Obtaining aromatic hydrocarbon from methane

FIELD: oil and gas industry.

SUBSTANCE: invention refers to method for obtaining aromatic hydrocarbons from methane, namely natural gas. Invention relates to method of methane transformation into aromatic hydrocarbons according to which methane-containing initial material and granular catalytic material is supplied to reaction zone, the operation of reaction zone is provided in reaction conditions effective for transformation of at least part of methane into aromatic hydrocarbons with associated deposition of carbonaceous material on granular catalytic material that causes its (catalytic material) de-activation. At least part of de-activated granular catalytic material is removed from reaction zone and heated till the temperature approximately from 700°C to 1200°C by direct or indirect contact with gaseous combustion products obtained by combustion of auxiliary fuel. Then heated granular catalytic material is regenerated by hydrogen-containing gas in conditions effective for transformation of at least part of deposited carbonaceous material into methane and regenerated granular catalytic material is returned back into reaction zone.

EFFECT: improvement of methane aromatisation method.

26 cl, 6 dwg, 10 ex, 7 tbl

 

The technical field to which the invention relates.

The present invention relates to a method for producing aromatic hydrocarbons from methane, in particular from natural gas.

The level of technology

Aromatic hydrocarbons, in particular benzene, toluene, ethylbenzene and xylenes, are important chemical products of mass production in the petrochemical industry. Currently aromatic compounds most often receive a variety of methods from starting materials based on crude oil, including catalytic reforming and catalytic cracking. However, as the world's supply of raw materials based on crude oil decreases, the necessity of finding alternative sources of aromatic hydrocarbons increases.

One possible alternative source of aromatic hydrocarbons is methane, which is the main component of natural gas and biogas. World reserves of natural gas are constantly replenished, and now natural gas fields offer more than oil fields. Because of the problems associated with transporting large volumes of natural gas, most of the natural gas produced along with oil, in particular in remote areas, flared. Therefore, attractive m is a mode of natural gas processing is the conversion of the contained alkanes directly into heavier hydrocarbons, such as aromatic hydrocarbons, provided that they can be overcome the attendant technical difficulties.

A significant part of the methods proposed in the present time for the conversion of methane to liquid hydrocarbons, includes a first transformation, or conversion of methane into synthesis gas, a mixture of N2and WITH the. However, the production of synthesis gas is associated with high capital costs and is energy intensive, resulting in preferred ways, which do not require the generation of synthesis gas.

Was offered a number of alternative methods of direct conversion of methane to heavier hydrocarbons. One such method involves the catalytic oxidative interaction of methane to olefins and subsequent catalytic conversion of these olefins to liquid hydrocarbons, including aromatic hydrocarbons. For example, in patent US 5336825 described two-step method for oxidative conversion of methane into hydrocarbons that are in the temperature range of the boiling gasoline fraction and includes aromatic hydrocarbons. At the first stage in the presence of free oxygen using a modified rare earth metal catalyst oxide of alkaline earth metal at a temperature of from 500 to 1000°C, the methane into ethylene and small amounts of the olefins With 3and C4. Then the ethylene and heavier olefins formed in the first stage, over an acidic solid catalyst, comprising pentasyllabic zeolite with a high content of silicon dioxide is converted into liquid hydrocarbons in the temperature range of the boiling gasoline fraction.

However, these methods of oxidative interaction inherent problem lies in the fact that their implementation involves conducting vysokotekhnologicheskikh and potentially dangerous combustion reaction of methane and is accompanied by the emission impact on the environment of carbon oxides in large quantities.

A potentially attractive destination processing methane directly into heavier hydrocarbons, in particular ethylene, benzene and naphthalene, is dehydroaromatization or reducing interaction. This method generally involves contacting methane with a catalyst comprising such a metal, such as rhenium, tungsten and molybdenum deposited on a zeolite, such as ZSM-5, at high temperature, in particular from 600 to 1000°C. the Catalytically active form of the metal often is the elemental form with zero valency, carbide or oxycarbide.

For example, in patent US 4727206 describes a method for liquids-rich aromatic hydrocarbons, the introduction of metal is at a temperature of from 600 to 800°C in the absence of oxygen in contact with a catalytic composition, includes aluminum silicate with a molar ratio of silica to alumina of at least 5:1, and in the aluminosilicate added gallium or its compound and the metal of group VIIB of the periodic table of elements or its connection.

In addition, in patent US 5026937 described by way of aromatization of methane comprising feed stream source materials containing hydrogen (molar content of 0.5%) and methane (molar content of 50%), in a reaction zone containing at least one layer of a solid catalyst comprising ZSM-5, gallium and phosphorus-containing alumina, in terms of transformations, including a temperature of from 550 to 750°C., an absolute pressure below 10 bar (1000 kPa) and the volumetric gas flow rate from 400 to 7500 h-1.

Further, in patent US 6239057 and US 6426442 described a method of producing hydrocarbons from a higher number of carbon atoms, such as benzene, hydrocarbons with a low number of carbon atoms, such as methane, the introduction of the latter into contact with a catalyst containing a porous carrier, such as ZSM-5, which dispersed rhenium and a metal promoter, such as iron, cobalt, vanadium, manganese, molybdenum, tungsten, or their mixture. After impregnation of the carrier rhenium and a metal promoter, the catalyst was activated by treatment with hydrogen and/or methane at a temperature of from about 100 to 800°C for the straps from about 0.5 to 100 hours It is also argued that the addition of CO or CO2in the methane source material increases the yield of benzene and stability of the catalyst.

However, successful application of restorative interaction for obtaining aromatic compounds on an industrial scale requires solving a number of serious complicated tasks. For example, the recovery process of interaction and endothermic, and limited in thermodynamic terms. Thus, if the process does not provide significant additional heating, the cooling effect during the reaction leads to a decrease in the reaction temperature sufficient to greatly reduce the reaction rate and the overall thermodynamic transformation.

Moreover, the process is accompanied by formation of carbon and other non-volatile materials which accumulate on the catalyst surface, causing a decrease in its activity and potentially unwanted shifts selectivity. In addition, typical process, the high temperatures of the active metal compounds (Mosx, WCxetc) on the catalyst surface can migrate to form agglomerates or change phase, which again leads to undesirable deterioration in the degree of conversion and selectivity of action. Therefore it is necessary conducts the frequent oxidative regeneration of the catalyst, to remove carbon and other non-volatile materials that have accumulated on its surface, and to redistribute active metal compounds. However, depending on the composition of the catalyst regeneration in an oxidizing atmosphere may have some unwanted side effects. For example, the metal on the catalyst surface can move from a catalytically active element or carbon-rich state in less active oxidized state. The catalyst after regeneration can be more active in terms of deposits of coke (coke formation) and the associated formation of hydrogen. Thus, there is interest in the development of restorative processes of interaction, in which regeneration of the catalyst is carried out in non-oxidizing conditions.

For example, in Japanese patent publication Kokai 2003-26613 from 29.01.2003 described method for the production of aromatic hydrocarbons and hydrogen from a lower hydrocarbon, containing at least 60 mol.% methane, in the presence of a catalyst, such as molybdenum, tungsten or rhenium on the media ZSM-5, characterized in that the catalyst is periodically and alternately transferred from the working cycle, in which the catalyst is in contact with the lower hydrocarbon, in the regeneration cycle, in which the catalyst to takeroot with hydrogen. Typically, the duty cycle is from 1 to 20 minutes, preferably from 1 to 10 minutes, and the regeneration cycle is from 1 to 30 minutes, preferably about 5 to 20 minutes.

In addition, in international publication WO 2006/011568 from 02.02.2006 described method for the production of aromatic hydrocarbons and hydrogen from a raw gas containing a lower hydrocarbon, characterized in that the mixture of the raw material gas and the hydrogen-containing gas is introduced into contact with the catalyst, such as metroselect based on molybdenum and/or rhodium at high temperature such as 750°C, and periodically during regeneration cycles, the flow of the above-mentioned raw material gas is stopped, and the flow of hydrogen containing gas continue. Allegedly, the continuous supply of hydrogen during the operating cycle, and during the regeneration cycle reduces the ratio of the duration of regeneration for the duration of the working cycle.

In the application publication US 2003/0083535 disclosed method flavoring metadatareader source material, characterized in that between the reactor system and the recovery system are circulating catalyst dehydroaromatization, and for the regeneration of various parts of the catalyst at different times enter into contact with different regenerative gases, including O2H2and H2O. the Percentage of catalyst, to the pulsing with each regenerative gas regulate to save in the reactor system and the recovery system mode heat balance. This reactor system includes a fluidized bed of catalyst in the Elevator-the reactor and regenerator system includes a second fluidized bed of catalyst contained in the reactor with a stationary fluidized bed. After passing through the regeneration system of hot regenerated catalyst in the reactor recycle system for the conveyor system, which may include the replacement vessel for increasing the activity of the regenerated catalyst to the introduction of the catalyst in the fluidized bed in contact with a stream of a reducing gas containing hydrogen and/or methane.

Disclosure of inventions

The basis of the invention is the development of an improved method for the aromatization of methane, in which regeneration of the catalyst is hydrogen-rich recycle gas.

One object of the present invention is a method for converting methane to heavier hydrocarbons containing aromatic hydrocarbons, including:

(a) feeding to a reaction zone metadatareader the source material and the granular catalytic material

(b) maintaining in the reaction zone under conditions effective to convert at Myung is our least part of the methane in more high molecular weight(s) hydrocarbon(s) with concomitant deposition of carbonaceous material on granular catalytic material, causing deactivation of the catalytic material,

(C) removing from the reaction zone at least a portion of the deactivated granular catalytic material

(g) heating at least part remote from the reaction zone of granular catalytic material to a temperature of from about 700°to about 1200°C, carried out by direct or indirect contact with the gaseous combustion products produced by combustion of additional fuel,

(d) regenerative heated part of the granular catalytic material is a hydrogen-containing gas in a regeneration zone under conditions effective to convert at least part of the deposited carbonaceous material to methane, and

(e) returning at least part of the granular catalytic material from step (d) into the reaction zone.

Another object of the present invention is a method for converting methane to heavier hydrocarbons containing aromatic hydrocarbons, including:

(a) feeding to a reaction zone metadatareader the source material and the granular catalytic material

(b) maintaining in the reaction zone under conditions effective to convert at least part of the methane in more high molecular weight(s) hydrocarbon(s) with the deposition of carbonaceous materials on the utilizatori, causing the deactivation of the catalyst,

(C) removing from the reaction zone at least a portion of the deactivated granular catalytic material

(g) heating the first part remote from the reaction zone of granular catalytic material to a temperature of from about 700°to about 1200°C, carried out by direct or indirect contact with the gaseous combustion products produced by combustion of additional fuel,

(d) regenerative heated first part of granular catalytic material is a hydrogen-containing gas in the first regeneration zone under conditions including a first pressure and effective to convert at least part deposited on granular catalytic material of the carbonaceous material to methane,

(e) moving the second part remote from the reaction zone of granular catalytic material from the stage (in) stage or (d) in the second regeneration zone,

(g) regenerative second part of granular catalytic material is a hydrogen-containing gas in the second regeneration zone under conditions including a second pressure, different from the first pressure effective to convert at least part deposited on granular catalytic material of the carbonaceous material to methane, and

(C) returning at least cha is ti regenerated granular catalytic material from step (d) and at least part of the granular catalytic material from step (g) into the reaction zone.

In one embodiment of the invention, the reaction zone is a reaction zone with a moving (moving) layer of the catalyst and can work with a reverse temperature profile. The source material may also contain at least one of the following substances: CO, CO2N2N2O and/or O2. Remote granular catalytic material can be heated in stage (d) to a temperature of from about 800°to about 1000°C., in particular from about 850°to about 950°C. the Conditions in stage (d) regeneration may include an absolute pressure of at least 100 kPa, in particular from about 150 to 700 kPa. Conditions on stage (W) regeneration may include an absolute pressure of at least 500 kPa, in particular from about 1000 kPa to about 5000 kPa. Stage (d) and (g) regeneration can be done in a separate regeneration zones.

In one embodiment of the invention the conditions of the reaction occurring in the reaction zone at the stage (b)are non-oxidizing. At the stage (b) reaction conditions in the reaction zone can include a temperature from about 400°to about 1200°C., an absolute pressure from about 1 kPa to about 1000 kPa and volumetric gas flow rate from about 0.01 to h-1approximately 1000 h-1.

In one embodiment, the Khujand is the implementation of the invention, granular catalytic material is a catalyst dehydrocyclization containing the metal or its compound on an inorganic carrier. Granular catalytic material may include at least one of the following substances: molybdenum, tungsten, rhenium compound of molybdenum, a compound of tungsten, a compound of zinc and a compound of rhenium, ZSM-5, silica or aluminum oxide.

Brief description of drawings

Figure 1 presents the scheme of the method of conversion of methane to heavier hydrocarbons in accordance with the first variant implementation of the invention.

Figure 2 presents the scheme of the method of conversion of methane to heavier hydrocarbons in accordance with the second variant implementation of the invention.

Figure 3 shows a graph of the dependence of the yield of methane from the temperature at termoregulirovanija processing hydrogen zakoksovanie Mo/ZSM-5 catalyst of example 6, carried out in the reactor of ideal displacement when the dynamics of the temperature increase, constituting 5°C/min

Figure 4 presents a plot of the coke content (wt.%) from the time the regeneration when regenerierung hydrogen zakoksovanie Mo/ZSM-5 catalyst of example 7 at 850°C and different partial pressures of hydrogen.

On figa and 5B shows edit the appropriate degree of conversion of methane (%) and selectivity for benzene and toluene at elevated temperature regenerating from 875°C. to 925°C.

The implementation of the invention

In the context of the present description the term "more high molecular weight(s) hydrocarbon(s)" means hydrocarbon(s)containing more than one carbon atom per molecule; the oxygenate containing at least one carbon atom in the molecule such as ethane, ethylene, propane, propylene, benzene, toluene, xylenes, naphthalene and/or methylnaphthalene; and/or organic(s) compound(s)containing at least one carbon atom and at least one non-hydrogen atoms, for example methanol, ethanol, methylamine and/or ethylamine.

In the context of the present description, the term "aromatic hydrocarbon (HC)" means a molecule containing one or more aromatic rings. Examples of aromatic hydrocarbons are benzene, toluene, xylenes, naphthalene and methylnaphthalene.

The concept of "coke" and "carbonaceous material" as used in this description interchangeably to refer to carbon-containing materials, in which the reaction conditions are essentially non-volatile solids with low hydrogen content relative to the carbon content (in particular, with the value of the molar ratio N:less than 0.8, most likely less than 0.5). They may include crystalline graphite, graphite sheet or plate materials, GRAFITO the s fragments, amorphous carbon or other carbon-containing structure, in which the reaction conditions are essentially non-volatile solid materials. When it comes to solid coke, more solid coke or refractory or difficult-to-thermal processing of coke, have in mind the types of coke, which, due to either its structure or location is harder to remove by using a reagent (usually oxygen or hydrogen)used for the conversion of coke to gaseous materials.

In the context of the present description, the term "decontamination" of the catalyst means the loss over time of catalytic activity and/or selectivity. The catalyst is deactivated when its catalytic activity by at least 1% lower in the other case, by at least 5% lower in the other case, by at least 10% lower in the other case, by at least 15% lower in the other case, by at least 20% lower in the other case, by at least 25% lower in the other case, by at least 30% lower in the other case, by at least 35% lower in the other case at least 40% lower in the other case, by at least 45% lower in the other case, by at least 50% lower in the other case, by at least 55% lower in the other case, by at least 60% lower in the other case, by at least 65% lower in the other the om case, at least 70% lower, in another case, at least 75% lower in the other case, by at least 80% lower in the other case, by at least 85% lower in the other case, by at least 90% lower in the other case, by at least 95% or below in the other case, by at least 100% lower than the catalytic activity of fresh catalyst or regenerated catalyst. Not limiting myself to any theory, the deactivation of the catalyst can be considered as a phenomenon in which changes in the structure and/or the condition of the catalyst, which leads to the loss of active sites on the catalyst surface and thus causes deterioration of the catalyst efficiency. So, for example, deactivation of the catalyst may occur as a result of coke formation (coking), blocking active sites or dealuminated aluminosilicate molecular sieve as a result of treatment with water vapor.

