US20110303581A1

US20110303581A1 - Process for preparing benzene

Info

Publication number: US20110303581A1
Application number: US13/140,458
Authority: US
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: 2011-12-15
Also published as: WO2010069487A3; EP2379677A2; DE102008064276A1; WO2010069487A2

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 in 5 to 12 serial reaction zones under adiabatic conditions.

Description

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 in 5 to 12 serial reaction zones under adiabatic conditions.
Naphtha is an untreated petroleum distillate from the refining of petroleum or natural gas, and one of the products typically recovered from it is benzene. Benzene in turn is a key starting material for many further petrochemicals.
For instance, benzene is used in the Chemical Industry for the synthesis of numerous compounds, such as, for example, aniline, styrene, nylon, synthetic rubber, plastics, detergents, insecticides, dyes and numerous further substances. Also obtained, by substitution, are numerous aromatics, such as, for example, phenol, nitrobenzene, aniline, chlorobenzene, hydroquinone and picric acid.
A further product, though one which has now taken a back seat for environmental reasons, is the use of the benzene as a fuel for internal combustion engines operating in accordance with the Otto cycle.
The key reactions in the preparation of benzene from naphtha are set out in the formulae below (I to IV). Formula (I) relates to the conversion of cyclohexane as a fraction of the naphtha, to form benzene, and is very endothermic, and formula (II) relates to the conversion of hexane to form cyclohexane, which can be reformed in turn into benzene in accordance with formula (I). The secondary reactions of formulae (III) and (IV), which may likewise take place during the preparation of benzene, are shown as well, and are especially exothermic reactions.
C₆H₁₂
C₆H₆+3.H₂ (I)
C₆H₁₄
C₆H₁₂+H₂ (II)
C₆H₁₂+2.H₂→2.C₃H₈ (III)
C₆H₁₄+H₂→2.C₃H₈ (IV)
The reaction according to formula (I), like that according to formula (II), is equilibrium-limited.
The benzene obtained from the reaction according to formula (I) forms a key starting product for further reaction to give, for example, the abovementioned products.
The controlled supply of heat in processes for obtaining benzene is important, since the position of the equilibrium in the aforementioned reaction according to formula (I) is heavily dependent on the temperature of the reaction zone, and it is therefore possible to control the yields and/or selectivities with respect to benzene by this means. In particular it is therefore possible at least partly to suppress the unwanted secondary reactions of formulae (III-IV).
An uncontrolled drop in temperature as a result of the endothermic reaction according to formula (I) may therefore promote the formation of more or less large quantities of cyclohexane (according to formula I by back-reaction) and/or propane (according to formula III) and/or hexane (according to formula II by back-reaction), which is a disadvantage for the subsequent 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 in the course of the process controlled at a level which allows rapid conversion with a minimization of the secondary reactions.
Accordingly, EP 0 601 398 A1 discloses a key influence on yield and conversion to target product exerted in the preparation of BTX aromatics (benzene, toluene and xylene) by the temperature level and by the catalyst employed. In accordance with the disclosure content of EP 0 601 398 A1, the temperature level at which the reaction is to be performed is essentially determined by the nature and composition of the naphtha used, which is typically characterized by its boiling point. This underlines the importance of precise temperature control in such processes.
EP 0 601 398 A1 also discloses how it is now customary to perform catalytic reformation processes in a plurality of reactors, connected in series, containing catalysts in the form of a fixed bed.
EP 0 601 398 A1 discloses an isothermal procedure using a salt bath by means of which the temperatures of approximately 500° C. that are disclosed in the process are brought about in the reaction zone. An adiabatic regime is not disclosed. The catalyst used in the process disclosed in EP 0 601 398 A1 is composed of a support material, which is preferably alumina, on which there is located a layer of a platinum group metal with a promoter metal from group WB of the Periodic Table of the Elements.
The process as disclosed in EP 0 601 398 A1 is disadvantageous on account of the fact that the isothermal procedure disclosed is extremely complicated and therefore very expensive. In the production of industrial chemicals in particular, which includes the production of benzene, however, even slight process disadvantages have severe consequences for the economics of the process as a whole, as is also disclosed by EP 0 601 398 A1.
