MXPA96003982A - Production of methylami - Google Patents

Abstract

The present invention relates to: An improvement is presented in a process for the catalytic production of methylamines from methanol and HN3, or methanol, a mixture of methylamines and NH3, in a gas phase on a bed of zeolite catalyst, where the improvement is that the catalyst bed is divided into two or more sub-beds, and the difference between the inlet and outlet temperatures of each catalyst sub-bed is maintained within the range of about 58 ° C to about 70 ° C while the reaction is carried out.

Description

PRODUCTION OF METHODS BACKGROUND OF THE INVENTION HAMPO DF THE INVENTION The present invention relates to a process for the production of meth 113m? Using the reaction ca aliti vapor phases between metapol and ammonia. MORE SPECIFICALLY, the present invention relates to a process for the production of methyl films using a zeolite as a colorant. By means of said method it is possible to use more effectively the characteristics of high selectivity for the dimethylane of the zeolitic catalysts. La une fc i lamina is an important intermediate chemical product as initial material for various solvents, pharmaceutical products, rubber chemicals, tensides and the like. RELATED TECHNIQUES Typically, the sheet is produced by allowing methanol, in a gas phase, to react with ammonia at an elevated temperature of approximately 400 ° C) in the presence of a solid acid catalyst, such as alamin or siliceous alamin, capable of causing dehydration. and inaction. In addition to dimethylamine < to. Following "DMA"), onometi sheet (below "MMA"> and tri eti sheet (hereinafter "TMA") are also produced as by-products of this reaction.The demand for this sheet of subproducts is much lower than The demand for DMA, for this reason, after their separation from the product of the reaction, these products are recycled into the reaction system and used again. However, TMA forms a complicated azeotropic system together with ammonia, MMA and DMA, in such a way that it is equiped with a very complicated large-scale distillation process. of the energy consumed in the process> DMA recovery is extremely high Examples of the process of ra < cvp > zra are presented with details, eg, - = &n "Manuf cturing Process Charts, Revised Edition "'publi cited by abushili 1-a? sha Kagat ^ u i ogyo-Sha on April 25, 1978). To reduce the cost of producing DMA and to decrease the size of the equipment, it is essential to suppress, as far as possible, the formation of by-product methylamines, especially the formation of TMA, and promote 1? DMA training. However, the selectivities for all three metals are usually determined conventionally in the conventional amorphous solid acid catalyst described above, for example, alumina or siliceous alumina. Under typical reaction conditions, the formation ratio of TMA is considerably greater than the proportion of DMA formation. For example, in the case where the reaction temperature is 400 ° C and the ratio between ammonia and methanol at the entrance of a reactor is 1: 1 weight ratio), 1? equilibrium weight ratio of the amines produced, calculated thermically is MMA; DMA? TMA = 0.284: 0.280: 0.436. In this case, the selectivity for DMA defined by the following equation (1) is only 28. M. Selectivity for DMA Oi) = (2 x WD / DMW) / (M / MMW + 2: <WD / DMW + 3: <T / TM) 100 fl > where WM, WD and WT are the proportions by weight of MMA, DMA, and TMA produced, respectively, and MMW, DMW and TMW represent the molecular weights of MMA, DMA and TMA, respectively. For this reason, it is necessary to continue separating a large amount of MMA and TMA and recirculate these two membranes together with a large amount of ammonia in the reaction system in such a way that the reaction can follow profitably for DMA from the point of view of balance. In recent years, vain zeolite catalysts have been proposed in order to solve the aforementioned problem. For example, the catalysts described in the following publications and patents may be mentioned: Japanese Patent Laid-Open No. 69846/1981 which refers to zeolite A; It publishes ions of Tapanese Open Patent No. 1 8708/1979 and 69846/1983 that refer to FU-1; North American Patent No. 4,082,805, which refers to ZSM-5, 7SM-11 and ZSM-21; Japanese Patent Laid-Open No. 113747/1981 which refers to ferperite and eponite; He publishes ions of Japanese Open Patent No. 178951/1986 and 8358/1988 which refer to rho zeolite, 7K-5 and chahazite; Japanese Open Patent Application No. 254256/1986 which refers to a catalyst having a selectivity for DMA m < improved as obtained by the treatment of a specific zeolite with tetraet 11 ostium 1 or similar; Japanese Open Patent Publication No. 002740/1995 which refers to ordinate modified by a sililating agent; It publishes ions of Japanese Open Patent Nos. 46846/1981, 