In the context of the present description, the term reactor with a moving (moving) layer" means a zone or vessel by contacting solid particles and gaseous flows in such a way that the expenditure rate (U) of gas below the velocity required for the pneumatic entrainment of solid particles in the form of liquefied phase, with the aim of preserving a layer of solid particles with porosity less than 95%. A reactor with a moving bed can work in several dir is the max current, including the mode of settling movement (subsidence) or mode of movement of the compacted layer (U<Umf), the mode with sparging bubbles (Umf<U<Umb), the mode with channel - and pornobratva (Umb<U<Uc), mode, transition to turbulent and turbulent fluidization (Uc<U<Utr) and a mode with a high flow velocity (U>Utr). These different regimes of fluidization described, for example, in the works Kunii, D., Levenspiel, O., Chapter 3, Fluidization Engineering, 2nd edition, Butterworth-Heinemann, Boston, 1991, and Walas, S.M., Chapter 6, Chemical Process Equipment, Butterworth-Heinemann, Boston, 1990.

In the context of the present description, the term "support layer" is used to identify a zone or vessel where the particles (grains) in contact with gaseous flows in such a way that the expenditure rate (U) of gas is less than the minimum velocity required for fluidization of solid particles, the minimum velocity of fluidization (Umf), U<Umfin at least part of the reaction zone and/or zone or vessel operating at a speed which is above a minimum velocity of fluidization while maintaining the gradient properties of the gas and/or solids (such as temperature, composition of the gas or solids etc) along the axis from the bottom up in the reactor layer when applying the internals of the reactor in order to minimize reverse paramesh the of gas and solids. Description minimum velocity of fluidization see, for example, Chapter 3 in "Fluidization Engineering", D.Kunii and O.Levenspiel, 2nd edition, Butterworth-Heinemann, Boston, 1991, and Chapter 6 work, "Chemical Process Equipment" S.M. Walas, Butterworth-Heinemann, Boston, 1990, the contents of which are fully integrated into the description by reference.

In the context of the present description, the term "reactor fluidized bed" is used to identify a zone or vessel, where the solid particles are in contact with gaseous flows in such a way that the expenditure rate (U) of gas sufficient for fluidization of solid particles (i.e. above the minimum velocity of fluidization Umfand below the velocity required for the pneumatic entrainment of solid particles in the form of liquefied phase, with the aim of preserving a layer of solid particles with porosity less than 95%. Used in the present description the term "cascading fluid layers" is used to denote sequential location separate fluidized beds so that this may be the gradient properties of the gas and/or solids (such as temperature, composition of the gas or solids, pressure and so on) as you move solids or gas from one fluidized bed with the change of level of the cascade to another layer. A graph of the minimum velocity of fluidization is given, for example, in the above published is the R works Kunii and Walas.

In the context of the present description the term "Elevator-reactor" refers to an area or receptacle (such as a vertical cylindrical pipe)used for purely vertical transport of solid particles in the regimes of fluidization with high flow rate (fast fluidization or fluidization with pneumatic transport. Regimes of fluidization with high flow rate and fluidization with pneumatic characterized by a consumable velocity (U) of gas greater than the transport speed (Utr). Regimes of fluidization with high flow rate and fluidization with pneumatic transport is also described in the above published works Kunii and Walas.

In the context of the present description indications heating indirect contact with the gaseous combustion products include heat transfer through the surface heat transfer and/or application of the fluid (gas, liquid or solid), which is heated gaseous products of combustion and which gives its heat to the granular catalytic material.

In the present invention proposes a method for aromatic hydrocarbons by contacting metadatareader source material, usually with H2H2O, O2WITH and/or CO2with granular catalyst dehydrocyclization in the reaction C is not in the conditions, effective to convert the methane to aromatic hydrocarbons and hydrogen. As reaction on the catalyst accumulates coke, thereby lowering the activity or selectivity of the catalyst, and hence, from the reaction zone continuously or periodically take part zakoksovanie catalyst and direct it to a separate regeneration zone where zakochany the catalyst is introduced into contact with a hydrogen-containing regenerating gas. Since the reaction of dehydrocyclization is endothermic, SamAccountName the catalyst withdrawn from the reaction zone, direct or indirect contact with the gaseous combustion products produced by combustion of additional fuel, bring the heat to raise its temperature to the required temperature of the regenerate, is usually component of from about 700°to about 1200°C. Then the part is heated zakoksovanie catalyst may be recycled to the reaction zone to provide heat for the reaction of dehydrocyclization, and the residue is heated catalyst is introduced into contact with a hydrogen-containing regeneration gas in the regeneration zone under such conditions that at least part of the coke deposited on the catalyst, converted to methane. Next, the regenerated catalyst back to the reaction zone.

In one Varian is E. the invention, the regenerative spend removal from the reaction zone of the two or more portions zakoksovanie catalyst, supply of heat to those portions of the catalyst and contacting the heated portions of the catalyst with a hydrogen-containing gas in a separate regeneration zones, working in such conditions that the partial pressure of hydrogen in at least two regeneration zones.

In addition, the invention provides a method of utilization of hydrogen produced as a by-product of the reaction of dehydrocyclization and, in particular, the method of turning at least part of the hydrogen into more valuable products.

Source material

In the proposed invention the method can be used any metadatabase source material, but in General the proposed method is intended for use as a source material gas. Other suitable matenadarani source materials include materials derived from such sources as coal seams, landfills, fermentation of agricultural or municipal waste and/or gaseous streams neftepererabativaushi enterprises.

Containing methane source materials, such as natural gas, as a rule, in addition to methane contains carbon dioxide and ethane. Ethane and other aliphatic hydrocarbons which may be contained in the original material on stage dehydrocyclization can be, of course, is not asheni in the target aromatic products. In addition, as discussed below, carbon dioxide can also be turned into useful aromatic products either directly on the stage of dehydrocyclization or indirectly by conversion to methane and/or ethane at a stage of reduction of hydrogen content.

In the proposed invention the method containing methane threads also usually contain nitrogen and/or sulfur-containing impurities that before using these threads can be deleted or their number can be reduced to low concentrations. In one of the embodiments of the invention, the source material supplied to the stage of dehydrocyclization, contains less than 100 ppm million, for example less than 10 ppm million, in particular less than 1 ppm million, each of nitrogen and sulphur compounds.

To facilitate the reduction of coke formation in addition to methane in the source material supplied to the stage of dehydrocyclization, you can add at least one of the following substances: hydrogen, water, carbon monoxide and carbon dioxide. These additives can be introduced in a separate jointly supplied raw material streams, or they may be present in the methane stream, for example, when methane flux receive as derived from natural gas containing carbon dioxide. Other sources of carbon dioxide can be in order to carry flue gases, installation of LNG, hydrogen unit, ammonia plant, picoline installation and paleoamerican installation.

In one embodiment of the invention, the source material supplied to the stage of dehydrocyclization contains carbon dioxide and comprises about 90 mol.% about to 99.9 mol.%, in particular from about 97 mol.% to about 99 mol.% methane and from about 0.1 mol.% up to about 10 mol.%, in particular from about 1 mol.% up to about 3 mol.% CO2. In another embodiment of the invention, the source material supplied to the stage of dehydrocyclization contains carbon monoxide and includes from about 80 mol.% about to 99.9 mol.%, in particular from about 94 mol.% to about 99 mol.%, methane and from about 0.1 mol.% to about 20 mol.%, in particular from about 1 mol.% until about 6 mol.% WITH. In yet another embodiment of the invention, the source material supplied to the stage of dehydrocyclization contains water vapor and comprises about 90 mol.% about to 99.9 mol.%, in particular from about 97 mol.% to about 99 mol.% methane and from about 0.1 mol.% up to about 10 mol.%, in particular from about 1 mol.% up to about 5 mol.% water vapour. In one embodiment of the invention, the source material supplied to the stage of dehydrocyclization, contains hydrogen and comprises from about 80 mol.% prima is but to 99.9 mol.%, in particular from about 95 mol.% to about 99 mol.% methane and from about 0.1 mol.% to about 20 mol.%, in particular from about 1 mol.% up to about 5 mol.% of hydrogen.

The source material supplied to the stage of dehydrocyclization may also contain heavier hydrocarbons than methane, including aromatic hydrocarbons. Such heavier hydrocarbons may be returned to the process from the stage of reducing the amount of hydrogen added as a separate jointly supplied raw materials or may be present in the methane stream, for example in the case when the original material is in the form of natural gas contains ethane. To heavier hydrocarbons are returned to the process from the stage of reduction of hydrogen content, typically include monocyclic aromatic compounds and/or paraffins and olefins, mainly containing 6 or less, in particular 5 or less, for example 4 or less, usually 3 or less carbon atoms. Usually the source material supplied to the stage of dehydrocyclization, contains less than 5 wt.%, in particular less than 3 wt.% hydrocarbons With3+.

Dehydrocyclization

At the stage of dehydrocyclization proposed in the invention method metadatabase source material is introduced into contact with the granular catalyst dehydrocyclization in the conditions,typically non-oxidizing and as a rule - reduction, effective for the conversion of methane to heavier hydrocarbons, including benzene and naphthalene. The main result reactions are:

Monoxide and/or carbon dioxide which may be present in the source material, increases the activity and stability of catalyst, promoting reactions such as:

but negatively affects the balance, enabling the flow parallel to the resulting reaction such as:

.

In the proposed invention the method can be used any catalyst dehydrocyclization, effective for the conversion of methane into aromatic compounds, although typically, the catalyst includes a metal component, in particular a transition metal or its compound on an inorganic carrier. In a suitable embodiment of the invention the metal component is present in amount from about 0.1 wt.% to about 20 wt.%, in particular from about 1 wt.% up to about 10 wt.%, by weight of the total catalyst. Typically, the metal present in the catalyst in the form of free items or in the form of carbide.

Suitable metal which is a mini-components of the catalyst include calcium, magnesium, barium, yttrium, lanthanum, scandium, cerium, titanium, zirconium, hafnium, vanadium, niobium, tantalum, chromium, molybdenum, tungsten, manganese, rhenium, iron, ruthenium, cobalt, rhodium, iridium, Nickel, palladium, copper, silver, gold, zinc, aluminum, gallium, silicon, germanium, indium, tin, lead, bismuth and transuranic metals. These metal components may be present in the form of free elements or compounds of metals, such as oxides, carbides, nitrides and/or phosphides, and they can be used independently or in combination. As one of the metal components can also be used platinum and osmium, but they are generally not preferred.

The inorganic carrier may be either amorphous or crystalline and may constitute, in particular, oxide, carbide or nitride of boron, aluminum, silicon, phosphorus, titanium, scandium, chromium, vanadium, magnesium, manganese, iron, zinc, gallium, germanium, yttrium, zirconium, niobium, molybdenum, indium, tin, barium, lanthanum, hafnium, cerium, tantalum, tungsten or other transuranic elements. In addition, the carrier may be a porous material, such as microporous crystalline material or mesoporous material. In the context of the present description, the term "microporous" refers to pores with diameter less than 2 nm, whereas the term "mesoporous the" refers to pores with a diameter from 2 to 50 nm.

Suitable microporous crystalline materials include silicates, aluminosilicates, titanosilicates, alumophosphate, metallophosphates, kriminaalreformi or mixtures thereof. Such microporous crystalline materials include materials with frames types MFI (for example, ZSM-5 and silicalite), MEL (for example, ZSM-11), MTW (for example, ZSM-12), TON (for example, ZSM-22), MTT (e.g., ZSM-23), FER (for example, ZSM-35), MFS (for example, ZSM-57), MWW (e.g., MCM-22, PSH-3, SSZ-25, ERB-1, ITQ-1, ITQ-2, MCM-36, MCM-49 and MCM-56), IWR (for example, ITQ-24), KFI (for example, ZK-5), BEA (e.g., zeolite beta), ITH (for example, ITQ-13), MOR (e.g., mordenite), FAU (for example, zeolites X, Y, ultrastabilized Y and dealuminated Y), LTL (for example, zeolite L), IWW (for example, ITQ-22), VFI (for example, VPI-5), AEL (for example, SAPO-11), AFI (e.g., ALPO-5) and AFO (SAPO-41), as well as materials such as MCM-68, EMM-1, EMM-2, ITQ-23, ITQ-24, ITQ-25, ITQ-26, ETS-2, ETS-10, SAPO-17, SAPO-34 and SAPO-35. Suitable mesoporous materials include MCM-41, MCM-48, MCM-50, FSM-16 and SBA-15.

Examples of preferred catalysts include molybdenum, tungsten, zinc, rhenium, as well as their connections and combinations on ZSM-5, silica or aluminum oxide.

The metal component may be dispersed on an inorganic carrier by any means well known in the art, such as coprecipitation, impregnation on capacity, evaporation, conventional impregnation, spray the drying, Sol-gel, ion exchange, chemical vapor deposition, diffusion, and physical mixing. In addition, the inorganic carrier may be modified by known methods such as treatment with water vapor, acid washing, washing with caustic soda and/or processing of silicon-containing compounds, phosphorus-containing compounds and/or elements or compounds of elements of groups 1, 2, 3 and 13 of the Periodic table of elements. Such modifications can be used to change the surface activity of the media, as well as to hamper or enhance access to any internal porous structure of the media.

In some embodiments of the invention in the reaction dehydrocyclization, in addition to the catalytic granular material may be non-catalytic granular material. Non-catalytic granular material can be used as material for energy transport (heat) in the system and/or to fill the space, depending on the needs of maintaining the required hydrodynamic conditions. Non-catalytic granular material may form the particulates without binders or particles can be bound inorganic binder, such as clay, silicon dioxide, aluminum oxide, zirconium dioxide or other oxide, m is metal, used to help preserve the physical integrity of the particles. In a preferred embodiment, the particles have an essentially spherical shape. Examples of acceptable non-catalytic granular material are silicon dioxide, aluminum oxide, ceramics and silicon carbide with a small specific surface area.

Stage dehydrocyclization carry out the introduction of metastorage source material in contact with the catalyst dehydrocyclization in the reaction zones with one or more fixed bed, moving layers or fluid layers. Usually the source material into the reaction zone of the or each reaction zone is introduced into contact with a moving bed of catalyst dehydrocyclization, where the source material moves in counterflow to the direction of movement of the catalyst dehydrocyclization. In one embodiment of the invention, the reaction zone comprises a reactor with a deposited layer, under which imply vertically positioned reactor, in which the granular catalyst enters the top of the reactor or near it, and moves under its own weight with the formation of the catalyst layer, while the source material is fed into the reactor at the base of the reactor or near it and moves upward through the catalyst bed.

In the embodiment, the image is placed using the deposited layer, the movement of the catalyst dehydrocyclization in the reaction zone is almost free of fluidization. In the context of the present description the term "practically free from the fluidization means that the average speed of the gas flow in the reactor below the minimum velocity of fluidization. The concept of "practically free from fluidized" in the context of the present description also means that the average speed of the gas flow in the reactor is less than 99%, in particular less than 95%, usually less than 90%, even less than 80%, the minimum velocity of fluidization. When the reaction zone or each reaction zone operates with a deposited layer of granular catalytic material and/or any granular non-catalytic material is characterized by an average size of particles constituting from about 0.1 mm to about 100 mm, in particular from about 1 mm to about 5 mm, for example from about 2 mm to about 4 mm, In some embodiments of the invention, at least 90 wt.% granular catalytic material and/or at least 90 wt.% granular non-catalytic material has a particle size of about 0.1 mm to about 100 mm, in particular from about 1 mm to about 5 mm, for example from about 2 mm to about 4 mm

Alternatively, the reaction dehydrocyclization carried out in several series-connected reactors, fluidized bed, in which the granular catalyst to Cascadia the t in one direction from one reactor to the next adjacent reactor in this series, while the source material is passed in the opposite direction through the reactor and between them. When each reaction zone works fluidized bed catalytic granular material and/or any non-catalytic granular material is characterized by an average size of particles constituting from about 0.01 mm to about 10 mm, in particular from about 0.05 mm to about 1 mm and in particular from about 0.1 mm to about 0.6 mm In some embodiments of the invention, at least 90 wt.% catalytic granular material and/or at least 90 wt.% non-catalytic particulate material has a particle size of about 0.01 mm to about 10 mm, in particular from about 0.05 mm to about 1 mm, for example from about 0.1 mm to about 0.6 mm

Typically, the ratio of the mass flow catalytic granular material plus any non-catalytic granular material to the mass flow of the hydrocarbon starting material in the reaction zone of dehydrocyclization or each reaction zone of dehydrocyclization is from about 1:1 to about 100:1, in particular from about 1:1 to about 40:1, in particular from about 5:1 to 20:1.

