Title of Invention	PROCESS FOR THE DEHYDROGENATION OF ETHYLBENZENE TO STYRENE
Abstract	A process is described for the dehydrogenation of ethylbenzene lo styrene in a fluid - bed reactor-regenerator system, which uses a catalyst based on iron oxide supported on a modified alumina and promoted with further metal oxides.

Full Text

The present invention relates to a process for the dehydrogenaticn of ethylbenzer.e to styrer.e in a fluid-bed reactor/regenerator system, in the presence of a catalyst based on an iron oxide and further promoters, selected, e.g., from metal oxides such as alkaline ox¬ides, earth-alkaline metal oxides and/or oxides of the metals of the group of lanthanides, supported on a modi¬fied alumina. Styrene is an important intermediate which can be used in the preparation of plastic materials and rub¬bers . More specifically, styrene is used for the produc¬tion of polystyrenes (GF?S crystals, high impact HIPS and expandable EPS) ( acryionitriie-styrene-butaciene (A3S) and styrene-acryicr.ir.riie (SAX; copolymers and The dehydroger.ation. reaction cf ethyibenzene to styrer.e has a few particular characteristics which should be taken in asccunt for the technological design. The first lies in the fact that the reaction is ccntrolled by thermodynamic equilibrium and ccr.ssequentiy the conversion per passage is not total. The dehydroce- naticn decree increases wi- a r;s= m tne temoerature ana witn a oeorease in tne tota_ pressure, tne reaction taking place at a constant pressure, with an increase in the volume, in order to obtain economically acceptable conversions, it is therefore necessary to carry cut the reaction at temperatures generally ranging from 540 to j630°C. The use of high temperatures however stimulates side reactions characterized by a greater activation en¬ergy with respect to the dehydrogenation value. As a re¬sult of this, more or less significant quantities cf by¬products mainly consisting cf toluene, benzene, coke and light products are formed together with the main prod¬uct. heat equal to 28 Kcal/moles of styrene corresponding to 270 Kcal/kg of styrene produced. The high heat required and the high thermal levels at which it must be exchanged greatly influence the technological design. The technologies at present commercialized (Final Badger and Lummus/UOP Classic SM processes) satisfy the demands imposed by the thermodynamics of the reaction by means of processes which use a bulk catalyst prevalently based on iron oxide and promoted with alkalis, and which comprise the use of: -several adiabatic reactors in series, with interme- diate heating steps at a temperature ranging from 540°C to 630°C and with contact times in the order of tenths of a second; -radial flow reactors which operate under vacuum at a pressure ranging from 30.39 to 50.65 Kpa (absolute pascal); and -water vapor which is fed with the charge to be de-hydrogenated. Water is the main component in the charge fed to the reactor. The typical molar concentration is 90%, even if higher concentrations are often adopted to lengthen the chemical life of the catalyst. The vapor has the function of: - reducing the partial pressure of the products and therefore favorably shifting the thermodynamic equilib¬rium; - contributing to the removal of the coke which is de¬posited or. the surface of the catalyst, there being no regeneration cf the catalyst with air; - supplying the heat necessary for the dehydrogenatior. of the ethvlbenzen.e; - slowing down the aging of the catalyst. Operating with these technologies, conversions ranging from 60 to 65= are obtained, with a selectivity to styrene higher than 90S, by weight. These processes however have the following disad-vancages: - use cf large quantities cf vapor (K.0/E3 = 9.0-10 mo¬lar) superheated with temperatures higher than 700°C; this impels the use of superheating ovens and therefore high investment costs; - aging of the catalyst and consequently the necessity of substituting it after 13-36 months of operation; this involves stopping the unit and consequently interrupting the production for the period necessary for substituting the catalvst; por heat and not that of the latent heat: -carrying out the reaction under vacuum (average absolute pressure of 40.52 Kpa (absolute pascal) ) and therefore in extremely diluted phase in EB; the partial EB pressure is on an average equal to 4.052 Kpa (absolute pascal). It has now been found that it is possible to over- come these drawbacks by means of a process which uses a fluid-bed reactor/regenerator system and a cataiyst based on iron oxide supported on a microspheroidai alumina modified with silica and further metal oxides as promoters. The process of