210050/1984, 049340/1983 and 9510/1994 which refer to ordinate whose selectivity to DMA is improved by other methods of modifying ion; and U.S. Patent No. 3,384,667 which relates to zeolite X, Y and L, levin ta, analcita, chabazite, gmelinite, epomta, ptilotite, ferpepta, cl inopt 11 ol J ta and the like. Unlike conventional amorphous catalysts, such as, for example, siliceous alumina, all of the aforementioned zeolite catalysts provide sele tities for DMA greater than the thermo-maximum equilibrium value. For example, Japanese Open Patent Publication No. 210050/1984 presents a method for producing selectively DMA, using mordenite. According to the procedure, when a 1: 1 ratio by weight of ammonia and methanol is subjected to a reaction which is carried out, for example, by the use of mordenite binders having various cationic compositions at a temperature of reaction of 27 3 36 ° C, selectivity for DMA of approximately 5 and approximately 60% is obtained. These selectivities correspond to values of approximately 2.0 to approximately 3.0 when they become DMA equilibrium factors defined by the following equation '2): DMA equilibrium factor =' selection for DMA) / (the c 1 vity for DMA in thermodynamic equilibrium at the same temperature of the reaction) '2) Ad ma, almost all catalysts of zeolite described above provide DMA equilibrium factors of 1.2 to 4.0, and many of these factors are within the range of 1.5 to 3.5 When such a zeolite catalyst is employed for a continuous process for the production of ethylene Inas, the concentration of DMA in the exit gas of a reactor is vu < =? > & high because the selectivity for DMA of the catalyst is a. As a result, the amount of a recycling material to be returned from the recovery process to the reactor decreases. Accordingly, the total amount of materials that are fed from the reactor into the recovery process can be decreased. This effect can be shown by comparing circulation regimes per unit process defined by the following equation (3): Circulation regime by unit process = 'total flow regime of materials to be fed from the reactor to the recovery process (l-gmol / hour)) / 'quantity of DMA manufactured' 1 * gm l / hour)) '3) The unit flow rate can be adjusted by controlling the amount of materials to be recycled from a recovery system to a reactor , especially, the amount of ammonia. The degree of recycling of ammonia correlates with the atomic ratio N / C (the ratio between the number of nitrogen atoms and the number of carbon atoms), that is, the atomic ratio between N and C contained in all these materials that They are fed from the reactor to the recovery process. In order to reduce the load in the recovery process, it is necessary to reduce as much as possible the circulation regime per unit process, that is, decrease the N / C ratio. Nevertheless, it is unfavorable to drastically reduce the N / C ratio from the point of view of the impurities produced as byproducts, s im i l a r. The above-mentioned catalytic converters are charged because they can reduce the flow rate by unitary process without drastically reducing the N / C ratio compared to conventional catalysts of the controlled type by equilibrium. i co. From this point of view, it is necessary that the N / C ratio is generally 1.0 or more, preferably 1.3 or more »For example, in the case where -about 25 mol / hour is produced, approximately 65 mol / time, and approximately 10 mol / hour of MMA, DMA, and TMA, respectively, in the N / C ratio of 2.0 using zeolite catalysts that have several DMA equilibrium factors, the flow rates per unit process They are the following. Regimen of equilibrium factor Circulation by DMA catalyst Unitary process Alumina siliceous 1.0 18 'c unven i onal) Zeolite U) 1.5 13 Zeolite (2) 2.2 10 In this case, the essential purpose of the use of a zeolite catalyst is to reduce as much as possible the rate of circulation by unit process, so that this regime has a value lower than 18, which is the typical circulation regime per unit process when the conventional catalyst is used. However, it was found that when the catalyst? Zeolite is placed in an isolated reactor that has been used with the conventional catalyst, and when a reaction is initiated at a temperature l a. As low as possible (inlet temperature of the catalyst bed: 250-26 ° C) to prevent the formation of coke materials, the catalyst is rapidly deactivated as will be seen later in comparative examples, causing a very serious problem in practical use . The de-ivation of the catalyst can be indicated by means of a degeneration constant (rho) according to the definitions represented by the following equations' 4) and < 5): ist = 0. exp (-rhat) '4) where rho: constant of degenerate ion, M; s constant reaction rate when t days have passed, 10: constant reaction rate when the reaction starts; and