each reaction zone. For this purpose, the process gas is passed after exiting a reaction zone by one or more of the above heat exchange zones, which are behind the respective reaction zones are located. These can be embodied as heat exchange zones in the form of heat exchangers known to the person skilled in the art, such as, for example, tube bundle, plate, annular groove, spiral, finned tube, micro heat exchanger. The heat exchangers are preferably microstructured heat exchangers.

In the context of the present invention, microstructured means 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 μm and 5 mm. The hydraulic diameter is calculated as 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, the heating of the process gas takes place in the heat exchange zones by a condensation of a heat transfer medium.

Within this particular embodiment, it is preferable in the heat exchangers, which include the heat exchange zones, on the side of the heating medium to carry out a condensation, preferably partial condensation.

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

The execution of a condensation is particularly advantageous because in this way the achievable heat transfer coefficient to the process gases from the heating medium is particularly high and thus efficient heating can be achieved.

Performing a partial condensation is particularly advantageous because the release of heat by the heating medium thereby no longer results in a change in temperature of the heating medium, but only the gas / liquid balance is shifted. This has the consequence that over the entire heat exchange zone, the process gas is heated to a constant temperature. This in turn safely prevents the occurrence of radial temperature profiles in the flow of process gases, thereby improving control over the reaction temperatures in the reaction zones and, in particular, preventing the formation of local overheating by radial temperature profiles.

In an alternative embodiment, instead of a condensation / partial condensation, it is also possible to provide a mixing zone upstream of the inlet of a reaction zone in order to standardize the radial temperature profiles, if appropriate during the heating, in the flow of the process gases by mixing transversely to the main flow direction. In a preferred embodiment of the process, the successively connected reaction zones are operated at an average temperature increasing or decreasing from reaction zone to reaction zone. This means that within a sequence of reaction zones, the temperature can be both increased and decreased from reaction zone to reaction zone. 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 to those skilled in the art from the above-described residence time and the process gas quantities in each case. The mass flow rates of process gas which can be carried out according to the process according to the invention, based on the catalyst mass used (also called WHSV, weight hourly space velocity), is usually between 28 and 42 h.sup.- ¹ , preferably between 30 and 40 h.sup.- ¹ , particularly preferably between 33 and 38 h ^"1 .

The maximum exit temperature of the process gas from the first reaction zone is usually in the range of the inlet temperature, since the reactions according to the formulas (DI) and (IV) are exothermic reactions. They may also, especially at the exit from the last reactions in which a large amount of benzene has already been formed and therefore the particular endothermic reaction according to formula (I) loses influence, in a range of 770 to 820 K, preferably from 775 to 795 K, more preferably from 780 to 785 K.

The subsequent reaction zones can be freely determined by the person skilled in the art according to the erfϊndungsgemäßen method by the following measures with respect to their inlet temperature.

The control of the temperature in the reaction zones preferably takes place by at least one of the following measures: dimensioning of the adiabatic reaction zone, control of the heat supply between the reaction zones, addition of further process gas between the reaction zones, molar ratio of the reactants / excess of hydrogen used, addition of secondary constituents, 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 the catalyst activity can also be carried out by dilution with inert materials or carrier material.

1 kg / h to 50 kg / h, preferably 5 kg / h to 30 kg / h, more preferably 10 kg / h to 20 kg / h of benzene can be produced 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 erfϊndungsgemäßen and preferred embodiments of the new method. In particular, the interaction 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 of the by-products formed, such as carbon dioxide.

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

FIG. 1 shows reactor temperature (T) and molar mass flow of benzene (U) over a length (L) of 11 m at reaction zones with respective downstream heat exchange zones (according to Example 1), the lengths of the heat exchange zones being assumed idealized to zero, since here no conversion to be held.

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

Examples

Gaseous naphtha and hydrogen are fed to the process as process gases in a molar ratio of 7.77. The process is operated in a total of six fixed beds of rhenium and platinum with 0.29 wt .-% on an alumina support, ie in six reaction zones.

Each after a reaction zone is a heat exchange zone in which the exiting process gas is reheated before it enters the next reaction zone.

The absolute inlet pressure of the process gas directly in front of the first reaction zone is 25 bar. The length of the fixed catalyst beds, ie the reaction zones, varies from reaction zone to reaction zone, starting from 0.15 m in the first reaction zone up to 6 m in the sixth reaction zone. The exact lengths of the reaction zone are summarized in Table 1. The activity of the catalyst used is not variable across the reaction zones. There is no addition of process gas before the individual reaction zones. The WHSV is 35 h ^"1 .

Table 1: Lengths of reaction zones

Reaction zone length

[#] [m]

1 0.15

2 0.35

3 1

4 1.5

5 2

6 6

Σ 11.0

The results are shown in Fig.l. In this case, the continuous length of the reaction zones is listed on the x-axis, so that a spatial course of the developments in the process is visible, wherein the heat exchange zones are neglected. 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. Due to the idealized assumption of the length of the heat exchange zones to 0 m, discontinuities occur with regard to the temperature profile. On the right y-axis is the cumulative molar flow of benzene in the process gas indicated over the reaction path. The course of the same over the reaction path is shown as a thin solid line.

It can be seen that the inlet temperature of the process gas before the first reaction zone about

775K. Due to the essentially endothermic reaction to benzene under adiabatic conditions, the temperature in the first reaction zone drops to about 760 K, before the

Process gas is heated in the downstream heat exchange zone back to the aforementioned 775 K. By endothermic adiabatic reaction, the temperature drops in the second

Reaction zone to about 750 K. The sequence of cooling by endothermic, adiabatic reaction and heating is continued with changed outlet temperatures after the respective reaction zones, the inlet temperature is set in the heat exchange zones back to the desired 775 K.

There is obtained a conversion of cyclohexane and hexane of about 60%. The space-time yield achieved, based on the mass of catalyst used, is about 15 kg _benzo / kg _cat h.

Claims

claims

1. A process for the preparation of benzene from naphtha in the presence of hydrogen in an endothermic, heterogeneous catalytic gas phase reaction, characterized in that it comprises 5 to 12 successive reaction zones with adiabatic conditions.

2. The method according to claim 1, characterized in that the conversion takes place in 6 to 10, preferably 6 to 8 successive reaction zones.

3. The method according to claim 1 or claim 2, characterized in that the inlet temperature of the entering into the first reaction zone process gas 740 to 790 K, preferably from 750 to 780 K, more preferably from 755 to 775 K.

4. The method according to claim 1 to 3, characterized in that the absolute pressure on

Input of the first reaction zone between 10 and 40 bar, preferably between 15 and 35 bar, more preferably between 20 and 30 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 between 0.5 and 30 s, preferably between 1 and 20 s, more preferably between 5 and 15 s.

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

7. The method according to claim 6, characterized in that the catalysts are present as monoliths.

8. The method according to claim 7, characterized in that the monolith channels having a diameter of 0, 1 to 3 mm, preferably from 0.2 to 2 mm, particularly preferably from 0.5 to 1.5 mm.

9. Process according to claims 1 to 6, characterized in that the catalysts are present in beds of particles having average particle sizes of 1 to 10 mm, preferably 2 to 8 mm, particularly preferably 4 to 7 mm.

10. 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.

11. The method according to claim 10, characterized in that after each reaction zone at least one, preferably a heat exchange zone is located, through which the process gas is passed.

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

13. The method according to claim 12, characterized in that there is a heat insulation zone around each reaction zone.