the present invention has consider- abie economic advantages, in particular: -thermal profile of the reactor favorable for the reaction thermodynamics; -the heat is directly transferred to the reaction by the regenerated catalyst, superheating ovens are there- fore not required for the thermal exchange and the strong remixing of the fluidized bed prevents the formation of hot spots which would lower the selectivity; -the possibility of recycling the hydrogen; -the plant can be run with great flexibility in terms of actual productive capacity with respect to that projected; -the dehydrogenation reaction and the regeneration take place in physically separated zones; this avoids the mixing of hydrocarbon streams with oxygen streams; - the process is carried cut at a pressure which is at¬ mospheric or slightly higher; as a result, there are no ' air infiltrations frtrr. the outside in the reaction 2one; - the molar concentration of the inert gas/ethylbenzene in the feeding is r.uch lower with respect to commercial technologies; - it is not necessary to effect any specific treatment for reducing the emissions of gas pollutants; and - the possibility of operating without water vapor with¬ out there being any chemical deterioration in the catalyst. Japanese patent application 7-323,439 discloses a process for the dehydrocer.aticn of ethylbenzene in the presence of a catalyst which consists of alumina, carry¬ing a complex of pouassiun ferrate and possibly rare earth metal oxides, modified by basic metallic oxide ad¬dition. Said catalyst shows activity when operated in the 'presence of water but no data are given in said pat¬ent application about the performances of the catalyst in the absence of water, as co-feed of ethyibenzer.e, nor the effect of ageing is detailed. Surprisingly, it was found that through a partial modification of the alumina carrier cy silica it was possible to improve signifi- cantly the catalytic performances in dehydrogenation yield, with evident advantages. In the same time, the mechanical resistance of the cataWst itself is improved by silica modification making it more suited to fluid bed operations. Furthermore, the catalyst is also able to operate with nitrogen other than with water. In accordance with this, the present invention relates to a process for debydrogenating ethylbenzene to styrene which essentially consists in: (a) reacting ethylbenzene mixed with an inert product, in a fluid-bed reactor, in the presence of a catalytic system consisting of iron oxide and promoters supported on alumina modified with 0.01-10% by weight of silica and operating at a temperature ranging from 400 to 700CC, at a total pressure of 10.13 to 303.9 Kpa (absolute pascal) and with a GHSV space velocity ranging from 50 to 10,000 h-1 (normal liters of the mixture ethylbenzene and inert gas/h x liter of catalyst); and (b) regenerating the catalyst in a regenerator by burning the coke deposited on its surface at a temperature exceeding 400°C. The catalytic system used in the process of the present invention consists of: (1) 1-60% by weight, preferably 1-20%, of iron oxide; (2) 0.1-20% by weight, preferably 0.5-10%, of at least one alkaline or alkaline earth metal oxide; (3) 0-15% by weight, preferably 0,1-7% of a second pro¬moter consisting of at least one rare earth oxide; (4) the complement to 100 being a carrier consisting of a microspheroidal alumina with a diameter selected from those in delta, theta phase or their mixtures, in theta + alpha phase or delta + theta + alpha phase, modified preferably with 0.09-51 by weight of silica. The carrier has an average particle diameter and particle density such that the final produce can be classified as Group-A according to Geidart (Gas Fluidi-zaticr. Technology, D. Geidart, John Wiley £ Sons) and a surface area of less than 150 nr/g (BET). Alkaline metal preferably used as first promoter in the present invention is potassium. Preferred second promoters belonging to the rare earth metals are cerium, lantar.iun and praseodymium. An example of catalyst according to the present in¬vention consists of: (1) 5-50% by weight of iron oxide; (2) 0,5-10° by weight of a promoter expressed as oxide; (3) the complement to 1C0 being a carrier consisting of a microspheroidal alumina with a diameter ranging from 50 to 7 0 microns selected from those in delta, theta phase or their mixtures, ir. theta - alpha phase or delta + theta + alpha phase, modified preferably with 0.08-3?o by weight of silica. The process for preparing the catalytic system described above can be essentially carried out by means of the following steps: - preparation of solutions based, on derivatives of the components cf the catalytic