The possibility of an adiabatic regime is disclosed by J. Ancheyta-Juarez et al. in “Modeling and simulation of four catalytic reactors in series for naphtha reforming” in Energy & Fuels (2001) 15: 887-893.
Thus J. Ancheyta-Juarez et al. disclose how it can be advantageous to perform the reaction of naphtha to form (among other products) benzene in three to four, especially four, serial reaction zones, with the possibility of intermediate cooling between the aforementioned reaction zones.
The yields of benzene (A₆) which can be achieved by means of the process presented in the disclosure by J. Ancheyta-Juarez et al. are very low, with a fraction of only about 4 mol % as a proportion of the reaction product, thus making the process disadvantageous.
EP 1 251 951 (B1) discloses an apparatus and the possibility of conducting chemical reactions in the apparatus, the apparatus being characterized by a cascade of mutually contacting reaction zones and heat exchanger devices which are integrated materially with one another. The process to be conducted therein is characterized, therefore, by the contact of the various reaction zones with a respective heat exchanger device, in the form of a cascade. A disclosure as to the possibility of using the device and the process for preparing benzene is absent.
It remains unclear, then, as to how, on the basis of the disclosure content of EP 1 251 951 (B1), a reaction of this kind is to be carried out by means of the apparatus and of the process performed therein. In particular there is no disclosure of a process comprising endothermic reactions.
Furthermore, for reasons of consistency, it must be assumed that the process disclosed in EP 1 251 951 (B1) is performed in an apparatus which is identical or similar to the disclosure relating to the apparatus. The result of this is that, as a consequence of the extensive contact between the heat exchange zones and the reaction zones, as per the disclosure, a significant amount of heat takes place as a result of thermal conduction between the reaction zones and the adjacent heat exchange zones.
The disclosure concerning the oscillating temperature profile can only be understood, then, to mean that the temperature peaks found here would be sharper in the absence of this contact. A further indicator of this is the exponential increase in the disclosed temperature profiles between the individual temperature peaks. These indicate that there is a certain heat sink present in each reaction zone, with a marked but limited capacity, which is able to reduce the temperature increase in said zone. It is never possible to rule out a certain dissipation of heat (by radiation, for example); however, in the case of a reduction in the possible dissipation of heat from the reaction zone, a linear or degressive temperature profile would be suggested, since there is no subsequent metering of reactants and hence, after their consumption by exothermic reaction, the reaction would become increasingly slow and hence the heat produced would go down.
EP 1 251 951 (B1) thus discloses multi-stage processes in cascades of reaction zones from which heat is taken off in an undefined quantity by means of thermal conduction. The process disclosed, therefore, is not adiabatic and is disadvantageous insofar as precise temperature control of the reaction is impossible. This applies especially to the undisclosed possibility of an endothermic reaction in the reaction zones.
On the basis of the prior art, therefore, it would be advantageous to provide a process for preparing benzene that can be carried out in simple reaction apparatus and that enables precise, simple temperature control of the endothermic process, thereby allowing high conversions in conjunction with very high purities of the product, while meeting desired yields and/or selectivities. Said simple reaction apparatus would be readily transposable to the industrial scale, and inexpensive and robust in all sizes.
As has just been shown, there have to date been no suitable processes apparent that allow this for the endothermic, catalytic gas phase reaction of naphtha to form benzene.
The object is therefore that of providing a process for endothermic, catalytic gas phase reaction of naphtha to benzene which can be carried out with precise temperature control in simple reaction apparatus and which as a result allows high conversions in conjunction with high product purities.
It has surprisingly been found that a process for preparing benzene from naphtha in the presence of hydrogen in an endothermic, heterogeneously catalytic gas phase reaction, characterized in that it comprises 5 to 12 serial reaction zones with adiabatic conditions, is able to achieve this object.
In connection with the present invention, benzene is a process gas substantially comprising benzene. The benzene may also comprise fractions of hydrogen and further hydrocarbons.