The reaction dehydrocyclization is endothermic, and therefore as the course of the reaction temperature in each reaction zone dehydrocyclization and usually tends to decrease from the maximum temperature to the minimum temperature. Acceptable conditions for stage dehydrocyclization include a maximum temperature of from about 700°to about 1200°C., in particular from about 800°to about 950°C, and minimum temperature from about 400°to about 800°C., in particular from about 500°to about 700°C. However, as discussed below, to reduce the temperature drop during the reaction of dehydrocyclization in this reaction down the heat, and therefore in some configurations it may be possible to reduce the difference between the maximum and minimum temperatures essentially to zero. In another embodiment of the invention, by introducing into the reaction dehydrocyclization heated catalyst, you can create a reverse temperature profile, i.e. the reaction temperature at the outlet for the process gas is higher than the temperature at the inlet for the process gas.

In one embodiment of the invention for creating a reverse temperature profile in the system for the reaction of dehydrocyclization movement of the source material and granular catalyst dehydrocyclization organize in a counter, so that, despite the endothermic nature of the reaction dehydrocyclization, the difference between the reaction temperature for gaseous exhaust stream at the output of the system for the reaction of dehydrocyclization and the temperature is Oh reaction for metadatareader source material on the login for the reaction of dehydrocyclization is at least +10°C, in particular at least +50°C, for example at least +100°C and even at least +150°C.

In any case, since the reaction of dehydrocyclization is endothermic, granular catalytic material enters the system for the reaction of dehydrocyclization at the first, high temperature, typically from about 800°to about 1200°C., in particular from about 900°to about 1100°C, and out of the reaction system at a second, lower temperature, typically from about 500°to about 800°C., in particular from about 600°to about 700°Sabsa the temperature difference between the granular catalytic material after passing through the reaction zone is at least 100°C.

Other conditions that are generated in the reaction of dehydrocyclization, generally include a pressure of from about 1 kPa to about 1000 kPa, in particular from about 10 kPa to about 500 kPa, for example from about 50 kPa to about 200 kPa, and the volumetric gas flow rate of approximately from 0.01 h-1approximately 1000 h-1in particular from about 0.1 h-1approximately 500 h-1in particular from about 1 h-1approximately 20 h-1. In a suitable embodiment of the invention stage of dehydrocyclization carried out in the absence of O2.

The main components of the flow, okadama what about the stage with dehydrocyclization, are hydrogen, benzene, naphthalene, carbon monoxide, ethylene and unreacted methane. This waste stream typically contains at least 5 wt.%, in particular at least 10 wt.%, for example at least 20 wt.%, preferably at least 30 wt.%, more aromatic rings than the original material.

Benzene and naphthalene are separated from the stream of exhaust from the stage of dehydrocyclization, for example, by solvent extraction and subsequent separation into fractions, and these substances can get into the product flow. However, as discussed below, before or after extraction of the products of at least part of these aromatic components may be subjected to alkylation of more valuable materials such as xylenes. Moreover, as discussed below, proposed in the invention method provides for the utilization of hydrogen produced as a by-product of the reaction of dehydrocyclization, in particular the transformation of at least part of the hydrogen into more valuable products.

Regenerative catalyst

The reaction dehydrocyclization characterized by a tendency to the deposition of coke on the catalyst, and means for maintaining the activity of the catalyst dehydrocyclization at least part of the catalyst must be regenerated continuously or periodically. Usually e the CSO reach the outlet from the reaction zone or each reaction zone portion of the catalyst, performed on a periodic or continuous basis, and move this part of the catalyst in a separate regeneration zone. In the regeneration zone zakochany the catalyst dehydrocyclization enter into contact with a hydrogen-containing gas under conditions effective to convert at least part of the deposited thereon of carbonaceous material to methane. Typically, the hydrogen-containing gas does not contain significant quantities of methane or other hydrocarbons; hydrocarbons, typically less than 20 mol.%, in particular less than 18 mol.%, less than 15 mol.%, less than 10 mol.%, less than 8 mol.%, less than 6 mol.%, less than 4 mol.% or less than 2 mol.%. In one embodiment of the invention, the hydrogen required for the regenerate receive, at least partially, from a hydrogen-containing stream, the exhaust from the reaction of dehydrocyclization.

In a suitable embodiment of the invention the conditions regenerating include a temperature approximately ranging from 700°C. to about 1200°C., in particular from about 800°to about 1000°C., in particular from about 850°to about 950°C. and an absolute pressure of at least 100 kPa, in particular from about 150 kPa to about 5000 kPa. Usually zakochany the catalyst dehydrocyclization removed from the reaction zone or from each reaction zone, reducing isoamsa lower temperature, than what is optimal for the regenerate, and then remove the catalyst was first heated to the desired temperature regenerating direct or indirect (mediated) by contacting the combustion gases produced by the combustion of additional fuel. The heating is carried out in the heating zone, which may be in the same vessel, and a regeneration zone, or which may be in a vessel separate from the regeneration zone.

Under "additional fuel source" refers to the fact that the fuel source is physically separated from the catalyst, and therefore is not a coke formed on the catalyst as a by-product of the reaction of dehydrocyclization. As a rule, an additional source of fuel includes a hydrocarbon, such as methane, in particular, an acceptable fuel source is natural gas used as a source of material sent to the process. In a suitable embodiment of the invention in the heating zone of the support oxygen-poor atmosphere, allowing combustion of hydrocarbon fuels to heat the first part of the catalyst is formed synthesis gas, which can then be used for additional hydrocarbon product and/or fuel. In addition, in the case of a direct those whom lapertosa to the catalyst dehydrocyclization the use of oxygen-poor atmosphere prevents oxidation of the metal carbides, present in the catalyst, and minimizes the average partial pressure of water vapor, thereby reducing hydrothermal aging of the catalyst.

In another embodiment of the invention suitable additional fuel source is hydrogen, in particular of the hydrogen generated as a by-product of the reaction of aromatization.

If the catalyst dehydrocyclization heated by direct, or directly, zakochany the catalyst withdrawn from the reaction zone, it is advisable to enter into contact directly with the burning fuel source in the heating zone. In another embodiment of the invention the source of fuel is burned in a separate combustion zone, gaseous products of combustion formed in the combustion zone, is directed into the heating zone for heating the catalyst. In yet another embodiment of the invention the catalyst dehydrocyclization may be heated by indirect heat exchange, for example, using gaseous products of combustion for heating the inert medium (gas, liquid or solid) or the heat transfer surface and then contacting zakoksovanie catalyst with a heated inert atmosphere or surface heat transfer.

In one practical embodiment of the invention, the heater h is is elongated, and zakochany catalyst is passed through the heating zone from the entrance, located at one end of the heating zone or near it, to the outlet located at the other end of the heating zone or near it, and supply heat to the first part of the catalyst is carried out in many places, distributed along the length of the heating zone. Thus you can distribute the heat input to the catalyst along the length of the heating zone, thereby reducing to a minimum temperature on the catalyst surface and internal gradients.

If the first part of the catalyst is heated by direct contact with the burning fuel source in the heating zone, the gradual heating of the catalyst can be achieved by feeding essentially all of the additional fuel into the inlet end of the heating zone, and then feeding this heating zone of oxygen-containing gas with a certain step in a specified set of locations distributed along the length of the heating zone. In another embodiment of the invention, the entire oxygen-containing gas required for the combustion of additional fuel can be fed into the input end of the heating zone, and additional fuel with a certain step in a specified set of locations distributed along the length of the heating zone.

If the first part is utilizator heated by direct contact with hot gaseous combustion products, formed in a separate combustion zone, gradual heating of the catalyst can be achieved by the flow of hot gaseous combustion products in a specified number of sites distributed along the length of the heating zone.

In one embodiment of the invention the heating zone is a vertical pipe, and during the stage of re-heating the first portion of catalyst is passed upward through a vertical pipe. In practice, the heating zone may include several vertical tubes connected in parallel. In another embodiment of the invention the heating zone may include a movable layer of a specified catalyst.

In one embodiment of the invention zakochany the catalyst dehydrocyclization removed from the reaction zone is divided at least into two parts, which is heated as described above, and then sent to a separate regeneration zone operating under different pressures. For example, one regeneration zone operates at an absolute pressure of at least 100 kPa, in particular from about 150 kPa to about 700 kPa, as described above, and the other regeneration zone operates at an absolute pressure of at least 500 kPa, in particular from about 1000 kPa to about 5000 kPa. Thus, it was found, for which it is demonstrated in the examples, what is regenerative at higher pressure provides a more rapid removal of coke, and the destruction of more refractory coke. However, it was also found that removal of coke at higher pressures require more expensive equipment. For this reason, it may be desirable part of the coke is removed at a lower partial pressure of hydrogen of less expensive equipment, and another part of the coke is at a higher partial pressure of hydrogen on more expensive equipment.

Regeneration zone or each regeneration zone can be a reactor operating as a reactor with a fluidized bed, fluidized bed, settling layer, Elevator reactor or a combination thereof. In practice, each regeneration zone can include multiple reactors, in particular multiple-lift reactors connected in parallel or multiple reactors connected in series, in particular Elevator-reactor followed by a reactor with a deposited layer. After regenerating the catalyst back to the reaction zone.

In another embodiment of the invention, in particular, if the reaction dehydrocyclization carried out in a reactor with a fixed bed, regenerative can be performed without removing the catalyst from the reaction zone, temporarily interrupting the udachu metadatareader source material into the reaction zone, heating the reaction zone to a temperature of regenerating comprising from about 700°to about 1200°C, direct and/or indirect contact with the gaseous combustion products produced by combustion of additional fuel, regenerating granular catalytic material is a hydrogen-containing gas and then resuming the flow metadatareader source material into the reaction zone. Of course, that the heating of the reaction zone to a temperature of the regenerate can be done prior to the interruption metadatareader source material.

Re-heating of the catalyst

Since the reaction of dehydrocyclization is endothermic, the reaction should bring warm. In the proposed invention the method is conveniently achieved by removal of the catalyst from the reaction zone either on a periodic or continuous basis, the supply of heat to the catalyst and the subsequent return of heated catalyst back to the reaction zone. Because the stage of the regenerate hydrogen, described above, also involves the heating of the catalyst and the subsequent return of the hot regenerated catalyst back to the reaction zone, one possible way of applying heat to the reaction dehydrocyclization is in the process of regenerating.

In another embodiment, invented the I part of the heat or the heat necessary to maintain the reaction dehydrocyclization, can be provided as a separate stage of the re-heating of the catalyst. In this embodiment, the portion of the catalyst allowed for the reaction zone is moved to a separate heating zone, where the catalyst is again heated by direct or indirect contact with the hot gaseous products of combustion resulting from the combustion of additional fuel source. Then the heated catalyst back to the reaction zone with regenerierung hydrogen or without it.

Re-carbonization catalyst

It must be borne in mind that heating of the catalyst dehydrocyclization for the purposes of regenerating and/or to transfer the heat back into the reaction dehydrocyclization may expose the catalyst to high temperature oxidizing conditions, especially if heating of the catalyst involves direct contact with the hot gaseous combustion products. The result is present in the catalyst dehydrocyclization metals such as rhenium, tungsten or molybdenum, during the heating stage may move from their catalytically active element or carbide forms at the oxide substances. Thus, before returning to the reaction zone is regenerated and/or re-heated catalyst may be moved to the treatment area is utilizator, separate from the regeneration zone, the heating zone and the reaction zone, where the catalyst is in contact with a carburizing gas containing at least one hydrocarbon selected from methane, ethane, propane, butane, isobutane, hexane, benzene and naphthalene. In some cases, the carburizing gas may also contain at least one of the following substances: CO2WITH H2N2Oh and inert diluents. In another embodiment of the invention carburizing gas may be a mixture of hydrogen and at least one of co and CO2. Moreover, you may need the introduction of the catalyst into contact successively with many different carburizing gases, each of which comprises a hydrocarbon selected from methane, ethane, propane, butane, isobutane, hexane, benzene and naphthalene, or a mixture of hydrogen and at least one of co and CO2.

Normally, the maximum temperature in the treatment area of the catalyst is from about 400°to about 1100°C., in particular from about 500°to about 900°C, and minimum temperature ranges from 300°C to 500°C. typically, the treatment area of the catalyst operates at an absolute pressure comprising from 10 to 100 psi (69 to 690 kPa), in particular in the range from 15 to 60 pounds/CVD im (from 103 to 414 kPa). Usually, the average length of stay of catalytic particles in the treatment zone of the catalyst is in the range from 0.1 to 100 minutes, for example in the range from 1 to 20 minutes under these conditions carbonizing gas interacts with the metal oxide materials on the catalyst surface, returning the metal in its catalytically active element or carbide form. In addition, the carburizing gas is able to interact with the active centers of the surface of the catalyst carrier, reducing their tendency to form coke in the reaction zone of dehydroaromatization.

To maintain the temperature required for carburizing regenerated catalyst, heat can lead to catalyst and/or carbonizing gas before stage carburizing or during its implementation. So, for example, heat to the catalyst can be fed indirect heating, the introduction into contact with the hot gases coming from the reaction zone or heating zone, introducing into contact with a hot gaseous stream originating from a process of carbonization, or by mixing with the heated catalyst from the heating zone. Heat convenient to take to a carbonizing gas with the outside of the furnace or heat exchanger or through the heated catalyst from the heating zone.

The treatment area of the catalyst m who can work as a reactor with a fluidized bed, the fluidized bed reactor, a reactor with a deposited layer, the lift-reactor or circulating-lift reactor. In one embodiment, the treatment area of the catalyst consists of a reactor with a deposited layer. In another embodiment of the invention, the treatment area of the catalyst includes a single reactor with a fluidized bed with internal pneumatic valves to prevent back mixing, or more fluidized bed reactor, connected in series, and the regenerated catalyst kaskaderom between adjacent reactors. In any case, the contact of the catalyst with a gas in the processing zone of the catalyst contribute by organizing movement of the regenerated catalyst and carbonizing gas in the processing zone of the catalyst in opposite directions. Using such a counter-current movement in the treatment area of the catalyst may be formed of a temperature profile, whereby carbonization of the regenerated catalyst is initially at a low temperature, but as the catalyst through the layer temperature nagarajuna increases.

In some cases it may be desirable to stage carburizing first enter the heated aregenerally the catalyst into contact with a rich H2flow is La partial or full recovery of the metal component of the catalyst. Also, it may be desirable to subject carburized catalyst subsequent treatment with hydrogen and/or carbon dioxide to remove any excess carbon, which could secede on the catalyst at the stage of carbonization.

In practice, as the reaction dehydrocyclization in the process add fresh catalyst dehydrocyclization for loss of catalyst due to mechanical wear or decontamination, and despite the existence of many means of adding fresh catalyst, in order to avoid destruction of the catalyst fresh catalyst is usually desirable to add to the scope of the process operating at a temperature below the maximum temperature in each reaction zone of dehydrocyclization. In one embodiment of the invention the fresh catalyst dehydrocyclization added to the process by the introduction into the treatment area of the catalyst, due to which fresh catalyst is introduced into contact with the carburizing gas before moving into the reaction zone for introducing into contact with matenadarani source material. In another embodiment of the invention, the catalyst can be added to areas of lower temperature reactor system with an inverse temperature profile.

