RU2678980C1 - Aromatising non-aromatic hydrocarbons - Google Patents

Abstract

FIELD: chemistry.SUBSTANCE: invention relates to two variants of the system for producing aromatic hydrocarbons and two variants of the method. One of the options of the system includes: (a) flavouring, configured to (i) receive a specified amount of ≥1 million standard cubic feet per day [MSCFD] (29000 NM3D) of raw materials for flavouring, where the raw material for flavouring includes C2 hydrocarbons and non-aromatic C3 + hydrocarbons and has a predetermined concentration of non-aromatic C2 + hydrocarbons in the range from 15 to 90 mol. % per mole of raw aromatization, and (ii) flavouring (A) at least part of the non-aromatic C3 + hydrocarbons of the raw material for flavouring and (B) at least part of the hydrocarbon C2 raw materials for aromatization with obtaining the outgoing reaction stream containing molecular hydrogen, non-aromatic hydrocarbons and aromatic hydrocarbons; (b) product heat exchanger in fluid communication with the flavouring agent, configured to extract at least the first and second products from the outgoing reaction stream, where the first product includes at least part of the aromatic hydrocarbons leaving the reaction stream and the second product includes a tail gas containing (i) at least part of the molecular hydrogen of the outlet reaction stream; and (ii) at least part of the non-aromatic hydrocarbons of the outlet reaction stream; and (c) raw material pretreatment unit configured for (A) receiving ≥2 MSCFD (57000 NM3D) gaseous raw materials and (B) obtaining raw materials for flavouring from gaseous raw materials, where the gaseous raw materials comprise from 0 to 98 mol. % methane, from 1 to 50 mol. % C2 hydrocarbons and from 1 to 40 mol. % non-aromatic C3 + hydrocarbons, where the installation for pre-processing of raw materials includes (i) at least one heat transfer step with (A) input in the fluid communication with the pipeline for the raw materials and (B) the output, the heat transfer stage is configured to transfer heat from the gaseous raw material in order to condense part of the gaseous raw material at a fixed dew point, and (ii) at least one vapour-liquid separator, where the vapour-liquid separator has an inlet in fluid communication with the heat transfer step, at least one drum-separator and the first and second outputs, where: (A) separator inlet is configured to receive gaseous feedstock and condensed portions from the heat transfer step, (B) drum separator is configured to separate the gaseous feedstock and the condensed portions to produce a bottom stream, including a flavouring feed, and a top stream, including separated gaseous hydrocarbons, (C) first separator outlet is configured to transfer the separated gaseous portion from the pretreatment unit, and (D) second separator outlet is configured to transfer a predetermined amount of flavouring raw material to flavouring, and where (iii) the first stage of heat transfer installation for pre-processing includes at least one heat exchanger in indirect thermal contact with the flow of the refrigerant, (iv) the first heat exchanger is adapted to indirectly transfer the first, second and third parts of heat transferred from the gaseous feedstock, where the first part is transferred to the refrigerant, the second part is transferred to the separated gaseous hydrocarbons and the third part is transferred to the raw material for flavoring, and (v) the second heat transfer stage is configured to remove some heat from the refrigerant flow, where the amount of heat varies by no more than +/-50 %, whereas the condensed part condensed from the gaseous raw material has a volume change of ≤+/-10 %.EFFECT: present invention does not require the adaptation of reactors and related equipment to the composition of a specific C2-C9 hydrocarbon source and can be carried out using standardized modules that are easily transported to remote locations.21 cl, 1 tbl, 3 dwg

Description

CROSS REFERENCE TO RELATED APPLICATIONS

[0001] This claims claims priority over US Patent Application Serial No. 62 / 232.609, filed September 25, 2015, and 62 / 253.268, filed November 10, 2015, under the name AROMATIZATION OF NON-AROMATIC HYDROCARBONS; and European Patent Application No. 15195311.4, filed November 19, 2015, which are all fully incorporated herein by reference.

[0002] The following related applications are also fully incorporated herein by reference: US Patent Application Serial No. 62 / 234.262, filed September 29, 2015; European Patent Application No. 15195314.8, filed November 19, 2015; US Patent Application Serial No. 62 / 234.240, filed September 29, 2015; European Patent Application No. 15197698.2, filed December 3, 2015; US Patent Application Serial No. 62 / 247.795, filed October 29, 2015; European Patent Application No. 15197700.6, filed December 3, 2015; US Patent Application Serial No. 62 / 248.374, filed October 30, 2015; European Patent Application No. 15197702.2, filed December 3, 2015; US Patent Application Serial No. 62 / 253.268, filed November 10, 2015; US Patent Application Serial No. 62 / 326.918, filed April 25, 2016; European Patent Application No. 16175163.1, filed June 20, 2016; US Patent Application Serial No. 62 / 299.730, filed February 25, 2016; European Patent Application No. 16167395.9, filed April 28, 2016; US Patent Application Serial No. 62 / 313.288, filed March 25, 2016; European Patent Application No. 16173587.3, filed June 8, 2016; US Patent Application Serial No. 62 / 313.306, filed March 25, 2016; European Patent Application No. 16173980.0, filed June 10, 2016; US Patent Application Serial No. 62 / 298.655, filed February 23, 2016; and European Patent Application No. 16167672.1, filed April 29, 2016.

FIELD OF TECHNOLOGY

[0003] The systems and methods described herein relate to aromatization of non-aromatic hydrocarbons, equipment and materials used in such aromatization, and the use of such aromatization, for example, to enrich natural gas.

BACKGROUND

[0004] The improvement of production methods has led to an increase in the availability of hydrocarbon gases of geological origin, for example, natural gas. Natural gas mainly consists of a mixture of one or more hydrocarbon compounds with the number of carbon atoms in the range from C ₁ to C ₉ , and may additionally contain non-hydrocarbon compounds such as CO ₂ and / or H ₂ Ѕ. Natural gas is used for a wide variety of purposes, for example, generating electricity, transport fuel, heating residential and industrial facilities and cooking, and as a raw material for the commercial production of general-purpose reagents such as ethylene.

[0005] Often natural gas is produced from a geological formation that is at some distance from where it is consumed. This stimulated the construction of natural gas pipelines transporting compressed natural gas from producers to consumers. Some consumers of natural gas, such as operators of electric generators, pay particular attention to the calorific value of natural gas from the pipeline. When natural gas is too low in calories, power generators are inefficient. Too high calories can lead to equipment damage, for example, as a result of overheating of the combustion stages and expansion of the generator.

[0006] In order to avoid difficulties, pipeline operators have developed and approved specifications for the natural gas entering the operator’s pipeline. The calorific value, which is a measure of the calorific value of natural gas, is one of the typical technical characteristics of a pipeline. Another characteristic is the Wobbe index, which is closely related to the energy of combustion of natural gas. Two different natural gases, obtained from different sources and having different compositions, but with essentially the same Wobbe index, will have almost the same calorific value when burned under substantially the same conditions. For example, a pipeline operator may require that natural gas entering its pipeline have a calorific value of 36.07 MJ / cm ³ to 41.40 MJ / cm ³ ) and a Wobbe index of 49.01 MJ / cm ³ to 52, 22 MJ / cm ³ ).

[0007] The calorific value and Wobbe index of natural gas are highly dependent on the type and relative amount of hydrocarbon gas compounds. As a rule, hydrocarbon compounds having a greater number of carbon atoms provide increased heat of combustion. Therefore, when natural gas does not meet the minimum requirements for calorific value and the Wobbe index, propane enrichment can be used to obtain a gas suitable for the pipeline. Most often, natural gas exceeds the technical requirements for calorific value and the Wobbe index. In these cases, the natural gas producer can extract part of the relatively high heat content of the hydrocarbon gas compounds, for example, an excess of ethane and / or an excess of propane that goes beyond what is necessary to achieve the minimum regulated values of the calorific value and the Wobbe index.

[0008] Fractionation and regulation of the dew point are common methods for producing a natural gas suitable for a gas pipeline from a natural gas source that does not meet the technical requirements for calorific value and the Wobbe index for the pipeline. For example, natural gas with an excess Wobbe index and / or calorific value can be fractionated to produce a bottoms fraction consisting mainly of C ₅₊ hydrocarbons, a side stream containing mainly C ₃ and C ₄ hydrocarbons, and an overhead stream containing mainly methane and ethane. The bottom fraction is usually taken away, for example, for mixing with other liquid hydrocarbons. In places with equipment for this, the side stream can be compressed and transported as a LPG. At remote locations, the lateral flow is usually reintroduced into the reservoir. Fractionation is designed to produce a methane / ethane mixture that meets the requirements of pipeline transportation. Since the composition of natural gas and the corresponding pipeline transportation requirements vary by location, the fractionation conditions cannot be easily standardized. Partly due to this difficulty, fractionation is usually applied when natural gas is in the form of crude natural gas discharged from a natural gas well (“gas from a gas well”), in particular that has a lower C ₃₊ hydrocarbon content.

[0009] Dew point control is typically used when natural gas has a higher C ₃₊ hydrocarbon content, for example, when natural gas is associated. Associated gas is a form of natural gas that is commonly found in oil fields, for example, dissolved in oil or as a free “gas dome” above oil in a reservoir. To regulate the dew point, a mixture of gaseous and liquid hydrocarbons discharged from the reservoir is fed to the vapor-liquid separation stage. The bottom product containing crude oil is diverted. An overhead stream containing associated gas is also diverted, at least part of which is fed to the dew point control stage. The dew point control is carried out by cooling the associated gas to a sufficiently low temperature at a sufficiently high pressure to condense a sufficient part of it so that the remaining part of the vapor phase meets the requirements of pipeline transportation in this place. The liquid phase is separated from the vapor phase, for example, in a drum separator. Suitable natural gas for the pipeline, containing mainly methane and ethane, is withdrawn from the separator drum as an overhead stream. The bottoms stream, containing mainly C ₂₊ hydrocarbons, is also diverted. It is customary to convert at least part of the bottom stream into aromatic hydrocarbons, which significantly reduces the difficulties associated with transporting the bottoms stream from remote locations. See, for example, PCT published patent application No. WO 2015 / 084518A1, which is incorporated herein by reference in its entirety.

[0010] One of the aromatization methods disclosed in US Pat. No. 4,8555.522 involves the conversion of C ₂ , C ₃ and C ₄ hydrocarbons with increased selectivity for aromatic hydrocarbons. The method uses a dehydrocyclization catalyst comprising (a) an aluminosilicate with a molar ratio of silicon to aluminum of at least 5 and (b) a mixture of (i) Ga and (ii) at least one rare-earth metal. Reference material describes the implementation of the conversion by aromatization at a space velocity (LHSV) in the range from 0.5 to 8 h ⁻¹ , temperature ≥ 450 ° C (for example, from 475 ° C to 650 ° C), pressure from 1 bar to 20 bar and contact time of the raw material from 1 to 50 seconds. One of the difficulties observed with this method is that the accumulation of catalytic coke gradually reduces the yield of aromatic hydrocarbons. One way to overcome this difficulty involves cyclic reaction, for example, by conducting a reaction in at least two reactors. The first reactor performs aromatization (interaction mode), while the second reactor performs coke removal (regeneration mode), and vice versa. To reduce the amount of coke accumulation, a shortening of the period of time (“cycle time”) is used during which the fixed-bed reactor is operated in the interaction mode before switching to the regeneration mode. However, shortening the cycle time reduces the yield of aromatic hydrocarbons formed in the process, especially when a significant amount of coke is accumulated. In these cases, it may be necessary to carry out the regeneration mode over a relatively long time interval compared to the aromatization mode. Another way to overcome this difficulty involves the use of fluidized catalyst beds in the reaction stages. However, this usually requires special catalysts and technological conditions, and it is difficult to put into practice, especially in remote places.

[0011] Other difficulties encountered in the implementation of the generally accepted method of aromatization of natural gas arise as a result of changes in the composition of natural gas produced in various places. Some natural gas is relatively rich in ethane; other natural gas contains much less. When carrying out the conventional aromatization of C ₂ -C ₄ hydrocarbons, it is observed that the conditions necessary for aromatizing C ₂ usually lead to excessive conversion of C ₃ and C ₄ hydrocarbons to coke. In addition, conditions optimized for the aromatization of C ₃ and C ₄ hydrocarbons tend to result in a lower yield of aromatics from C ₂ hydrocarbons _of natural gas. This situation is exacerbated by the wide variation in the ethane content in natural gas produced from various reservoirs.

[0012] One way to overcome this is to specifically tailor the aromatization method so that it is compatible with a particular reservoir. For example, US Patent No. 8.835.706 describes the aromatization of ethane-propane feed in two reactors operating in a complex series / parallel connection. The feedstock for the process, obtained from natural gas by the cryogenic separation of methane, is subjected to interaction at the first stage when operating under conditions that maximize the conversion of propane to aromatic compounds. In a second step, ethane is reacted to produce additional aromatics. The second stage is carried out under conditions that maximize the conversion of ethane to aromatic hydrocarbons. The method can be specially adapted for a specific source of natural gas, for example, by optimizing the appropriate dimensions of the first reactor, the second reactor and the accompanying stages of combining and separation. However, this approach limits the applicability of the method to natural gas with a relatively narrow compositional range.

[0013] Therefore, there is a need for improved methods for aromatizing non-aromatic C ₂ -C ₉ hydrocarbons, such as hydrocarbons extracted from natural gas. In particular, there is a need for aromatization methods that make it possible to work with a relatively long cycle time and to avoid the difficulties of a conventional method. Aromatization methods that do not require the adaptation of reactors and related equipment to the composition of a specific source of C ₂ -C ₉ hydrocarbons, and, in particular, methods that can be carried out using standardized modules that are easily transported to remote locations, are even more desirable.

SUMMARY OF THE INVENTION

[0014] The invention mainly relates to the enrichment of gaseous hydrocarbons of various compositions obtained from various sources at various geographical locations. Each of the sources includes methane, C ₂ hydrocarbons and non-aromatic C ₃₊ hydrocarbons, but contains substantially different relative amounts of these components. It was found that controlling the inverse dependence of the volumetric flow rate of such gaseous hydrocarbons from the source on the concentration of hydrocarbons C _{2+ of the} hydrocarbon source before entering the separation stage provides a separated stream of C ₂₊ hydrocarbons with the required characteristics. The characteristics of the separated stream of C ₂₊ hydrocarbons are both a practically constant volumetric flow rate and a practically constant concentration of C ₂₊ hydrocarbons. Unexpectedly, it was found that these characteristics are achievable for a wide range of gaseous hydrocarbons, which can be obtained at a wide variety of geographical points in a wide range of ways.