In connection with the present invention, further hydrocarbons are compounds present in the form of process gas composed of carbon, hydrogen and possibly oxygen. Essentially, however, such hydrocarbons are composed of carbon and hydrogen. Such hydrocarbons are typically either those which are introduced into the process of the invention as further constituents of the naphtha, or those which are formed as a result of secondary reactions in the course of the process of the invention, as for instance by the reactions according to formulae (III and IV).
Non-exhaustive examples of hydrocarbons which are introduced into the process of the invention as further constituents of the naphtha are naphthalene, isopentane and toluene, for instance.
Non-exhaustive examples of hydrocarbons which are formed in the course of the process of the invention by secondary reactions, as for instance by the reactions according to formulae (III and IV), are hexane, cyclohexane and propane, for instance.
Naphtha identifies a mixture of hydrocarbons in the form of a process gas, as is general knowledge to the person skilled in the art. In connection with the process of the invention, naphtha is preferably a mixture of hydrocarbons substantially comprising cyclohexane.
In connection with the present invention, hydrogen is a process gas which substantially comprises hydrogen. This hydrogen may be formed, for instance, by the reactions according to formulae (I and II), or else may be supplied as process gas to the process.
The supplying of hydrogen as a process gas into the process of the invention is preferred. With particular preference, preheated hydrogen is supplied as process gas to the process.
Such supplying of hydrogen in particular is advantageous in that it allows the hydrogen to be used as a heat transfer medium in the process, 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 identification “substantially” refers, in connection with the present invention, to a mass fraction and/or a molar fraction of at least 80%.
The naphtha used in the process of the invention, its constituents, the hydrogen, the benzene and also the products of the process of the invention are also referred to below collectively as process gases.
It follows from this that the entire process of the invention is performed in the gas phase. If substances used in the process, such as the hydrocarbons, for instance, are not in gaseous form at room temperature (23° C.) and ambient pressure (1013 hPa), it can be assumed below that, before or during their use in the process of the invention, such substances will be converted into the gas phase by an increase in temperature and/or reduction in pressure.
Besides the substantial components of the process gases, they may also comprise secondary components. Non-exhaustive examples of secondary components which may be present in the process gases are argon, nitrogen and/or carbon dioxide, for instance.
In accordance with the invention the implementation of the process under adiabatic conditions means that substantially neither heat is actively supplied nor heat withdrawn from the outside to/from the reaction zone. It is common knowledge that complete insulation from ingress or egress of heat is possible only by complete evacuation, with the possibility of heat transfer by radiation being ruled out. In connection with the present invention, therefore, adiabatic means that no measures are taken to supply or remove heat.
In one alternative embodiment of the process of the invention, however, heat transfer may be reduced, for example, by insulation using conventional insulating means, such as polystyrene insulants, for example, or else by sufficiently large distances from heat sinks or heat sources, the insulation means being air.
An advantage of the adiabatic regime of 5 to 12 serial reaction zones in accordance with the invention as compared with a non-adiabatic regime is that in the reaction zones there is no need to provide means for heat removal, a fact which results in a considerable simplification in construction. As a result, in particular, there are simplifications in the manufacturing of the reactor and also in the scaleability of the process, and there is an increase in the reaction conversions.
A further advantage of the process of the invention is the possibility of very precise temperature control, as a result of the narrow staggering of adiabatic reaction zones. By this means it is possible in each reaction zone to set and control a temperature which is advantageous in the progress of the reaction.
The catalysts used in the process of the invention are typically catalysts composed of a material which as well as its catalytic activity for the reaction according to formula (I) is characterized by sufficient chemical resistance under the conditions of the process and also by a high specific surface area.
Catalyst materials which are characterized by such chemical resistance under the conditions of the process are, for example, catalysts comprising platinum and/or rhenium.
Preferred catalyst materials are composed of equal weight fractions of rhenium and platinum.