system; - dispersion of the solutions on carriers as defined above; - drying cf the solids obtained; - calcination of the dried solids at a temperature rang¬ing from 500 to 900°C. The dispersion of the catalyst components on the carrier can be carried out using conventional techniques such as impregnation, ion exchance, vapor deposition or surface adsorption. The "incipient wetness" impregnation technique is preferably used. According to a preferred embodiment, the catalyst is prepared by: (a) addition cf an aliquot of the promoters to the car¬rier; i'o) drying at 100-150°C and, optionally, calcination cf the dried solid at a temperature not exceeding 900°C; (c) dispersion of the iron oxide and remaining aliquot of the promoters on the modified carrier (a); (d) drying at !00-150°C and calcination of the dried solid at a temperature ranging from 500 to 900°C. Steps c) and d) can be repeated several times. Nitrogen, methane, hydrogen or water vapor can be used as the gaseous inert product, in a volumetric ratio inert gas/ethylbenzene ranging from 1 to 6, preferably from 2 to 4. Methane and nitrogen are preferably used. ' According to a further embodiment of the process of the present invention, the ethylbenzene can be co-fed to the reactor with a paraffin selected from ethane, pro¬pane, isobutane, in order to obtain the contemporaneous dehydrogenation of the co-fed products to give styrene and the corresponding olefins respectively. In particular when the ethylbenzene is fed with ethane, the process can be carried out as described in in U.S. patent 6,031,143. According to our further embodiment of the process of the present invention, ethylene can be recycled to an alkylation unit together with a stream of benzene to give ethylbenzene. In the reactor-regenerator system, the catalyst circulates continuously, in fluidized state, between the reactor and regenerator, thus allowing the process to be carried out in continuous. The heat necessary for the reaction is provided by . the regenerated catalyst which reaches the reactor at a temperature higher than the average reaction temperature. The catalyst is maintained in a fluicized state in the reactor by the reagent mixture (inert gas/ethylbenzer.e), which enters the catalytic bed from below, by means of an appropriate distribution system." The reacted gas, after passing through a system of cyclones or another powder separation system, leaves the reactor f ron above. The gas can then be sent to a heat exchanger fcr the preheating of the feeding and subse¬quently to the separation section where the styrene pro¬duced is recovered, whereas the non-reacted charge is recycled to the behydrogenaticn reactor and the reaction by-products (light hydrocarbons ar.d hydrogen- are recov¬ered and used in the regenerator as fuel gas. The catalyst moves in fluidized state in the reac¬tor, in countercurrent with respect to the gas phase. It enters the catalytic bed from above, through a distribu¬tor which disperses it equally on the surface of the bed, and it leaves the reactor from below, passing by gravity into a desorption zone where the moving 'and de¬sorption of the intraparticle gas take place, nitrogen or methane being introduced iron below, so that the moved or desorbed gas re-enters the reactor, thus avoid¬ing losses in reagents or products. It is preferable to operate in the fluid-bed reac¬tor as follows : - at a temperature ranging from 450 to 650°C in relation to the desired reaction; the temperature is maintained within the cre-selected values by regulating the flow-rate of the regenerated catalyst; - at a pressure which is atmospheric or slight higher; - a' a GKSV space velocity ranging from 100 and 1000 h" ", preferably rrom I5C to 200 h~:; and - with a residence time of the catalyst in the fluid bed ranging from 5 to 31 minutes, and in the cescrption zone from 0.2 to 1C nmut.es. According to an embodiment cf the process cf the present invention, grids can be horitontally arranged inside the reactor, at a distance cf 20 to 200 cm from each other, and with a free area ranging from 10 to 90%, preferably front 20 to 40°o. The purpose of the grids is to prevent the gas and catalyst fro- re-ir.ixi.ng, so that the cas flow inside the reactor resembles a plug-flow. The use cf these grids allows maximisation cf the con¬version of ethylbenzene and selectivity to styrene. The selectivity of the reaction can be further im¬proved by the longitudinal thermal profile which is es¬tablished along the catalytic bed, with the maximum tem¬perature in the upper part -where the regenerated cata¬lyst arrives and the minimum temperature in the lower part. The temperature difference along the bed prefers- bly