[0015] This discovery led to the development of a system and method that includes at least two main functions: a flavoring function and a product recovery function. The aromatization function involves the production of gaseous raw materials for aromatization with (i) a given volumetric flow rate and (ii) a given concentration of non-aromatic hydrocarbons C ₂₊ . The aromatization function provides an exit reaction stream including aromatic hydrocarbons obtained by dehydrocyclization of at least a portion of the non-aromatic hydrocarbons C _{2+ of the} aromatization feed. The product recovery function provides the separation of at least a portion of aromatic hydrocarbons from the effluent. The method is advantageous in that it makes it possible to work with relatively long cycle times and to avoid the difficulties of the conventional method. Since the aromatization function provides for aromatization a predetermined volumetric flow rate of the flow of raw materials for aromatization having a predetermined concentration of C ₂₊ hydrocarbons, the method can be implemented in modules configured for specific functions, for example, using modules for individual preliminary processing of raw materials, aromatization and product recovery, configured to perform these functions. Modules, which can be of a given size and capacity, can easily be transported to various remote places, for example, for processing natural gas.

DESCRIPTION OF DRAWINGS

[0016] The advantages of this method will be better understood when referring to the following detailed description and the accompanying drawings, where:

[0017] FIG. 1 is a block diagram of a modular system for producing a product including aromatic hydrocarbons.

[0018] FIG. 2 is a simplified flow diagram illustrating some aspects of the invention.

[0019] FIG. 3 is a flow chart depicting a flavoring comprising two dehydrocyclization steps.

DETAILED DESCRIPTION

[0020] Although some aspects of the invention are disclosed in more detail, the invention is not limited to them, and this description does not exclude other aspects within the true essence and scope of the attached claims.

[0021] Some aspects of the systems and methods described herein relate to hydrocarbon conversion using a flavoring agent, recuperator, and optional pretreatment unit. For the conversion of hydrocarbons, a gaseous feedstock containing C ₂ hydrocarbons and non-aromatic C ₃₊ hydrocarbons is used. The gaseous feedstock is treated in a plant for pre-processing the feedstock, providing a feedstock for flavoring for flavoring, the gaseous feedstock for flavoring has (i) a predetermined volumetric flow rate and (ii) a predetermined concentration of non-aromatic C ₂₊ hydrocarbons. Preliminary processing of raw materials can be carried out in stages, for example, by removing heat from gaseous raw materials at the stage of heat transfer in order to condense part of the gaseous raw materials. The separation step can be used to separate the condensate from the remainder of the gaseous feed which contains gaseous hydrocarbons. The aromatization feed is obtained by evaporating at least a portion of the condensate. Evaporation of the condensate can be carried out at any convenient point, usually before the dehydrocyclization reaction. The form of heat transfer used is not critical and can be direct, indirect or a combination thereof. The separation form is not critical and may include one or more fractionation columns, one or more separator drums, one or more membranes, or a combination thereof. A generally accepted method of heat transfer and separation may be used, but the invention is not limited to this. For gaseous feedstocks obtained from various sources at different geographical locations, the condensation rate of the gaseous feedstocks can be adjusted so that the pretreatment gives a predetermined volumetric flow rate of the feedstock for flavoring, while the feedstock for flavoring will have a predetermined concentration of non-aromatic C ₂₊ hydrocarbons. Preferably, the predetermined values are practically constant, for example, no significant changes were required for the practical implementation of the invention using gaseous feedstocks extracted at different geographical locations.

[0022] This can be achieved, for example, by one or more of the following methods: (i) reducing the volumetric flow rate of the gaseous feedstock for pretreatment in geographical areas where the gaseous feedstock has a higher concentration of non-aromatic C ₂₊ hydrocarbons, (ii) increasing the volumetric flow rate of the gaseous feedstock in geographical areas where the gaseous feedstock has a lower concentration of non-aromatic C ₂₊ hydrocarbons, (iii) a decrease in the amount of heat removed from the gaseous feedstock, in geographic areas where gaseous feedstock has a higher concentration of non-aromatic C ₂₊ hydrocarbons, (iv) an increase in the amount of heat removed from gaseous feed, in geographic areas where gaseous feed has a lower concentration of non-aromatic C ₂₊ hydrocarbons, (v) increase a feed gas pressure in geographical areas where the raw material gas has a lower concentration of non-aromatic C ₂₊ hydrocarbons, and (vi) a feed gas pressure reduction on geographical areas x, where the raw material gas has a higher concentration of non-aromatic C ₂₊ hydrocarbons.

[0023] Since the given volumetric flow rate and the given concentration of non-aromatic hydrocarbons C _{2+ are} practically independent of the gaseous feed used, the capacities (for example, hydraulic power) for the pre-treatment, interaction, and extraction functions can be specified, in the sense that these capacities can be almost the same for any geographical location where suitable gaseous feedstocks may be available. Advantageously, this allows the standardization of pre-processing, interaction, extraction and auxiliary functions. For example, one or more of these functions can be modulated, for example, a function can be implemented in one or more modules of standardized size, shape and / or power. Conventional separators, flavors, and recuperators can be adapted to convert gaseous feedstocks to a particular geographic location. Unlike the usual case, the pretreatment unit, flavoring and recuperator of the invention are standardized. The higher the power required in a particular geographical location, the less standardized pre-treatment plants, flavorings and recuperators can be used (for example, fewer standard modules for one or more of these functions). In places where a higher conversion power is needed, you can use more flavors, standardized pre-treatment plants and recuperators (for example, more standardized modules for one or more of these functions). In other words, one of the advantages of the invention is that it compensates for changes in the required conversion power by increasing or decreasing the power in discrete, standardized amounts, rather than requiring adaptation by continuously changing the power, as in the prior art.

[0024] The aromatization feed is reacted in the presence of a catalytically effective amount of at least one dehydrocyclization catalyst under dehydrocyclization conditions. The reaction converts at least a portion of the non-aromatic hydrocarbons of the aromatization feed to aromatic hydrocarbons and optionally other cyclic hydrocarbons. The number of reactors used to carry out the reaction is not critical, nor is the number of catalyst beds in each reactor. When the dehydrocyclization reaction is carried out in a number of reactors, a portion of the aromatic hydrocarbons produced in the first reactor can be removed before the remainder of the first reactor effluent is reacted in one or more additional reactors. In other aspects, the effluent of the first reactor may be reacted in one or more additional reactors with little or no separation of aromatic hydrocarbons from the first effluent of the reaction. The presence of aromatic hydrocarbons of the first product and methane in the second stage is less harmful for the additional production of aromatic hydrocarbons than could be expected in accordance with the recommendations of the prior art.

[0025] The preprocessing, interaction, extraction, and auxiliary functions are further explained in more detail. For the avoidance of doubt as to the meaning of certain terms used in the present description and the appended claims, the following definitions are given.

Definitions

[0026] Two or more elements are “close” to each other when they are spatially close, regardless of whether the spatial connection places one element under another, above another, or next to another element. Elements of a certain size and / or shape (for example, physical components) may be close to each other and / or close to elements that may be of an indefinite size and / or shape (for example, certain chemical interactions).

[0027] As used herein, “substantially” or other formulations indicating a degree are relative modifiers intended to indicate a tolerance for a characteristic that has been modified in this way. They are not intended to limit the absolute value or characteristic that they modify, but rather have more physical or functional characteristics than the opposite, and preferably approach or almost correspond to such physical or functional characteristics. To convert the units of measurement to normal meters ³ per day (используютNM ³ Dʺ) for gaseous composition and m ³ per day (ʺM ³ PD) for liquid composition, conversion factors from the Petroleum Engineering Handbook, Vol. VII, Indexes and Standards, 169, Society of Petroleum Engineers, 2007.

[0028] The term ʺC _n углеводород hydrocarbons means hydrocarbons with n carbon atom (s) in a molecule, where n is a positive integer. The term ʺC _{n +} ʺ means hydrocarbons containing at least n carbon atom (s) in the molecule. The term ʺC _n- углеводород hydrocarbons means hydrocarbons containing at most n carbon atoms (s) in a molecule. The term “hydrocarbons” means a class of compounds containing hydrogen bound to carbon and encompasses (i) saturated hydrocarbons, (ii) unsaturated hydrocarbons, and (iii) mixtures of hydrocarbons, including mixtures of hydrocarbon compounds (saturated and / or unsaturated), including mixtures of hydrocarbon compounds with different values of n. A “gaseous” composition is basically a vapor phase composition.

[0029] The terms "alkanes" and "paraffin hydrocarbons" mean essentially saturated compounds containing only hydrogen and carbon, for example, those containing ≤1% (molar) unsaturated carbon atoms. As an example, the term alkane encompasses C ₂ -C ₂₀ linear, iso and cycloalkanes. Aliphatic hydrocarbons mean hydrocarbons that practically do not contain hydrocarbon compounds with carbon atoms arranged in one or more rings.

[0030] The terms “unsaturated compound” and “unsaturated hydrocarbon” refer to one or more C ₂₊ hydrocarbon compounds that contain at least one carbon atom directly bonded to another carbon atom by a double or triple bond. The term "olefin" refers to one or more unsaturated hydrocarbon compounds containing at least one carbon atom directly bonded to another carbon atom by a double bond. In other words, an olefin is a compound that contains at least one pair of carbon atoms, where the first and second carbon atoms of the pair are directly linked by a double bond. The terms "aromatic compounds" and aromatic hydrocarbons mean hydrocarbon compounds containing at least one aromatic nucleus.

[0031] The term “periodic table” means a periodic system of elements, as depicted on the inside cover of The Merck Index, Twelfth Edition, Merck & Co., Inc., 1996.

[0032] The term "reaction zone" or "reactor zone" means a place inside the reactor, for example, a certain volume inside the reactor to carry out a specific reaction. The reactor or reaction step may span one or more reaction zones. More than one reaction can be carried out in a reactor, reactor stage or reaction zone. For example, a reaction step may include a first zone for conducting the first and second reactions and a second zone for carrying out a third reaction, where the first reaction (e.g., dehydrocyclization) may be the same or different from the second reaction, and the third reaction (e.g., selective oxidation) be the same or different from the second reaction.

[0033] "Dehydrocyclization" means the removal of hydrogen from a non-cyclic hydrocarbon and its cyclization to produce, for example, one or more cycloparaffins, cycloolefins and aromatic hydrocarbons. The reaction can be carried out according to one or more of the following points: (i) one step, which includes both dehydrogenation and cyclization; (ii) two steps, for example, dehydrogenation followed by cyclization of the dehydrogenated intermediate; and (iii) three or more steps, for example, dehydrogenation of normal paraffin, cyclization of an olefin intermediate and further dehydrogenation (aromatization) of a cycloolefin intermediate. Dehydrocyclization (including any dehydrogenation carried out in connection with dehydrocyclization) is “non-oxidative”, which means that the reaction is carried out practically without the oxidative combination of hydrocarbons, intermediate hydrocarbons (if any) or a dehydrocyclization product.

[0034] The term "selectivity" refers to the preparation (by weight) of a particular compound in a catalytic reaction. For example, the phrase “light hydrocarbon conversion reaction has 100% selectivity for aromatic hydrocarbons” means that 100% of the light hydrocarbons (by weight) that are converted in the reaction are converted to aromatic hydrocarbons. As applied to a particular reactant, the term "conversion" means the amount of reactant (by weight) used in the reaction. For example, when said reagent is C ₄ paraffinic hydrocarbons, 100% conversion means that 100% C ₄ paraffinic hydrocarbons are consumed in the reaction. The yield, by weight, determines the conversion product and selectivity.

[0035] The term "modularization" means the organization of technological equipment and related pipelines, valves, instrumentation, monitoring and control systems in compact and transportable units. Such a block or a plurality of them may be supported by one or more mobile platforms. Modularity effectively provides flexibility of technological parameters. Production capacity is easily increased, for example, by the addition of additional modules.

[0036] Some aspects of the invention, which include the processing of hydrocarbons of geological origin, will now be described in more detail with reference to FIG. 1. The invention is not limited to these aspects, and this description is not intended to exclude other aspects within the broader scope of the invention, such as those where gaseous feeds are obtained as a result of chemical processes and / or purification, such as one or more coking processes, catalytic cracking, catalytic reforming and / or steam cracking.

[0037]FIG. 1 schematically illustrates a systemone hundred for converting non-aromatic hydrocarbons contained in gaseous feeds into aromatic hydrocarbons. Hydrocarbon source102 provide, for example, a mixture of gaseous and liquid hydrocarbons, such as a mixture of natural gas and crude oil, extracted from the reservoir. Hydrocarbons from a hydrocarbon source are fed to a separator104 to separate at least crude natural gas and liquid, the liquid contains, for example, crude oil. Crude natural gas is diverted through the pipeline116. A fluid, which may contain entrained solids, liquids, and gases, for example entrained water, organic compounds, metals, etc., is withdrawn through a conduit 106 to a stabilization tower 108. Stabilization of a fluid involves removing various contaminants from the fluid, such as impurities water, light hydrocarbons and gases. Typically, a sufficient amount of these components is evaporated or otherwise removed to obtain a stable fluid flow.110that complies with the technical requirements for crude oil, for example, requirements related to saturated vapor pressure, water content and precipitation. In addition to the stabilized fluid flow, the steam flow is also separated from the fluid entering the stabilization tower 108. Separated steam, typically including non-aromatic hydrocarbons C_9-withdraw from the stabilization column through the pipeline112. Pipeline 116 Natural Gas Combined with Pipeline Separated112,receiving gaseous raw materials which is diverted through line 118 to the pre-treatment unit 114.

[0038] At least two streams are diverted from the pretreatment unit 114. The first stream is an aromatization feed that is discharged at a given volumetric flow rate and which has a predetermined concentration of non-aromatic C ₂₊ hydrocarbons. The second stream, which consists mainly of methane and ethane, has been found to correspond to the calorific value and Wobbe index of various natural gas pipelines. The first stream is diverted through conduit 120 (for example, as condensed liquids and / or steam) to flavor 124 . The second stream is diverted through conduit 115 , for example, to one or more storages; for transportation, for example, by pipeline; for further processing, for example, cooling / compression to produce compressed gas and / or LNG; re-injection into the source of hydrocarbons 102 , etc. In addition to the first and second streams, typically a third liquid stream containing hydrocarbons is optionally obtained and diverted from the pretreatment unit via line 122. The third stream can be generated in those aspects where the pretreatment unit includes at least one fractionation column ( for example, one or more distillation columns), wherein the third stream may be, for example, a bottoms stream of a fractionation column.

[0039] The aromatization feed is diverted via conduit 120 to flavor 124, to convert at least a portion of the non-aromatic C ₂₊ hydrocarbon feed to aromatic hydrocarbons. The effluent reaction stream is led off through line 126. The effluent reaction stream includes (i) non-aromatic hydrocarbons and (ii) aromatic hydrocarbons and molecular hydrogen, at least a portion of the aromatic hydrocarbons and molecular hydrogen (and, optionally, a portion of the non-aromatic hydrocarbons) are formed during dehydrocyclization .

[0040] The effluent reaction stream is diverted via conduit 126 to a recuperator 128 to recover at least a portion of the first product and tail gas from the effluent reaction stream. The first product includes at least a portion of the aromatic hydrocarbons from the effluent. Tail gas comprises (i) at least a portion of the molecular hydrogen of the effluent reaction stream and (ii) at least a portion of the non-aromatic hydrocarbons of the effluent reaction stream. Tail gas, which typically includes mainly molecular hydrogen, methane and ethane, is vented through conduit 130, for example, to one or more storage facilities; for transportation, such as by pipeline; for further processing, for example, cooling / compression to produce compressed gas and / or LNG; re-injection into the source of hydrocarbons 102 , etc. Since it has been established that tail gas corresponds to the calorific value and Wobbe index of various natural gas pipelines, in some aspects (not shown) at least a portion of the tail gas is combined with at least some of the second stream of the pre-treatment unit, for example, for transporting combined flow through the pipeline. In aspects schematically illustrated in FIG. 1, tail gas is vented through conduit 130 to hydrogen separator 142.