Control of hydrogen content

Since hydrogen is one of the main compo is having the exhaust from the stage of dehydrocyclization thread this waste stream after extraction of aromatics may be processed to reduce the hydrogen content in the exhaust stream before returning unreacted methane on stage dehydrocyclization to achieve maximum utilization of the source material. Stage reduction of hydrogen content, as a rule, involves the interaction of at least part of the hydrogen contained in the exhaust from the stage of dehydrocyclization the flow of oxygen-containing materials such as CO and/or CO2with the receiving water and the second exhaust stream having a lower content of hydrogen in comparison with the first exhaust (from the stage of dehydrocyclization) thread. Appropriate ways of reducing the hydrogen content is described below and in international application PCT/US2005/044042 (publication WO/2006/068814)filed by the authors of the present invention to December 2, 2005

In a suitable embodiment, the stage of reduction of hydrogen content includes (I) mahanirvana and/or tonirovanie, (II) the Fischer-Tropsch process, (III) synthesis of alcohols With1-C3in particular methanol and other oxygenates, (IV) synthesis of light olefins, paraffins and/or aromatics by methanol or dimethyl ether as an intermediate product and/or (V) selective combustion of hydrogen. To get the biggest win of Eristavi can be done sequentially; for example, initially, there may be performed the Fischer-Tropsch process with receiving stream enriched in hydrocarbons2+, with subsequent mahanirvana to achieve a high degree of conversion of N2.

Usually at the stage of reducing the content of hydrogen and hydrocarbons are formed, as described below, and in this case, after separation of simultaneous water at least part of the hydrocarbons, it is advisable to return to the stage of dehydrocyclization. For example, if the hydrocarbons produced at the stage of reduction of hydrogen content, include paraffins and olefins, part of the hydrocarbons returned to the stage of dehydrocyclization usually includes paraffins or olefins with six or less carbon atoms, in particular with five or less carbon atoms, for example with four or less carbon atoms or with three or less carbon atoms. If the hydrocarbons produced at the stage of reduction of hydrogen content, include aromatic compounds, it is reasonable that part of the hydrocarbons returned to the stage of dehydrocyclization included monocyclic aromatic substances.

Mahanirvana/tonirovanie

In one embodiment of the invention stage reduction of hydrogen content comprises the reaction of at least part of the hydrogen present in about the walking with the stage of dehydrocyclization thread with carbon dioxide to form methane and/or ethane in accordance with the following reactions result:

In an expedient embodiment of the invention used carbon dioxide is part of the natural gas stream, as a rule, the same stream of natural gas that is used as the source material supplied to the stage of dehydrocyclization. When carbon dioxide is part metadatareader flow, the ratio of CO2:CH4in this thread it is advisable to maintain in the range of about from 1:1 to about 0.1:1. A mixture containing carbon dioxide stream and the exhaust from the stage of dehydrocyclization flow, it is advisable to provide a gaseous source materials into the inlet of the jet pump.

At a stage of reduction of hydrogen content with the capture of methane or ethane, as a rule, is used the molar ratio of N2:CO2close to the stoichiometric proportions required for the target reaction 6 or reactions 7, though, if you want to get containing CO2or containing H2the second waste stream, in this stoichiometric ratio can be made small changes. Stage reduction of hydrogen content with the capture of methane or ethane expediently carried the ü in the presence of a bifunctional catalyst, containing a metal component, in particular a transition metal or its compound on an inorganic carrier. Suitable metal components include copper, iron, vanadium, chromium, zinc, gallium, Nickel, cobalt, molybdenum, ruthenium, rhodium, palladium, silver, rhenium, tungsten, iridium, platinum, gold, gallium and combinations thereof, and connections. The inorganic carrier may be an amorphous material such as silicon dioxide, aluminum oxide or silicon dioxide/aluminum oxide or similar to those listed for catalyst dehydroaromatization. In addition, the inorganic carrier may be a crystalline material, such as microporous or mesoporous crystalline material. Suitable porous crystalline materials include aluminum silicates, alumophosphate and kriminaalreformi listed above for catalyst dehydrocyclization.

Stage reduction of hydrogen content with the capture of methane and/or ethane may be conducted in a wide range of conditions, including a temperature from about 100°C. to about 900°C., in particular from about 150°C. to about 500°C., for example from about 200°to about 400°C., a pressure from about 200 kPa to about 20000 kPa, in particular from about 500 kPa to about 5000 kPa, and the volumetric gas flow rate from about 0.1 h-1approximately 10000 h-1in , particularly from about 1 h -1approximately 1000 h-1. The values of the degree of conversion of CO2as a rule, are in the range from 20 to 100%, and more preferably 90%, in particular more than 99%. This exothermic reaction can be conducted in several layers of the catalyst heat between the layers. In addition, to obtain the maximum possible kinetic velocities, the process in the first (first) flow layer (layers) can be performed at higher temperatures, and to obtain the maximum possible thermodynamic transformation process in the last (recent) thread layer (layers) can be performed at lower temperatures.

The main products of this reaction are water and, depending on the molar ratio of N2:CO2, methane, ethane and higher alkanes together with some unsaturated hydrocarbons containing two or more carbon atoms. In addition, some preferred partial hydrogenation of carbon dioxide to carbon monoxide. After removal of water, methane, carbon monoxide, any unreacted carbon dioxide and heavier hydrocarbons can be sent directly to the stage of dehydrocyclization for more aromatic products.

The process of Fischer-Tropsch

In yet another embodiment of the invention stage reducing the amount of water the ode involves the reaction of at least part of the hydrogen, present in the exhaust from the stage of dehydrocyclization stream with carbon monoxide according to the method of Fischer-Tropsch obtaining paraffins and olefins With2-C5.

the Fischer-Tropsch process is well known in the art (see, for example, US patents 5348982 and US 5545674, the contents of which are incorporated into this description by reference). This process typically involves the reaction of hydrogen and carbon monoxide in a molar ratio of about 0.5:1 to about 4:1, in particular from about 1.5:1 to about 2.5:1, at a temperature of from about 175°C. to about 400°C., in particular from approximately 180°To approximately 240°C. and at a pressure from about 1 bar to about 100 bar (100 to 10,000 kPa), in particular from about 10 bar to about 40 bar (1000 to 4000 kPa), in the presence of a catalyst Fischer-Tropsch, usually supported on a carrier or used without substrate element of group VIII base metal such as Fe, Ni, Ru, Co, promoter or without it, for example, ruthenium, rhenium, hafnium, zirconium, titanium. The media, when used, can serve as refractory oxides of metals, such as metals of group IVB, i.e. titanium dioxide, zirconium dioxide or silicon dioxide, aluminum oxide or silicon dioxide/aluminum oxide. In one embodiment of the invention, the catalyst comprises a non-conversion catalyst, n is an example cobalt, or ruthenium, in particular cobalt, rhenium or zirconium as a promoter, in particular cobalt and rhenium deposited on a silicon dioxide or titanium dioxide, usually titanium dioxide.

In yet another embodiment of the invention the catalyst for synthesis of hydrocarbons includes a metal such as si, Cu/Zn and Cr/Zn, media ZSM-5, and the process is conducted with obtaining significant quantities of monocyclic aromatic hydrocarbons. An example of such a process is described in the work of Jose Erena Study of Physical Mixtures of Cr2O3-ZnO and ZSM-5 Catalysts for the Transformation of Syngas into Liquid Hydrocarbons; Ind. Eng. Chem Res. 1998, 37, 1211-1219, the contents of which are incorporated into this description by reference.

Emit liquid Fischer-Tropsch process, i.e. With5+and heavier hydrocarbons to separate light gases such as unreacted hydrogen and CO, hydrocarbons With1-C3or4and the water. Then heavier hydrocarbons can be selected as products or aimed at the stage of dehydrocyclization for more aromatic products.

The presence of carbon monoxide required for the reaction of the Fischer-Tropsch process, can be fully or partially provided by the carbon monoxide present in metadatareader source material or supplied together with him and formed as a by-product at the stage of dehydrocyclization. the ri necessary, you can generate additional amounts of carbon monoxide, feeding the carbon dioxide contained, for example, in natural gas, the catalyst conversion, resulting in a reverse reaction of the conversion of water gas is obtained carbon monoxide:

and the following reaction:

CH4+H2O↔CO+3H2

In yet another embodiment of the invention stage reduction of hydrogen content comprises the reaction of at least part of the hydrogen present in the exhaust from the stage of dehydrocyclization stream with carbon monoxide to obtain alcohols With1-C3in particular methanol. Obtaining methanol and other oxygenates from synthesis gas is well known and is described, for example, in patents US 6114279, US 6054497, US 5767039, US 5045520, US 5254520, US 5610202, US 4666945, US 4455394. US 4565803, US 5385949, the contents of which are incorporated into this description by reference. Used synthesis gas typically has a molar ratio of hydrogen (H2) to oxides of carbon (CO+CO2) in the range of about 0.5:1 to about 20:1, in particular in the range from about 2:1 to about 10:1, and carbon dioxide is not necessarily present in an amount of not more than 50 wt.% on the total weight of the synthesis gas.

The catalyst used in the methanol synthesis process usually includes the oxide of at least one element selected from the group comprising copper, silver,zinc, boron, magnesium, aluminum, vanadium, chromium, manganese, gallium, palladium, osmium and zirconium. In an expedient embodiment of the invention the catalyst is a catalyst based on copper, for example in the form of copper oxide, optionally in the presence of oxide of at least one element selected from silver, zinc, boron, magnesium, aluminum, vanadium, chromium, manganese, gallium, palladium, osmium and zirconium. In an expedient embodiment of the invention, the catalyst contains copper oxide and an oxide of at least one element selected from zinc, magnesium, aluminum, chromium and zirconium. In one embodiment of the invention, the methanol synthesis catalyst selected from the group consisting of copper oxides, zinc oxides and oxides of aluminum. In a more preferred embodiment of the invention, the catalyst contains oxides of copper and zinc.

The process of methanol synthesis can be performed in a wide range of temperatures and pressures. Acceptable temperature in the range from about 150°to about 450°C., in particular from about 175°C. to about 350°C., for example from about 200°to about 300°C. Acceptable pressure in the range from about 1500 kPa to about 12500 kPa, in particular from about 2000 kPa to about 10000 kPa, in particular from about 2500 kPa to about 7500 kPa. Flow rate of feed the gas varies depending on the type of the process pursued, but usually the volumetric rate of gas flow by passing the gas through a layer of catalyst is in the range from about 50 h-1approximately 50000 h-1in particular from about 250 h-1approximately 25000 h-1in particular from about 500 h-1approximately 10000 h-1. This exothermic reaction can be carried out either in a fixed or fluid layers, including multiple layers of catalyst, with removal of heat between the layers. In addition, to obtain the maximum possible kinetic velocities, the process in the first (first) flow layer (layers) can be performed at higher temperatures, and to obtain the maximum possible thermodynamic transformation process in the last (recent) thread layer (layers) can be performed at lower temperatures.

The resulting methanol and/or other oxygenates may go on sale as a standalone product, can be used for alkylation of aromatic compounds formed at the stage of dehydrocyclization to more valuable products, such as xylenes, or can be used as source material for more low molecular weight olefins, particularly ethylene and propylene. The conversion of methanol to olefins is a well-known process and is described, for example, in patent US 4499327, the content is of which is hereby incorporated into this description by reference.

Selective combustion of hydrogen

In yet another embodiment of the invention stage reduction of hydrogen content includes selective combustion of hydrogen, which is a process in which hydrogen is mixed stream interacts with oxygen to produce water or water vapor without substantial interaction in the flow of hydrocarbons with oxygen to form carbon monoxide, carbon dioxide and/or oxygenated hydrocarbons. Usually selective combustion of hydrogen is carried out in the presence of oxygen-containing particulate material, such as a mixed metal oxide, releasing part of the bound oxygen to hydrogen.

One suitable method of selective combustion of hydrogen is described in the patent US 5430210, the content of which is incorporated into this description by reference, and comprises contacting under reaction conditions a first stream comprising hydrocarbons and hydrogen and a second stream containing oxygen, with separate surfaces of a membrane that is impermeable to not containing oxygen gas and containing a metal oxide, a selective towards hydrogen burning, and the selection of product selective combustion of hydrogen. This metal oxide, as a rule, is a mixed oxide of bismuth, indium, antimony, thallium and/or zinc.

In the patent U 5527979, the content of which is incorporated into this description by reference, describes how clean catalytic oxidative dehydrogenization alkanes with getting alkenes. This method involves simultaneous equilibrium dehydrogenization alkanes to alkenes and selective burning the resulting hydrogen for carrying out equilibrium reactions dehydrogenization with further formation of alkenes. Thus, in particular, the source alkanoyl material digitalout over the catalyst equilibrium dehydrogenization in the first reactor, and then the exhaust from the first reactor flow together with oxygen is directed to the second reactor, containing a catalyst of a metal oxide, which is used for catalysis of the selective combustion of hydrogen. The catalyst equilibrium dehydrogenization may include platinum, as a catalyst for the selective combustion of the metal oxide may include bismuth, antimony, indium, zinc, thallium, lead and tellurium or a mixture.

In patent application US 2004/0152586, which was published on August 5, 2004, the contents of which are incorporated into this description by reference, describes a method of reducing the hydrogen content in the exhaust from the cracking installation thread. In this method using a catalytic system comprising (1) at least one solid acidic cracking component and (2) at least the Dean component selective combustion of hydrogen on the metal base, consisting essentially of (a) a combination of metals selected from the group comprising: I) at least one metal from group 3, and at least one metal from groups 4-15 of the Periodic table of elements; (II) at least one metal from groups 5-15 of the Periodic table of elements and at least one metal from at least one of the groups 1, 2 and 4 of the Periodic table of elements; (III) at least one metal from groups 1 and 2, at least one metal from group 3, and at least one metal from groups 4-15 of the Periodic table of elements; and (IV) two or more metals from groups 4-15 of the Periodic table of elements; and (b) oxygen and/or sulfur, and oxygen and/or sulfur are chemically bound and both within and between the metals.

The reaction of the selective combustion of hydrogen in accordance with the present invention is usually carried out at a temperature in the range from about 300°C. to about 850°C., and the pressure being in the range of from about 1 ATM to about 20 ATM (100 to 2000 kPa).

Selection/processing of aromatic products

In addition to hydrogen, other products stage dehydrocyclization are benzene and naphthalene. These products usually can be distinguished from the flow of exhaust from the stage of dehydrocyclization, by solvent extraction and subsequent separation into fractions, and then immediately send in sales is as chemical products for mass production. In another embodiment of the invention a portion of the benzene and/or naphthalene or all of the benzene and/or naphthalene can be alkilirovanii with obtaining, for example, toluene, xylenes and alkylnaphthalenes, and/or can be subjected to hydrogenation to obtain, for example, cyclohexane, cyclohexene, dihydronaphthalene (benzylchloride), tetrahydronaphthalene (tetralin), hexahydronaphthalen (dicyclohexano), octahydronaphthalene and/or decahydronaphthalene (decalin). Suitable methods of alkylation and hydrogenation are described below and in more detail in the applications filed by the authors of the present invention, namely in international applications PCT/US 2005/043523 (publication WO/2006/068800), filed December 2, 2005, and PCT/US 2005/044038, filed December 2, 2005

Alkylation of aromatic compounds

Alkylation of aromatic compounds such as benzene and naphthalene, is well known in the art and typically involves the reaction of an olefin, alcohol or alkylhalogenide with aromatic substances in gaseous or liquid phase in the presence of an acid catalyst. Suitable acid catalysts include zeolites with medium-sized pores (i.e. permeability index (Constraint Index)defined in the patent US 4016218 and comprising from 2 to 12), including materials, with frames types MFI (for example, ZSM-5 and silicalite), MEL (EmOC is emer, ZSM-11), MTW (for example, ZSM-12), TON (for example, ZSM-22), MTT (e.g., ZSM-23), MFS (for example, ZSM-57) and FER (for example, ZSM-35, and ZSM-48, and zeolites with larger pores (i.e. permeability index lower than 2), such as materials with frames types WEAH (e.g., zeolite beta), FAU (for example, ZSM-3, ZSM-20, zeolite X, Y, ultrastabilized Y and dealuminated Y), MOR (e.g., mordenite), MAZ (for example, ZSM-4), MEI (for example, ZSM-18) and MWW (e.g., MCM-22, PSH-3, SSZ-25, ERB-1, ITQ-1, ITQ-2, MCM-36, MCM-49 and MCM-56).