reaction rate constant k = FPT / POV.ln '1/1-O (5) where F: Methane feed rate 1 Gas constant Rj, T: Reaction temperature: i n, PO; initial partial pressure of methane], V: volume of catalyst,; -: conversion rate of methanol. When a bed reactor is used as is the case in the previous process, it is necessary, from the commercial point of view, that a catalyst can be used continuously for at least one year, preferably for 2 or more years. If a catalyst has an initial catalytic activity sufficiently high for omerial use, it can be used until the catalytic activity proceeds rapidly. half of the initial activity. In this case, the decay constant corresponding to a catalyst life of one year is approximately 0.0021. However, in the case where a zealite catalyst is used with a conventional type reactor, any zeolite catalyst shows a deterioration constant 10 times or more than the aforementioned threshold value, as will be seen later in the examples comparative Therefore, it is extremely difficult to continuously use a zeolitic catalyst for commercial scale production. Accordingly, a main object of this technical field is to develop a process for the production of methylas, wherein a zeolite catalyst can be used with in uue for a prolonged period of time. SUY OF THE INVENTION It is an object of the present invention to provide a process for the production of eleven amines using a zeolite catalyst. In said process, the deactivation of the zeolite catalyst is suppressed in such a way that the zeolite catalyst can be employed with inuamente for a prolonged period of time. We carry out in-depth studies to develop the procedure described above for the production of metal sheets, using a zeolitic catalyst. As a result, it was found that when a reaction is performed in such a way that a zeolite catalyst bed is divided into two or more sections, in series or in parallel, and that the difference between the inlet and outlet temperatures of each bed of catalyst is kept inside? a specific range while the reaction is performed, the life of the catalyst can be drastically improved, and therefore the above object can be achieved. The present invention has been achieved based on this discovery. Accordingly, the present invention relates to a process for the production of methylamines, comprising the contacting of methanol and ammonia, or methane, a mixture of ethylamines and ammonia, or a mixture of methylamines and ammonia in a gas phase, with a bed of a catalytic converter of real? t3, where the catalyst bed is divided into two or more submamas connected in a parallel CD and the difference between the inlet and outlet temperatures of the catalyst is maintained. from the range of approximately 5"C to approximately 70nC while the reaction is taking place BRIEF DESCRIPTION OF THE DRAWINGS In the drawings, Figure 1 is a diagramatic flow sheet showing a procedure for the production of metal sheeting. n the present invention; and Figure 2 is a diagrammatic flow sheet showing another process for the production of metal sheets according to the present invention. DETAILED DESCRIPTION OF THE PREFERRED MODALITY OF INVENTION The present invention will now be explained with reference to the accompanying drawings. A process in which a catalyst bed is divided into three sections or subwebs connected in series as shown in Figure 1 can be mentioned as an embodiment of the present invention. The initial materials, methanol and ammonia (line (1)), are joined with recycling material 'line (4); gas and / or liquid) consisting essentially of unreacted ammonia, MMA and TMA, which come from the process of t ecuperac i? p 'R1 *. Marble pits, which are in the gaseous state,, e al l 'in line' 2)) to the first 'SI' section of the catalyst bed at a predetermined air temperature, vé5 of the evaporation steps, heating nto, etc. A zealite-3-tale ion that has a DMA equilibrium factor of 1.2 or more, preferably 1.5 or more, is used as a catalyst. Specific examples are such a zeolite catalyst in 1 order, c habacite, levtnite, ze lite rh, zeolite A, ru-1, epanite, ZSM-5, ZSM-11, 7SM-21, Zl-5 and montmop 11 oni, and zeolites obtained by modifying the ion of the aforementioned zeolites. Among these, mordemta, chabacita and modified ordenita are preferred. In the first section 'YES', the initial materials and the recycling material react between them. The outlet gas of the first section (line <2) -t) is fed to the second section 'S2) of a catalyst bed and the outlet gas (line (2) -2) of the second section is fed to the next section 'S3). The unreacted initial materials and the recycling matepal contained in the exit gas react with each other in each section of the tasting bed. The reaction is carried out by controlling the inlet and outlet temperatures of each catalyst bed in such a way that the difference between them can be maintained within the range of approximately 5''C? Approximately 70 ° C, preferably ipro; from 10 ° C to approximately 5 (", 0C, > higher degree of lending < i > approximately 0 ° C to approximately 40 ° C. greater than 70 ° C would cause a shorter life and catalyst, and if lower than 5 ° C, re-dissolve in temperature to obtain the conversion rate of 1 cm. For example, it is desirable for the inlet temperature of the catalyst slug to be as close as possible to the minimum temperature to initiate the reaction at a reasonable rate. It is desirable to control the inlet temperature appropriately. 