[0041] As shown in FIG. 1, a first product including recovered aromatic hydrocarbons is discharged through conduit 138. Valves denoted by V1 and V2 can be used to (i) divert all or part of the first product from the process for storage, transportation, or further processing and / or (ii) combining all or part of the first product with a liquid stream 110.

[0042] The tail gas, in addition to molecular hydrogen, methane and ethane, may also contain sulfur compounds such as mercaptan and / or hydrogen sulfide. As a rule, this is the case when the raw material for aromatization includes organosulfur compounds, H ₂ S, or both. If necessary, tail gas can be treated (not shown in FIG. 1) to remove at least some of the sulfur compounds that may be present. Sulfur removal can be performed before the introduction of tail gas into the pipeline 130 and / or after it. Conventional sulfur removal technology may be used, but the invention is not limited to this. Although sulfur removal is optional, it is usually done when tail gas includes ≥ 4 ppm H ₂ S, for example ≥ 1000 ppm, such as ≥ 1%, especially in aspects where at least some of the tail gas is subjected to additional dehydrocyclization.

[0043] All or some of the tail gas can be used to generate electricity on site, for example, by burning tail gas in one or more gas turbines, the gas turbine is used to power one or more electric generators of the power unit 140 . Electricity can be used in the process of, for example, heating the flavor 124 , for example, to provide heat for a (usually endothermic) dehydrocyclization reaction. Alternatively, or additionally, the generated electricity can be transmitted to one or more electrical networks. Alternatively or in addition to a gas turbine, one or more internal combustion engines can be used to burn tail gas, for example, to power an electric generator.

[0044] In the specific aspects illustrated in FIG. 1, molecular hydrogen is separated from the tail gas prior to entering the power unit 140, with at least some of the separated molecular hydrogen being discharged through a pipe 144 to the power unit 140. Typically, the separated molecular hydrogen is transferred to the power unit 140 together with other components of the tail gas such as methane and / or ethane. Such a combined stream is called a hydrogen-rich stream because it has a higher molecular hydrogen content than the tail gas from which it was separated. After separation of molecular hydrogen, the tail gas residue is a hydrogen depleted stream. The transfer of the hydrogen-rich stream from the hydrogen separator 142 to the power unit 140 may be desirable, for example, when the tail gas has too low BTU content to efficiently generate electricity, or if less electricity is desired. Conventional hydrogen separation technology may be used in hydrogen separator 142, but the invention is not limited to this. Typically, the hydrogen enriched stream contains ≥ 50 wt. % molecular hydrogen tail gas, for example, ≥ 75 wt. %, such as ≥ 90 wt. % Typically, a hydrogen depleted stream contains ≥ 25 wt. % non-aromatic hydrocarbons of the tail gas, for example, ≥ 50 wt. %

[0045] Alternatively or in addition to the electricity that can be obtained by burning tail gas (or by burning a tail gas component such as molecular hydrogen) in power unit 140 , the thermal energy generated by such combustion, or part of it, can be transferred to flavor 124 through heat channel 146 . Thus, it is possible, for example, to provide heating for dehydrocyclization. For this purpose, one or more recovery steam generators may be used, including those typically adapted to service gas turbines. The generated steam can be used as a coolant in the heat channel 146 , for example, to heat the flavor 124 . Alternatively or additionally, a portion of the separated molecular hydrogen of conduit 144 , or some portion of the tail gas of conduit 130, or both, may be used to heat the flavor 124 . For example, one or more of these streams can be burned in a heater or chemically reacted in an exothermic reaction to generate heat for the aromatization reaction (or reactor (s)).

[0046] Continuing with respect to FIG. 1, the hydrogen-depleted stream can be transferred via line 150 to an optional compressor 152. All or some of the compressed hydrogen-depleted stream can be used in a method, for example, to recycle to a pre-treatment unit 114 via lines 154 and 112. This may be desirable when the compressed hydrogen depleted stream contains a sufficient amount of non-aromatic C ₂₊ hydrocarbons, which may be useful to increase the yield of aromatic hydrocarbons in the process. Alternatively or in addition, the valves designated V ₃ and V ₄ can be used to divert all or part of the hydrogen depleted stream from the process, for example, for one or more storage facilities, further processing, transportation, sale, re-injection into the reservoir collector etc.

[0047] The block diagram of FIG. 1 should not be interpreted as indicating that system 100 includes all components shown. Some related aspects within the broader scope of the invention are not shown in FIG. 1. For example, although it is generally undesirable to do so, within the scope of the invention, burn all or part of the pipeline flows 150 or 154 on one or more flares, for example, during technological disruptions. Other aspects that are not shown in FIG. 1 do not include an optional pretreatment unit. These aspects, as a rule, use a source of hydrocarbons, which is able to provide raw materials for aromatization with a predetermined volumetric flow rate meeting the technical requirements and a predetermined concentration of non-aromatic C ₂₊ hydrocarbons meeting the technical requirements.

[0048] Part or all of the system 100 can be built as a modular system that can be quickly moved and installed in place, used, then disassembled and moved to a new place when it is no longer needed.

[0049] Some gaseous feedstocks suitable for the production of technical specifications for flavoring feedstocks will now be described in more detail. The invention is not limited to this gaseous feedstock, and this description is not intended to preclude the use of other gaseous feedstocks within the broader scope of the invention.

Gaseous feed

[0050] A gaseous feed, which in the aspects illustrated in FIG. 1, typically introduced into the pretreatment unit 114 via line 118, typically includes one or more non-aromatic C ₂ -C ₉ hydrocarbon compounds, for example, one or more light hydrocarbon compounds (i.e., C ₂ -C ₅ ) such as one or more light paraffinic hydrocarbon compounds. For example, gaseous feedstock 118 may include ≥ 1 wt. %, based on the mass of gaseous feedstock 118, one or more (i) paraffinic C ₂ -C ₉ hydrocarbons, (ii) aliphatic C ₂ -C ₉ hydrocarbons, (iii) aliphatic paraffinic C ₂ -C ₉ hydrocarbons, (iv) light paraffinic hydrocarbons, (v) light aliphatic hydrocarbons and (vi) light aliphatic paraffin hydrocarbons; for example, ≥ 10 wt. %, or ≥ 25 wt. %, or ≥ 50 wt. %, or ≥ 75 wt. %, or ≥ 90 wt. %, or ≥ 95 wt. % Optionally, the gaseous feed further comprises a diluent. The diluent, when present, is usually found in gaseous feedstocks in an amount of ≤ 60 wt. %, based on the mass of gaseous feedstock 118 , for example, ≤ 50 wt. %, such as ≤ 40 wt. % or ≤ 30 wt. %, or ≤ 20 wt. %, or ≤ 10 wt. % A feed component is a diluent when it is practically non-reactive under specific reaction conditions in the presence of a specific dehydrocyclization catalyst, for example methane, molecular nitrogen and inert atomic gases such as argon.

[0051] The gaseous feed generally contains C ₃ and / or C ₄ hydrocarbons, for example, (i) ≥ 10 wt. % propane, such as ≥ 40 wt. % or ≥ 60 wt. %, and / or (ii) ≥ 5 wt. % butanes, such as ≥ 40 wt. % or ≥ 60 wt. % Although the gaseous feed may contain C ₅₊ hydrocarbons, the amount of C ₅₊ hydrocarbons, when present, is small, for example, ≤ 30 wt. %, such as ≤ 10 wt. % or ≤ 01 wt. % Typically, gaseous feed contains ≤ 15 wt. % saturated hydrocarbons With ₆₊ , for example, ≤ 5 wt. %

[0052] The gaseous feed may contain methane, for example, ≥ 1 wt. % methane, such as ≥ 10 weight. % or ≥ 20 wt. %, or ≥ 60 wt. % Despite the fact that methane is a diluent, i.e. it usually does not react with the production of aromatic hydrocarbon or catalytic coke in the presence of the specified dehydrocyclization catalyst under given reaction conditions; its presence is useful. It is believed that this benefit is due to at least a partial decrease in the partial pressure of C ₂ -C ₉ hydrocarbon compounds, which is achieved when the gaseous feedstock further comprises methane. It has been established that a decrease in the partial pressure of C ₂ -C ₉ hydrocarbons, in particular the partial pressure of C ₂ -C ₅ hydrocarbons, reduces the amount of catalytic coke formed under the conditions of this dehydrocyclization process. Typically, gaseous feedstocks include a total of ≤ 10 wt.% Impurities such as CO, CO ₂ , H ₂ S, and a total of mercaptan; for example ≤ 1 wt. % or ≤ 0.1 wt. % Although not typically found in gaseous feedstocks extracted from crude natural gas, some gaseous feedstocks include molecular hydrogen, for example, ≥ 1 wt. % molecular hydrogen based on the mass of gaseous feedstock 118, such as ≥ 5 wt. %

[0053] The gaseous feed includes ethane, typically ≥ 1 wt. %, based on the weight of the gaseous feed, for example, ≥ 5 wt. % or ≥ 10 wt. %, such as in the range of 10 wt. % up to 40 wt. % Suitable gaseous feedstocks include gaseous feedstocks containing a significant amount of ethane, namely> 50 wt. %, such as ≥ 75 wt. % or ≥ 90 wt. %, or ≥ 95 wt. % One of the options for a typical gaseous feedstock includes (i) ≥ 10 wt. % ethane, such as in the range of 10 wt. % up to 40 wt. %; and also includes (ii) from 1 wt. % up to 40 wt. % methane, (iii) from 20 wt. % up to 50 wt. % propane and (iv) from 20 wt. % up to 50 wt. % butanes. In other aspects, the amount of ethane in the gaseous feed is <1 wt. %, for example, ≤ 0.1 wt. % or ≤ 0.1 wt. %

[0054] Although the gaseous feed may contain unsaturated C ₂₊ hydrocarbons, such as unsaturated C ₂ -C ₅ hydrocarbons, the amount of these unsaturated compounds is usually ≤ 20 wt. %, for example, ≤ 10 wt. %, such as ≤ 1 wt. % or ≤ 0.1 wt. %, or in the range from 0.1 wt. % up to 10 wt. % The gaseous feed 118 may be substantially non-aliphatic hydrocarbon free. In particular, the gaseous feedstock 118 may be substantially free of aromatic hydrocarbons, where substantially free in this context means <1 wt. % based on the mass of gaseous feedstock 118 , such as ≤ 0.1 wt. % or ≤ 0.01 wt. %, or ≤ 0.001 wt. %

[0055] In some aspects, the source of the gaseous feed comprises natural gas, for example, raw natural gas (“raw gas”). Natural gas is (i) a mixture comprising hydrocarbons, (ii) mainly in the vapor phase at a temperature of 15.6 ° C and pressure 1.016 bar (absolute), and (iii) withdrawn from the geological formation. Natural gas can be obtained, for example, from one or more oil fields, coal deposits and shale deposits. One suitable crude natural gas comprises from 3 mol.% To 70 mol.% methane, from 10 mol .% And 50 mol.% Ethane, 10 mol.% To 40 mol.% Propane and 5 mol.% To 40 mol.% Of butanes, and from 1 mol.% To 10 mol.% Of total hydrocarbon FROM₅-FROM₉. In some aspects, ≥ 50 wt. % of the gaseous feed includes natural gas, such as raw natural gas, for example, ≥ 75 wt. % or ≥ 90 wt. %, or ≥ 95 wt. %

[0056] As disclosed with respect to FIG. 1, all or part of the gaseous feed can be a stream exiting a gas-liquid separator 104 that separates the gaseous overhead stream (into conduit 116 ) from the hydrocarbon source 102 . Any form of hydrocarbon source can be used as feedstock 102 , as long as it contains a sufficient amount of non-aromatic C ₂₊ hydrocarbons to produce a gaseous feedstock and a feedstock for flavoring. For example, a hydrocarbon source 102 may include hydrocarbons obtained from field processing plants, refineries, chemical plants, and the like.

[0057] In some aspects, the gaseous feedstock includes one or more gases from (i) gas obtained from a natural gas well (“gas well”, unassociated or “dry” gas), (ii) natural gas obtained from a condensate well (“condensate well gas”) and (iii) gas from the borehole (“wet” or “associated” gas). For example, associated gas may contain from 0 mol. % to 95 mol. % methane, from 5 mol. % up to 50 mol. % ethane, from 2 mol. % to 40 mol.% propane, from 0.1 mol. % up to 30 mol. % iso-butane, from 1 mol. % up to 30 mol. % n-butane and from 0.05 mol. % up to 25 mol. % iso-pentane. Table 1 includes typical ranges of the composition of the raw gas (mol%) and, in parentheses, the typical average composition (mol%) of some raw gases that may occur in these streams.

Table 1: Examples of Raw Material Compositions

Associated gas Dry gas Condensate well gas Component CO ₂ 0-50 (0.63) 0-25 (0) 0-25 (0) N ₂ 0-50 (3.73) 0-25 (1.25) 0-25 (0.53) H ₂ s 0-5 (0.57) 0-5 (0) 0-5 (0) CH ₄ 0-80 (64.48) 0-97 (91.01) 0-98 (94.87) C ₂ H ₆ 5-20 (11.98) 2-10 (4.88) 1-5 (2.89) C ₃ H ₈ 2-10 (8.75) 0.5-5 (1.69) 0.1-5 (0.92) iso-butane 0.1-5 (0.93) 0.05-1 (0.14) 0.1-5 (0.31) n-butane 1-5 (2.91) 0.05-2 (0.52) 0.05-2 (0.22) iso-pentane 0.05-2 (0.54) 0.01-1 (0.09) 0.1-1 (0.09)

[0058] One suitable gaseous feedstock includes ≥ 75 wt. % associated gas, based on the mass of the gaseous feed, for example, ≥ 90 wt. % or ≥ 95 wt. % In the extraction of natural oil, the absence of effective means for normal transportation, for example, the lack of means for liquefying natural gas and / or main pipelines, can cause associated gas to remain inside or near the reservoir. The remaining associated gas may be uneconomical for transportation and therefore flared, which is undesirable. In addition, even in places where there are main pipelines, associated gas can be excluded from the pipeline, since it usually exceeds one or more technical requirements for pipeline transportation, for example, ≤ 12 wt. % ethane, ≤ 5 wt. % propane, ≤ 2 wt. % butanes, the Wobbe index from 49.01 MJ / cm ³ to 52.22 MJ / cm ³ ) and the calorific value from 36.07 MJ / cm ³ to 41.40 MJ / cm ³ ).

