EP2288587A2

EP2288587A2 - Method for producing phthalic acid anhydride

Info

Publication number: EP2288587A2
Application number: EP09757195A
Authority: EP
Inventors: Evin Hizaler Hoffmann; Leslaw Mleczko; Ralph Schellen; Stephan Schubert; André Hoffmann
Original assignee: Bayer Technology Services GmbH
Current assignee: Bayer Intellectual Property GmbH
Priority date: 2008-05-29
Filing date: 2009-05-26
Publication date: 2011-03-02
Also published as: WO2009146812A2; WO2009146812A3; US20110130573A1; DE102008025843A1

Abstract

The present invention relates to a method for the production of phthalic acid anhydride by catalytic gas phase oxidation of o-xylol with oxygen, wherein the reaction is carried out in 5 to 60 serially arranged reaction zones under adiabatic conditions, and a reactor system for carrying out said method

Description

Process for the preparation of phthalic anhydride

The present invention relates to a process for the preparation of phthalic anhydride by catalytic gas phase oxidation of O-xylene with oxygen, wherein the reaction is carried out in 5 to 60 successive reaction zones under adiabatic conditions, and a reactor system for carrying out the method.

Phthalic anhydride is generally subject to the catalytic influence of metal oxide catalysts e.g. Vanadium pentoxide prepared from gaseous O-xylene and oxygen in an exothermic catalytic reaction according to formula (I):

The phthalic anhydride prepared by the reaction of formula (I) is often used as a starting material for the preparation of plasticizers (usually phthalate) or as a raw material for the production of synthetic resins for the surface coatings of wood. In addition, it is a raw material in the production of dyes or color pigments based on the phthalocyanines. Another technically important reaction of phthalic anhydride is the anthraquinone.

The removal and use of the heat of reaction is important in carrying out phthalic anhydride synthesis. An uncontrolled increase in temperature can lead to permanent damage to the catalytic converter. In addition, at high temperatures, there is the possibility of side reactions to produce more or less large quantities of maleic anhydride, as well as carbon monoxide and carbon dioxide. It is therefore advantageous to maintain the temperature of the reaction zones in the process in a controlled manner at a level which allows for rapid conversion while minimizing side reactions and / or catalyst deactivation.

US Pat. No. 6,380,399 B1 discloses a process for the preparation of phthalic anhydride from O-xylene or naphtha, in which at least three catalyst stages which form a

Reaction zone can be summarized in a fixed bed reactor sales at

Temperatures between 300 ° C and 400 ° C is performed. The proportion of oxygen the process gases at the beginning of the process is between 10 and 21% by volume. Based on this, the concentration of O-xylene or naphtha is at least 70 g / Nm ³ . The catalyst in the presence of which the conversion is carried out preferably comprises vanadium oxide and titanium oxide. It is further disclosed that the reaction zones are cooled. No separate heat exchange zones are disclosed.

The process disclosed in US Pat. No. 6,380,399 B1 is disadvantageous because the reaction zones are actively cooled and an attempt is made to prevent overheating of the reaction zones by adjusting the catalyst activity by dilution with non-catalytically active material. The fact that the reaction zones are cooled directly means that due to the geometry of the reaction device and the heat exchange medium used, only a maximum heat exchange rate can be achieved, which can not be easily changed during operation of the process. In turn, this results in it, e.g. By feeding in too much starting material in the context of a valve misalignment, operating states can occur in which overheating of the process in the "hot spots" can no longer be intercepted, which leads at least to a loss of the target product, but can also lead to a loss of catalyst activity and thus to the necessity of its renewal, or even destruction of the device in which the reaction zone is located, lead.

US Pat. No. 6,774,246 B2 discloses a process similar to the disclosure of US Pat. No. 6,380,399 B1, in which phthalic anhydride is prepared from O-xylene or naphtha in two

Reaction zones is prepared, wherein the reaction zones comprise fixed beds, which are cooled. The total pressure of the process can be between 0.1 and 2.5 bar, while the temperatures at which the process gases enter the process can be between 300 ° C and 450 ° C. The process gas may comprise 1 to 100 mol% oxygen and have a concentration of O-xylene or naphtha between 60 and 120 g / Nm ³ . As in US Pat. No. 6,380,399 B1, no heat exchange zones separated from the reaction zones are disclosed.