2 ° C to approximately 350 ° C, in general, preferably from about 22 ° C to about 330 ° C, and preferably from about 230 ° C to about 310 ° C It is preferred that the exhaust gas from the catalyst bed be fed to the next catalyst bed after cooling In this way, the inlet temperature of the catalyst bed can be easily adjusted, and the operation to maintain the temperature of the catalyst bed can also be carried out easily. The exhaust gas can be cooled either by means of a condenser 'O which uses a cooling medium as air, ga =. Nitrogen or steam, or by means of a 3-color generator. At the outset, the cold in the exit line, where the direct feed of the recycling material and / or a part of the materials is distributed. and / or 03) to the outlet gas The cooling of the outlet gas of the last form is espe- lly from the viewpoints of the heat recovery and the reduction of the cost of the gas. The gas from s li a (line (3)) of last section 'S3) exchanges heat with the feed line of the initial materials and the like, and then it is fed to the recovery process (P). recovery process, the sheets are separated respectively, by means of a plurality of distillation columns, of the ammonia without reac a to r, water produced by the reaction, etc., and recovered. The unreacted ammonia, MMA and TMA are returned to the reactor as 'line' recycling material 4)). It is preferable that the catalyst bed be divided into 2 to 10 sections, preferably 2 to 7 sections. When the effects of the present invention, the cost of the equipment, and the operating characteristics are taken into consideration, it is especially preferable to divide the catalyst bed into 3 to 5 sections / sub-beds. Furthermore, in relation to the reaction of each catalyst sub-level, it is preferable to control the reaction in such a way that the difference between the rates of conversion of methanol to the heat and to the entry of each sub-layer in of methanol conversion and the entry of pr = mer =? sub-rows are within the range of W <;% to h?%, d preference 15 to%. When the conversion of methane] e = less than 1 * A to the total number of catalyst sub-masses-would probably be: highly elevated, resulting in apparatus complexity, whereas when the melunol conversion is greater than é, * Á , the catalyst1 life would probably be shorter. The relation N / C, viz. the ratio between the number of nitrogen atoms and the number of carbon atoms, in the catalyst bed is from 0.8 to 3., preferably from 1.0 to 2.5, and with greater preferably, from 1.2 to 2.2. An N / C ratio of less than .8 would result in the production of byproducts in high quantity, and an N / C ratio greater than 3.0 would result in an increase in the amount of recycle and the size of the appliance used. The pressure of the reaction is, in general terms, from the normal pressure to 200 at. In addition, a process presented in Figure 2 can be mentioned as another embodiment of the present invention. In this process a multi-tube reactor is used < R), where vain reaction tubes are installed in parallel, and the gas is circulated outside the tubes. The initial materials, methane, and ammonia 'line'! > > , they are united with the recycling material consisting essentially of unreacted ammonia1, MMA and TMA 'líe (, which comes from the recovery process (P) .These materials, which are in a gaseous state, "line" 2)) are fed at a predetermined temperature * to the multi-tube reactor from the lower part of the cover e) of the same, by means of the evaporation steps, c lent and - _? im? lares. This gas is subjected to heat exchange, through the tube walls, with the gas passing through the catalyst beds placed in the reaction tubes, which flow in the opposite direction. Then, the gas is circulated in the upper part of the reactor, and is introduced from the upper part of the tubes, towards the catalytic masses (Sl ... Sn), thus causing a reaction. After exchanging the heat of the g -.s fed in this way with gas in the coating of the reactor, it is fed as reaction product gas' line (3) to the next step, the recovery step. ), from the bottom of the reactor. In the recovery process, the sheets of unreacted ammonia, water produced by the reaction, etc. are separated respectively by a plurality of distillation columns and recovered.