[0059] Since methane is not detrimental to the process and is useful in at least some aspects, the methods described herein can at least partially eliminate the need for costly and inefficient cryogenic plants for methane separation. Typically, the production of gaseous feedstocks from a hydrocarbon source does not include (i) bringing the gaseous feedstock, hydrocarbon source, or any intermediate thereof to a temperature of ≤ -37 ° C, for example ≤ -46 ° C, such as ≤ -60 ° C. Some aspects of the methods do not include cryogenic treatment, for example, cryogenic separation of methane is not used.

[0060] Therefore, the methods are particularly advantageous for remote and economically backward areas where (i) the absence of cryogenic plants for methane separation limits the usefulness of conventional methods for aromatizing natural gas, (ii) the absence of pipelines or natural gas production infrastructure can lead to significant amounts of light hydrocarbons are flared or combusted as fuel, and (iii) the residual associated gas remains unused in a remote area due to shortages pipeline capacity or breach of one or more of the requirements for in-service main pipeline.

Gaseous feed pretreatment

[0061] The gaseous feedstock is pretreated to obtain a suitable flavoring feedstock. Some forms of pretreatment schematically illustrated in FIG. 2 will now be described in more detail. The invention is not limited to these aspects, and this description is not intended to exclude other forms of pretreatment within the broader scope of the invention, for example, those using fractional distillation. The design features shown in FIG. 2, which essentially perform the same function as in FIG. 1 are indicated under the same number.

[0062] The pretreatment unit 114 uses the gaseous feedstock supplied through line 118 to produce a flavoring feed intended for flavoring 124 in a predetermined amount of ≥ 1 MSCFD (29,000 NM ³ D), for example, ≥ 10 MSCFD (290,000 NM ³ D), such as, for example, in the range from 2 MSCFD (57.000 NM ³ D) to 80 MSCFD (2.3⋅10 ⁶ NM ³ D) or in the range from 10 MSCFD (290.000 NM ³ D) to 30 MSCFD (860.000 NM ³ D ) The aromatization feed has a predetermined concentration of non-aromatic C _{2 +} hydrocarbons in the range of 15 mol. % to 90 mol. % per mole of raw materials for flavoring, as a rule, from 55 mol. % to 80 mol. % In most cases, the gaseous feed is supplied to the pretreatment unit in an amount ranging from 10 MSCFD (290.000 NM ³ D) to 150 MSCFD (4.3⋅10 ⁶ NM ³ D).

[0063] In the aspects illustrated in FIG. 2, the pre-treatment includes (i) transferring heat from the gaseous feedstock, at least in the first heat transfer step 201, to partially condense the gaseous feedstock at the set dew point. The upper stream is separated from the partially condensed gaseous feed in the drum separator 203, the upper stream is discharged through a pipe 115 . The upper stream usually has characteristics within the limits established by the technical requirements for many gas pipelines, for example, ≤ 12 wt. % ethane, ≤ 5 wt. % propane, ≤ 2 wt. % butanes, and the Wobbe index is from 49.01 MJ / cm ³ to 52.22 MJ / cm ³ ) and the calorific value is from 36.07 MJ / cm ³ to 41.40 MJ / cm ³ ).

[0064] The aromatization feed is withdrawn from the separator drum 203 in the form of a bottoms stream through a conduit 120. In order to provide the aromatization feed to the aroma 124 with a predetermined concentration of non-aromatic hydrocarbons C ₂₊ and a predetermined volumetric rate expiration, separation in the drum 203 is usually carried out at a given dew point in the range from -40 ° F (-40 ° C) to -20 ° F (-28.9 ° C), typically at an absolute pressure in the range of about 300 psi inch (2068 kPa) up to 2000 psi inch (13.800 kPa), e.g. from 400 psi inch (2760 kPa) up to 700 psi inch (4830 kPa).

[0065] In the aspects shown in FIG. 2, the first heat transfer step includes at least one heat exchanger in indirect thermal contact with a stream of refrigerant circulating in channel 204 . For example, the refrigerant may include propane. The second heat transfer stage 202 is used to remove heat from the refrigerant. Typically, the second heat transfer stage includes at least one heat exchanger and at least one refrigerant compressor. The refrigerant compressor may be powered by shaft power received from the gas turbine of power unit 140 (shown in FIG. 1), for example. Typically, the heat exchanger of the heat transfer step 202 indirectly transfers heat from the refrigerant to at least one heat transfer medium (not shown) that is removed.

[0066] In some aspects not shown in FIG. 2, the first heat transfer stage 201 indirectly transfers the first, second, and third part of the total amount of heat from the gaseous feed. The first part is transferred to the refrigerant, and the second part is transferred to the separated gaseous hydrocarbon. The third part is transferred to the bottoms stream of the separator drum 203, which at least partially vaporizes the bottoms stream, providing raw materials for flavoring the pipe 120 in gaseous form. These aspects are particularly effective. Energy is needed (usually provided at least by the refrigerant compressor) to remove heat from the circulating refrigerant in the second heat transfer stage. Most of this energy is spent on condensing the gaseous feedstock into the bottoms stream, but some energy is spent on cooling the gaseous overhead stream, which is removed through line 115. Indirectly transferred heat from the gaseous overhead stream to the refrigerant in the first heat transfer stage reduces the refrigerant compressor power requirements. It has been successfully established that such a configuration makes it possible to reduce or eliminate changes in the refrigerant compressor power requirements that would otherwise be considered during operation of the method using the first gaseous source from the first geographical area and the second gaseous source in the second geographical area, the first and second gaseous sources having a different molar ratio (ʺMʺ) of C ₂ hydrocarbons to non-aromatic C ₃₊ hydrocarbons. For gaseous feed materials with a higher M, the amount of refrigerant compressor power used to cool the gaseous overhead stream is comparatively greater than for the gaseous overhead stream with a lower M. Therefore, the heat transfer in the first stage of heat transfer from the gaseous overhead stream to the refrigerant automatically compensates for the compressor power refrigerant that would otherwise be needed or unnecessary due to a change in the gaseous feed.

[0067] When operating in accordance with these aspects, the amount of heat transferred from the gaseous feed to the refrigerant in the first heat transfer step usually changes by no more than +/- 50% (essentially equivalent to changing the compressor power by no more than + / - 50%) with a very wide variety of gaseous feedstocks obtained from a very wide range of hydrocarbon sources. Typically, this change is achieved when the condensed portion of the gaseous feed has a volume change of ≤ +/- 25%, for example, ≤ +/- 10% or +/- 5%. A further reduction of this change, for example, by no more than +/- 40%, such as no more than +/- 30%, can be achieved provided that (A) the specified amount of raw material for flavoring is in the range of 5 MSCFD (140.000 NM ³ D) to 35 MSCFD (1⋅10 ⁶ NM ³ D), (B) the amount of gaseous feed is in the range from 30 MSCFD (860.000 NM ³ D) to 140 MSCFD (4⋅10 ⁶ NM ³ D) and (C) the gaseous feed has a molar ratio of C ₂ hydrocarbons to non-aromatic C ₃₊ hydrocarbons in the range of 0.5 to 1.5. This effect provides a significant advantage in standardizing the cooling power of the pre-treatment plant, for example, since modularization is easier to achieve when the cooling power of the pre-treatment plant is practically independent of the choice of gaseous feed.

[0068] In some aspects, pre-processing is provided to advantageously reduce fluctuations in the capacity of the recuperator, as otherwise might be required when switching from the first gaseous feed to the second gaseous feed. By dehydrocyclization of C ₂ hydrocarbons, the aromatization raw materials produce less aromatic hydrocarbon compounds (by molarity) than by dehydrocyclization of non-aromatic C ₃₊ aromatics. Therefore, for a gaseous feedstock having a higher M value, a relatively less power recuperator will be needed than for a lower M value. It has been found that this effect can be at least partially mitigated by using pretreatment to produce feed for aromatization having a given concentration of C ₃₊ hydrocarbons within a relatively narrow range. For example, it has been found that when the aromatization feed comprises C ₂ hydrocarbons in an amount corresponding to a range of 10 moles. % to 35 mol. % and C ₃₊ hydrocarbons in an amount corresponding to a range of 40 mol. % up to 50 mol. %, this first product is recovered in an amount that changes by no more than +/- 25%. Further reduction of this change can be achieved, for example, by no more than +/- 20%, such as no more than +/- 15%, when (A) the specified amount of raw material for flavoring is in the range of 5 MSCFD (140,000 NM ³ D) up to 35 MSCFD (1⋅10 ⁶ NM ³ D), (B) the amount of gaseous feedstock ranges from 30 MSCFD (860.000 NM ³ D) to 140 MSCFD (4⋅10 ⁶ NM ³ D) and (C) the gaseous feed has a molar ratio of C ₂ hydrocarbons to non-aromatic C ₃₊ hydrocarbons in the range of 0.5 to 1.5. This effect gives a significant advantage, for example, in standardizing the capacity of the recuperator, since modularization is easier to achieve when the capacity of the recuperator is practically independent of the choice of gaseous feed.

[0069] The installation for pre-processing of raw materials 114 can be built on a primary set of modules, where the separation of blocks into separate modules is based on the amount of flow. For example, drum separator 203 and support devices can be designed as a single unit for small-scale applications, for example, less than about 10 million standard cubic feet per day (MSCFD) (290,000 NM ³ D), about less than 5 MSCFD (140,000 NM ³ D ) or approximately less than 1 MSCFD (29.000 NM ³ D).

[0070] For larger applications, for example, about more than 20 MSCFD (570,000 NM ³ D), about more than 35 MSCFD (1 NM10 ⁶ NM ³ D), or about more than 50 MSCFD (1,4⋅10 ⁶ NM ³ D) , a first heat transfer step 201, a second heat transfer step 202, and a drum separator 203 can be divided into two or more modules. A typical unit installation has a gross mass of ≤ 200,000 pounds (91,000 kg), for example, ≤ 80,000 pounds (36,300 kg), such as in the range from 10,000 pounds (4,500 kg) to 200,000 pounds (91,000 kg) or from 50,000 pounds (22,700 kg ) up to 150,000 pounds (68,000 kg); width ≤ 15 feet (4.6 m), for example, ≤ 102 inches (2.6 m); and a height of ≤ 15 feet (4.6 m), for example, ≤ 14.6 feet (4.45 m). It will be apparent to those skilled in the art that these designs are just one example, and any number of different designs and systems, like these, can be used for a plant for pre-treatment of raw materials 114 .

Flavoring

[0071] Some forms of flavor 124 will now be described in more detail with reference to FIG. 2. The invention is not limited to these forms of flavoring, and this description is not intended to exclude other forms of flavoring within the broader scope of the invention. Flavor 124 may be contained in a single module or, alternatively, be distributed across multiple modules, for example, to house reactors and related equipment. In the aspects shown in FIG. 2, one reactor 224 is used . The feedstock for flavoring the pipe 120 may be heated to the reactor temperature while passing through a series of heat exchangers, 226 , 228 , 236 , for example, for heat exchange with the exit reaction stream 126. To achieve the final temperatures used to maintain the endothermic dehydrocyclization reaction, the feedstock for the reactor in conduit 120 can pass through the heater 230. The heater 230 can receive heat from the hot flue gas discharged through the conduit 146, such as a hot flue gas generated by combustion in turbomash nach unit 140. Alternatively, or in addition to (e.g., during start-up) the heater may be one of the devices, which utilizes combustion. Some of the tail gas in conduit 130 can be fed into the heater to provide fuel for combustion, at least until the start is advanced enough to transfer the hot flue gas from power unit 140 to the heater through conduit 146. Heater 230 may also be used to heat the circulation stream 232 from the reactor 224 to maintain the temperature of the reactor. It is understood that the flame heater 230 may include one or more heat recovery units (HRUs) associated with flue gas from a power unit 140, for example, one or more steam recovery generators. An optional hot oil stream may be circulated through line 234 to provide heat to the components of system 100 and / or other systems.

[0072] Reactor 224 uses a dehydrocyclization process to convert at least some of the non-aromatic C ₂₊ aromatics feed aromatics to aromatics. In this form of dehydrocyclization, the aromatization feed obtained from conduit 120 is heated in heat transfer steps 226 and 228 by transferring heat from the effluent from the reaction stream in conduit 126 to the aromatization feed. The heated aromatization feed is reacted in the presence of a catalytically effective amount of at least one dehydrocyclization catalyst located in at least one reaction zone operating under catalytic dehydrocyclization conditions. The reaction converts at least some of the non-aromatic hydrocarbons C ₂ -C ₉ feedstock for aromatization into aromatic hydrocarbons and molecular hydrogen. Typically, the dehydrocyclization catalyst includes ≥ 10 wt. % molecular sieve component and ≥ 0.005 wt. % dehydrogenating component.

[0073] When the molecular sieve component and the dehydrogenating component together include less than 100 wt. % catalyst, ≥ 90 wt. % of the catalyst residue may include a matrix component, such as ≥ 99 wt. % residue. The catalyst typically includes a molecular sieve component in an amount of ≥ 20 wt. %, based on the weight of the catalyst, for example, ≥ 25 wt. %, such as in the range of 30 wt. % to 99.9 wt. % In some aspects, the molecular sieve component includes aluminosilicate, for example, ≥ 90 wt. % of at least one aluminosilicate. The aluminosilicate may be an unsubstituted aluminosilicate, a substituted aluminosilicate, or a combination thereof. For example, aluminosilicate may be in the form where at least a portion of its natural metal is replaced, for example, by ion exchange, with another suitable metal (usually a metal cation) of groups 1-13 of the periodic table. Typically, aluminosilicate includes zeolite aluminosilicate, for example, ≥ 90 wt. % of at least one zeolite based on the weight of aluminosilicate. The term zeolite includes those zeolites in which at least part of the aluminum is replaced by another trivalent metal, such as gallium or indium.

[0074] The molecular sieve component typically includes ≥ 90 wt. % of one or more of these molecular sieves, for example, ≥ 95 wt. % In some aspects, the molecular sieve component includes at least one of the zeolite molecular sieves, for example, ≥ 90 wt. % zeolite, such as ≥ 95 wt. %, based on the weight of the molecular sieve component. Although the molecular sieve component may practically consist or even consist of zeolite, in alternative aspects, the zeolite (s) are present in the molecular sieve component in combination with another (e.g. non-zeolite) molecular sieve. The zeolite can be a zeolite in hydrogen form, for example, can be synthesized in the form of an alkali metal zeolite, but then converted from an alkaline form to a hydrogen form. Typically, zeolite has an average pore size and a permeability index of 2-12 (as defined in US patent No. 4.016.218). Examples of suitable zeolites include ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-48, including both mixtures thereof and intermediates such as ZSM-5 / ZSM- eleven. Optionally, the zeolite includes at least one plurality of pores of substantially the same size passing through a molecular sieve, where the average geometric cross-sectional dimensions of each of the plurality of pores are> 5 Å or> 5.3 Å, for example, ≥ 5.4 Å, such as ≥ 5 , 5 Å, or in the range from 5 Å to 7 Å, or from 5.4 Å to 7 Å. Suitable are ZSM-5 and / or ZSM-12, in particular H-ZSM-5. For example, the molecular sieve component may include ≥ 90 wt. % (A) ZSM-5 and / or (B) ZSM-12, based on the weight of the molecular sieve component, for example, ≥ 95 wt. % H-ZSM-5. In some aspects, the molecular sieve has a relatively small crystal size, for example, fine crystalline ZSM-5, implying ZSM-5 with a crystal size of ≤ 0.05 micrometers (μm), such as for example in the range from 0.02 μm to 0.05 μm. Fine crystalline ZSM-5 and a method for determining the crystal size of a molecular sieve are described in US patent No. 6.670.517, which is fully incorporated herein by reference.