Accordingly, the process disclosed in US 6,774,246 B2 is disadvantageous for the same reasons as the process disclosed in US 6,380,399 B1. In addition, the use of a maximum of two reaction zones allows an even lower temperature control. 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 for the synthesis of phthalic anhydride from gaseous oxygen and O-xylene. Thus, it 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. 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 there is no subsequent addition of educts and thus after exothermic Abreaktion the reaction slower and thus would reduce the heat of reaction generated. 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. Based on the prior art, it would therefore be advantageous to provide a method that can be carried out in simple reaction devices and that 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 O-xylene with oxygen to phthalic anhydride, as described above, neither suitable reactors have yet been described, nor are suitable processes disclosed which permit them.

It is therefore an object to provide a process for the catalytic gas-phase oxidation of O-xyI ol and oxygen to phthalic anhydride, which is feasible under precise temperature control in simple reaction devices and thereby high conversions at high purities of the product allowed, the heat of reaction either in favor of Reaction or otherwise can be used.

It has surprisingly been found that a process for the preparation of phthalic anhydride from O-xylene and oxygen in the presence of heterogeneous catalysts, characterized in that it comprises 5 to 60 successive reaction zones with adiabatic conditions, is capable of solving this problem.

O-xylene in the context of the present invention refers to a process gas which is introduced into the process according to the invention and which comprises O-xylene. The proportion of O-xylene on the process gases supplied to the process is usually between 0.8 and 10 mol%, preferably between 1 and 7 mol%.

Oxygen, in the context of the present invention, denotes a process gas which is introduced into the process according to the invention and which essentially comprises oxygen. Preferably, oxygen is ambient air and therefore comprises a proportion of about 20% by volume of oxygen.

In addition to the essential components of the process gases O-xylene and oxygen, these may also include secondary components. Non-exhaustive examples of minor components which may be included in the process gases include argon, nitrogen and / or carbon dioxide. Generally, in the context of the present invention, process gases are understood as gas mixtures which comprise oxygen and / or o-xylene and / or phthalic anhydride and / or secondary components.

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, 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 60 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 single reaction zone to increase the conversion in a controlled manner.

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 customary

Catalysts consisting of a material which, in addition to its catalytic activity for the reaction according to formula (I), has sufficient chemical resistance among the

Conditions of the process, as well as by a high specific surface area ,

Marked are. Catalyst materials characterized by such chemical resistance under the conditions of the process are, for example, catalysts comprising mixed oxides of vanadium, as well as oxides and / or salts of the elements selected from the list of elements containing Nb, Sb, P, K, Na, Cs, Rb and Mo include. Preferred are catalysts comprising mixed oxides of vanadium with P and Rb. These catalysts can be applied to support materials. Such support materials usually include alumina, silica and / or titania. Preferred are support materials of titanium dioxide.

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 10 m ² / g, preferably of at least 20 m ² / g.

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

Preferred are the forms of fixed bed and fluidized 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 4 to 7 mm. Also preferably, the catalyst is in a fixed bed arrangement in monolithic form. Particularly preferred in a fixed bed arrangement is a monolithic catalyst comprising mixed oxides of vanadium and phosphorus supported on titania.

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.

A monolithic catalyst with channels of the specified diameter is particularly advantageous, since this explosion protection can be ensured. This is done by absorbing the enthalpy through the wall of the monolith, thus suppressing the propagation of flames.

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 area of the catalyst material compared to the process gases oxygen and O-xylene 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 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 30, more preferably 7 to 20 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 described, for example, in Ullmann's Encyclopedia of Industrial Chemistry (Fifth, Completely Revised Edition, VoI B4, pages 95-104, pages 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. Especially with thin fixed beds can in the flow direction before the

Catalyst beds technical devices for uniform gas distribution can be attached. 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 method, the inlet temperature of the process gas entering the first reaction zone is from 10 to 490 ° C., preferably from 150 to 480 ° C., particularly preferably from 300 to 470 ° C.