These catalysts may be applied on support materials. Such support materials typically comprise alumina and/or titanium dioxide. Preference is given to alumina support materials.
Particularly preferred catalysts are composed of rhenium and platinum applied at the same weight fraction on an alumina support. Methods of producing such catalysts are general knowledge to a person skilled in the art, from EP 0 601 398 A1, for instance.
Specific surface area in connection with the present invention identifies the surface area of the catalyst material which can be reached by the process gas, based on the mass of catalyst material employed.
A high specific surface area is a specific surface area of at least 1 m²/g, preferably of at least 10 m²/g.
The catalysts of the invention are located in the reaction zones in each case and may be present in all conventional presentation forms, e.g. fixed bed, moving bed.
The presentation form is preferably that of a fixed bed.
The fixed bed arrangement comprises a catalyst bed in the actual sense, i.e. loose, supported or unsupported catalyst in any desired form, and also in the form of suitable packings. The term catalyst bed as used herein also encompasses coherent regions of suitable packings on a support material or structured catalyst support. Examples of such would include ceramic honeycomb supports for coating, having comparatively high geometric surface areas, or corrugated layers of metal wire mesh with catalyst granules, for example, immobilized thereon. In connection with the present invention, the presence of the catalyst in monolithic form is viewed as a special form of packing.
Where a fixed bed arrangement of the catalyst is used, the catalyst is preferably in beds of particles having average particles sizes of 1 to 10 mm, preferably 2 to 8 mm, more preferably of 3 to 7 mm.
Likewise with preference the catalyst in the case of a fixed bed arrangement is in monolithic form. In the case of a fixed bed arrangement particular preference is given to a monolithic catalyst which comprises the aforementioned metals, rhenium and platinum, in equal weight fractions on an alumina support.
Likewise particularly preferred is a fixed bed arrangement having particle beds, having average particles sizes of 1 to 10 mm, preferably 2 to 8 mm, more preferably of 3 to 7 mm, the particles being alumina particles to which the aforementioned metals, rhenium and platinum, have been applied in equal weight fractions.
If a catalyst is used in monolithic form in the reaction zones, then, in a preferred development of the invention, the catalyst present in monolithic form is provided with channels through which the process gases flow. Typically the channels have a diameter of 0.1 to 3 mm, preferably a diameter of 0.2 to 2 mm, more preferably of 0.5 to 1.5 mm.
Where a moving bed arrangement of the catalyst is used, the catalyst preferably takes the form of loose beds of particles, of the kind already 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 area of the catalyst material in relation to the process gases and it is therefore possible to achieve a high conversion rate. Accordingly, the limitation of mass transport in the reaction by diffusion can be minimized. At the same time, however, the particles are also not so small that increased pressure drops occur disproportionately when the gases flow through the fixed bed. The ranges of particle sizes specified in the preferred embodiment of the process, comprising a reaction in a fixed bed, are therefore an optimum between the achievable conversion from the reactions according to formulae (I and II) and the pressure drop generated when the process is implemented. Pressure drop is coupled directly with the necessary energy, in the form of compressor output, and so a superproportional increase in the latter would result in an uneconomic process regime.
In one preferred embodiment of the process of the invention the conversion takes place in 6 to 10, more preferably 6 to 8, serial reaction zones.
A preferred further embodiment of the process 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 one particularly preferred further embodiment of the process each reaction zone is followed by at least one, preferably exactly one, heat exchange zone through which the process gas emerging from the reaction zone is passed.
These reaction zones may be disposed either in one reactor or, in divided form, in two or more reactors. The arrangement of the reaction zones in one reactor leads to a reduction in the number of apparatuses used.
The individual reaction zones and heat exchange zones may also be arranged in one reactor or, in divided form, in any desired combinations of reaction zones with heat exchange zones in two or more reactors.
Where reaction zones and heat exchange zones are present in one reactor, then, in an alternative embodiment of the invention, there is a heat insulation zone located between these zones in order to allow the adiabatic operation of the reaction zone to be maintained.