The unreacted ammonia, MMA and TMA are returned with recycling material (lynx '1)) < * \ reactor. L .- > i. On the operating modes for this mode, the tilt-up sub-machines routed in parallel can -j > -r l., my n or similar to the cspd u. i o is p ra I aforementioned modality with the catalyst subweaves connected in series unless the con ditions d * = »the previous mode interfere with the conditions of the present. ? fashion 1 One of the zeolites as described above can be used as > . tie 11 z * * for the reaction. It is desirable that the difference between the inlet and outlet temperatures of 1-catalyst bed be small. Accordingly, it is preferable to carry out the reaction maintaining the above difference within the range of about 5 * C to about 70 ° C, preferably between about 5 ° C and about 30 ° C, and with a greater degree of preference between about 5 ° C. C and approximately 20 ° C. To prevent the formation of impurities such as coke, it is desirable that the inlet temperature of the catalyst bed be close to the minimum temperature to initiate the reaction at a reasonable rate. inlet generally from about 200 ° C to about 350 ° C, preferably from about 220 ° C to about 330 ° C, with a greater degree of preference from about 23 ° C to about 310 ° C. The number of the reaction tubes is 2 or more, preferably 10 or more, with greater preference, preferably 5 or mec .. In the case in which this process is extensively carried out, the tube number of r The ratio is several hundred to approximately 2.00, when this depends on the production area. The total conversion rate of methanol in the catalyst layers is 80 * / »or more, with a higher degree of preference 85 * 4 or more. The ratio N / C (the ratio between the number of nitrogen atoms and the number of carbon atoms) in the catalyst bed is from 0.8 to 3.0, preferably from 1.0 to 2.5, with a higher degree of preference from 1.2 to 2.2. . The reaction pressure is, in general terms, of normal pressure at 200 atm. The present invention will now be specifically explained with reference to the following examples. However, the present invention is not limited to the following examples.

E TEMP The catalysts used Zl: orderite of form H treated with steam at 300 ° C during hours in granules with a diameter of approximately 5 mm.

I1? Z2; atica form morclenite containing 0.7 * 4 er > weight of f and 0.7 * 4 by weight of Ca in granules of a diameter of approx. i adamen 5 mm. Z3: mixture of ehabacite and eponite in granules of an apiro diameter; 5 mm. A: amorphous silica alumina containing approximately 70 * 4 silica in granules of an approximate diameter of 5 mm. Examples 1 to 5 Two or three reaction tubes installed in series as shown in Figure 1 (one isolated reactor having a total capacity of approximately 20 m3) were filled with a zeolite catalyst. The gas between the sub-zones SI and S2 was used by the liquid and gas recycling material from the recovery process (P), and the gas between the sub-zones S2 and S3 was cooled by the material of recycling (liquid and gas) co or air. Methanol and ammonia were fed at rates of 1 μg / kg / hr and 87 μg / hr, respectively. Therefore, a continuous reaction was carried out (for 2 weeks to a month) to obtain MMA, DMA and TMA with production regimes of 20 kgmol / hour, 60 l-gmal / hour and 7 kg or w / w, respectively. The decay constant of the catalyst, presented in Table 1, was obtained from the change in methanol conversion rate, in accordance with equations (4) and (5) before m nc i added. Examples 6 to 8 By the use of a production process in the Q reaction tubes, each with a diameter of approximately 2.54 centimeters, were connected in seine, and the gas in the space between the tubes was cooled, a continuous reaction experiment was carried out for 2 weeks to a month). Since this procedure did not involve recovery or recycling processes, the following composition was prepared and used as the initial material assuming that both recovery and recycling had been done in advance. Finally, for the thermal dissipation by quantity of unit gas in the process to be equal to the thermal dissipation of the previous examples, the surface of the reactor was kept hot. It was confirmed in examples »3 and 6 that the processes in these examples were similar to each other. Initial material composition 'molar percentage) Ammonia Methanol MMA TMA Examples 6 and 7 60 30 8 2 Example 8 59 31 7 3 The results obtained in the previous examples 1 to 8 appear in table 1.