In a further preferred embodiment of the method, the absolute pressure at the inlet of the first reaction zone is between 1 and 10 bar, preferably between 1.1 and 3 bar, more preferably between 1.2 and 1.5 bar.

In yet another preferred embodiment of the method, the residence time of the process gas in all reaction zones together is between 0.05 and 25 s, preferably between 0.1 and 10 s, particularly preferably between 0.15 and 3 s.

The O-xylene and the oxygen 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 one or more of the reaction zones following the first reaction zone as required into the x-xylene and / or oxygen in the process gas. In addition, the temperature of the conversion can be controlled via the supply of gas between the reaction zones.

In a preferred embodiment of the method according to the invention, the process gas is particularly preferred after at least one of the reaction zones used cooled after each reaction zone. 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 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, steam is generated during cooling of the process gas in the heat exchange zones by the heat exchanger.

Within this particular embodiment, it is preferable in the heat exchangers including 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 equilibrium 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 safely prevents the occurrence of temperature profiles in the flow of Process gases, whereby the control over the reaction temperatures in the reaction zones is improved and in particular the formation of local overheating is prevented by temperature profiles.

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

In a preferred embodiment of the process, the reaction zones connected in series are operated at an average temperature increasing or decreasing from reaction zone to reaction zone. This means being within one

Result of reaction zones, the temperature of reaction zone to reaction zone can both rise and fall. 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 mass flows of product gas (phthalic anhydride) which can be carried out according to the invention by the process, and from which the amounts of process gas to be used, are usually between 0.01 and 35 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 400 ° C to 520 ° C, preferably from 420 ⁰ C to 510 ° C, more preferably from 430 ⁰ C to 500 ⁰ C. The control of the temperature in the

Reaction zones are 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 starting materials / excess of oxygen 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 the catalyst activity can also be carried out by dilution with inert materials or carrier material. 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.

0.01 kg / h to 1 kg / h, preferably 0.02 kg / h to 0.75 kg / h, more preferably 0.03 kg / h to 0.4 kg / h per 1 kg of catalyst with the inventive method phthalic anhydride.

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 inventive 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 in the by-products formed, such as maleic anhydride and CO ₂ .

Another object of the invention is a reactor system for the reaction of O-xylene and oxygen to phthalic anhydride, characterized in that there are leads (Z) for a process gas comprising O-xylene and oxygen or for at least two process gases, of which at least one O-xylene and comprising at least one oxygen and comprises 5 to 60 successive reaction zones (R) in the form of fixed beds of a heterogeneous catalyst, wherein 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 on and Derivatives 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 30, preferably 7 to 20 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 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 numerals being used in the figures:

Z: supply line (s)

R: reaction zone (s) I: heat insulation zone (s)

W: heat exchange zone (s)

2 shows reactor temperature (T), O-xylene conversion (U) and phthalic anhydride selectivity (Y) over a number of 8 reaction zones (S) with downstream heat exchange zones (according to Example 1).

3 shows reactor temperature (T), O-xylene conversion (U) and phthalic anhydride selectivity (Y) over a number of 12 reaction zones (S) with downstream heat exchange zones (according to Example 2).

4 shows reactor temperature (T), O-xylene conversion (U) and phthalic anhydride selectivity (Y) over a number of 18 reaction zones (S) with downstream heat exchange zones (according to Example 3).