In one embodiment, this method of benzene are separated from the stream of exhaust from the stage of dehydrocyclization, and then alkylate the olefin, such as ethylene, obtained as a by-product at the stage of reduction of hydrogen content using tonirovania/mahanirvana. Typical conditions for vapor-phase alkylation of benzene with ethylene include a temperature of approximately 650°F to about 900°F (343 to 482°C)gauge pressure of from about atmospheric to about 3000 psig (100 to 20800 kPa), volumetric gas flow rate (OS) ethylene comprising from about 0.5 h-1about to 2.0 h-1and the molar ratio of benzene to ethylene of 1:1 to 30:1. Liquid phase alkylation of benzene with ethylene can be carried out at a temperature from 300°F to 650°F (150 to 340°C), With gauge Yes the relocation to about 3000 psig (20800 kPa), when OS ethylene comprising from about 0.1 h-1approximately 20 h-1and when the molar ratio of benzene to ethylene of 1:1 to 30:1.

Atilirovanie benzene should be carried out in at least partially liquid phase conditions using a catalyst comprising at least one of zeolite beta, zeolite Y, MCM-22, PSH-3, SSZ-25, ERB-1, ITQ-1, ITQ-2, ITQ-13, ZSM-5 MCM-36, MCM-49 and MCM-56.

Atilirovanie benzene can be performed at the venue of the process of dehydrocyclization/reduction of hydrogen content, or benzene for conversion to ethylbenzene can be transported to another region. Then the resulting ethylbenzene can be sent in a sale, be used as a precursor, for example upon receipt of styrene, or isomerizate by methods well known in the art, with the production of mixed xylenes.

In another embodiment, this method of alkylating agent is a methanol or dimethyl ether (DME) and is used for the alkylation of benzene and/or naphthalene, emitted from the exhaust from the stage of dehydrocyclization flow, obtaining toluene, xylenes, methylnaphthalene and/or dimethylnaphthalene. If methanol or dimethyl ether is used for the alkylation of benzene, it is advisable to carry out in the presence of a catalyst, linking the zeolite, such as ZSM-5, β zeolite, ITQ-13, MCM-22, MCM-49, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-48, modified by treatment with water vapor so that he had a diffusion parameter for 2,2-Dimethylbutane comprising from about 0.1 to 15-1when the measurement at a temperature of 120°C and a pressure of 2.2-Dimethylbutane 60 Torr (8 kPa). This method is selective in obtaining para-xylene and described, for example, in patent US 6504272, the content of which is incorporated into this description by reference. If methanol is used for the alkylation of naphthalene, it is advisable to carry out in the presence of a catalyst comprising ZSM-5, MCM-22, PSH-3, SSZ-25, ERB-1, ITQ-1, ITQ-2, ITQ-13, MCM-36, MCM-49 or MCM-56. This method can be used for selective receipt of 2,6-dimethylnaphthalene, and it is described, for example, in patents US 4795847 and US 5001295, the contents of which are incorporated into this description by reference.

If proposed in the invention method as alkylating agent is methanol or DME, this agent can be introduced into the process as a separate source material or it may, at least partially, be formed in situ by adding containing carbon dioxide gaseous source material, such as a natural gas stream, a portion of the stream of exhaust from the stage of dehydrocyclization, or this whole thread. In particular, the exhaust with the adiya's dehydrocyclization stream before any allocation of aromatic components can be sent back into the reactor conversion and to carry out the reaction containing carbon dioxide starting material in the context of rising the content of carbon monoxide in the exhaust stream, i.e. such reactions as the above reactions 5 and 8.

In addition, the reactor inverse conversion can be sent methane and CO2and/or water vapor to obtain synthesis gas, which can then be mixed with a part of the exhaust from the stage of dehydrocyclization flow to regulate the relations H2/CO/CO2depending on the needs for phase alkylation.

Typically, the reactor inverse conversion contains a catalyst comprising a transition metal on a carrier, such as Fe, Ni, Cr, Zn aluminum oxide, silicon dioxide or titanium dioxide, and working conditions, including a temperature of approximately 500°C. to about 1200°C., in particular from about 600°to about 1000°C., for example from about 700°to about 950°C. and a pressure of about 1 kPa to about 10000 kPa, in particular from about 2000 kPa to about 10000 kPa, for example about 3,000 kPa to about 5000 kPa. Space velocity of the gas may vary depending on the type of process method, but usually the volumetric rate of gas flow through the catalytic layer is in the range from about 50 h-1approximately 50000 h-1in particular from about 250 h-1approximately 25000 h-1more preferably about about the 500 h -1approximately 10000 h-1.

Then the exhaust from the reactor inverse conversion of the flow can be directed to the alkylation reactor operating under conditions ensuring the reactions as the following:

Suitable for such a reactor alkylation conditions include a temperature of approximately 100 to approximately 700°C, a pressure of from about 1 to about 300 atmospheres (100 to 30000 kPa) and the volumetric feed rate of the raw material component for aromatic hydrocarbon approximately from 0.01 h-1approximately 100 h-1. Suitable catalyst includes a molecular sieve having a permeability index from 1 to 12, such as ZSM-5, usually together with one of the metals or oxides of metals such as copper, chromium and/or zinc oxide.

If the alkylation catalyst comprises a molecular sieve, last expedient to modify to change its diffusion characteristics so that the predominant isomer of xylene, obtained by the reaction of 11, was a pair of xiap. Acceptable means of diffusion modifications include treatment with water vapor and precipitation ex situ or in situ compounds of silicon, coke, metal oxides such as MgO, and/or R on the surface or the mouths of the pores of the molecular sieve. Also preferred is an introduction to molecular sieve active metal in such a way as to ensure saturation of the more highly reactive substances, such as olefins, which may be formed as by-products and which otherwise would cause deactivation of the catalyst.

Then the exhaust from the reactor alkylation stream can be directed in section separation, in which the aromatic products were first separated from the hydrogen and other low molecular weight materials, suitable extraction solvent. Further aromatic products can be divided into a benzene fraction, a toluene fraction, fraction C8and a heavy fraction comprising naphthalene and alkylated naphthalenes. Then the aromatic fraction C8directed the process of crystallization or adsorption to separate valuable p-xylene component, and the remaining mixed xylenes can either go on sale as a finished product, or can be in the path of isomerization for an additional amount of p-xylene. Toluene fraction can either be removed as a ready-to-sell products, or to return to the alkylation reactor or be directed to the installation of the disproportionation of toluene, in particular in the installation of selective disproportionate the Oia toluene, for an additional amount of p-xylene.

Hydrogenation of aromatic compounds

In addition to stage alkylation or instead of at least part of the aromatic components in the exhaust from the stage of dehydrocyclization flow can be gidrirovanii with the production of useful products, such as cyclohexane, cyclohexene, dihydronaphthalene (benzylchloride), tetrahydronaphthalene (tetralin), hexahydronaphthalen (DICYCLOHEXYL), octahydronaphthalene and/or decahydronaphthalene (decalin). These products can be used as fuels and chemical intermediates, and in the case of tetralin and decalin - and as a solvent for extraction of aromatic components from the stream of exhaust from the stage of dehydrocyclization.

The hydrogenation is advisable, but not necessary, to carry out after separation of the aromatic components from the stream of exhaust from the stage of dehydrocyclization, and it is advisable to use a portion of the hydrogen formed in the reaction of dehydrocyclization. Appropriate ways hydrogenation of aromatic compounds is well known in the art, and typically use a catalyst comprising Ni, Pd, Pt, Ni/Mo or sulfatirovanne Ni/Mo supported on a carrier of alumina or silicon dioxide. Suitable for the process of hydrogenation working the conditions include temperature, approximately from 300°F to about 1000°F (150 to 540°C), in particular from about 500°F to about 700°F (260 to 370°C), gauge pressure of from about 50 to about 2000 psig (445 to 13890 kPa), in particular from about 100 to about 500 psig (790 to 3550 kPa), and the volumetric feed rate of approximately 0.5 h-1approximately 50 h-1in particular from about 2 h-1approximately 10 h-1.

To obtain materials suitable for polymerization or other subsequent chemical transformations may also be desirable partial hydrogenation, allowing you to leave the product one or more olefinic carbon-carbon bonds. Suitable methods partial hydrogenation are well known in the art, and typically use a catalyst comprising a noble metal, preferably ruthenium supported on metal oxides such as La2O3-ZnO. Can also be used in homogeneous catalytic systems with noble metals. Examples of how partial hydrogenation described in patents US 4678861, US 4734536, US 5457251, US 5656761, US 5969202 and US 5973218, the content of which is fully incorporated into this description by reference.

An alternative method of hydrogenation involves hydrocracking is afterinvoke component at low pressure to get alkyl benzenes, carried out over such a catalyst, as sulfatirovanne Ni/W or sulpicianus Ni deposited on amorphous aluminosilicate or zeolite, such as zeolite X, zeolite Y or zeolite beta. Suitable for hydrocracking low-pressure operating conditions include a temperature of approximately 300°F to about 1000°F (150 to 540°C), in particular from about 500°F to about 700°F (260 to 370°C), gauge pressure of from about 50 to about 2000 psig (445 to 13890 kPa), in particular from about 100 to about 500 psig (790 to 3550 kPa), and the volumetric feed rate of raw materials, component from about 0.5 h-1approximately 50 h-1in particular from about 2 h-1approximately 10 h-1.

Further, the invention is discussed in more specifically with reference to the accompanying drawings and the following non-limiting examples.

Figure 1 presents a simplified scheme of the reactor dehydrocyclization with heater catalyst and one regenerator catalyst for converting methane to aromatic hydrocarbons in the first embodiment of the invention. In this embodiment of the invention the reactor dehydrocyclization contains vertically positioned reactor 11 with a deposited layer, in which through the inlet 12 located near the top of the reactor 11, flows in load, the addition and the regenerated granular catalyst, and through the outlet 13 located near the base of the reactor 11, enters the flow of the chilled zakoksovanie catalyst. Methane source material (raw material) 15 is introduced into the reactor 11 near its base. Typically, the heated catalyst into the reactor 11 at a temperature of about 850°C, and cooled zakochany catalyst leaves the reactor at a temperature of about 650°C.

Chilled zakochany the catalyst moves under the action of gravity (gravity) from exit 13 to the base of the vertical pipe of the heater 16, where the catalyst is fond of mixture of air and methane fuel held in a vertical pipe through the manifold 17. The catalyst is transported upward through a vertical pipe to the heater 16 air-methane mixture during its passage through the vertical pipe to the heater 16 is heated by the combustion of methane. Mixture flowing in a vertical tube heater 16 through the manifold 17, contains enough fuel to heat the catalyst to the desired reaction temperature, but not enough oxygen. Therefore, in the vertical pipe heater 16 through multiple inlet ports 18 distributed along the length of the vertical pipe (for simplicity in figure 1 indicated only one inlet opening 18, but in reality their numbers which can be much larger), introduce additional air, which means that heating of the catalyst occurs gradually, as the movement of the catalyst upward through a vertical pipe to the heater 16. Usually upon reaching the top of the vertical pipe of the heater 16, the temperature of the catalyst is approximately 850°C.

Coming out of the top of the vertical pipe of the heater 16, the heated catalyst passes into the separator 19, where the solid granular catalyst is separated from the gaseous products of combustion and is then directed into the hopper 20, and then into the upper end of a vertical strut 21 of the catalyst. Gaseous combustion products discharged from the separator through the outlet 22 and then, before they go on heat recovery, they are directed to the cyclones 22 (not shown) for removal of catalyst dust. When using air as the medium of combustion in the heater 16 facing gaseous combustion products typically contain 40-70 wt.% N2, <1 wt.% O2, 1-30 wt.% H2, 2-20 wt.% WITH, 1-20 wt.% CO2and 5-25 wt.% H2O.

The speed of movement of catalyst through the system and the internal diameter of the riser 21 is such that the heated catalyst accumulates in the riser 21, forming a densely Packed layer, moving under the action of gravity through the riser 21 and entering the regenerator 23, connec the config to the lower end of the strut 21. Depending on the height of the riser 21 and the pressure in the regenerator can be maintained at a level necessary for efficient regeneration of the catalyst is hydrogen. For example, calculations have shown that to achieve the absolute pressure in the regenerator is about 700 to 1000 kPa using a catalyst Mo/ZSM-5 circulating so that the absolute pressure in the separator 19 is approximately equal to 140 kPa, the height of the riser should be approximately 100 to 300 feet (30 to 91 m).

The catalyst from the riser 21 enters the regenerator 23 through the inlet 24 located near the top of the regenerator 23 and moves down through the regenerator 23 against the flow of hydrogen 25 injected into the regenerator 23 near its base. The regenerated catalyst leaves the regenerator 23 through the outlet 26, also located near the base of the regenerator 23, and then through another vertical pipe 28 is moving in the separating vessel 31, before re-enroll in the reactor 11 through the inlet 12. This vertical pipe 28 and the separating vessel 31 are designed for separation of hydrogen trapped by the catalyst, and if necessary, to facilitate the movement of catalyst through a vertical pipe 28, to the base of the vertical pipe 28 can build up methane 28.

Figure 2 illustrates the implementation of the method of converting methane to aromatic hydrocarbons in the Orom embodiment of the invention, in which the reactor dehydrocyclization equipped with a heater catalyst, and the first and second catalyst regenerators, connected in series and operating at different pressures. In this embodiment of the invention the reactor dehydrocyclization also contains a vertically positioned reactor 111 with a deposited layer, through which the inputs 112, 113, located near the top of the reactor 111 is a flow of heated and regenerated granular catalyst, and which through the outlet 114, located near the base of the reactor 111, enters the flow of the chilled zakoksovanie catalyst. Methane source material 115 is introduced into the reactor 111 near its base. Typically, the heated catalyst into the reactor 111 at a temperature of about 850°C, and cooled zakochany catalyst leaves the reactor at a temperature of about 650°C.

Chilled zakochany the catalyst moves under the influence of gravity from the outlet 114 to the base of the vertical pipe of the heater 116, where the catalyst is fond of mixture of air and methane fuel held in a vertical pipe through a manifold 117. The catalyst is transported upward through a vertical pipe to the heater 116 air-methane mixture during its passage through the vertical pipe podogrevatelya heated by the combustion of methane. Mixture flowing in a vertical tube heater 116 through a manifold 117, contains enough fuel to heat the catalyst to the desired reaction temperature, but not enough oxygen. Therefore, in the vertical pipe heater 116 through multiple inlet ports 118, distributed along the length of the vertical pipe (for simplicity in figure 2 indicated only one inlet 118, but in reality their number may be much higher), introduce additional air, which means that heating of the catalyst occurs gradually, as the movement of the catalyst upward through a vertical pipe to the heater 116. Usually upon reaching the top of the vertical pipe of the heater 116, the temperature of the catalyst is approximately 850°C.

Coming out of the top of the vertical pipe of the heater 116, heated catalyst passes into a cyclone separator 119, where the solid granular catalyst is separated from the gaseous products of combustion and then fed into the upper portion of the first regenerator 121, operating at relatively low pressure. The first regenerator 121 is located vertically above the second regenerator 122 operating at relatively high pressure, and is connected with the second regenerator 122 through a vertical riser 123 for the catalyst is. In the first regenerator 121 near its base is blown in a stream of hydrogen 125, which moves up through the first regenerator 121. The height of the riser 123 determines the pressure difference between the first and second regenerators 121, 122.

In the second regenerator 122 near its base is blown in a stream of hydrogen 124 which moves upwards through the second regenerator 122, strut 123 for catalyst and the first regenerator 121 towards the stream of catalyst moving under the action of gravity from the first regenerator 121 in the second regenerator 122. Hydrogen regenerates zakochany catalyst as it moves through the regenerator, and a portion of the regenerated catalyst is returned to the reactor 111 through the inlet 112 of the first regenerator 121, and the remainder of the regenerated catalyst is returned to the reactor 111 through the inlet 113 of the second regenerator 122.

Below follows a more detailed description of the invention with reference to the following non-limiting examples.

Example 1

Example 1 demonstrates the use of jointly supplied raw materials (such as H2, CO2and H2O) to achieve the reduced speed coking of the catalyst Mo/ZSM-5 during dehydrocyclization methane with obtaining mainly benzene. Example 2 demonstrates that by reducing the number of coke, otlozhiv is gosia on the catalyst during contact with a hydrocarbon, it is possible to maintain high efficiency after several cycles of contact with hydrocarbons and regenerating and that a lower rate of coking per cycle is converted to lower the resulting rate of coke deposition for several cycles.