[0075] In other aspects, the molecular sieve component comprises at least one molecular sieve of the MCM-22 family, for example, MCM-22 alone or in combination with other molecular sieves, such as one or more of these zeolites. The MCM-22 family includes those molecular sieves that have a diffraction pattern that includes the maximum values of interplanar spacing d at 12.4 ± 0.25, 6.9 ± 0.15, 3.57 ± 0.07, and 3.42 ± 0, 07 Angstrom. The X-ray diffraction data used to characterize the material was obtained by standard methods using a K-alpha copper doublet as incident radiation and a diffractometer equipped with a scintillation counter and connected to a computer as a collection system. Examples of suitable molecular sieves of the MCM-22 family are described in US Pat. Nos. 4,954,325; 4,439,409 (PSH-3); 4.826.667 (SSZ-25); 6.077.498 (ITQ-1); 5.250.277 (MCM-36); 5.236.575 (MCM-49) and 5.362.697 (MCM-56). Others include UZM-8, ERB-1, as described in European Patent No. 0293032, and ITQ-2, as described in International Patent Publication No. WO 97/17290. Mixtures of molecular sieves of the MCM-22 family can also be used.

[0076] When the molecular sieve component includes at least one aluminosilicate, the silica: alumina ratio in the aluminosilicate (essentially the same as the atomic ratio Si: Al ₂ ) is usually ≥ 2, for example in the range of 5 to 100. Silica ratio: alumina is intended to display the atomic ratio of Si: Al ₂ in a rigid anionic framework of crystalline aluminosilicate. Alternatively or in addition, the catalyst can be made more resistant to deactivation (and increase the yield of aromatic hydrocarbons) by incorporating trivalent phosphorus into the molecular sieve component. When used, the amount of trivalent phosphorus is usually ≥ 1 wt. % based on the weight of the molecular sieve component. For example, when the molecular sieve component includes aluminosilicate, the atomic ratio of trivalent phosphorus: aluminum can be in the range from 0.01 to 1. Zeolite with a higher ratio of silica: alumina can be used when lower catalyst acidity is desired, for example, in the range of 44 to 100, such as from 50 to 80 or 55-75.

[0077] In addition to the molecular sieve component, the catalyst includes ≥ 0.005 wt. %, based on the weight of the catalyst, the dehydrogenating component, for example, at least one dehydrogenating metal. The dehydrating component may include one or more neutral metals selected from groups 3-13 of the periodic table, such as one or more of Ga, In, Zn, Cu, Re, Mo, W, La, Fe, Ag, Pt and Pd, and / or one or more of the oxides, sulfides and / or carbides of these metals. For example, the dehydrogenating component may be Ga, Zn, or a combination thereof, optionally supported on a catalyst comprising ZSM-5 as a molecular sieve component.

[0078] Typically, the dehydrogenating component includes ≥ 90 wt. % of one or more of these dehydrogenating metals and / or their oxides, for example, ≥ 95 wt. % or ≥ 99 wt. % For example, the dehydrogenating component may include ≥ 90 wt. % (A) Ga and / or (B) Zn, including their oxides. Typically, the catalyst includes ≥ 0.01 wt. % dehydrogenation component, calculated on the weight of the catalyst, for example, ≥ 0.1 wt. % dehydrogenating component, such as ≥ 0.5 wt. % or ≥ 1 wt. %

[0079] It will be apparent to those skilled in the art that when a dehydrogenating component comprises one or more metals of greater catalytic dehydrogenating activity, for example, Pt and / or Pd, a smaller amount of the dehydrogenating component, for example, in the range of 0.005 wt. % to 0.1 wt. %, based on the weight of the catalyst, such as from 0.01 wt. % to 0.6 wt. % or from 0.01 wt. % to 0.05 wt. % When the dehydrogenating component includes one or more metals of lower dehydrogenating activity, for example, one or more Ga, In, Zn, Cu, Re, Mo and W, a larger amount of the dehydrogenating component will be required, for example, in the range from 0.05 wt. % up to 10 wt. %, based on the weight of the catalyst, such as from 0.1 wt. % up to 5 wt. % or from 0.5 wt. % to 2 wt. %

[0080] The dehydrating component may be provided on the catalyst, in or near the catalyst by any method, for example, by conventional methods such as impregnation or ion exchange. At least a portion of the dehydrogenating metal may also be present in the crystalline framework of the molecular sieve. Regarding a typical catalyst, (i) the dehydrogenating component comprises ≥ 95 wt. % (A) Ga and / or (B) Zn, and (ii) the first molecular sieve component comprises ≥ 95 wt. % H-ZSM-5.

[0081] In some aspects, the dehydrogenating component comprises ≥ 99 wt. % of one or more of Ga, Zn, and In, and the molecular sieve component includes ≥ 99 wt. % ZSM-5 zeolite, which is impregnated with a dehydrogenating metal component and / or ion-exchanged with a dehydrogenating metal component. For example, the catalyst may include Ga-filled and / or In-filled H-ZSM-5, Ga-substituted and / or In-substituted H-ZSM-5, H-gallosilicate of the ZSM-5 type and H-galloaluminosilicate of the ZSM type -5. Optionally, the catalyst includes (i) tetrahedral aluminum and / or gallium, which is present in the zeolite framework or lattice, and / or (ii) octahedral gallium or indium, which is not present in the zeolite framework, but is present in zeolite channels in the vicinity of protic acid zeolite centers. Without reference to any theory or model, it is believed that tetrahedral or scaffold Al and / or Ga determines the acid function of the catalyst; octahedral or non-scaffold Ga and / or In determines the dehydrogenation function of the catalyst. Although the zeolite is usually impregnated or ion-exchanged with a dehydrogenating metal, other forms of the zeolite can be used, such as an H-galloaluminosilicate of the ZSM-5 type, with framework (tetrahedral) atomic ratios of Si / Al and Si / Ga from about 10: 1 to 100: 1 and from 15: 1 to 150: 1, respectively, and non-frame (octahedral) Ga from about 0.5 wt. % to 0 wt. %

[0082] In addition to the molecular sieve component and the dehydrogenating component, the catalyst may further include an optional matrix component, for example, one or more inorganic binders. The amount of matrix component is not critical. If present, the amount of the matrix component is usually in the range of 0.01 times the mass of the molecular sieve component to about 0.9 times the weight of the molecular sieve component, for example, in the range of 0.02 to 0.8. The matrix component may include active materials, such as zeolites of synthetic or natural origin. Alternatively, or in addition, the matrix component may include clays and / or oxides such as alumina, silica, silica-alumina, zirconia, titanium dioxide, magnesium oxide, or mixtures thereof, and other oxides. The matrix component may include materials of natural origin and / or materials in the form of gelatin precipitates or gels, including mixtures of silica and metal oxides. Clays can also be incorporated into oxide-type binders to modify the mechanical properties of the catalyst or to facilitate the manufacturing process.

[0083] Alternatively or in addition, the matrix component may include one or more substantially inactive materials. Inactive materials conveniently serve as diluents to control the amount of conversion so that the products can be obtained economically and systematically without the use of other means to control the reaction rate. Alternatively or in addition to any trivalent phosphorus added to or filled into the molecular sieve component, the matrix component may optionally include trivalent phosphorus, for example, to reduce acidity of the catalyst. It will be apparent to those skilled in the art that reducing the acidity of the catalyst reduces the amount of catalytic coke produced during the catalytic conversion of light hydrocarbons of feedstock to aromatic hydrocarbons. Suitable trivalent phosphorus containing matrices are described in US Pat. No. 5,026,937, which is incorporated herein by reference in its entirety. The matrix component is optional. In some aspects, the catalyst is practically free of matrix, for example, contains ≤ 1 wt. % matrix, for example, ≤ 0.1 wt. % In particular, the catalyst may practically not contain a binder, for example, contains ≤ 1 wt. % binder, such as ≤ 0.1 wt. % For example, the molecular sieve component of the catalyst may include ≥ 95 wt. % self-linked molecular sieve, for example, ≥ 95 weight. % self-bound ZSM-5, and in particular fine crystalline H-ZSM-5.

[0084] The catalyst may be a catalyst that has undergone one or more treatments, for example, a selective treatment to increase selectivity for the desired aromatic hydrocarbon compounds, such as para-xylene. For example, selection can be carried out before introducing the catalyst into the reactor and / or in situ, in the reactor, for example, bringing the catalyst into contact with a selection agent, such as at least one organosilicon compound, in a liquid carrier, and then calcining the catalyst at a temperature from 350 ° C to 550 ° C. This selection process can be repeated two or more times and changes the diffusion characteristics of the catalyst so that the formation of para-xylene becomes preferable to the formation of other xylene isomers. This selection process is described in detail in US patent No. 5.633.417 and 5.675.047.

[0085] Typically, the catalyst has a surface area, measured by physical absorption of nitrogen, in the range of 100 m ² / g to 600 m ² / g, for example, in the range of 200 m ² / g to 500 m ² / g. When the catalyst contains aluminosilicate, which includes trivalent phosphorus, the atomic ratio of trivalent phosphorus: aluminum, as a rule, is in the range from 0.01 to 0.5. For example, the catalyst may contain ≥ 10 wt. % trivalent phosphorus-modified alumina, such as ≥ 15 wt. %, or in the range of 10 wt. % up to 20 wt. %

[0086] During dehydrocyclization, at least a portion of the non-aromatic C ₂₊ hydrocarbons of said feedstock for flavoring the pipe 120 is converted to aromatic hydrocarbons, molecular hydrogen, and optionally additional non-aromatic hydrocarbons. The conditions of catalytic dehydrocyclization may include exposure to raw materials of temperature in the range from 400 ° C to 650 ° C, pressure in the range from 100 kPa to 2200 kPa. Typically, the conditions for catalytic dehydrocyclization additionally include hourly average feed rate (WHSV) ≥ 0.1 h ^-1 . Most commonly, catalytic dehydrocyclization conditions include temperatures in the range of 500 ° C to 625 ° C, and absolute pressures in the range of 30 psi. inch (207 kPa) up to 80 psi inch (522kPa). WHSV may be in the range of 0.1 h ^-1 to 20 h ^-1 . Typically, the WHSV of C ₂₊ hydrocarbons (ʺ WHSV C ₂₊ ʺ) in said aromatization feed relative to the dehydrocyclization catalyst is in the range from 0.1 h ^-1 to 20 h ^-1 , for example, from 0.2 h ^-1 to 5 h ^-1 or from 0.3 h ^-1 to 1 h ^-1 . WHSV C ₂₊ means the hourly consumption of C ₂₊ hydrocarbons (in grams per hour) exposed to the second catalyst per gram of dehydrocyclization catalyst. The reaction is usually endothermic. Typically, the average temperature difference in the reaction zone is ≤ 600 ° C, most often in the range from 20 ° C to 200 ° C, for example, in the range from 50 ° C to 150 ° C.

[0087] The effluent of the pipeline 126 from the reactor 224 can be cooled by heat exchange with the feed stream for the reactor of the pipe 120 through the heat transfer steps 226 and 228 . Further cooling may be provided by the flow of the reaction stream through line 126 through other coolers, for example, air cooler 236 . The effluent reaction stream is withdrawn from flavor 124 through a conduit 126 to a recuperator 128, for example, to extract at least some of the aromatics from the effluent reaction stream.

[0088] When the flavoring is used using the specified raw materials for flavoring under the indicated conditions, dehydrocyclization usually results in (i) ≥ 25 wt. % conversion of C ₂₊ raw materials for aromatization, for example, ≥ 25%, such as ≥ 50% or ≥ 75%; and (ii) C ₃₊ conversions ≥ 50%, for example, ≥ 75% or ≥ 90%, or ≥ 95%, or ≥ 98%. The exit reaction stream typically contains (i) unreacted feed, including an unconverted diluent, and (ii) the vapor-phase and / or liquid-phase products of the dehydrocyclization reaction and other reaction paths, the product phase being that which exists at the outlet of the lowest downstream zone dehydrocyclization reactions of reactor 224 . Typically, the reaction stream does not include solid (e.g., catalytic coke) and / or semi-solid (e.g., catalytic coke precursors) products of dehydrocyclization and other reaction paths, since they usually remain in the first stage until they are removed in the mode regeneration.

[0089] The effluent reaction stream typically contains ≥ 10 wt. % aromatic hydrocarbons, molecular hydrogen, ≥ 1 wt. % ethane, from 1 wt. % up to 40 wt. % methane, ≤ 2 wt. % propane and ≤ 1 wt. % butanes. The amount of ethane in the reaction product may be, for example, ≥ 5 wt. %, such as ≥ 10 wt. %, in particular, when dehydrocyclization is carried out at a relatively high temperature within the specified temperature range. Since ethane is usually formed by the hydrogenolysis of non-aromatic C ₃₊ hydrocarbons for aromatization during dehydrocyclization, the effluent reaction stream usually contains ethane and methane even when the aromatization feed contains <1 wt. % ethane and / or <1 wt. % methane.

Recuperator

[0090] Some forms of recuperator 128 are schematically illustrated in FIG. 2. Typically, the recuperator includes at least one separation stage 129 for separating at least tail gas, which is discharged through conduit 130, and the first product, which is discharged through conduit 138, from the effluent of the conduit 126.

[0091] In the aspects illustrated in FIG. 2, the separation stage 129 includes at least one drum separator for separating tail gas and the first product from the effluent from the reaction stream. The invention encompasses recuperators with one or more separation stages, and separation stages having one or more drum separators. Any suitable separation method can be used at separation stage 129 , including conventional separation technologies such as membrane separators, distillation columns, etc.

[0092] The first product typically contains ≥ 50 wt. % aromatic hydrocarbons exiting the reaction stream, for example, ≥ 75 wt. %, such as ≥ 90 wt. % or ≥ 95 wt. % The extraction of the first product from aromatic hydrocarbons from the effluent is typically carried out with a recovery in the range of 200 barrels per day (ʺBPDʺ) (3.2 M ³ PD) to 10.000 BPD (160 M ³ PD), for example, from 500 BPD (7.9 M ³ PD) up to 5000 BPD (79 M ³ PD), such as from 1000 BPD (16 M ³ PD) to 4000 BPD (64 M ³ PD). Tail gas typically contains ethane and molecular hydrogen, and typically additionally contains methane. For example, tail gas may contain 1 wt. % ethane, for example, ≥ 5 wt. %, such as ≥ 10 wt. % In a specific aspect, the tail gas comprises (A) ≥ 75 wt. % ethane from the exit reaction stream, for example, ≥ 90 wt. %, such as ≥ 95 wt. %; and (B) ≤ 10 wt. % aromatic hydrocarbons from the reaction product, for example, ≤ 5 wt. %, such as ≤ 1 wt. % Typically, the tail gas further includes (a) ≥ 10 wt. % molecular hydrogen from the effluent, for example, ≥ 25 wt. %, such as ≥ 50 wt. % or ≥ 95 wt. %, and / or ≥ 10 wt. % methane from the effluent, for example, ≥ 25 wt. %, such as ≥ 50 wt. % or ≥ 95 wt. %

Optional second dehydrocyclization step

[0093] The aromatization of the aromatization feed is not limited to one reactor zone, or to one reactor, or to one reaction step. The division of the reaction between several zones and / or reactors creates the possibility of more efficient production of cyclic products, for example, with higher yields and longer periods before coking. In addition, dividing the reactor into several zones and several vessels reduces the difficulties associated with designing a modular system with a higher throughput, for example, with a throughput of 50 MSCFD (1.4 × 10 ⁶ NM3D) or more.