In addition it is possible for certain of the serial reaction zones to be replaced or supplemented, independently of one another, by one or more parallel reaction zones. The use of parallel reaction zones makes it possible in particular to exchange or supplement the zones during ongoing, continuous operation of the process overall.
Parallel and serial reaction zones may in particular also be combined with one another. With particular reference, however, the process of the invention features exclusively serial reaction zones.
The reactors used preferably in the process of the invention may be composed of simple vessels with one or more reaction zones, of the kind described, for example, in Ullmann's Encyclopedia of Industrial Chemistry (Fifth, Completely Revised Edition, Vol. B4, pages 95-104, pages 210-216), it being possible for heat insulation zones to be provided additionally between each of the individual reaction zones and/or heat exchange zones.
In one alternative embodiment of the process, therefore, there is at least one heat insulation zone located between a reaction zone and a heat exchange zone. Preferably there is a heat insulation zone located around each reaction zone.
The catalysts or fixed beds thereof are mounted in a conventional way on or between gas-permeable walls encompassing the reaction zone of the reactor. In the case of thin fixed beds in particular, technical devices for uniform gas distribution may be fitted upstream of the catalyst beds. These devices may be perforated plates, bubble trays, valve trays or other internals which, by generating a low but uniform pressure drop, produce uniform entry of the process gas into the fixed bed.
In one preferred embodiment of the process the entry temperature of the process gas entering one reaction zone is from 740 to 790 K, preferably from 750 to 780 K, more preferably from 755 to 775 K.
In a further preferred embodiment of the process the absolute pressure at the entry 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 the reaction zones together is between 0.5 and 30 s, preferably between 1 and 20 s, more preferably between 5 and 15 s.
The naphtha and, where appropriate, the hydrogen are preferably supplied only ahead of the first reaction zone. This has the advantage that the whole of the process gas is available for the accommodation of heat of reaction in all the reaction zones. Such a procedure additionally enables the space-time yield to be increased, or the mass of catalyst required to be reduced. It is, however, also possible to meter naphtha and, where appropriate, hydrogen into the process gas ahead of one or more of the reaction zones that follow the first reaction zone, if needed. The supply of these process gases between the reaction zone is an additional way of controlling the temperature of the conversion, if they are preheated.
In preferred embodiments of the process of the invention the molar ratio of hydrogen to hydrocarbons present in the naphtha is set in ranges from 3 to 9, preferably from 4 to 8, more preferably from 5 to 7 mol of hydrogen per mole of hydrocarbon in the naphtha.
The advantages of such supply of hydrogen have already been elucidated. They apply in particular in connection with the supply of an excess.
The person skilled in the art is aware of suitable means for determining the molar amounts of hydrocarbons in a process gas, such as naphtha. One non-exhaustive example is quantitative analysis by means of gas chromatography. If the molar composition of the naphtha process gas is known, the molar ratio of hydrogen to it can be set by simple setting of the volume flow ratio of the naphtha and hydrogen process gases.
In a further preferred embodiment of the process of the invention the process gas is heated after at least one of the reaction zones used, more preferably after each reaction zone. This is done by passing the process gas, following exit from a reaction zone, through one or more of the abovementioned heat exchange zones which are located downstream of the respective reaction zones. These zones may be configured as heat exchange zones in the form of heat exchangers known to the person skilled in the art, such as shell-and-tube, plate, annular-groove, spiral, ribbed-tube or micro-type heat exchangers, for example. The heat exchangers are preferably microstructured heat exchangers.
Microstructured in connection with the present invention 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 of between 50 μm and 5 mm. The hydraulic diameter is calculated from four times the flow-traversed cross-sectional area of the fluid-carrying channel, divided by the circumference of the channel.
In one particular embodiment of the process the process gas is heated in the heat exchange zones by condensation of a heat transfer medium.
Within this particular embodiment it is preferred to perform condensation, preferably partial condensation, on the side of the heating medium in the heat exchangers which constitute the heat exchange zones.