ranges frcm 15 to 65QC. In order to optimize the longitudinal thermal pro¬file, the regenerated catalyse car. be distributed at various heights of the catalytic bee.. The fluicized catalyst is subsequently sent to the regenerator through a pneumatic transport system con¬sisting cf: - a transport line with at least cr.e zone in which the catalyst moves downwards by the introduction of suitable quantities of gas at appropriate heights, and - a cone in which the catalyst moves upwards until it reaches the upper part cf the catalytic bed, by the in¬troduction of gas at the base of the raiser. The regenerator preferably has similar dimensions to those of the reactor to maintain the catalyst for a period sufficient fcr its regeneration. The regeneration of the catalyst is carried out bv the combusticn cf coke with air and oxygen, whereas its heating is effected with the use of methane, a fuel gas, or by-products of the dehydrogenation reaction, at a temperature higher than the average reaction tempera¬ture . The movement of the gas and solid ta! trcduced at suitable heights along the bed. The gas leaving the regenerator, substantially con¬sisting of nitrogen az-.c. combustion products, is passed through a system of cyclones, or other system, situated in the upper part c: the apparatus, to separate the en¬trained powders and is then sent to a heat exchanger to preheat the combustion air. Before teing discharged into the atmosphere, these gases can be treated with a filter system c: other de¬vices for reducing the powder content to a rev tenths of mg per Mm" of gas. In the regenerator, it is preferable to operate at atmospheric pressure or slightly higher, at a space ve¬locity ranging from 100 to 1,000 h~' and with a residence time cf the catalyst ranging from 5 to 60 minutes, pref¬erably from 20 to 40 minutes- The regenerated and reheated catalyst is sent to the reactor by means of a pneumatic system having the characteristics described above. The use of the reactor-regenerator system has the following advantages: - the possibility of keeping the operating parameters and catalytic performances constant for the whole tech- the regenerated catalyst: there is therefore no need for super-heating ovens for the thermal exchange and the strong re-mixing of the fluid bed prevents the formation of hot spots which would lower the selectivity: -the hydrogen can be recycled; -the process can be carried out in continuous without having to modify the operating parameters during the life of the plant; -the reaction and regeneration take place in physically separated zones so that the hydrocarbon streams do not mix with streams containing oxygen; -the molar concentration inert product/ethylbenzene in the feeding is much lower with respect to the commercial technologies. With reference to figure 1, a possible application of the reactor-regenerator scheme is provided, which uses the catalyst based on supported iron oxide. The liquid stream of ethylbenzene (1), consisting of fresh and recycled feeding, at room temperature and a pressure of 263.38 Kpa (absolute pascal), is vaporized in the evaporator (2). preheated to about 420°C in the gas-gas exchanger (3), mixed in a suitable mixer (4) with a stream (5) prevalently consisting of nitrogen and whose origin is described hereunder, and fed to the reactor (6) by means of an appropriate dist Ibutor situated in the lower part. The stream (7), effluent from the reactor at a temperature of 600°C. at a pressure of 135.74 Kpa (absolute pascal) essentially consisting of nitrogen, styrene, hydrogen and non-reacted ethylbenzene, undergoes a first cooling in the gas-gas exchanger (3) and a second cooling in the gas-gas exchanger (8), from which it flows at a temperature of 320°C. This stream then passes through the filter system (9) to eliminate the fine particles entrained and is subsequently cooled with water to a temperature of 40°C in the exchanger (10). The mixture becomes biphasic at this temperature as a result of the partial condensation of the hydrocarbon. The condensed stream (12) is recovered from the bottom in the phase separator (11), and is sent, like the gas stream (13), to the subsequent recovery and purification zone of the products (14), not shown in detail, where the following streams are recovered, using techniques known to experts in the field: -stream (15) consisting of pure styrene (product); -stream (16) consisting of ethylbenzene, which is recycled to the dehydrogenation; -stream (17) essentially consisting of nitrogen and hydrogen, containing light hydrocarbons; -stream (18) essentially consisting of benzene and toluene; -stream (19) consisting of heavy hydrocarbon by-products. The stream (17). after flushing stream (20), is heated in the gas-gas exchanger (21) up to a temperature of 550°C and fed to the regenerator (22) by means of the distributor (23) situated above the air inlet. The stream of air (24) is compressed in the compressor (25) and preheated to a temperature of 560°C in the gas-gas exchanger (26), before being fed to the regenerator (22). The stream (27) effluent from the regenerator, prevalently consisting of nitrogen and water vapor is subsequently cooled in die exchangers (21) and (26), passes through the filters (28) to eliminate the fine powders entrained and is cooled in the exchanger (29) at 40°C. The stream of condensed water (30) is separated in the vessel (31), whereas the remaining gas stream (32), still containing significant quantities of water vapor, is compressed in the compressor (33) at a pressure of 263.38 Kpa (absolute pascal) and is subsequently cooled in the exchanger (34) at such a temperature as to allow the almost complete condensation of the water present. The condensed stream (35) is removed from the bottom of the vessel (36), whereas the gas stream (37), after a part of it has been flushed (38), is heated in the gas-gas exchanger (8). The resulting stream (5) _s men t re a tec as described above. All the catalytic tests are carried out using a quartz micro-reactor in which abcut 50-1CO mi of cata¬lyst are charged. Z'r.e reactor is heated by an. electric oven in order to keep the catalytic bed at the desired temperature. The ethylbenser.e is fed to an evaporator by means of a dosing pump end is then mixed with the inert gas whose flow-rate is measured by means of a rotameter. 7he reaction mixture is preheated to 2 00°C and fed to the reactor from below through a calibrated septum which acts as gas distributor, thus fluidinir.g the cata¬lyst. A quartz expansion vase is assembled on the head of the reactor, which has the function of decelerating the effluent gas and making the fine catalyst particles fail back into the reactor. The expander and sampling lines are maintained at 200°C to avoid the condensation of styrene, non-reacted ethyibensene and any possible heavy by-prcducts. The catalytic cycle consists of: - a reaction phase, in which the ethylbenzene mixed with the ir.ert product cr with the paraffin, is fed to the reactor over a period of 10 minutes; - a stripping phase, into which nitrogen is passed for about 15 minutes to removed the products adsorbed on the catalyst; .- a regeneration phase, into which air is fed for about 4 5 minutes; and - a washing phase with nitrogen for about 20 minutes. The catalytic cycle was carried out continuously for 100 hours without having any loss of activity of the catalyst.. The dehydrogenation reaction is carried out at 560-650°C, whereas the regeneration is carried cut at 660°C. The overall space velocity, expressed as normal li¬ters of ethylbenzene plus normal liters of inert prod¬uct, is maintained at 300 ± 5 Nl/h/lt of catalytic bed. During the reaction and stripping phase, the efflu¬ent is cooled in a trap immersed in liquid nitrogen in which the non-reacted ethylbenzene, styrene and condens¬able by-products are condensed. The effluent from the trap is sent to a sack from which hydrogen, inert prod¬ucts and C3-C.H light hydrocarbons from cracking reac¬tions, are recovered. The liquid fraction is weighed and analyzed by gas-chromatcgraphy using an H? 5890 gas chromatocraph equipped with a C? WAX 10 capillary column. The dosing of the components is effected using an internal stan- dard. The gas recovered from che sack is analyzed by gas chromatography using the external standard procedure for the dosing of the components. The contents of the sack ere measured with a counter for the material balancing. The coke deposited on the surface of the catalyst is combusted with air and the effluent collected in a sack. The gas is then analyzed by gas chromatography to dose the concentration of CO., whereas the volume is measured to establish the quantity of coke formed curing the reaction. The following examples, whose sole purpose is to describe this invention in greater detail, should in no way be considered as limiting the scope of the inven¬tion. Example 1 A r.icrospheroidal pseudcboheir.ite to which silica (1.2% by weight) has been added, is prepared, with a particle diameter ranging from 5 to 300 \i, by the spray drying of a sol of hydrared alumina and Ludox silica. A sample of pseudobohemite is calcined at 450°C for 1 hour and then at 1150°C for 4 hours in a stream of dry air. The product obtained, consisting of 5, 9 and a transition alumina, has a specific surface of 34 irr/c and a porosity of 0.22 cc/g. 150 g of microspheroidal alumina