[0094] Some aspects of the invention regarding a flavoring agent with first and second reaction steps are illustrated schematically in FIG. 3. Design Features in FIG. 3, which essentially perform the same function as in FIG. 1 and FIG. 2 are indicated under the same number. Although two reaction steps can be carried out in one reactor vessel or in separate vessels for each stage, aspects of FIG. 3 include a second reaction step in a reactor 334 , which is different from a reactor 224 for a first reaction step. It was found that the total selectivity for one or more of (i) additional methane and (ii) additional ethane, and (iii) additional light C ₃₊ hydrocarbons in steps 1 and 2 is less than if steps 1 and 2 are combined into one step, when operating under the conditions specified for step 1 or step 2. Typically, in the case of aromatization feedstocks that include associated gas, aromatization (for example, steps 1 and 2 carried out together) is characterized by (i) a total conversion of C ₂ hydrocarbons C ₄ ≥ 30%, e.g., ≥ 40%, for example, ≥ 50%, and (ii) a selectivity in relation ii aromatics ≥ 30%, e.g., ≥ 40% such as, ≥ 50%. It is desirable that these results can be achieved in one pass.

[0095] As shown in FIG. 3, the tail gas in conduit 130A is heated in heat transfer stages 326, 336, and 328. Heater 330 provides additional tail gas heating. If necessary, a portion of the molecular hydrogen of the tail gas can be burnt to provide additional heat. When using molecular hydrogen burning, burning is usually selective burning of molecular hydrogen. The heated tail gas is fed to a second reactor 334 to dehydrocyclize at least a portion of the C ₂₊ tail gas hydrocarbons to produce a second effluent reaction stream containing additional aromatic hydrocarbons and additional molecular hydrogen, which are “additional” in the sense that they are additional to those obtained in the first reactor 224 . Although reactor 334 shows the presence of two layers of an inter-layer dehydrocyclization catalyst provided by heater 330 , the second reactor is not limited to such forms. For example, part of the molecular hydrogen generated in the top layer of reactor 334 can be burned to provide additional heat for the endothermic dehydrocyclization reaction in the downstream layer.

[0096] Heat is removed from the second effluent of the reaction stream at heat transfer stages 326, 336, and 328, and passed through a conduit 325 to a recuperator 128 to recover a second tail gas and a second product. As shown in FIG. 3, this separation may be performed in a second separation stage related to recuperator 128 . As in the case of the first and second stages of the reaction, which can be localized in separate modules of the flavor, the first and second stages of the recuperator can be localized in separate modules of the recuperator. This reduces the difficulties associated with the distribution of these functions among the standard-sized modules used for work in various geographical locations.

[0097] The second stage of flavoring, as a rule, converts ≥ 5 wt. % ethane tail gas, for example, ≥ 10 wt. %, such as ≥ 20 weight. % Since the first product usually contains ≤ 5 wt. % hydrocarbons C ₃₊ , for example, ≤ 2 wt. % propane and ≤ 1 wt. % butanes, almost no conversion of C ₃₊ hydrocarbons in reactor 334 occurs. This effect advantageously reduces the accumulation of coke catalyst in stage 2.

[0098] It has been unexpectedly discovered that a marked conversion of ethane of the first product to aromatic carbon can be achieved with less or little aromatic hydrocarbon removal of the first exiting reaction stream before it enters the second stage, even if the reaction temperature of the second stage exceeds the temperature of the first stage. Accordingly, in some aspects that are not shown in FIG. 3, the recuperator does not use the first separation stage 129 or uses the first separation stage of lower power, and / or uses various separation conditions. Although these aspects cover the diversion of substantially all of the first effluent from the reaction stream to the second reactor 334 , they also include, for example, setting the separation stage 129 to remove at least a portion of the molecular hydrogen of the first effluent reaction stream (and / or a portion of the methane of the first effluent reaction stream ) before the tail gas residue is introduced into the reactor 334 . Separation of methane may be desirable when the first effluent of the reaction stream otherwise contains more methane than is necessary to reduce the partial pressure of ethane in reactor 334 to a range that reduces coke accumulation and increases the yield of aromatic hydrocarbons.

[0099] As an example of such alternative aspects, feedstock for reactor 334 may include ≥ 90 wt. % of the reaction effluent from reactor 224, for example, ≥ 95 wt. %, for example, ≥ 99 wt. % Despite the fact that these aspects use a minor, if any, compositional separation between reactors 224 and 334 , within the scope of their volume it is possible to divide the entire first outgoing reaction stream into two or more streams with one stream including ≥ 50 wt. . % of the entirety of the first exiting reaction stream fed to the reactor 334 , for example, ≥ 75 wt. %, such as ≥ 90 wt. % or ≥ 95 wt. %; and one or more additional threads deduced from the process. Also within the scope of these aspects is the separation of at least a portion of the aromatic hydrocarbons of the first effluent from the reaction stream before entering reactor 334 . However, in these aspects, the portion of the first reaction effluent that is reacted under dehydrocyclization conditions in reactor 334 typically includes ≥ 50 wt. % aromatic hydrocarbons of the first exiting reaction stream, for example, ≥ 75 wt. %, such as ≥ 90 wt. % or ≥ 95 wt. %; and ≥ 50 wt. % molecular hydrogen of the first exiting reaction stream, for example, ≥ 90 wt. % or ≥ 95 wt. % It is also typical of these aspects that the portion of the first effluent of the reaction stream to be reacted under dehydrocyclization conditions in reactor 334 includes ≥ 50 wt. % ethane of the first exiting reaction stream, for example, ≥ 75 wt. %, such as ≥ 90 wt. % or ≥ 95 wt. %; and ≥ 50 wt. % of any other (i.e., in addition to ethane) non-aromatic hydrocarbons in the first exiting reaction stream, for example, ≥ 75 wt. %, such as ≥ 90 wt. % or ≥ 95 wt. % The inter-stage separation of these aspects, when used, is typically carried out at the separation stage 129 of the recuperator 128. Alternatively or in addition to the aspects illustrated by FIG. 2 and FIG. 3, heating may be provided for the dehydrocyclization reaction of reactors 224 and / or 334 by methods other than fire heaters. For example, in some aspects, at least part of the heat is provided by (i) electric heating, which can use the electricity generated by power unit 140, and / or (ii) a hot oil system or a steam system that recovers heat from the flue gas into the heat exchanger, the power unit 140 emitted during operation of the gas turbine. In these or other aspects, at least a portion of the heat is provided by burning at least a portion of the molecular hydrogen generated in the upstream and downstream layers of the kat izatora. The oxidizing agent for combustion can be introduced into the reactors, for example, in one or more of the following places: upstream of the catalyst bed, (ii) downstream of the catalyst bed, between upstream and downstream catalyst beds, downstream of the underlying catalyst bed. Typically, the combustion of molecular hydrogen, when used, is selective combustion of molecular hydrogen (ʺSHCʺ), for example, selective combustion of molecular hydrogen instead of hydrocarbons. SHC may be a catalytic SHC, for example, with an SHC catalyst placed close to the dehydrocyclization catalyst. Suitable molecular hydrogen combustion schemes are described in US Pat. No. 5,026,937, which is incorporated herein by reference in its entirety. The oxidizing agent for the combustion of molecular hydrogen or the selective combustion of molecular hydrogen may be, for example, atmospheric oxygen and / or molecular oxygen, including molecular oxygen in air. Alternatively or in addition, the oxidizing agent may be provided with one or more solid oxygen carriers (solid oxygen carriers or “SOC”), which can be replenished with oxidizing agent when the flavor is in a regeneration mode. In some aspects, the SOC is a composite with SHC. It was found that a decrease in the amount of molecular hydrogen in reactors increases the yield of aromatic hydrocarbons. It was also noted that, as a rule, it is desirable not to burn almost all molecular hydrogen in reactors, but to burn only a sufficient amount of molecular hydrogen to provide the heat required for endothermic dehydrocyclization. It has been found that a long time cycle is easier to achieve when at least some molecular hydrogen is present in the reactors.

[0100] Although the process conditions for dehydrocyclization in reactor 334 may be the same as in reactor 224 , they generally differ. Similarly, although the same dehydrocyclization catalyst can be used in reactors 224 and 334 , as a rule, they are different. Although FIG. 2 and 3 show the reactors 224 and 334 as containing two catalytic layers (an overlying layer and an underlying layer), the invention is not limited to reactors of this form. The scope of the invention includes reactors with one catalytic layer, as well as reactors with three or more layers. Reactors with moving catalyst beds and / or fluidized catalyst beds are within the scope of the invention, although typically reactors 224 and 334 contain at least one fixed catalyst bed. In the aspects illustrated in FIG. 3, each of the first and second reaction steps includes one reactor ( 224 and 334 ). However, the invention is not limited to the reaction steps carried out in one reactor, and in other aspects, additional reactors are used, for example, for additional dehydrocyclization. Flavor 124 may include more than two reaction steps. Each stage may have the same number of reactors, but this is not required. In some aspects, a two-reaction flavoring agent is used (as in the aspects illustrated in FIG. 3), the catalytic dehydrocyclization conditions in reactor 224 include exposing the feedstock to aromatize temperature T ₁ between 400 ° C and 630 ° C and pressure P ₁ , sufficient for dehydrocyclization. Typically, T ₁ ranges from 450 ° C to 605 ° C. Usually absolute P ₁ ≥ 20 psi inch (137.9 kPa), for example, ≥ 35 psi inch (241, .3 kPa), such as in the range of 35 psi inch (241.3 kPa) up to 300 psi inch (2070 kPa). For example, P ₁ may be in the range of 35 psi. inch (241.3 kPa) up to 300 psi inch (2070 kPa), for example, from 37 psi inch (255.1 kPa) up to 80 psi inch (522 kPa), or from 40 psi inch (275.8 kPa) up to 80 psi inch (522 kPa) or from 45 psi inch (310.2 kPa) up to 80 psi inch (522 kPa). Typically, the conditions for catalytic dehydrocyclization additionally include the hourly average feed rate (WHSV) in the range from 0.1 h ^-1 to 20 h ^-1 , for example, from 0.2 h ^-1 to 5 h ^-1 , such as from 0 , 3 h ^-1 to 1.0 h ^-1 . WHSV is based on the hydrocarbon content of C ₂₊ aromatization feed and means the hourly average hydrocarbon feed rate of C ₂₊ aromatization feed (in grams) per gram of the first catalyst. In a specific aspect, (i) the average temperature in any reaction zone in the reactor 224 (and in any catalyst bed located in the reaction zone of the reactor 224 ) is ≤ 600 ° C and (ii) the aromatization feed is not exposed to a temperature of ≥ 630 ° C at the entrance to the reactor 224 . Typically, the dehydrocyclization reaction in the reactor 224 is carried out at a relatively high pressure (relative to the pressure in the reactor 334 ), therefore, hydrogen is formed during the dehydrocyclization reaction at a relatively high partial pressure. It was expected that the high partial pressure of hydrogen will suppress the dehydrogenation of the feed, which will lead to a decrease in the yield of C ₆₊ hydrocarbons such as aromatic hydrocarbons. Surprisingly, this was not found to be the case: the relatively high partial pressure of hydrogen in reactor 224 was found to reduce the rate of accumulation of catalytic coke with little or no decrease in the yield of aromatic hydrocarbons. For example, when the dehydrocyclization conditions in reactor 224 include pressures P ₁ ≥ 35 psi. inch (241.3 kPa), the dehydrocyclization reaction can be maintained for a time period of ≥ 50 hours, for example, ≥ 100 hours, such as ≥ 200 hours or ≥ 500 hours, while reducing the yield of aromatic hydrocarbons by ≤ 10%, for example, by ≤ 5%, such as for ≤ 1%.

[0101] It will be apparent to those skilled in the art that, within the process parameters specified for reactor 224 , there are process conditions that, if selected, will result in maximum conversion of propane to ʺX _MP аромат aromatics. Unlike conventional multi-stage processes, reactor 224 is typically operated upon converting propane to aromatic hydrocarbons, which is less than X _MP . Instead, when the flavoring feed includes propane and / or when propane is formed in reactor 224 , the process conditions are usually chosen such that the initial (start of run) conversion of propane to aromatic hydrocarbons in reactor 224 ʺX _1P ʺ is ≤ 0.95⋅X _MP . Typically, ʺX _1P ʺ ≤ 0.90⋅X _MP , for example, ≤ 0.85⋅X _MP or ≤ 0.75 X _MP . It has been found that operating reactor 224 under conditions that provide X _1P ≥ X _MP causes excessive coking of the catalyst, usually resulting in shorter cycle times when operating with a fixed bed. The same effect is observed for the conversion of butanes to aromatic hydrocarbons, but with less sensitivity to changes in process conditions in reactor 224 . Instead of operating at a reactor 224 with a maximum conversion of butanes to aromatic hydrocarbons X _MB , these process conditions usually lead to the initial (start of run) conversion of butanes to aromatic hydrocarbons (ʺX _1B ʺ), which is less than X _MB , for example, X _1B ≤ 0.995⋅X _MB , such as ≤ 0.99⋅X _MB or ≤ 0.985⋅X _MB .