TABL 1 Temperature number Difference jube s l intro d of temp i I - N d • of * 1 tura * 2 / Example Catalyst catalyst < °) CC; C 1 Zt 2 2 0 40 2. 2 Zl 2 2 0 5 1.7 3 Zl 3 250 35 1.7 4 Z 2 250 60 5 Z2 3 250 IO 6 Zl 3 250 35 7 3 3 260 30 8 Z2 4 260 25 Equilibrium Rate Factor Constant equilibrium Example circulation * 3 'xl / 1000) of DMA 1 11 1.4? 9 1.9 T 9 1.0 4 8 1.9 5 8 0.9 6 9 1.0 7 11 1.2, or 8 -T 1 0.8 * 1 Each subcategory of ca tal i zador * 2 Between the input and output of each subca a of catalyst + Par process unit Comparative examples 1 and 2 An adiabatic reactor with a capacity of approximately 10 m3 was filled with the tasting! siliceous alumina hoist (A). To achieve this, methanol and ammonia were fed at rates of 12 t-gmol / hour and 65 kgmol / hour, respectively. Accordingly, a continuous reaction was carried out to produce MMA, DMA and TMA at production rates of 15 1-gmol / hour, 45 l-gmol / hour and 5 ^ g / hour, respectively. The decay constant of the catalyst in this case is shown in Table 2. E p p e rs 11 a 3 and 4 A continuous reaction was carried out in the same way as in Example 1, except that the gases »between the = Catalyst ufcicamas were not cooled. This lot can therefore correspond to the one in which an adiabatic reactor is used, or in comparative example 1, with a single sub-layer of catalyst. The deterioration constant of the catalyst in this case is shown in Table 2. Comparative Examples 5 to 7 A continuous reaction was carried out in the same manner as in Example 6, except that the gases between the catalyst subeamas were not cooled. Furthermore, in order that the thermal dissipation per unit gas quantity in this process was equal to the thermal dissipation in the previous comparative examples, the surface of the chlorine was heated. It was confirmed in Comparative Examples 4 and 5 that the procreses of these examples were similar to each other. These batches correspond to the batch in which an adiabatic reactor was used, or in comparative example 1, on a sol 3 subeama catalyst. The decay constant of the catalyst in this case is shown in Table 2. Composition of the initial material molar percentage Ammonia Metanal MMA TMA Comparative Example 60 30 8 2 - . The refracts obtained in the above comparative examples 1 to 7 are shown in table 2. Table 2 Difference number sub-temperature Temperature temperature Example of input ratio * 2 comparative catalyst catalyst of * 1 '° C) < ßC) 1 A 1 350 60 2 A 1 340 70 3 Zl 1 250 100 4 Zl 1 250 110 5 Z1 1 250 110 1 5 N Constant Factor Example / Regime of deterioration rho equilibrium comp a rat i vo C circul ci n # 3 f X 1/1 00) of DMA 1 2? ..00 1 188 0.8 1. 1.7 16. ß i. or 2. U 17 ^ .T 4 1.7 9 5 1.7 9 6 1.7 8 1 '.7 ~ 11 23 2. o * 1 C3da catalyst sub-layer1 * 2 Between input and output of each catalyst sub-layer + By unit process' Consideration - the case in which a catalyst cradle is divided into sections / sub-keys connected in senes). The case examples 1 and 2 present the case in which a reaction is carried out in a conventional adiabatic reactor, using siliceous alumina, a conventional thermo-regulating equilibrium type catalyst. The catalyst deterioration constant in this case was 0.8, which corresponds to a catalyst life of 2 years or more. The comparative examples 2 to 3 _'.-) are] case in which a reaction was performed with the same equipment as in comparative examples 1 and 2, using several zeolite catalysts. Even when the entry temperature of the catalyst crape was maintained at a level of 250 to 260 ° C, which is approximately the minimum temperature to initiate the reaction at a reasonable rate, to prevent side reactions such as the formation of In this case, the catalyst deterioration constant was surprisingly found at an extremely high level of 17 to 24%. The high value means that the catalyst life is only 1 or 2 months. It is a result very insa 1 sfac tor 10 from the point of practical view. In general termsThe zeolite catheters have extremely small pores in comparison with