The present invention will be further illustrated by the following Examples 1 to 3, without being limited thereto. Examples

Example 1:

In this example, the process gas flows over a total of 8 fixed catalyst beds of titanium dioxide coated with vanadium pentoxide, ie through 8 reaction zones. Each after a reaction zone is a heat exchange zone in which the process gas was cooled before it enters the next reaction zone. The process gas used at the beginning of the first reaction zone contains 0.94 mol% of O-xylene, 20.79 mol% of oxygen, and 78.27 mol% of inert gases (nitrogen, CO ₂ , argon). With a share of 0.94 mol% of O-xylene, the proportion is below the limit for the preservation of a potentially ignitable mixture (1 mol% with air), so that it is not necessary to pay attention to the exceeding of a potential ignition temperature. The absolute inlet pressure of the process gas directly in front of the first reaction zone is 1.5 bar. The length of the fixed catalyst beds, ie the reaction zones, is always 0.1 m, except for the last reaction zone whose length is 0.14 m. The activity of the catalyst used is not variable across the reaction zones. There is no metered addition of gas before the individual reaction zones. The total residence time in the system is 0.2 seconds.

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 O-xylene, and the selectivity of phthalic anhydride is given. The course of the conversion over the individual reaction zones is shown as a thick dashed line. The course of selectivity, as a thin solid line.

It can be seen that the inlet temperature of the process gas before the first reaction zone is about 420 ° C. Due to the exothermic reaction to phthalic anhydride under adiabatic conditions, the temperature in the first reaction zone rises to about 490 ° 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 420 ⁰ C. By exothermic adiabatic reaction, it rises again to about 49O ⁰ C. The sequence of heating and cooling continues. The inlet temperatures of the Process gas upstream of the individual reaction zones changes in the course of the process to a value of about 480 ° C.

There is obtained a conversion of O-xylene of 99 mol%. The selectivity is obtained at 85.4 mol%. The space-time yield obtained based on the mass of catalyst used is 0.19 kgp _hΛa i _säureanhyd ri _d / kgκ _at h.

Example 2:

In this example, the process gas flows through a total of 12 reaction zones, so over 12 fixed catalyst beds of titanium dioxide coated with vanadium pentoxide. Each after a reaction zone is a heat exchange zone in which the process gas is cooled and in which a further addition of O-xylene before reaction zones according to Table 1 is carried out. This metered addition is controlled in such a way that the ignition-critical limit of 1 mol% of O-xylene is not reached, so that care must be taken not to exceed a potential ignition temperature. A precise statement of the proportions of O-xylene based on the total amount supplied to the process is given in Table 1. TABLE 1 Allocation of the addition of O-xylene in accordance with Example 2

The process gas used at the outset, as well as the inlet pressure before the first reaction zone, is identical to that of Example 1. In the metered additions, streams of pure gaseous O-xylene are used to at least partially replace the amount consumed. The volume and mass flows of the metered additions thus result from the proportion according to Table 1, as well as the volume flow of the feed and the concentration of O-xylene in the feed. The length of the fixed catalyst beds, so the Reaction zones, is always 0.1 m, except for the last reaction zone whose length is 0.18 m. The activity of the catalysts is not variable across the reaction zones. This is achieved after the first 6 reaction zones oscillating in a temperature window between 415 ° C and 490 ° C. The residence time in the system is 0.3 seconds in total.

It can be seen that the inlet temperature of the process gas before the first reaction zone is about 420 ° C. Due to the exothermic reaction to phthalic anhydride under adiabatic conditions, the temperature in the first reaction zone rises to about 490 ° C, before the process gas is cooled in the downstream heat exchange zone again. The cooling is also achieved here by the added at a lower temperature O-xylene. The inlet temperature before the next reaction zone is about 415 ⁰ C. By exothermic adiabatic reaction, it rises again to about 490 ° C. The sequence of heating and cooling continues. The inlet temperature of the process gas upstream of the individual reaction zones changes significantly from the 7th reaction zone. Here, further heating up to about 46O ⁰ C is allowed before the last reaction zone.