[0102] The dehydrocyclization conditions in reactor 334 typically include a temperature of T ₂ in the range of 450 ° C to 700 ° C and an absolute pressure of P ₂ ≤ 35 psi. inch (241.3 kPa). Typically, T ₁ ≤ 0.9⋅T ₂ , for example, T ₁ ≤ 0.85⋅T ₂ , such as T ₁ ≤ 0.8⋅T ₂ . The pressure in reactor 334 is generally less than the pressure in reactor 224 , for example, P ₂ ≤ 0.95⋅P ₁ , such as P ₂ ≤ 0.90⋅P ₁ or P ₂ ≤ 0.85⋅P ₁ , or P ₂ ≤ 0.8⋅P ₁ . Typically, reaction conditions in reactor 334 include T ₂ in the range of 500 ° C to 675 ° C and P ₂ ≤ 34 psi. inch (abs.) (234.4 kPa), e.g. ≤ 32 psi inch (abs.) (220.6 kPa), such as ≤ 30 psi inch (abs.) (207 kPa), or in the range of 10 psi inch (abs.) (68.9 kPa) up to 35 psi inch (abs.) (241.3 kPa) or from 12 psi inch (abs.) (82.8 kPa) up to 34 psi inch (abs.) (234.4 kPa). Typically, the reaction in the reactor 334 is carried out at a WHSV of C ₂₊ hydrocarbons of said first product with respect to the second catalyst in the range from 0.1 h ^-1 to 20 h ^-1 , for example, from 0.2 h ^-1 to 5 h ^{- 1} or from 0.3 h ^-1 to 1 h ^-1 . According to the WHSV of reactor 224 , the WHSV of reactor 334 is based on the C ₂₊ hydrocarbon content and means the hourly consumption of C ₂₊ hydrocarbons (in grams) introduced into the reactor 334 per gram of the second catalyst. It will be apparent to those skilled in the art that T ₁ and T ₂ are average temperatures in the reaction zone, or more specifically, in a catalyst bed located in the reaction zone. The average temperature is calculated by adding the input temperature of the zone to the output temperature of the zone, and then dividing the sum by 2. P ₁ and P ₂ are not average pressures. Instead, they correspond to the inlet pressure at said reactor, for example, reactor 224 for P ₁ and reactor 334 for P ₂ . It has been found that, unlike reactor 224, increased reaction pressure in reactor 334 (i) reduces the yield of aromatic hydrocarbons and (ii) increases the rate of accumulation of catalytic coke. Absolute pressure P ₂ ≤ 35 psi inch (241.3 kPa), it is usually necessary to achieve 334 in the reactor cycle time ≥ 50 hours.

[0103] Contrary to expectations, it is harmful to operate reactor 334 at a temperature sufficient to maximize ethane conversion, typically T ₂ > 700 ° C. In this case, there is a decrease in selectivity for a given aromatic hydrocarbon product. Typically, the average temperature in any reaction zone in the reactor 334 (and in any catalytic bed located in the reaction zone in the reactor 334 ) is ≤700 ° C. Typically, the flavoring feed is not exposed to a temperature of ≥700 ° C at the inlet to the reactor 334 .

[0104] It has also been found that an active reactor 334 at a temperature> 700 ° C can lead to chemical conversion of the dehydrogenating component of the catalyst and loss of catalytic dehydrocyclization activity, especially when the dehydrogenating component contains one or more Zn oxides. Without reference to any theory or model, it is believed that the use of temperatures> 700 ° C leads to the conversion of the oxide form to the form of metallic zinc, the vapor pressure of which is higher than that of the oxide form. Thus, the loss of catalytic dehydrocyclization activity is explained, at least in part, by the evaporation of Zn from the catalyst. Typically, the total conversion of ethane in the second stage is ≤ 60%, for example, ≤ 50, such as ≤ 40%, or in the range from 25% to 60%, or from 30% to 55%, or from 30% to 50%, or from 30% to 40%.

[0105] At least to reduce the rate of accumulation of catalytic coke, typically the dehydrocyclization conditions in reactor 224 are chosen to convert (i) ≥ 50 wt. %, for example, ≥ 75 wt. %, such as ≥ 90 wt. % non-aromatic hydrocarbons With ₃₊ raw materials for aromatization, but (ii) ≤ 50 wt. %, for example, ≤ 25%, such as ≤ 10% or ≤ 1% of C ₂ hydrocarbons for flavoring. Since feedstock for reactor 334 typically contains less non-aromatic C ₃₊ hydrocarbon compounds than feedstock for flavoring, reactor 334 is typically operated under more stringent conditions to achieve at least some conversion of ethane to aromatic hydrocarbons. It will be apparent to those skilled in the art that, within the ranges of process parameters indicated for reactor 334 , process conditions exist that, when used, can result in maximum conversion of ethane to ʺX _ME ʺ aromatics. Unlike conventional multi-stage processes, reactor 334 typically does not run on X _ME . Instead, the process conditions are usually chosen so that the initial (start of run) conversion of ethane to aromatic hydrocarbons in the 334 ʺX _2E ʺ reactor is less than X _ME , for example, ≤ 0.9⋅X _ME . In most cases, X _2E ≤ 0.85⋅X _ME , for example, ≤ 0.8⋅X _ME or ≤ 75⋅X _ME . It was found that the working step two under conditions providing X _2E ≥ X _ME leads to excessive coking of the catalyst, which inevitably entails a shortened cycle time when working with a fixed bed. Since ethane is actually converted in reactor 334 , X _{2E is} greater than zero. Since ethane can be formed (and usually is) in reactor 224 , the ethane conversion in reactor 224 (X _1E ) may be less than zero. As a rule, when working in these conditions using the specified raw materials X _2E greater than X _1E . Usually [1- (X _1E / X _2E )] is ≤ 5, for example, ≤ 2, for example ≤ 1, or in the range from 0.5 to 5, or from 0.6 to 2, or from 0.6 to 1 .

[0106] In specific aspects, the first stage reactor 224 comprises at least two fixed layers of the first catalyst, and the second stage reactor 334 contains at least two fixed layers of the second catalyst. In particular, in these aspects, the aromatization feed reacts in the presence of a first catalyst in a reactor 224 when exposed to a temperature of T ₁ in the range of 450 ° C. to 605 ° C and an absolute pressure of P ₁ in the range of 37 psi. inch (abs.) (255.1 kPa) up to 100 psi inch (abs.) (689.5 kPa), with a space velocity (WHSV) in the range from 0.1 h ^-1 to 20 h ^-1 , such as from 0.25 h ^-1 to 2 h ^-1 . This part of the first effluent reaction stream, which reacts to the second stage of reactor 334, is subjected to a temperature of T ₂ in the range of 500 ° C. to 675 ° C. and an absolute pressure of P ₂ ≤ 32 psi. inch (220.6 kPa), with a space velocity (WHSV) in the range from 0.1 h ^-1 to 20 h ^-1 , such as from 0.25 h ^-1 to 2 h ^-1 , and usually less than volume the rate used in the reaction of the aromatization feedstock in reactor 224 . More specifically, T ₁ ≤ 0.9⋅T ₂ and P ₂ ≤ 0.9⋅P ₁ , with P ₂ in the range of 1 psi. inch (abs.) up to 32 psi inch (abs.) (220.6 kPa), for example, from 5 psi inch (abs.) (34.5 kPa) up to 30 psi inch (abs.) (206.8 kPa) or from 5 psi inch (abs.) (34.5 kPa) up to 29 psi inch (abs.) (199.9 kPa).

[0107] The first and second catalysts may be selected from among the dehydrocyclization catalysts described in the relevant aspects, illustrated by the single-stage flavoring in FIG. 2. In specific aspects, the first catalyst includes a molecular sieve component, a dehydrogenating component, and optionally a matrix component. The molecular sieve component contains ≥ 90 wt. % aluminosilicate in hydrogen form, aluminosilicate with a permeability index in the range of 2-12 (for example, modified with trivalent phosphorus H-ZSM-5) and has a silica to alumina ratio in the range of 50 to 80. The molecular sieve component may include ≥ 95 wt. . % self-bound fine crystalline H-ZSM-5. The dehydrating component includes ≥ 90 wt. % of at least one Ga oxide. The matrix component includes ≥ 75 wt. % alumina, silica and combinations thereof. The catalyst optionally includes ≥ 1 wt. % phosphorus, for example, in the form of modified trivalent phosphorus H-ZSM-5 and / or due to the inclusion of trivalent phosphorus in the matrix component.

[0108] In specific aspects, the second catalyst includes a molecular sieve component, a dehydrogenation component, and optionally a matrix component. The second catalyst can be selected from among the same catalysts that are provided for use in step 1 (first catalysts), although, as a rule, the second catalyst has a higher acidity than the first catalyst. For example, the molecular sieve component includes ≥ 90 wt. % aluminosilicate in hydrogen form, aluminosilicate with a permeability index in the range from 2-12 (for example, crystalline, H-ZSM-5). Typically, (i) the second catalyst has a silica: alumina ratio in the range of from 3 to 60, for example, from 10 to 40, such as from 15 to 35, and (ii) the catalyst comprises <0.01 wt. % phosphorus.

[0109] In these aspects, the dehydrogenation component of the second catalyst may include ≥ 90 wt. % of at least one Zn oxide. The matrix component, when used, includes ≥ 90 wt. % alumina, silica and combinations thereof. The second catalyst typically includes <0.01 wt. % phosphorus.

[0110] A second reaction effluent is formed by the reaction of this portion of the first reaction effluent 334 from the reactor with said second catalyst under the indicated conditions. The second effluent reaction stream is discharged as a second effluent reaction stream through a conduit 325. Typically, the second effluent reaction stream comprises (i) ≥ 0.5 wt. % additional aromatic hydrocarbons, (ii) additional molecular hydrogen and, optionally, methane and ethane. The second exiting reaction stream, as a rule, includes ≥ 1 wt. % of the total amount of aromatic hydrocarbons, for example, the total amount of (i) additional aromatic compounds and (ii) some of the aromatic compounds of the effluent reaction stream, which enter the second stage but are not converted. Typically, the second effluent of the reaction stream has a total aromatic hydrocarbon content of ≥ 5 wt. %, based on the mass of the second exiting reaction stream, such as ≥ 10 wt. %, or in the range of 1 wt. % up to 95 wt. %, or from 10 wt. % up to 75 wt. % The content of additional aromatic hydrocarbons of the second exit reaction stream is usually ≥ 1 wt. %, for example, ≥ 2 wt. %, such as ≥ 10 wt. % or ≥ 25 wt. %, or ≥ 50 wt. %, or ≥ 75 wt. % The method provides the desired BTX product. Despite the fact that T _{2 is} greater than T ₁ , it was found that the second exit reaction stream is characterized by an unexpected increase in the desired xylene isomers and an unexpected decrease in the yield of less desirable C ₁₁₊ aromatic hydrocarbons.

[0111] Light hydrocarbons, such as ethane, may be present in the second effluent. Methane, when present in the aromatization feed, is usually not converted to other hydrocarbon compounds under the conditions specified for step 1 and / or step 2; therefore, when no intermediate separation of methane is used, substantially all of the methane feed is typically present in the second effluent stream. The second effluent of the reaction stream usually includes additional methane, namely, that which is formed in step 2, for example, by hydrogenolysis of C ₂₊ hydrocarbons, and, optionally, that which is formed in step 1, when the intermediate separation of methane is reduced or not used.

Product recovery from the reaction stream leaving the second stage

[0112] Although the reaction stream leaving the second stage can be recovered in the separation stage 129 of the recuperator 128, it is usually used in the second separation stage, for example, in the stage 329 , as shown in FIG. 3. The second separation stage 329 can be configured to operate essentially the same way and, in practice, the same components can be used on it as on the separation stage 129 . For example, in the second separation stage 329, tail gas, which is vented through conduit 130B, and (ii) a second product that is vented through conduit 338. can be separated from the second effluent from the second effluent, and the second product typically contains ≥ 50 wt. % aromatic hydrocarbons of the second exit reaction stream, for example, ≥ 75 weight. % or ≥ 90 wt. % The recovery of the second product from the aromatic hydrocarbons of the second effluent is usually carried out with a recovery in the range from 200 BPD (3.2 M ³ PD) to 10.000 BPD (160 M ³ PD), for example, from 500 BPD (7.9 M ³ PD ) up to 5000 BPD (79 M ³ PD), such as from 1000 BPD (16 M ³ PD) to 4000 BPD (64 M ³ PD). Optionally, the second product is combined with the first product, as shown in FIG. 3. The separated tail gas of conduit 130B typically contains molecular hydrogen, methane, and unreacted ethane.

Dehydrocyclization Catalyst Regeneration

[0113] In some aspects, the flavor (including, for example, a first step and / or a second step; and / or additional steps, if used) may contain one or more additional reaction zones, for example, additional reactors (not shown). These additional reactors can be essentially the same as the rest, namely, almost the same configuration of the layers and with almost the same amount of essentially the same catalyst. Typically, stage reactors operate in a reactive (dehydrocyclization) mode, while additional reactors operate in a regeneration mode, and vice versa. Continuous or semi-continuous operation can be carried out at each stage, for example, by alternating reactors and additional reactors connected in series in reaction and regeneration modes. In the aspects illustrated in FIG. 3, the work in each of the first and second stages is usually carried out in dehydrocyclization mode for an average cycle time of ≥ 50 hours, for example, ≥ 75 hours, for example ≥ 100 hours or ≥ 120 hours, when aromatic compounds yield, which is ≥ 95% of the yield achieved at the beginning of the dehydrocyclization regimen, for example, ≥ 90%, such as ≥ 75%.

[0114] Typically, the dehydrocyclization catalyst used in the steps is regenerated at a temperature of ≤ 700 ° C. It was found that exceeding this temperature during regeneration causes dealumination of the catalyst and / or loss of structure, which leads to an undesirable loss of acidity of the catalyst. The regeneration of the catalyst for any of these catalysts is usually carried out using processes that limit the maximum temperature to which the catalyst is exposed during regeneration, to approximately a temperature of up to 750 ° C, typically up to 650 ° C. Conventional catalyst regeneration methods can be used, for example, exposing the catalyst to an oxidizing agent, such as air or oxygen in air, for a time sufficient to remove at least a portion of the catalytic coke, but the invention is not limited to this. When carrying out the dehydrocyclization steps 1 and / or 2 in a series of series-connected adiabatic fixed-bed reactors, a suitable regeneration method involves circulating a stream of regeneration gas containing a limited amount of oxygen, which limits the magnitude of the exothermic effect in which coke is burned from the catalyst. As a rule, in the place where the regeneration gas enters the first (highest upstream of the regeneration gas stream) reactor, for example, at the inlet of the reactor, the regeneration gas is exposed to a temperature of ≤ 350 ° C, for example, ≤ 325 ° C, as e.g. ≤ 300 ° C. If necessary, the oxidizer content in the regenerating gas can be reduced to reduce the risk of exceeding the maximum temperature.

Power unit

[0115] Some forms of a power unit that use tail gas as feed as feed 130B will now be described in more detail with reference to FIG. 3. The invention is not limited to these forms, and this description is not intended to exclude other forms, such as forms that use raw materials containing tail gas transported through conduit 130A .

[0116] As shown in FIG. 3, tail gas is compressed in a compressor 160 and then fed to a combustion stage in a gas turbine 161 , which also includes a compression stage and an expansion stage. The expansion and compression stages are connected through a rotating shaft 166. The oxidizing agent of the pipe 162 , for example, oxygen, air or oxygen in the air, is compressed in the compression stage before the combustion stage. During combustion, high-pressure exhaust gases are formed that pass through the expansion stage, creating mechanical power (power on the shaft) that rotates the shaft 166 . The shaft 166 is rotatably coupled to an electric generator 164 that produces electricity 165. Low pressure flue gas 163 still conserves a significant amount of thermal energy. Accordingly, heater 230 may include one or more HRUs. Optionally, these aspects include the combustion of a portion of the tail gas (e.g., in conduits 130A and / or 130B) in a heater 230 .