are impregnated, using the "incipient wetness" procedure, with 33 ml of an aqueous solution containing 7.8 g of KNO; (titer 99.5p0) in deionized water, maintained at a temperature of 25°C. The impregnated product is dried at 30CC for 1 night and then calcined, in a stream of dry air, at 650°C for 4 hours in a stream of dry air. The concentra¬tion of potassium, expressed as oxide, with respect to the calcined product, is equal to 2.4 % by weight. An impregnating solution is then prepared by dis¬solving in 2 3 ml cf deionized water: 56.2 g of Fe (NOJ ;-9H:0 (titer 991 by weight) and 6.7 g of KNO, (titer 99.5°* by weight). The solution, heated to 50QC to complete the dissolution of the salts, is maintained at this temperature" for the whole duration of the impregna¬tion. The alumina modified with potassium oxide (153.6 g) is impregnated with an aliquot (4 3 g) of impregnating solution, dried at 120°C for 4 hours and impregnated again with the remaining aliquot of impregnating solu¬tion. 1'r.e impregnated product is dried at 12C°C for a night and finally calcined at 70C°C fcr 4 hours. The weight composition cf the formulate is the fol- lowing: 6.6% Fe.-O;-, 4°D KO and carrier the complement to 100. The formulate was nested in the dehydrogenacicn re¬action of ethylbenzer.e to styrene and the average re¬sults, after a test run of 100 hours, are indicated in table 1. Example 2 150 g of micrcspheroidai alumina obtained as de¬scribed in example I, are impregnated with a solution containing: 56.3 g of Fe (NO,; -9H_C (titer 99° by weight; and 14.2 g of KNG, (titer 99.5CQ by weight]. The impregnation, drying and calcination are car¬ried cut with the same procedure described in example 1. The weight composition cf the formulate is the fol¬lowing: 6.6% Fe;Oi, 4 o K_0 and carrier the complement to 100. The dehydrogenation average results of e thy lb-en-zene, after a test run of 100 hours, are indicated in table 1. Example 3 The same procedure is adopted as in example 2, but using an impregnating solution containing: 55.2 g of Fe (NO-.;-.-9H.-0 (titer 99°D by weight! and 6.7 g 0- KNO-(titer 99.5°=, by weight). 1'r.e weight composition cf the formulate is the fol- lowing: 6.6% Fe.-O-, 1.9% K-0 and carrier the complement to 100. The dehydroger.ation average results of ethylben-zene, after a test run of 100 hours, are indicated "in table 1. E.xamp_le 4 The same procedure is adopted as in example 2, but using an impregnating solution containing: 53.9 g of Fe(NG,.) v9H..O (titer 99% by weight) and 2.3 q of KNO-, (titer 99.5% by weight). The weight composition of the formulate is the fol¬lowing : 6.5% Fe:-0.., 0.3% K.0 and carrier the complement to IOC. The dehydrogenation average results of ethylben-zene, after a test run of 100 hours, are indicated in table 1. Example 5 The same procedure is adopted as in example 2, but using an impregnating solution containing: 93 .'l g of Fe (NO;! i-9H?0 (titer 99% by weight) and 14.3 g of KNC^ (titer 99.5% by weight) and at a temperature of 60°C. The impregnation is carried out in three subsequent steps using 45 g of the mother impregnating solution in each step. The first aliquot is added en alumina alone 'which is then dried at 120&C for A hours after impregnation. The dried product is then impregnated with a further 45 g cf nother solution and dried at 120°C. This treatment is repeated twice. The weight composition cf the formulate is the fol¬lowing: lC3Dd and carrier the complement to 100. The dehycrogenaticn average results of ethylben-zene, after a test run of 100 hours, are indicated in table 1. Example 6 ~. carrier having a surface area cf ICC rtr/g has been prepared by calcining the same pseudcbohemite containing silica of example 1 at 1060°C. 200 g of such a carrier are impregnated with a so¬lution containing 37. C5" g of Fe (KO;) 9K:0 (titer 99° b.w.) and 17.23 g of KNO. (titer 99.5ao b.w.), 2.97 g of Ce(NOj) J.6H.-0 and 2.93 g of La (NO,),. 6K;0 at a temperature of 60°C. The impregnation is carried out in a unique step. The impregnated material is dried at 120°C for 4 hours, then calcined at 750°C for 4 hours. The weight composition cf the formulate is the fol¬lowing: 5.0°. and carrier the complement to 1C0. The cehyarogenation results of ethylberiiene, during a test run of 150 hours, are shown in table 2. Example 7 (Comparative) In order to demons trace the promoting effect of silica in the carrier, a sample has been prepared ac¬cording to the ssme procedure of example 6 but based on a carrier, free of silica, having a surface area of 104 -r/g. The dehydrogenation average results cf ethylben-:ene, during a test run cf 163 hours, are indicated in Table 2. Example 8 The contemporaneous dehydrogenation