The catalysts of Mo/ZSM-5 was prepared by two methods: (1) impregnation on the capacity of the carrier NH4+-ZSM-5 (having a ratio Si/Al225) the required amount of the solution heptamolybdate ammonium, followed by drying at 120°C for 2 hours and the final calcination at 500°C for 6 hours in a stream of air, and (2) by grinding in a ball mill, molybdenum oxide with the carrier NH4+-ZSM-5 (having a ratio Si/Al225) for 2 hours followed by calcination at 500°C for 5 h in air flow. The molybdenum content (in terms of metal in wt.%) varied by changing the concentration of heptamolybdate ammonium in the impregnating solution or the amount of molybdenum oxide, added to the milled mixture.

Catalytic testing of the prepared catalysts of Mo/ZSM-5 was carried out using a vibrating microbalance with tapering element (Engl. Tapered-Element Oscillating Microbalance). The TEOM are presented)that allow precise and fast response time to determine changes in the weight of the catalyst during the reaction. The catalyst (after Prokaeva the Oia) alloy preformed, to grind and sieved to select particles with a particle size of 20-40 mesh. Approximately 0.10 g sifted particles of catalyst was loaded into the sample holder microbalance TEOM are presented (total sample: 0,20 cm3and using the pads of the quartz fiber is condensed with obtaining a fixed catalyst layer. Functional properties of the catalyst when dehydrocyclization of methane to benzene was determined at 800°C and at an absolute pressure of 20 psi (138 kPa) in experiments a-B and at an absolute pressure of 14.7 psi (101 kPa) in experiments In a-D listed in table 1, using a source material containing these jointly submitted materials (CO2N2O, H2), Ar (Ar is used as an internal standard), the rest - CH4at a given flow rate (methane). Otheriwse from the reaction stream was analyzed using a mass spectrometer to determine the concentrations of methane, benzene, naphthalene, hydrogen and argon. The rate of coke deposition on the catalyst (i.e. the formation of hard carbon deposits, not uletuchivalsja from the surface of the catalyst) was determined directly on the weight changes observed using a microbalance. The values of the functional properties of the catalyst (for example, performance on benzene, the conversion of methane, electively on benzene) are cumulative or average values over a period of time, beginning with the submission of methane and ending at the time instant when the output of benzene was reduced to <2%. The results are shown in table 1.

Table 1
ExperimentMo, wt.%Jointly submitted materialsThe average rate of coking, wt.%/hPerformance benzene, grams of benzene per gram of catalystThe conversion of methane grams of CH4per gram of catalystThe selectivity for benzene, wt.%
And4,6No5,20,500,9354
B4,62%CO210% of H22,21,31of 2.2159
In2,7No8,60,230,61 38
G2,73% H2O6,00,340,8540
D2,720% H2the 3.80,260,4953

As shown in table 1, when the content of Mo is 4.6% for experiments a and B, the average rate of coke deposition decreased from 5.2 wt.% per hour in the absence of materials supplied together with methane (i.e. when the composition of the feed stream: 95% SN4and 5% Ar) to 2.2 wt.% in the hour when filing jointly with methane 2% CO2and 10% N2(i.e. when the composition of the feed flow: 2% CO210% of H2, with 83.6% of CH4and 4.4% Ar). Performance benzene increased from 0.5 g of benzene obtained per gram of catalyst in the absence of a jointly submitted materials to 1.3 grams of benzene per gram of catalyst. This shows that the presence of CO2and H2as supplied together with methane materials can significantly improve the performance of the catalysts of Mo/ZSM-5 benzene by reducing speeds coke formation. Experiments C, D and E show the effect of filing a joint N2Oh and the 2the catalyst containing 2.7% of Mo on ZSM-5. The average rate of coking, as has been decreased from 8.6 wt.%/h in the absence of a jointly submitted materials to 6.0 wt.%/h when combined with the methane feed 3% water vapor and to 3.8 wt.%/h when combined with the methane supply 20% N2. In both cases, it was found that by adding together the feed material productivity and selectivity to benzene was increased.

Example 2

The catalysts of Mo/ZSM-5 prepared by the methods described above in example 1. The catalysts synthesized in the state was subjected to cyclic aging, which consisted of (1) the period of contact with the hydrocarbon, when the catalyst was affected by methane CH4supplied at 800°C and at an absolute pressure of 14.7 psi (101 kPa) together with the specified materials for 5 min at the specified flow rate in CH4and (2) follow-up period, regenerating, when the catalyst was heated up to 850°C and at an absolute pressure of 14.7 psi (101 kPa) in gaseous H2at a rate of 10°C/min and kept at 850°C for a specified time at a volumetric rate of gas supply 9000 cm [normal]/g catalyst/hour, After the stage of regenerating the catalyst before re submission of methane starting material was cooled to 800°C. Perret is cyclic aging the catalyst synthesized in state pre-carbonized by heating the catalyst in a gaseous mixture of 15% CH 4-H2from 150°C to 800°C at 5°C/min and kept at 800°C for 1 hour. After the catalysts have missed over a specified number of cycles contact with hydrocarbons and regenerating, the spent catalysts were removed and tested for functional properties on the microbalance TEOM are presented.

Approximately 0.10 g of spent catalyst was loaded into the sample holder microbalance TEOM are presented and the use of the linings of the quartz fiber is condensed with obtaining a fixed catalyst layer. Functional properties of the catalyst when dehydrocyclization of methane to benzene was determined at 800°C and at an absolute pressure of 14.7 psi (101 kPa) for experiments a-D and at an absolute pressure of 20 psi (138 kPa) for experiments D-W in table 2, using a mixture of starting material of 95% SN4and 5% Ar (Ar is used as an internal standard) and the volumetric feed rate of the raw material corresponding to the supply of 4 g of CH4per gram of catalyst per hour. Otheriwse from the reaction stream was analyzed using a mass spectrometer to determine the concentrations of methane, benzene, naphthalene, hydrogen and argon. The definition of coke on the catalyst was performed using termoregulirovanija oxidation of spent catalyst in a thermogravimetric analyzer (TGA).

the table 2 compares accumulated on the catalyst coke after a few cycles of contact of the catalyst with hydrocarbons and regenerating the catalyst and the residual capacity of spent catalysts by the benzene. Experiments a-C demonstrate the functional properties of the catalyst of 2.7% Mo/ZSM-5 after 0, 20 and 40 cycles. Experiment D shows the functional properties of the catalyst of 2.7% Mo/ZSM-5 after 40 cycles. Experiments D-W demonstrate the functional properties of the catalyst 4.6% of Mo/ZSM-5 after 0, 50, 389 cycles. When comparing experiments C and D shows that the addition of 20% H2as the material to be applied together with methane, reduced resulting deposition of coke on the catalyst to about 55%, 2.0 wt.% (in the case together without feed material) to 0.9 wt.% (in collaboration with the methane supply 20% N2). It was found that when combined with the methane supply 20% N2apart from reducing coke formation performance benzene after 40 cycles significantly, by about 25%, increased (0.25 g benzene/g of catalyst), when compared with the case of absence in conjunction feed material (0.20 g benzene/g of catalyst). Experiments D-W show that the resulting accumulation of coke on the catalyst can be minimized, if as a co-feed material to add 2% CO2. It was found that after 50 cycles, the total deposition of coke on the catalyst was only 0.15 wt.%, significantly less (approximately 93%) coke deposition on the catalyst was observed after 40 cyclo is in the absence of jointly supplied material (i.e. 2.0 wt.%), even though the volumetric feed rate of the source material during contact with a hydrocarbon increased from 1 to 4. After 50 cycles was observed no performance degradation of benzene. After 389 cycles measured the deposition of coke on the catalyst amounted to about 0.5 wt.%, indicating that the total accumulation of coke on the catalyst during long-term cyclic aging can be minimized.

Example 3

Examples 3 and 4 are designed to show the effect of temperature regenerating the output of the benzene by dehydrocyclization methane for several cycles.

The catalyst, containing about 4 wt.% Mo on ZSM-5, used for flavoring source material containing 86,65 mol.% CH4, 1.8 percent2H6at 0.9% CO2, 0,45% N2and 10 mol.% Ar. To calculate the output values for benzene waste stream was analyzed by gas chromatography. The catalyst is alternately used in cycles of reaction and regenerating, and the ratio of reaction time to the time of regeneration was 1:1. Conditions and reactions, and regeneration included a temperature of 800°C and a gauge pressure of about 7 psi (obstawianie 149 kPa). The output values of benzene after approximately 15 minutes from the beginning of the reaction and in a few moments time is after the start of the experiment are listed below:

Time, hThe yield of benzene, % of the original material (carbon)
4,75of 10.05
10,757,78
14,756,39

Example 4

The process described in example 3 was repeated, but with each regeneration cycle at a temperature of 850°C and a gauge pressure of about 7 psi (abs. pressure 149 kPa). Reaction conditions and the ratio of reaction time to time regeneration remained the same as in example 3. The output values of benzene after approximately 15 minutes from the beginning of the reaction and in a few moments time after the start of the experiment are listed below:

Time, hThe yield of benzene, % of the original material (carbon)
to 4.6211,82
10,30to 9.91
14,957,98

This shows that the output values of benzene in example 4 (where the temperature of the regenerating amounted to 850°C) is higher than in example 3 (where the temperature is ur regenerating amounted to 800°C).

Example 5

This example is intended to show that regeneration at higher pressures is advantageous from the viewpoint of improving the selectivity of the catalyst, i.e. improving the selectivity for benzene and decrease the selectivity in respect of the formation of coke.

The catalyst, containing about 4 wt.% Mo on ZSM-5, used for flavoring source material containing 86,65 mol.% CH4, 1.8 percent2H6at 0.9% CO2, 0,45% N2and 10 mol.% Ar. To calculate the output values for benzene waste stream was analyzed by gas chromatography. The catalyst is alternately used in cycles of reaction and regenerating, and metadatabase source material was applied to the catalyst during each 20-minute cycle reactions, and hydrogen was fed to the catalyst during each 40-minute regeneration cycle. Within the first 90% of each cycle of the reaction in metastasi source material as a co-feed material was added 20% H2but in the final 10% of each cycle of the reaction is the addition was stopped. Gauge pressure of the reaction was approximately 7 psi (abs. pressure 149 kPa), and the reaction temperature was increased during each cycle of the reaction from initial values of about 700°C to a final value of about 800°Stredney temperature during each cycle Regener the tion was equal to approximately 850°C, the maximum temperature was 860°C. At the beginning of the test experiment gauge pressure during each regeneration cycle was 34 psi (abs. pressure 335 kPa), and stable values of selectivity and yield of benzene are listed below. After a specified time gauge pressure during each regeneration cycle was reduced to approximately 7 psi (abs. pressure 149 kPa) and after about 11 hours after the change of pressure of the regeneration of the newly measured values of the selectivity and yield of benzene, which are presented below. This shows that at higher pressure regeneration selectivity for benzene and yield of benzene were higher and the selectivity for Cox - below.

Gauge pressure, psi:
347
The yield of benzene, % of original carbon10,78,9
The selectivity of the converted carbon in % from the original:
Benzene:65,0 57,1
Coke:13,119,5

Example 6

This example is intended to show that when you use containing coke catalyst Mo/HZSM-5 dehydrocyclization methane at high pressure H2you can delete more coke than at low pressure H2.

Catalyst 5 wt.% Mo/ZSM-5 prepared by grinding in a ball mill, molybdenum oxide with the carrier NH4+-ZSM-5 (having a ratio Si/Al225) for 2 hours followed by calcination at 500°C for 5 hours in air flow. The catalyst was pre-carbonized in a gaseous mixture containing 15% of CH4and 85% H2at 800°C and a flow rate of 1.2 for 1 hour in a quartz reactor of ideal displacement. The conversion of methane was performed at a temperature of 800°C, flow rate of 1.2 and an absolute pressure of 20 psi (138 kPa) to dehydrocyclization of methane to benzene using a source material containing 95 vol.% CH4and 5% vol. Ar (argon was used as internal standard). The reaction of the conversion of methane was stopped after 1 hour that the catalyst is in contact with the hydrocarbon. Then, after cooling to the ambient temperature of the reactor took a sample of the catalyst and this sample was pushing the congestion measured quantity of coke, which was about 8.9 wt.% [(weight of coke)/(total mass of coke and Mo/ZSM-5)].

Approximately 100 mg containing coke sample of catalyst was placed in the nozzle tubular reactor with a diameter of ¼ " (6.35 mm). The reactor temperature was raised from ambient temperature to about 1000°C, increasing with a speed of 5°C./min in a stream of H2submitted with a flow rate of 50 ml/min. of methane and hydrogen in the exhaust stream of the reactor was analyzed by the method of GC-MS in real time. The amount of coke gas with hydrogen, was calculated based on the concentration of methane in the exhaust flow and the flow of hydrogen; other carbon-containing compounds except methane was not observed.

Figure 3 shows the change in the amount of methane (mmol)emitted from the reactor in a linear increase in the temperature of the reactor, starting from the ambient temperature, at two different partial pressures of hydrogen. From the graph it is seen that when the absolute partial hydrogen pressure of 20 psi (138 kPa), the maximum concentration of methane is achieved at a temperature of about 850°C. At higher absolute partial pressure of hydrogen (75 psi [517 kPa]), the temperature of the maximum concentration decreased to about 750°C. the Magnitude of the peak concentration of the emitted methane at an absolute pressure of 75 psig much what about the above, than 20 psi, which means a much higher intensity of removal of coke at high pressures. More importantly, a significant decrease in the maximum temperature of methane at high pressure versus low pressure means that at a given temperature regeneration with hydrogen not all Cox remote at high pressure, can be removed at low pressure.

Table 3 summarizes the values of the quantity of coke (mass percentage) on the sample of catalyst depending on the temperature of the reactor during termoregulirovanija processing hydrogen (H2). The figures show that at any given temperature the amount of coke remote at higher pressure, much higher than the number of coke, remote at low pressures. For example, after TO-H2at an absolute pressure of 20 psi remained approximately 2.6 wt.% coke, while at an absolute pressure of 75 psi remained only 0.3 wt.% Cox.

td align="center"> 8,9
Table 3
Temperature, °CCoke, wt.%
20 psi75 psi
2978,9
3508,98,9
3918,98,8
4338,98,8
4758,88,8
5178,78,7
5598,68,5
6018,58,2
6438,37,8
6848,07,1
7267,66,0
7686.94.8
8106,03.7
8525,12,8
8944,2 2,1
9363.51,6
9773,01,1
10192,70,5
10302,60,3

Example 7

The catalyst of 2.7 wt.% Mo/ZSM-5 prepared by impregnation on the capacity of the carrier NH4+-ZSM-5 (having a ratio Si/Al225) a solution of heptamolybdate ammonium, followed by drying at 120°C for 2 hours and the final calcination at 500°C for 6 hours in air flow. The catalyst was placed in the nozzle reactor in the form of a fixed layer. The reactor temperature was raised to 800°Spri this temperature 800°C in the reactor was filed methane, maintaining contact methane with a catalyst for 20 min at space velocity of 1.2 and the absolute partial pressure of methane, which comprised about 15 psig (103 kPa). After 20 minutes of contact of the catalyst with methane feed methane source material turned off and for the gasification of coke on the catalyst included the supply of hydrogen. Duration gasification hydrogen accounted for 40 min at 800°SETI two stages repeated 39 times, after CEG is at the end stage gasification hydrogen sample of catalyst removed from the reactor. Measured the amount of coke on the sample of the catalyst, which amounted to about 8.7 wt.% [(weight of coke)/(total mass of coke and Mo/ZSM-5)].

Approximately 100 mg containing coke sample of catalyst was placed in the nozzle tubular reactor with a diameter of ¼ " (6.35 mm). The reactor temperature was raised from ambient temperature to approximately 850°C, increasing with a speed of 5°C./min in a stream of helium, was served with a flow rate of 50 ml/min After the temperature of the reactor was stabilized at 850°C. through the reactor missed hydrogen submitted with a flow rate of 50 ml/min. of methane and hydrogen in the exhaust stream of the reactor was analyzed by the method of GC-MS in real time. The amount of coke gas with hydrogen, was calculated based on the concentration of methane in the exhaust flow and the flow of hydrogen.