Partial condensation in connection with the present invention means a condensation in which the heating medium used is a substance in the form of a gas/liquid mixture, and in which, following heat transfer in the heat exchanger, this substance is still in the form of a gas/liquid mixture.
Performing a condensation is particularly advantageous since it means that the coefficient of heat transfer to the process gases from the heating medium that can be achieved becomes particularly high and hence that it is possible to achieve efficient heating.
Performing a partial condensation is particularly advantageous because it means that the delivery of heat by the heating medium no longer results in a temperature change to the heating medium, but instead merely shifts the gas/liquid equilibrium. As a result of this, the process gas is heated towards a constant temperature over the entire heat exchange zone. This in turn reliably prevents the incidence of radial temperature profiles in the flow of the process gases, thereby improving the control of the reaction temperatures in the reaction zones and, in particular, preventing the development of instances of local overheating as a result of radial temperature profiles.
In an alternative embodiment, instead of a condensation/partial condensation, it is also possible to provide a mixing zone before the entry of a reaction zone, in order to unify any radial temperature profiles formed in the course of heating in the flow of the process gases by mixing transverse to the main flow direction.
In one preferred embodiment of the process the succession of reaction zones are operated with an average temperature rising or falling from reaction zone to reaction zone. This means that, within a sequence of reaction zones, the temperature may be made both to rise and to fall from reaction zone to reaction zone. This can be brought about, for example, by means of control of the heat exchange zones inserted between the reaction zone. Further possibilities for setting the average temperature are described below.
The thickness of the flow-traversed reaction zones may be the same or different and is a function of laws which are general knowledge to the person skilled in the art and relate to the above-described residence time and the volumes of process gas processed in each case. The mass flows of process gas that can be processed in accordance with the invention, relative to the mass of catalyst used (and also called WHSV, Weight-Hourly Space Velocity), is typically between 28 and 42 h⁻¹, preferably between 30 and 40 h⁻¹, more preferably between 33 and 38 h⁻¹.
The maximum exit temperature of the process gas from the first reaction zone is typically in the region of the entry temperature, since the reactions according to formulae (III) and (W) are exothermic reactions. Particularly in the case of exit from the final reactions, in which a large amount of benzene has already been formed and therefore the especially endothermic reaction according to formula (I) is losing its influence, these temperatures may also be situated within a range from 770 to 820 K, preferably from 775 to 795 K, more preferably from 780 to 785 K.
The person skilled in the art is able freely to determine the entry temperature of the subsequent reaction zones by means of the measures below in accordance with the process of the invention.
Temperature control in the reaction zones is accomplished preferably by at least one of the following measures: sizing 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 reactants/excess of hydrogen used, addition of secondary constituents, especially nitrogen, carbon dioxide, ahead of and/or between the reaction zones.
The composition of the catalysts in the reaction zones of the invention may be the same or different. In one preferred embodiment the same catalysts are used in each reaction zone. An advantageous alternative is to use different catalysts in the individual reaction zones.
Thus it is possible in particular in the first reaction zone, when the concentration of the reactants is still high, to use a less active catalyst, and to increase the activity of the catalyst from reaction zone to reaction zone in the further reaction zones. The catalyst activity can also be controlled by dilution with inert materials and/or support material.
With the process of the invention it is possible to prepare 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 per kg of catalyst.
The process of the invention is therefore distinguished by high space-time yields, in conjunction with a reduction in apparatus sizes and also with a simplification of the apparatus and/or reactors. This surprisingly high space-time yield is made possible through the interaction of the inventive and preferred embodiments of the new process. The interaction of staggered adiabatic reaction zones with interposed heat exchange zones and the defined residence times, in particular, allows precise control of the process and the resulting high space-time yields, and also a reduction in the by-products formed, such as carbon dioxide, for instance.