There is obtained a conversion of 99 mol% of the O-xylene used, calculated from the remaining mass output of the last reaction zone. The selectivity is about 84.7 mol% with respect to phthalic anhydride. The space-time yield obtained based on the mass of catalyst used is 0.25 kgphth _a i _acid a _nhydπd / kg _cat h. Example 3:

In this example, the process gas flows over a total of 18 fixed catalyst beds in the form of monoliths with channel diameters of monoliths of 1 mm, which are coated with a catalyst comprising vanadium pentoxide on a support of titanium dioxide, ie by 18 reaction zones. Each after a reaction zone is a heat exchange zone in which the process gas was cooled before it enters the next reaction zone. The process gas used at the beginning of the first reaction zone contains 2 mol% of O-xylene, 20.56 mol% of oxygen and 77.44 mol% of inert gases (nitrogen, CO ₂ , argon). With a content of 2 mol% O-xylene, the proportion is above the limit for the preservation of a potentially ignitable mixture (1 mol% with air), so that attention must be paid to exceeding a potential ignition temperature (450 ° C). After the seventh reaction zone enough O-xylene has already been degraded, so that higher temperatures are allowed in the other reaction zones. The absolute inlet pressure of the process gas directly in front of the first reaction zone is 1.5 bar. The length of the fixed catalyst beds, ie the reaction zones, is always 0.5 m, except for the last reaction zone with a length of 0.74 m. The activity of the catalyst used is not variable across the reaction zones. There is no metered addition of gas before the individual reaction zones. The total residence time in the system is 1, 6 seconds.

It can be seen that the inlet temperature of the process gas before the first reaction zone is about 400 ° C. Due to the exothermic reaction to phthalic anhydride under adiabatic conditions, the temperature in the first reaction zone rises to about 440 ° C, before the process gas is cooled in the downstream heat exchange zone again. The inlet temperature before the next reaction zone is about 394 ° C. By exothermic adiabatic reaction increases again to about 440 ° C. The sequence of heating and cooling continues until the exit of the sixth reaction zone. In the following heat exchange zone, less cooling takes place so that the inlet temperature of the seventh reaction zone is about 435 ° C. By exothermic adiabatic reaction, this rises to about 490 ⁰ C. The sequence of heating and cooling continues, with a slow rise in the inlet temperatures is up to 475 ° C at the beginning of the last reaction zone is tolerated.

There is obtained a conversion of O-xylene of 99 mol%. The selectivity is obtained at 84.4 mol%. The space-time yield achieved based on the mass of catalyst used is 0.034 kgp _ht haisäu _rean hy _d rid / kgκath.

Claims

claims

1. A process for the preparation of phthalic anhydride from O-xylene and oxygen in the presence of heterogeneous catalysts, characterized in that it comprises 5 to 60 successive reaction zones with adiabatic conditions.

2. The method according to claim 1, characterized in that the conversion takes place in 6 to 30, preferably 7 to 20 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 10 to 490 ° C, preferably from 150 to 480 ° C, more preferably from 300 to 470 ° C.

4. The method according to claim 1 to 3, characterized in that the absolute pressure at the entrance of the first reaction zone is between 1 and 10 bar, preferably between 1.1 and 3 bar, more preferably between 1.2 and 1.5 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 zone between 0.05 and 25 s, preferably between 0.1 and 10 s, more preferably between 0.15 and 3 s.

6. The method according to any one of the preceding claims, characterized in that the catalysts mixed oxides of vanadium, and oxides and / or salts of the elements selected from the list of elements containing Nb, Sb, P, K, Na, Cs,

Rb and Mo include.

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

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

9. The method according to claim 8, 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.

10. The method according to any one of claims 1 to 6, characterized in that the catalysts are in fluidized bed arrangement.

11. The method according to claims 1 to 7 or 10, 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 from 4 to 7 mm.

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 there is at least one heat insulation zone between a reaction zone and a heat exchange zone.

15. The method according to claim 14, characterized in that around each

Reaction zone is a heat insulation 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 O-xylene and oxygen or for at least two process gases, of which at least one O-xylene and at least one oxygen and 5 to

Comprises 60 successive reaction zones (R) in the form of fixed beds of a heterogeneous catalyst, wherein 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 on and Derivatives for the process gases are connected and include the supply and discharge lines for a cooling medium.