Figure 4 shows the change in the content of the coke on the catalyst (mass fraction in%) depending on the time of regenerating the hydrogen at 850°C and different partial pressures of hydrogen. From the above graph shows that the initial (before 40 min) the rate of gasification of coke was much higher at higher pressures, which is consistent with the results shown in example 6. In industrial applications it is desirable that the time of regenerating the hydrogen is as small as possible. In the advantages the society significantly faster regeneration to remove coke at high pressures becomes extremely noticeable.

Table 4 shows the content of coke (wt.%) at different times of the regenerate. The data show that the amount of coke remaining on the catalyst after 36 min regenerating hydrogen at absolute partial pressure of hydrogen 21 psi (145 kPa) and a temperature of 850°C, 7.7 wt.%, approximately 31% higher than the amount of coke remaining on the catalyst (5.9 wt.%) after 36 min regenerating hydrogen at the same temperature and the absolute partial hydrogen pressure of 105 psi (724 kPa).

Table 4
Min. timeAbsolute pressure, psi
214575105
08,78,78,78,7
98,38,07,77,3
188,17,67,0
277,97,26,66,2
367,76,96,45,9
467,56,76,2the 5.7
557,46,56,05,6
647,26,45,9of 5.4
737,16,3the 5.75,3
827,06,15,65,2
916,96,05,65,1
1006,86,05,1
1096,75,9of 5.45,0
1186,65,85,3a 4.9

It is also important to note that at a constant temperature (850°C) after an initial more rapid decrease in the content of the coke gasification of coke is very slow. Curves content of coke began to level off after about 100 min of being in hydrogen. Pay attention to the significant differences between the amount of coke remaining on the catalyst at different pressures regeneration at the end of the process regenerating the hydrogen for 2 hours. For example, the amount of coke remaining after about 2 hours regenerating the hydrogen at absolute partial pressure of hydrogen 21 psi (145 kPa), 6.6 wt.%, which is 35% higher than the amount of coke remaining in the absolute partial hydrogen pressure of 105 psi (724 kPa). At an absolute pressure of 21 psi to achieve the level of coke, similar to those obtained at 105 psi, may require many hours of extra time (regeneration. In practice, these results about the mean, not all the coke is removed at high pressures, can be removed at low pressures in the same temperature conditions regenerating hydrogen.

Example 8

The catalysts of Mo/ZSM-5 prepared by impregnation on the capacity of the carrier NH4ZSM-5 (with the ratio of Si/Al2equal to 28) the required amount of the solution heptamolybdate ammonium, followed by drying at 120°C for 2 hours and the final calcination at 500°C for 6 hours in air flow. The nominal target molybdenum content (mass percentage of the metal in the total mass of the catalyst) was 2.7 wt.%; small fluctuations in the content of molybdenum in the conclusions are not affected. Each sample of the catalyst Mo/ZSM-5 (after calcination) alloy preformed, to grind and sieved to select particles with a particle size of 30-60 mesh. Catalytic testing of catalysts of Mo/ZSM-5 was carried out in a quartz reactor containing a nozzle for receiving the fixed layer using linings of the quartz fiber.

Functional properties of the catalyst when dehydrocyclization of methane to benzene was determined at different temperatures using as starting material a mixture of 95 wt.% CH4and 5 wt.% argon (argon was used as internal standard) at a space velocity of the feedstock (methane) 1,2 h-1. All experimental the data were obtained at an absolute pressure of 138 kPa (20 psig) and at the same pressure has also carried out all the simulations. Otheriwse from the reaction stream was analyzed using a mass spectrometer and a gas chromatograph to determine the concentration of methane, benzene, toluene, ethylene, naphthalene, hydrogen and argon. The rate of coke deposition on the catalyst (i.e. the formation of hard carbon deposits, not uletuchivalsja from the surface of the catalyst) was determined by carbon balance. Adding N2in the source material at two temperatures (750 and 800°C) and at a concentration of respectively 6 and 20 mol.%, received additional data.

For the purposes of the experiment in example 8 obtained experimental data were reduced to two variables: Sat.BTNand Sat down.Cox. The Size Of Villages.BTNrepresents the average selectivity on the basis of the molar concentration of carbon, is defined by the sum of the moles of carbon in products: benzene, toluene and naphthalene, divided by the moles of carbon contained in the methane reacted. Sat.Coxrepresents the average selectivity on the basis of the molar concentration of carbon determined by the amount of moles of carbon remaining in the reactor divided by the moles of carbon contained in the methane reacted. The Amount Of Villages.BTNand Sat down.Coxnot equal 100% due to the formation of other small products, predominantly ethylene. Since it is often difficult which returns the exact experimental thermodynamic data transformations, to find the values Pravr.Nlused distributed on a commercial basis software (PROII/6,0 Copyright 2003 Invensys Systems Inc.) for the simulation. Pravr.Nlis defined as the maximum thermodynamically attainable conversion of methane to benzene and hydrogen (i.e., the absence of model limitations, due to which there are no other products, such as coke, naphthalene, ethylene, and so on) at a given temperature and absolute pressure of 138 kPa (20 psi). The results of experiment and simulation are presented in table 5.

Table 5
Temp.,°Jointly submitted H2Sat.BTNSat.CoxPravr.Nl
% With the source material
6000%99%0,01%5%
6500% 98%0,1%8%
7000%96%1%12%
7500%85%9%17%
7506%89%5%
8000%68%24%23%
80020%84%8%
8500%45%46%29%
9000%20%71%37%

Of course, a quantitative measure of selectivity and conversion may vary due to different catalytic compositions, the application of owls the local feed materials (CO 2WITH H2O, H2About2, ethane, propane, etc), different working pressures and/or different volumetric flow rates of the raw material, but if the exact level described improvements may vary, the General direction is considered in the description of improvements still to be achieved. In addition, you must bear in mind that when performing discussed below modeling calculations assumed that directed into the reactor methane source material is always pre-heated to the same temperature (600°C), and in all cases we used a nominal feed rate of methane 100 kg/h is Also assumed that the temperature of the catalyst is sent to a reactor system with a movable layer, supported on one and the same meaning (850°C). The amount of catalyst required to maintain this temperature, was calculated for each configuration of the reactor. For simplicity, an assumption was made that thermal conductivity, coefficient of thermal diffusivity and the emissivity of the surface of the catalyst remained constant. The following table 6 lists the physical constants and properties of the catalyst used in calculations.

Table 6
The model parameters
The density of the catalyst particles1400kg/m3
The heat capacity of the catalyst1262J/kg·K
thermal conductivity of the catalyst0,4W/m·K
The coefficient of thermal diffusivity of the catalystof 2.26×10-7m2/s
The emissivity of the surface of the catalyst0,85

To enable simulation of various configurations of the reactor to Sat.BTNAnd sat down.Coxand Pravr.Nlbrought equation by obtaining for the above set of experimental points of polynomial equations by the criterion of best fit; the experimental point where the original material included N2to calculate the equations were not involved. The equations and the values of R2below:

Sat.BTN=(1,E-10)T4-(5,E-07)T3+(5,E-04)T2-(2,A-01)T+4,E+01

R2BTN=9,E-01

Sat.Cox=(-1,E-10)T4+(5,E-07)T3-(6,E-04)T2+(2,I-01)T-5,E+01

R2Cox=9,E-01

Pravr.Nl=(1,E-06)T2-(1,E-03)T+4,E-01

R2Nl=9,E-01,

where T is the temperature in °C.

In all examples R2is the coefficient of determination, which compares estimated and actual y values and ranges of values from 0 to 1. If it is equal to 1, then the sample excellent correlation is the difference between the calculated y value and the actual value y is absent. If the determination coefficient has a different extreme value is 0, the regression equation will not help in predicting y values. The variant used in the present description, based on the analysis of the decomposition of variance:

In the above definition

In other words, SSTis the full sum of squares, SSRis the regression sum of squares, a SSE- the sum of squared errors.

R2BTNis the coefficient of determination for the correlation Sat down.BTN,

R2Coxis the coefficient of determination for the correlation Sat down.Coxand

R2Nlis the coefficient of determination for the pair correlation is AI Pravr. Nl.

This set of equations was used to calculate the values of the output that was achieved for different configurations of the reactor, where the value of the OutputBTNwas defined as the product of Villages.BTN×Pravr.Nl, integrated over the profile of the temperature distribution in the reactor system, and the magnitude of the OutputCoxwas defined as the product of Villages.Cox×Pravr.Nl, integrated over the profile of the temperature distribution in the reactor system. Although it was determined and shown in table 5, which is a byproduct of H2improves the selectivity of the reaction, these equations are such an improvement of selectivity do not take into account, providing a conservative estimate on the level of improvement provided in the implementation of the proposed method.

The reactor with the transport of the catalyst or Elevator-reactor

Using the above equations for the reactor with the transport of the catalyst, or Elevator-reactor with adiabatic lowering of the temperature when the inlet temperature of 850°C, the required circulation rate of the catalyst to maintain the outlet temperature of 800°C was defined as 3211 kg per hour (kg/h) on the face of methane feed rate of 100 kg/h at 600°Spri were calculated for the following values of yield and selectivity:

Sat.BTN=51%

Sat.Cox=40%

you who ar BTN=12%

outputCox=8,9%

Δ treaction=-50°C (negative temperature 50°C),

where δ treactionis defined as the product temperature at the outlet of the reaction (i.e. the last temperature at which the catalytic reaction proceeds before the hydrocarbon product leaves the reactor system) minus the temperature of the hydrocarbon feedstock at the inlet of the reaction (i.e. the first temperature at which the catalytic reaction proceeds when the hydrocarbon feedstock enters the reactor system).

A reactor with a fixed catalyst bed

Modeling of comparable processes in potential systems with a fixed bed of the catalyst showed a worse characteristics than in the case of a reactor with the transport of the catalyst or Elevator-reactor, because in configurations with a fixed layer of the heat of reaction must be fed with matenadarani flow (because of heat in the reaction zone of a moving catalyst is not used). Therefore, for the operation of the reactor with a porous layer was required to heat metadatabase stream to a temperature far above the desired outlet temperature of 800°C, which thus led to the increase in the modulus exponent δ treactionamounting to -60°C or more in the module with a minus sign.

The reactor settling slo is m catalyst

In the case of modeling deposited catalyst layer with a reverse temperature profile and the differential between the temperature of the catalyst and the outlet temperature of the process, which constituted 50°C, the reactor was operating at the entrance at 620°C, and the output is at 800°C, the circulation rate of the catalyst was reduced to 717 kg/h, and the results of the reaction was improved:

Sat.BTN=89%

Sat.Cox=7%

OutputBTN=20%

OutputCox=1,5%

Δ treaction=+180°C

Example 9

Based on predicted when modeling the benefits of a reverse temperature profile, to confirm the results of the simulation was designed to install a laboratory scale. Although the model was focused on the work of the reaction system as a system with a deposited layer of the catalyst, laboratory reactor had a fixed bed of catalyst with a reverse temperature profile due to the use of external heaters. In all cases, the experimentally observed degree of conversion was lower degrees of conversion, predicted by the model. This may be due to the presence associated with interference experiment laboratory scale, such as bypassing the catalyst layer and/or the back of the mix, due to the hydrodynamic regime, where the work of the reactor laboratory is atomnogo scale.

The catalyst Mo/ZSM-5 prepared by grinding in a ball mill to 7.5 wt.% Mo (mass percentage of the metal in the total mass of the catalyst) in the form of Moo3with NH4ZSM-5 media (having a ratio Si/Al225) for 2 hours followed by calcination at 500°C for 5 hours in air. The catalyst alloy preformed, to grind and sieved to select particles with a particle size of 20-40 mesh. Catalytic testing of Mo/ZSM-5 catalyst was carried out in a quartz reactor with an inner diameter of 7 mm with a porous layer when the length of the layer to about 18 refer To all layers were of equal length, as a thinner layer of catalyst used inert silica particles (particle size 20-50 mesh).

Functional properties of the catalyst when dehydrocyclization of methane to benzene was determined using as starting material a mixture of 95 wt.% CH4and 5 wt.% argon (argon was used as internal standard). All experimental data of the reaction was obtained at an absolute pressure of 20 psi (138 kPa). For determination of the concentration of products otheriwse from the reaction stream was analyzed using a mass spectrometer.

For comparison was carried out ten separate experiments regarding the functional properties of the catalyst. In all experiments, the catalyst was activated heat is the W in the environment 15% vol. CH 4that 80% vol. H2, 5% vol. Ar with heating rate 5°C/min to 800°C and aged for 30 minutes followed ageing of the catalyst with five cycles of reaction and regenerating (also the same for all ten experiments). Each reaction part lasted for 20 min at 800°C in an environment of 95% vol. CH4and 5% vol. Ar used as source material in a volumetric feed rate of the starting material 1,4 h-1is determined by the CH4. Each regeneration part consisted of connection to the source of N2, heating up to 850°C with a 10 minute exposure and subsequent cooling again to 800°C (total time spent in the environment of H214 min). These ten experiments differed only their sixth reaction cycle, which was carried out in an environment of 95% vol. CH4and 5% vol. Ar used as source material for 4 hours Conditions for the sixth cycle was chosen for comparison of the influence of the temperature profile of the catalyst layer at different volumetric flow rates of the source material. In particular, experiments 1-5 were performed, varying the volumetric feed rate of the source material ranging from 0.25 to 8 h-1while maintaining the catalyst at isothermal conditions at 800°C. In contrast, the experiments 6-10 conducted in the same interval of values of flow rate, what about the linear temperature gradient in the catalyst bed, which was increased from 650°C at the inlet to the source material up to 800°C at the outlet for the products (reverse temperature profile). The results of determining the functional properties of the catalyst for the ten experiments during the reaction cycle No. 6 are summarized in table 7.

The results in table 7 show the obvious advantage of doing process with an inverse temperature profile, if most of the values of the volumetric feed rate of the source material was improved instantaneous selectivity for shorter periods of time process and consistently provided the prolongation of selectivity to benzene over extended periods of time process. It is possible to increase the total number of products in comparison with the isothermal layer at all volumetric flow rates of the source material.

Although the examples 8 and 9 relate to reactors with a deposited layer of the catalyst, a similar improvement in selectivity would be demonstrated for other reactor systems with equivalent inverse temperature profile.

As shown by simulation, improved values of yield, selectivity and speed of circulation of the catalyst. Furthermore, the reaction can generally be accomplished in one reaction zone, thus minimizing the need for the equipment. As an option, you can use two or more reaction zones.

As it is illustrated by examples 8 and 9, the configuration of the reactor with a deposited layer or other reactory system with the same inverse temperature profile make possible the conversion of methane to heavier hydrocarbons, such as aromatic compounds, with reduced losses due to aging/mechanical abrasion of the catalyst, improved efficiency and higher selectivity, i.e. with less coke formation than in configurations with a fixed layer and/or transport layer or upward movement of the layer in the vertical pipe configurations.

Example 10

This example is intended to show the effect of temperature regenerating on the stability of the degree of conversion of methane and selectivity for benzene and toluene at dehydrocyclization methane for several cycles.

The catalyst of 7.5 wt.% Mo/ZSM-5 prepared by grinding in a ball mill, molybdenum oxide with the carrier NH4+-ZSM-5 (having a ratio Si/Al225) for 2 hours followed by calcination at 500°C for 5 hours in air flow. The catalyst alloy preformed, to grind and sieved to select particles with a particle size of 20-40 mesh. Catalytic testing of catalyst Mo/ZSM-5 was carried out in a quartz reactor with a fixed bed is. A sample of catalyst weight of approximately 300 mg was placed in a tubular reactor with a diameter of 1/4" (6.35 mm). As a thinner layer of catalyst used inert silica particles (particle size 20-40 mesh).