The present invention is illustrated with reference to the figures, though without being restricted thereto.

FIG. 1 shows reactor temperature (T) and molar mass flows of benzene (U) over a length (L) of 11 m of reaction zones each with downstream heat exchange zones (in accordance with Example 1), the lengths of the heat exchange zones being assumed ideally to be zero, since no conversion is to take place here.

The present invention is further illustrated by the following example, but without being limited thereto.

EXAMPLES

Gaseous naphtha and hydrogen are supplied to the process as process gases in a molar ratio of 7.77. The process is operated in a total of six fixed catalyst beds of rhenium and platinum, each at 0.29% by weight, on an alumina support, in other words in six reaction zones.
After each reaction zone there is a heat exchange zone in which the exiting process gas is heated again before entering the next reaction zone.
The absolute entry pressure of the process gas directly ahead of the first reaction zone is 25 bar. The length of the fixed catalyst beds, and therefore of the reaction zones, varies from reaction zone to reaction zone, beginning from 0.15 m in the first reaction zone through to 6 m in the sixth reaction zone. The precise links of the reaction zone are summarized in Table 1. The activity of the catalyst used is unvarying over the reaction zones. No process gas is metered in ahead of the individual reaction zones. The WHSV is 35 h⁻¹.

TABLE 1

Lengths of the reaction zones

	Reaction	Length
	zone [#]	[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. 1. Varying the cumulative length of the reaction zones is plotted on the x-axis, so that it is possible to see a spatial course of the developments in the process; the heat exchange zones are disregarded. On the left-hand, y-axis, the temperature of the process gas is indicated. The temperature profile over the individual reaction zones is depicted as a thick, continuous line. As a result of the idealized assumption of the length of the heat exchange zones as 0 m, there are discontinuities in the temperature profile. On the right-hand y-axis the cumulative molar flow of benzene in the process gas over the reaction path is indicated. Its profile over said path is depicted as a thin continuous line.
It can be seen that the entry temperature of the process gas ahead of the first reaction zone is approximately 775 K. As a result of the substantially endothermic reaction to form benzene under adiabatic conditions the temperature in the first reaction zone drops to about 760 K, before in the downstream heat exchange zone the process gas is reheated to the aforementioned 775 K. As a result of endothermic adiabatic reaction, the temperature in the second reaction zone drops to about 750 K. The sequence of cooling as a result of endothermic, adiabatic reaction, and heating continues, with changes in exit temperatures, after the respective reaction zones, the entry temperature being re-established at the desired 775 K in each of the heat exchange zones.
A conversion of cyclohexane and hexane of approximately 60% is obtained. The space-time yield achieved, based on the mass of catalyst employed, is approximately 15 kg_benzene/kg_cath.

Claims

1. Process for preparing benzene from naphtha in the presence of hydrogen in an endothermic, heterogeneously catalytic gas phase reaction, which comprises 5 to 12 serial reaction zones with adiabatic conditions.

2. Process according to claim 1, wherein the conversion takes place in 6 to 10 serial reaction zones.

3. Process according to claim 1, wherein the entry temperature of the process gas entering the first reaction zone is 740 to 790 K.

4. Process according to claim 1, wherein the absolute pressure at the entry of the first reaction zone is between 10 and 40 bar.

5. Process according to claim 1, wherein the residence time of the process gas in all reaction zones is between 0.5 and 30 s.

6. Process according to claim 1, wherein the catalysts are present in a fixed bed arrangement.

7. Process according to claim 6, wherein the catalysts are present in the form of monoliths.

8. Process according to claim 7, wherein the monolith comprises channels having a diameter of 0.1 to 3 mm.

9. Process according to claim 1, wherein the catalysts are present in beds of particles having average particles sizes of 1 to 10 mm.

10. Process according to claim 1, wherein at least one reaction zone is followed by at least one heat exchange zone through which the process gas is passed.

11. Process according to claim 10, wherein each reaction zone is followed by at least one heat exchange zone through which the process gas is passed.

12. Process according to claim 1, wherein between a reaction zone and a heat exchange zone there is at least one heat insulation zone.

13. Process according to claim 12, wherein around each reaction zone there is a heat insulation zone.