cf ethylber.tene and ethane is carried out in the micro-reactor described above, at a temperature of 600°C, using the catalyst of example 1. Table 3 enclosed indicates the operating parameters and results obtained. We Claim 1. A process for the dehydrogenaiion of ethylbenzene to styrene which comprises: (a) reacting the ethylbenzene mixed with an inert gas in a fluid-bed reactor, in the presence of a catalytic system consisting of iron oxide and promoters supported on alumina modified with 0.01-10% by weight of silica and operating at a temperature ranging from 400 to 700°C. at a total pressure of 10.13 to 303.9 Kpa (absolute pascal) and with a GHSV space velocity ranging from 50 to 10,000 h"1. (normal liters of a mixture of ethylbenzene and inert gas/h x liter of catalyst); and (b) regenerating and heating the catalyst in a regenerator at a temperature exceeding 400°C. 2. The process according to claim 1, wherein the catalyst consists of: (1) 1-60% by weight of iron oxide; (2) 0-1-20% by weight of at least one alkaline or alkaline earth metal oxide; (3) 0-15% by weight of a second promoter consisting of at least one rare earth oxide; (4) the complement to 100 being a carrier consisting of a microspheroidal alumina selected from 25 those in delta, theta phase or their mixtures, in theta + alpha phase or delta + theta + alpha phase, modified with 0.0S-5°o by weight of silica. 3. The process according to claim 1 or 2, wherein the catalyst consists of: (1; 5-50* by weigh" of iror. oxide; (2) 0.5-10% by weight cf a metal promoter, ex¬pressed as oxide (2) the complement to 100 being a carrier consist¬ing of a micrcspheroidai 'alumina with a diameter ranging from 50 to 70 microns selected from those in delta, theta phase or their mixtures, in theta + alpha phase or delta + theta - alpha phase, modi¬fied preferably with O.C3-3°6 by weight of silica. 4. The process according tc any of the preceding claims, wherein the promoter is selected from alka¬ line, earth-alkaline metals or from the group of lanthanides. 5. The process according to claim 4, wherein the pro¬moter is potassium oxide. 6. The process according to claim 4, wherein the pro¬moters are potassium oxide, cerium oxide, lantanium oxide and p r a s e o d vm iurn oxide. 8. The process according to claim 1, wherein the inert gas is selected from nitrogen, methane, hydrogen and water vapor. 9. The process according to claim 8. wherein the inert gas is selected from nitrogen and methane. 10. The process according to claim 1, wherein the volumetric ratio inert gas/ethylbenzene ranges from 1 to 6. 11. The process according to claim 10, wherein the volumetric ratio ranges from 2 to 4. 12. The process according to claim 1. wherein the dehydrogenation reaction in step (a) is carried out at a temperature ranging from 450 to 650°C. at atmos- pheric pressure c: slightly higher, at a GHSV space velocity ranging from 100 tc 1,000 h"1 and with a residence time c: the catalyst ranging from 5 to 30 minutes. 13. The process according to claim 12, wherein the soace velocity ranges from 15C to 300 h~~ and the residence time ci the catalyst ranges frc^i 10 tc 15 minutes. 14. The process according to claim 1, wherein in step in) the regeneration c: the catalyst is carried cut with air or cxycer. whereas its heating is effected using methane, a fuel gas or by-products of the ce-hydroger.ation reaction, operating at a higher tem¬perature with respect to the average dehytirogena-tion temperature, at atmospheric pressure or slightly higher, at a space velocity ranging from 100 tc 1,000 r.~~ and with a residence time of the catalyst from 5 to 50 minutes. 15. The process according to claim 1, wherein in step a) the inert gas is a gaseous stream, essentially consisting of nitrogen recovered from the ccirJcus-tion products cf the regenerator. 16. The process according to claim 1, wherein in step a the ethyloenzene is fed into the reactor mixed with a paraffin selected frorr. ethane, propane or isobutane obtaining the contemporaneous dehydrogenation of the components of the mixture to give styrene and the corresponding olefins, respectively. 17. The process according to claim 16. wherein the ethylbenzene is fed 10 the reactor mixed with ethane obtaining the contemporaneous dehydrogenation of the components of the mixture to give styrene and ethylene respectively. 18. The process according to claim 17, wherein the ethylene is recycled to an alkylation unit together with a stream of benzene to give ethylbenzene. 19. A process for the dehydrogenation of ethylbenzene to styrene substantially as herein described, with reference to the accompanying drawings.

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