The catalyst synthesized in state pre-carbonized by heating the catalyst in a gaseous mixture of 15% of CH4-N2from 150°C to 800°C at 5°C/min and kept at 800°C for 0.5 hour. Then the catalyst was subjected to cyclic aging, which consisted of: (1) the period of contact with the hydrocarbon, when the catalyst was affected by hydrocarbon source material (molar composition: 13,4% H2, 78.2% of CH4group , 1.8%, 1,5%6H6, 0,2%2H4and 4.9% Ar) at 785°C., an absolute pressure of 23 psi (159 kPa) and flow rate 5 h-1CH4and (2) follow-up period, regenerating, where the catalyst was heated to about 850-925°C in an environment of gaseous H2at an absolute pressure of 100 psig (689 kPa) for a specified time at a volumetric rate of gas supply 34000 cm [normal]/g catalyst/hour, After the stage of regenerating the catalyst before re-feeding hydrocarbon starting material was cooled to 785°Sakagami flow reactor were analyzed by GC and MS in real time.

During the first 3670 cycles the catalyst was subjected to the action of hydrocarbon source material for 4-6 minutes at 785°C and regenerates in the environment of H 2within 6-12 minutes at 850°C or 875°C for each cycle. With increasing number of cycles the catalyst showed a more rapid deterioration of activity in relation to the degree of conversion of methane and selectivity to benzene and toluene. To restore the activity of the catalyst was regenerated during these cycles 8 times during 12-23 hours. Continuous regeneration could temporarily restore the catalyst, but its functional properties still deteriorated rapidly.

On figa and 5B shows the characteristic changes expressed in percentage the degree of conversion of methane and selectivity for benzene and toluene with increasing temperature regeneration with 875°C to 925°Spoke cycle 3610 the catalyst was regenerated in the environment of H2at 900°C for 20 hours, then the catalyst was returned in the cyclic process and in each cycle between cycles 3610 and 3670 he was subjected to the action of hydrocarbon source material for 4 minutes at 785°C and regenerates in the environment of H2within 12 minutes at 875°C. the Catalysts showed rapid degradation activity towards the conversion of methane and selectivity to benzene and toluene. After cycle 3670 the catalyst was regenerated in the environment of H2at 925°C for 20 hours, then the catalyst was returned in the cyclic process and in each cycle between cycles 3680 and 3800 was exposed uglev the portly source material for 4 minutes at 785°C and regenerates in the environment of H 2within 12 minutes at 925°C. In the case of the regenerate at a temperature of 925°C, the catalyst showed the stability of the degree of conversion of methane and selectivity for benzene and toluene, demonstrating the advantage of higher temperature for regenerating older catalyst.

Other embodiments of the inventions

1. The method of conversion of methane to heavier hydrocarbons containing aromatic hydrocarbons, including:

(a) feeding to a reaction zone metadatareader the source material and the granular catalytic material

(b) maintaining in the reaction zone under conditions effective to convert at least part of the methane in more high molecular weight(s) hydrocarbon(s) with concomitant deposition of carbonaceous material on granular catalytic material, causing the deactivation of the catalytic material,

(C) removing from the reaction zone at least a portion of the deactivated granular catalytic material

(g) heating at least part remote from the reaction zone of granular catalytic material to a temperature of from about 700°to about 1200°C, carried out by direct and/or indirect contact with the gaseous combustion products produced by combustion of additional fuel,

(d) Regener is of the heated part of the granular catalytic material is a hydrogen-containing gas in the first regeneration zone under conditions effective to convert at least part of the deposited carbonaceous material to methane, and

(e) returning at least part of the granular catalytic material from step (d) into the reaction zone.

2. The method according to claim 1, in which the reaction zone is a reaction zone with a movable layer.

3. The method according to claim 2, in which the reaction zone with a movable layer works with the inverse temperature profile.

4. The method according to any of the preceding paragraphs in which the source material also contains at least one of the following substances: CO, CO2N2H2O and/or O2.

5. The method according to any of the preceding paragraphs, in which remote from the reaction zone of granular catalytic material is heated in stage (d) to a temperature of from about 800°to about 1000°C., preferably from about 850°to about 950°C.

6. The method according to any of the preceding paragraphs, in which when heated in stage (g) the catalytic material is introduced into direct contact with an additional source of fuel or gaseous products of combustion.

7. The method according to any of the preceding paragraphs, in which the catalytic material is heated by direct contact with inert environment or heat exchange surface, the heated gaseous products of combustion.

8. The method according to any of the previous points is, where an additional source of fuel includes a hydrocarbon and/or hydrogen.

9. The method according to any of the preceding paragraphs, in which additional fuel source contains methane.

10. The method according to any of the preceding paragraphs, in which an additional source of fuel includes a hydrocarbon, the hydrocarbon is burned in oxygen-poor atmosphere to obtain synthesis gas.

11. The method according to any of the preceding paragraphs, in which an additional source of fuel contains hydrogen.

12. The method according to any of the preceding paragraphs, in which an additional source of fuel contains hydrogen generated as a by-product of the process of conversion of methane.

13. The method according to any of the preceding paragraphs, including:

(g) removing from the reaction zone an additional part of the granular catalytic material

(C) heating the additional part of the granular catalytic material by direct and/or indirect contact with the gaseous combustion products produced by combustion of additional fuel, and

(I) returning at least part of the granular catalytic material from step (C) into the reaction zone.

14. The method according to any of the preceding paragraphs, including:

(i) moving the second part remote from the reaction zone of granular catalytic material is from the stage (in) stage or (d) in the second regeneration zone,

(ii) regenerative second part of granular catalytic material is a hydrogen-containing gas in the second regeneration zone under conditions including a pressure different from the pressure in the first regeneration zone, and effective to convert at least part deposited on granular catalytic material of the carbonaceous material to methane, and

(iii) returning at least part of the regenerated granular catalytic material from step (g) into the reaction zone.

15. The method according to 14, in which the conditions of the regenerate (ii) include an absolute pressure of at least 500 kPa, preferably from about 1000 kPa to about 5000 kPa.

16. The method according to 14 or 15, in which the pressure in the second regeneration zone create by means of a vertical column of granular catalytic material.

17. The method according to any of the preceding paragraphs in which the terms regenerating stage (d) include an absolute pressure of at least 100 kPa, preferably from about 150 kPa to about 5000 kPa.

18. The method according to any of the preceding paragraphs, in which the pressure in the first regeneration zone to create through a vertical column of granular catalytic material.

19. The method according to any of the preceding paragraphs, in which the conditions of the reaction occurring in the reaction zone at the stage (b), include a temperature from about 400°to about 1200°C., an absolute pressure from about 1 kPa to about 1000 kPa and a space velocity of feed of the source material of about from 0.01 h-1approximately 1000 h-1.

20. The method according to any of the preceding paragraphs, in which granular catalytic material is a catalyst dehydrocyclization containing the metal or its compound on an inorganic carrier.

21. The method according to any of the preceding paragraphs, in which granular catalytic material includes at least one of the following substances: molybdenum, tungsten, rhenium compound of molybdenum, a compound of tungsten, a compound of zinc and a compound of rhenium, ZSM-5, silica or aluminum oxide.

22. The method of conversion of methane to heavier hydrocarbons containing aromatic hydrocarbons, including:

(a) filing metadatareader source material into the reaction zone containing granular catalytic material

(b) maintaining in the reaction zone under conditions effective to convert at least part of the methane in more high molecular weight(s) hydrocarbon(s) with concomitant deposition of carbonaceous material on granular catalytic material, causing the deactivation of the catalytic material,

(C) interruption metanode the containing source material into the reaction zone,

(g) heating the reaction zone to a temperature of regenerating factor of 700 to 1200°C, carried out by direct and/or indirect contact with the gaseous combustion products produced by combustion of additional fuel,

(d) regenerative located in the heated reaction zone of granular catalytic material is a hydrogen-containing gas under conditions effective to convert at least part of the deposited carbonaceous material to methane, and

(e) renewal filing metadatareader source material into the reaction zone.

23. The method according to item 22, in which the flow metadatareader interrupt source material after heating at stage (g).

24. The method according to item 22 or 23, in which the heating stage (d) is carried out in the presence of hydrogen-containing gas.

25. The method according to any of PP-24, in which the reaction zone is a reaction zone with a porous layer.

26. The method according to any of PP-25, in which the conditions regenerating stage (d) include an absolute pressure of at least 100 kPa.

Although the present invention is described and illustrated with reference to specific ways of its implementation, for the specialist obvious possibility of carrying out the invention, and embodiments other than those discussed above. For this reason, the definition and the true scope of protection of the present invention should refer only to the formula.

1. The method of conversion of methane to heavier hydrocarbons containing aromatic hydrocarbons, including:
(a) feeding to a reaction zone metadatareader the source material and the granular catalytic material,
(b) maintaining in the reaction zone under conditions effective to convert at least part of the methane in more high molecular weight(s) hydrocarbon(s) with concomitant deposition of carbonaceous material on granular catalytic material, causing the deactivation of the catalytic material,
(C) removing from the reaction zone at least a portion of the deactivated granular catalytic material,
(g) heating at least part remote from the reaction zone of granular catalytic material to a temperature of from about 700°to about 1200°C, carried out by direct or indirect contact with the gaseous combustion products produced by combustion of additional fuel,
(d) regenerative heated part of the granular catalytic material is a hydrogen-containing gas in the first regeneration zone under conditions effective to convert at least part of the deposited carbonaceous material to methane, and
(e) returning at least part of the granular catalytic material from step (d) into the reaction zone.

3. The method according to claim 2, in which the reaction zone with a movable layer works with the inverse temperature profile.

4. The method according to one of claims 1 to 3, in which the source material also contains at least one of the following substances: CO, CO2N2N2O and/or O2.

5. The method according to one of claims 1 to 3, which is remote from the reaction zone of granular catalytic material is heated in stage (d) to a temperature of from about 800°to about 1000°C.

6. The method according to one of claims 1 to 3, in which when heated in stage (g) the catalytic material is introduced into direct contact with an additional source of fuel or gaseous products of combustion.

7. The method according to one of claims 1 to 3, in which the catalytic material is heated by direct contact with inert environment or heat exchange surface, the heated gaseous products of combustion.

8. The method according to one of claims 1 to 3, in which an additional source of fuel includes a hydrocarbon and/or hydrogen.

9. The method according to one of claims 1 to 3, in which additional fuel source contains methane.

10. The method according to one of claims 1 to 3, in which an additional source of fuel includes a hydrocarbon, the hydrocarbon is burned in oxygen-poor atmosphere obtained with the eat synthesis gas.

11. The method according to one of claims 1 to 3, in which an additional source of fuel contains hydrogen.

12. The method according to one of claims 1 to 3, in which an additional source of fuel contains hydrogen generated as a by-product of the process of conversion of methane.

13. The method according to one of claims 1 to 3, including:
(g) removing from the reaction zone an additional part of the granular catalytic material,
(C) heating the additional part of the granular catalytic material by direct or indirect contact with the gaseous combustion products produced by combustion of additional fuel, and
(I) returning at least part of the granular catalytic material from step (C) into the reaction zone.

14. The method according to one of claims 1 to 3, including:
(i) moving the second part remote from the reaction zone of granular catalytic material from the stage (in) stage or (d) in the second regeneration zone,
(ii) regenerative second part of granular catalytic material is a hydrogen-containing gas in the second regeneration zone under conditions including a pressure different from the pressure in the first regeneration zone, and effective to convert at least part deposited on granular catalytic material of the carbonaceous material to methane, and
(iii) return on Myung is our least part of the regenerated granular catalytic material from step (g) into the reaction zone.

15. The method according to 14, in which the pressure in the second regeneration zone create by means of a vertical column of granular catalytic material.

16. The method according to one of claims 1 to 3, in which the conditions regenerating stage (d) include an absolute pressure of at least 100 kPa, preferably from about 150 kPa to about 5000 kPa.

17. The method according to one of claims 1 to 3, in which the pressure in the first regeneration zone to create through a vertical column of granular catalytic material.

18. The method according to one of claims 1 to 3, in which the conditions of the reaction occurring in the reaction zone at the stage (b)include a temperature from about 400°to about 1200°C., an absolute pressure from about 1 kPa to about 1000 kPa and a space velocity of feed of the source material of about from 0.01 h-1approximately 1000 h-1.

19. The method according to one of claims 1 to 3, in which the granular catalytic material is a catalyst dehydrocyclization containing the metal or its compound on an inorganic carrier.

20. The method according to one of claims 1 to 3, in which the granular catalytic material includes at least one of the following substances: molybdenum, tungsten, rhenium compound of molybdenum, a compound of tungsten, a compound of zinc and a compound of rhenium, ZSM-5, silica or aluminum oxide is tion.

21. The method of conversion of methane to heavier hydrocarbons containing aromatic hydrocarbons, including:
(a) filing metadatareader source material into the reaction zone containing granular catalytic material,
(b) maintaining in the reaction zone under conditions effective to convert at least part of the methane in more high molecular weight(s) hydrocarbon(s) with concomitant deposition of carbonaceous material on granular catalytic material, causing the deactivation of the catalytic material,
(C) interruption metadatareader source material into the reaction zone,
(g) heating the reaction zone to a temperature of regenerating factor of 700 to 1200°C, carried out by direct or indirect contact with the gaseous combustion products produced by combustion of additional fuel,
(d) regenerative located in the heated reaction zone of granular catalytic material is a hydrogen-containing gas under conditions effective to convert at least part of the deposited carbonaceous material to methane, and
(e) renewal filing metadatareader source material into the reaction zone.

22. The method according to item 21, in which the flow metadatareader interrupt source material after heating at stage (g).

23. The method according to item 21 or 22, in which the heating stage (d) is carried out in the presence of hydrogen-containing gas.

24. The method according to item 21 or 22, in which the reaction zone is a reaction zone with a porous layer.

25. The method according to item 21 or 22, in which the conditions regenerating stage (d) include an absolute pressure of at least 100 kPa.



 

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31 cl, 3 tbl, 2 dwg, 21 ex

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6 cl, 1 tbl, 4 ex, 1 dwg

FIELD: chemistry.

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2 ex

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The invention relates to the regeneration of the solid catalyst, which comprises the reaction product of a metal halide selected from the group comprising aluminum, zircaloy, tin, tantalum, titanium, gallium, antimony, phosphorus, iron, boron and their mixture, and bound surface hydroxyl groups of the refractory inorganic oxide and metal with zero valency selected from the group comprising platinum, palladium, Nickel, ruthenium, rhodium, osmium, iridium, and their mixture

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FIELD: petrochemical industry; methods of production of the ethylene non-saturated halogen-containing aliphatic hydrocarbons.

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36 cl, 1 ex, 9 dwg

FIELD: chemistry.

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2 ex

FIELD: oil and gas industry.

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6 cl, 1 tbl, 4 ex, 1 dwg

FIELD: technological processes.

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31 cl, 3 tbl, 2 dwg, 21 ex

FIELD: process engineering.

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EFFECT: ruling out separate activation step and activation at higher temperatures.

4 cl, 1 tbl, 9 ex

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26 cl, 6 dwg, 10 ex, 7 tbl

Catalysts // 2497590

FIELD: chemistry.

SUBSTANCE: invention relates to catalyst regeneration. Method of regeneration of waste powder-like, paraffin-containing catalyst of Fischer-Tropsch synthesis based on cobalt is described, where claimed method includes the following sequential processing: (i) de-waxing processing, (ii) oxidative processing with regulation of work temperature by discharge of heat from layer of catalyst particles with application of cooling device, which contains device for providing passage of cooling medium and cooling medium, passing through said device of passage providing, which ensure in such way heat conductive surfaces, located in and/or around catalyst layer, with obtaining oxidised particles of catalyst, and (iii) reduction processing. Re-application of regenerative catalyst is described.

EFFECT: increase of process efficiency.

15 cl, 9 dwg, 4 ex

FIELD: chemistry.

SUBSTANCE: invention relates to catalysis. Described is a method of regenerating one or more particles of a cobalt-containing Fischer-Tropsch catalyst in situ in a reactor tube, the method comprising steps: (i) oxidising the catalyst particle(s) at temperature of 20-400°C; (ii) treating the catalyst particle(s) for more than 5 min with a solvent; (iii) drying the catalyst particle(s); and (iv) optionally reducing the catalyst with hydrogen or any hydrogen-containing gas.

EFFECT: high catalyst activity.

10 cl, 4 tbl, 4 ex

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