the amorphous catalysts, so that their catalytic activities are easily affected by coke materials deposited on their surface. This is considered to be the reason why this result was obtained and the original factor. On the contrary, examples 1 to 8 are the case in which a reaction was carried out under the same conditions as in the comparative examples except that these examples were made in accordance with the process of the present invention: the zeolite catalyst bed it was divided into 2 to 4 sections / sub-areas connected in senes, and the difference between the inlet and outlet temperatures of each catalyst bed remained within the range of 25"C to 60" C while the reaction was carried out. As a result, it was surprisingly found that? the catalyst constant b or a 7 to 11. Such a value corresponds to a catalyst life of 1-2 years or more. Therefore ~ >It was confirmed that the life of the catalyst increases dramatically until it reaches a sufficiently long period so that the process of the present invention can be employed at the industrial scale of production. A multi-tube reactor of the heat exchanger type was used (Figure 2> where there were 6 reaction tubes with a diameter of 1.27 centimeters is one, in parallel) Each reaction tube in the reactor was filled with 25 ml of zeolite catalyst to form catalyst sub-beds. For this purpose, an initial material composition consisting of 61 * 4 molar-ammonia, 29 * 4 molar ethanol, 8 * 4 molar MMA was fed. 2 * 4 molar of TMA through the external reversal of the reactor at a rate of 3 percent per hour of 3O0 g / hour in the direction opposite to the flow of the gas that enters the subweaves by the line '2) and which passes through the catalyst subeamas (Sl ... Sn), thus realizing an area for 2 weeks. The difference between the entry and greet temperatures of the -nbc as of c < ? ta 1 i zaclor was 15 ßC * an average. The result is shown in Table 3. Table 7 Temperature Number Conversion sub-levels of input MeoH Example catalyst catalyst of * t í * C) () 9 71 6 270 1 Z2 6 270 95 11 Z3 6 280 94 N Rate regime constant / circulation deterioration equilibrium Example C * 2 (X1 / 1 00) of DMA 9 1.7 9 1.7 2.2 10 1.7 8 1.0 2.3 11 1.7 9 2.0 2.2 # 1 Each catalyst sub-bed * 2 By unit process Comparative Examples 8 and 9 An adiabatic reactor with a diameter of 2.54 centimeters was filled with 100 ml of zeolite catalyst. This was fed an initial material composition consisting of 61 * 4 molar ammonia, 29 * 4 molar methane, 8 * 4 molar MMA and 2 * 4 molar TMA from the bottom of the reactor to a regime of 2 0 g / hour, thus making a? 8 reaction during > . Weeks the difference between lai. The temperature of the entrance and of the exit from the bed in the room was 9 ° C on average. Fjein l s compare i os lO -a 12 An adi bient reactor with a diameter of 2.5-1 centimeter was used. with an external reverse. The inner tube of the reactor was filled with 10 ml of a zeolite catalyst. This was fed a material composition consisting of 61 * 4 molar ammonia, 29"/, molar methanol, 8 * 'molar MMA and 2 * 4 molar TMA through the external coating of the reactor at a rate of 200 g / hour in the direction opposite to the circulation of gas that passes through the old izadenr crama, thus performing a reaction for two weeks, the difference between the inlet and outlet temperatures of the catalyst bed was 70 * C on average The results obtained in the comparative examples 8 to 12 are presented in the tatnla 4. TABLE 4 Conversion Temperature Number Example of input subemphases MeDH comparative catalyst catalyst * 1 (* C) (* 4) 8 Zl 1 240 97 9 Z3 1 250 95 l? Zl 1 25 < ? 11 Z2 1 250 5 12 Z 1 260 9 No Fac or constant of / regime? deterioration rho balance E c it C c i r > : ru l ion * 2 (X1 / 1 00) of DMA 8 1. 7 1 1 9 1. 7 12 20 1. 1 1. 7 10 12 2.1 1111 11..77 99 9 2.2 12 1. 7 1 1 10 2.0 - * 1 Each catalyst sub-layer +2 Per-unit process' Cennsí derací ón - the case in which a catalyst bed is divided into secc ion.es/subcamas connected in parallel).

Comparative examples 8 and 9 are the case in which the metal sheets are synthesized in a single bed of a zeolite catalyst by the use of a small size equipment which was of the same type as a conventional adiabatic reactor 1 used for the production of sheet metal. Even when the recorded temperature of the catalyst cradle was maintained at 240 ° C to 260 ° C, which corresponds approximately to the minimum temperature to initiate the reaction at a reasonable rate, to prevent nuclear reactions, such as, for example, of 3 < > coke, the deterioration constant l catalyst was found at a high ex-ternal level of 2 IHJ¿. This corresponds to a catalyst life of only approximately one month, and is far from the level suitable for practical use. The comparative examples 11 and 11 are those in which a single bed of a zeolite futg catalyst was installed in the inner tube of the 2-tube reactor, and the LO reaction was followed? heat exchange of the gas fed with the gas passing through the catalyst bed flowing in the direction opposite to the gas fed. In this case, the deterioration constant of the catalyst is approximately 10. However, when this value improves slightly compared with the comparative examples 8 and 9, it is still far from the useful level for practical use. On the contrary, examples 9 and 11 are the case where a reaction is carried out, in accordance with the process of the present invention, by means of a multi-tube heat exchanger reactor employing a zecnl catalyst. It has been divided into 6 sections / sub-beds connected in parallel, with gas supply in heat exchange condition with the gas passing through the catalyst bed., flowing in the opposite direction to the feed gas. As a result, it was surprisingly found that the deterioration constant of the catalyst showed an improvement of 1 to 2, which is about 1/10 of that observed in 1 example. The hoist was greater than in the Comparative Examples. Such value corresponds to a catalyst life of 1 to 2 years. The current procedure can therefore be suitably used for production on an industrial scale. In accordance with the purpose of the present invention, the life of a zeolite catalyst for use in the production of melamine can be greatly improved, and the catalyst can continue to be used in production for a long time. period of time at a reaction temperature of 300 ° C or less. Accordingly, the production of metal sheets on an industrial scale can be carried out with prenvechen.

Claims (13)

1. PEIN I I ATIONS 1. In a procedure for the production of metal products, w involves the contact of methanol and ammonia or mei anol, a mixture of methylamines and ammonia, or uni ez. I d put sheets and ammonia, in a gaseous phase, with a bed of a zeolite catalyst, the improvement comprising the use of a catalyst? zeolite in such a way u the catalyst bed is divided into days or more ups connected in series and / or parallel, and the difference between the entry and exit temperatures of each sub-bed of? The catalyst is maintained within the range of approximately C to about 70 ° C while the reaction is being carried out.
2. 2. The process according to claim 1, wherein the zeolite catalyst is a catalyst that can provide a di methylamine a selectivity of? 1.2 times or more the thermal equilibrium value at a temperature at w the zeolite catalyst is used.
3. 3. The conferring procedure with claim 1, wherein the zeolite catalyst is selecerated into the group consisting of mordenite, chabazite, levinite, rho zeolite, zeolite A, FU-1, enonite, ZSM-5, ZSM-11, ZSM- 21, Zlí-5 and montmor i 1 loni t, and zeolites obtained by modifying those mentioned.
4. 4. The method of claim 1, wherein the inlet temperature of the catalyst sub-bed is within the range of about 2 ° C to about 350 ° C.
5. 5. The procedure in accordance with the rei dication 1, where the N / C ratio is a ratio between the number of atoms of nitrogen and the number of carbon atoms in the catalyst sub-bed is within the range of .8 to 3.0. with any of Claims 1 to 5, where the catalyst sub-chamber is connected in sequence 7. The procedure in accordance with rei indication 6, where the total number of sub-beds of the catalyst-divided bed is from 2 to 10. 8. The process according to claim 6, wherein a gas fed from a given catalyst sub-shell is cooled to the subsequent connected sequencer sub-chamber 9. The method according to claim 1. n 6, where a gas that has been subjected to the reaction is treated in such a way that the desired mixture formed is recovered from the gas and at least a part of the remaining gas is sprayed to the process and where a gas is fed. from a subcategory of catalyst given to the subcatalyst of subsequent catalyst connected in seine is cooled by the use of the ga or recycling liquid from the purification process, or a part of the initial mr-f pal, ammonia or methanol. 10. The method according to claim 6, wherein the reaction in each catalyst sub-frame is controlled in such a way that the difference between the rate of convolution of * methanol at the outlet and the entry of •. ada 3ut > The catalyst and the catalyst based on the methanol conversion rate at the entrance of the first catalyst sub-layer is found to be 1 * 4 3 60 * 4. 11. The process according to claim 1 to 5, wherein the catalyst bed is divided into sub-rows in parallel, inside a reactor. 12. The procedure in accordance with the rei in ication 11, where the methane conversion rate] is averaged 80 * 4 or more. 13. The procedure in accordance with claim 11, where the total number of sub-commas is from 2 to 2000.