[0117] The generated electrical power can be used locally or exported to the power grid. The amount of electricity produced may depend on the volumetric flow rate of the raw material for flavoring. In examples where the aromatization feed has a volumetric flow rate of about 50 MSCFD (1.4 × 10 ⁶ NM ³ D), the generated electric power can be about 150 megawatts (MW).

[0118] Although the present methods may be susceptible to various modifications and alternative forms, the above examples have been shown by way of example only. However, it should be understood once again that these methods should not be construed as being limited to the specific examples provided in this document. On the contrary, the present methods include all alternatives, modifications, and equivalents falling within the true spirit and scope of the appended claims.

Claims (86)

1. A system for producing aromatic hydrocarbons, including: (a) a flavoring configured to (i) receive a predetermined amount of ≥1 million standard cubic feet per day [MSCFD] (29000 NM ³ D) of a flavoring feed, where the flavoring feed comprises C ₂ hydrocarbons and non-aromatic C ₃₊ hydrocarbons and has a predetermined concentration of non-aromatic hydrocarbons C _{2 +} in the range from 15 to 90 mol.% per mole of raw materials for aromatization, and (ii) aromatizing (A) at least part of non-aromatic hydrocarbons C ₃₊ raw materials for aromatization and (B) at least part hydrocarbons With ₂ raw materials for flavoring with received using an exit reaction stream containing molecular hydrogen, non-aromatic hydrocarbons and aromatic hydrocarbons; (b) a product recuperator in fluid communication with a flavoring agent configured to recover at least the first and second products from the effluent reaction stream, where the first product comprises at least a portion of the aromatic hydrocarbons in the effluent reaction stream and the second product comprises a tail gas comprising (i) at least a portion of the molecular hydrogen of the effluent reaction stream; and (ii) at least a portion of the non-aromatic hydrocarbons of the effluent reaction stream; and (c) a feedstock pretreatment unit configured to (A) receive ≥2 MSCFD (57000 NM ³ D) gaseous feedstocks and (B) produce a flavoring feedstock from a gaseous feedstock, where the gaseous feedstock contains from 0 to 98 mol% methane , from 1 to 50 mol.% hydrocarbons C ₂ and from 1 to 40 mol.% non-aromatic hydrocarbons C ₃₊ , where the installation for preliminary processing of raw materials includes: (i) at least one heat transfer step having (A) an input in fluid communication with the feed pipe and (B) an output, the heat transfer step is configured to transfer heat from the gaseous feed in order to condense part of the gaseous feed at a specified dew point, and (ii) at least one vapor-liquid separator, where the vapor-liquid separator has an inlet in fluid communication with the heat transfer step, at least one drum separator and first and second outlets, where: (A) the separator inlet is configured to receive gaseous feed and condensed portions from the heat transfer step, (B) a drum separator is configured to separate gaseous feedstocks and condensed portions to produce a bottoms stream including a flavoring feedstock and an overhead stream including separated gaseous hydrocarbons, (C) a first separator outlet configured to transfer the separated gaseous portion from the pretreatment unit, and (D) the second output of the separator is configured to transfer a given amount of raw material for flavoring in the flavoring, and where (iii) the first heat transfer step of the pretreatment unit includes at least one heat exchanger in indirect heat contact with the refrigerant stream, (iv) the first heat exchanger is adapted to indirectly transfer the first, second and third parts of the heat transferred from the gaseous feed, where the first part is transferred to the refrigerant, the second part is transferred to the separated gaseous hydrocarbons and the third part is transferred to the raw material for flavoring, and (v) the second heat transfer step is configured to remove a certain amount of heat from the refrigerant stream, where the amount of heat changes by no more than +/- 50%, while the condensed part condensed from the gaseous feed has a volume change of ≤ +/- 10%. 2. The system of claim 1, wherein the system further includes a tail gas converter in fluid communication with a product recuperator, where the tail gas converter includes at least one molecular hydrogen separation step and at least one electric power generation step, where: (i) the molecular hydrogen separation step includes an inlet for receiving tail gas from a product recuperator and first and second outlets, (ii) a hydrogen separator in fluid communication with a molecular hydrogen separation step, wherein the hydrogen separator is configured to separate from the tail gas a hydrogen enriched stream comprising ≥75 wt.% molecular hydrogen tail gas and a hydrogen depleted stream comprising ≥25 wt.% non-aromatic tail gas hydrocarbons, (iii) the first output of the molecular hydrogen separation step is configured to transmit at least a portion of the molecular hydrogen enriched in a hydrogen stream to the electric power generation step, (iv) the power generation step includes at least one flame heater and / or at least one gas turbine having a rotary shaft connected to at least one electric generator, (v) a gas turbine is configured to burn at least a portion of the hydrogen-rich stream to generate shaft power; (vi) the generator is configured to convert at least a portion of the shaft power to electricity, where at least a portion of the electricity is used to supply heat to the flavor, (vii) a flame heater configured to burn at least a portion of the hydrogen-rich stream to produce heat, where at least a portion of the heat is supplied to the flavor; and (viii) the second outlet of the molecular hydrogen separation step is configured to transfer the raffinate from the tail gas converter. 3. The system according to any one of paragraphs. 1 and 2, where: (i) the flavoring includes at least one inlet of at least one reactor and at least one outlet, (ii) the inlet of the flavoring is configured to receive raw materials for flavoring and transferring the raw materials for flavoring to the reactor, (iii) the reactor is configured to contact the aromatization feed with at least one layer of a dehydrocyclization catalyst in a temperature range of 400 to 650 ° C, a pressure in the range of 100 to 2200 kPa, and a space velocity in the range of 0.1 to 20 hours ^-1 to obtain exiting the flavor stream, and (iv) the flavor outlet is configured to transfer at least a portion of the effluent reaction stream to a product recuperator. 4. The system according to any one of paragraphs. 1-3, where the pretreatment plant receives from 10 MSCFD (290000 NM ³ D) to 150 MSCFD (4.3⋅10 ⁶ NM ³ D) of gaseous feed, a predetermined amount of flavoring feed from 2 MSCFD (57000 NM ³ D) up to 80 MSCFD (2.3 × 10 ⁶ NM ³ D), and the product recuperator extracts ≥ 90 wt.% Of aromatic hydrocarbons from the effluent reaction stream with a recovery in the range of 200 BPD (3.2 M ³ PD) to 10000 BPD (160 M ³ PD). 5. The system according to any one of paragraphs. 1-4, where the gaseous feed contains associated gas, comprising from 0 to 95 mol.% Methane, from 5 to 50 mol.% Ethane, from 2 to 40 mol.% Propane, from 0.1 to 30 mol.% Is- butane, from 1 to 30 mol% of n-butane and from 0.05 to 25 mol% of isopentane. 6. The system according to any one of paragraphs. 1-5, where at least one of: installations for pre-treatment, flavoring and recuperator exists in the form of one or more mobile modules. 7. The modular system according to claim 6, where at least one of the mobile modules is based on one or more block installations, each block installation has a gross mass ≤100000 pounds (45360 kg), width ≤15 feet (4.6 m) and height ≤ 15 feet (4.6 m). 8. The method of conversion of hydrocarbons, including: (a) providing a predetermined amount of ≥1 million standard cubic feet per day [MSCFD] (29000 NM ³ D) of a flavoring feed comprising C ₂ hydrocarbons and C ₃₊ non-aromatic hydrocarbons, where the flavoring feed has a predetermined concentration of non-aromatic C _{2 +} hydrocarbons in the range from 15 to 90 mol.% per mole of raw materials for flavoring; (b) a reaction of a feedstock for aromatization in the presence of at least one dehydrocyclization catalyst under catalytic dehydrocyclization conditions to produce an exit reaction stream comprising (i) non-aromatic hydrocarbons and (ii) aromatic hydrocarbons and molecular hydrogen, where aromatic hydrocarbons and molecular hydrogen are formed when dehydrocyclization of at least a portion of the C ₂ hydrocarbons for aromatization and / or non-aromatic C ₃₊ hydrocarbons for aromatization; and (c) recovering at least the first and second products from the effluent reaction stream, wherein the first product comprises at least a portion of the aromatic hydrocarbons of the effluent reaction stream and the second product comprises a tail gas comprising (i) at least a portion of the molecular hydrogen of the effluent reaction stream and (ii) at least a portion of non-aromatic hydrocarbons from the effluent; and (d) obtaining a feed for aromatization by pretreating ≥2 MSCFD (57000 NM ³ D) of a gaseous feed, the gaseous feed contains from 0 to 98 mol% of methane, from 1 to 50 mol% of C ₂ hydrocarbons and from 1 to 40 mol. % non-aromatic hydrocarbons With ₃₊ ; where the pre-treatment includes (i) transferring heat from the gaseous feed at least in the first heat transfer step, in order to partially condense the gaseous feed at a fixed dew point, and (ii) separating the overhead stream, including gaseous hydrocarbons and bottoms, from the partially condensed feed, where the bottoms stream is a raw material for flavoring; Where (iii) the first heat transfer step includes at least one heat exchanger in indirect heat contact with the refrigerant stream, (iv) the heat exchanger indirectly transfers the first, second and third parts of the heat transferred from the gaseous feed, where the first part is transferred to the refrigerant, the second part is transferred to the separated gaseous hydrocarbons and the third part is transferred to the raw material for flavoring, and (v) the pre-treatment further includes the removal of the refrigerant stream from the first heat transfer step to the second heat transfer step and the removal of a certain amount of heat from the refrigerant to the second heat transfer step, where the amount of heat changes by no more than +/- 50% when the condensed part is condensed from the gaseous raw materials having a volume change of ≤ +/- 10%. 9. The method according to p. 8, where the gaseous feedstock includes associated gas containing from 0 to 95 mol.% Methane, from 5 to 50 mol.% Ethane, from 2 to 40 mol.% Propane, from 0.1 to 30 mol % iso-butane, from 1 to 30 mol.% n-butane and from 0.05 to 25 mol.% isopentane. 10. The method according to p. 8 or 9, where the gaseous feed is supplied in an amount varying from 10 MSCFD (290000 NM ³ D) to 150 MSCFD (4,310 ⁶ NM ³ D), the specified amount of raw material for flavoring varies in the range from 2 MSCFD (57000 NM ³ D) to 80 MSCFD (2.3 × 10 ⁶ NM ³ D) and step (c) involves recovering ≥ 90 wt.% of aromatic hydrocarbons from the aromatization stream with a recovery in the range of 200 BPD (3.2 M ³ PD) up to 10000 BPD (160 M ³ PD). 11. The method according to p. 8, where the partial condensation of the gaseous feed is carried out at an absolute pressure in the range from 300 psi. inch (2068 kPa) up to 2000 psi inch (13800 kPa), where the set dew point is in the range of -40 ° F (-40 ° C) to -20 ° F (-28.9 ° C). 12. The method according to p. 8, further comprising: (d) separating from the tail gas an extract comprising? 75 wt.% molecular hydrogen tail gas and a raffinate comprising? 90 wt.% non-aromatic hydrocarbon tail gas; (e) providing at least one gas turbine connected through at least one rotating shaft to generate power on the shaft of at least one electric generator; (f) burning at least a portion of the molecular hydrogen of the extract in a gas turbine to generate shaft power; (g) converting at least a portion of the shaft power into electricity, where at least a portion of the electricity is used to supply heat to the reaction of step (b). 13. The method of claim 12, wherein the second heat transfer step comprises at least one refrigerant compressor, wherein at least a portion of the refrigerant compressor power is obtained from a gas turbine and / or generator. 14. The method according to any one of paragraphs. 8-10, where: (i) the reaction of step (b) is carried out in at least one first reactor, (ii) the first reactor includes at least one layer of a dehydrocyclization catalyst, and (iii) the catalytic dehydrocyclization conditions in the first reactor include a first temperature T ₁ in the range of 400 to 650 ° C., a first pressure P ₁ in the range of 100 to 2200 kPa, and a first space velocity in the range of 0.1 to 20 h ⁻¹ . 15. The method according to p. 14, further comprising: (h) withdrawing at least a portion of the tail gas to a second reactor, where the second reactor contains at least one layer of a second dehydrocyclization catalyst, and (i) a tail gas reaction in a second reactor in the presence of a second dehydrocyclization catalyst under second dehydrocyclization conditions, which include a temperature T ₂ in the range of 500 to 675 ° C, an absolute pressure of P ₂ ≤ 30 psi. inch (207 kPa) and hourly average feed rate in the range from 0.1 to 20 h ^-1 , to obtain the product of the second reactor containing additional aromatic hydrocarbons, additional molecular hydrogen and additional non-aromatic hydrocarbons, where: (i) T ₁ is in the range of 450 to 605 ° C., P ₁ is in the range of 37 psi. inch (abs.) (255 kPa) up to 80 psi inch (abs.) (522 kPa), and (ii) T ₂ > T ₁ and / or P ₁ > P ₂ . 16. The method of claim 15, wherein: (i) the first catalyst comprises at least 10 wt.% of the first molecular sieve component and at least 0.005 wt.% of the first dehydrogenating component, (ii) the second catalyst comprises at least 10 wt.% of the second molecular sieve component and at least 0.005 wt.% of the second dehydrogenating component; (iii) each of the first and second molecular sieve components includes (i) one or more of MCM-22, ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM- 48, and (iv) each of the first and second dehydrogenating components includes one or more of Ga, Zn, Cu, Re, Mo, W, La, Fe, Ag, Pt, and Pd. 17. The method according to p. 15, further comprising the removal of molecular hydrogen from the tail gas to enter the second reactor. 18. The method of claim 8, wherein the overhead stream has a Wobbe index in the range of 49.01 to 52.22 MJ / cm ³ and a calorific value in the range of 36.07 to 41.40 MJ / cm ³ . 19. The method according to p. 11, where: (i) the aromatization feed comprises C ₂ hydrocarbons in an amount varying from 10 to 35 mol%, and C ₃₊ hydrocarbons in an amount varying from 40 to 50 mol%, and the first product is recovered in an amount changing by no more than +/- 25%, when (A) the specified amount of raw material for flavoring is in the range from 5 MSCFD (140,000 NM ³ D) to 35 MSCFD (1⋅10 ⁶ NM ³ D), (B) the amount of gaseous feedstock is in the range of 30 (860000 NM ³ D) MSCFD 140 MSCFD (4⋅10 ⁶ NM ³ D) and (C) a gaseous raw material has a molar ratio of C ₂ hydrocarbons to nonaromatic Carbs portly C ₃₊ in the range from 0.5 to 1.5. 20. A system for producing aromatic hydrocarbons, including: (a) at least one pipeline for raw materials intended to be transferred from a hydrocarbon source in an amount of ≥2 million standard cubic feet per day [MSCFD], a gaseous feed comprising C ₂ hydrocarbons and C ₃₊ non-aromatic hydrocarbons; (b) a pretreatment plant in liquid communication with a feed pipe, where a pretreatment plant is configured to produce a predetermined amount, ≥1 MSCFD, pretreated feedstock from a gaseous feedstock, the pretreated feedstock has a predetermined concentration of non-aromatic C ₂₊ hydrocarbons in the range from 55 to 80 mol.% per mole of pre-processed raw materials; (c) a flavoring in fluid communication with a feedstock pretreatment unit, wherein the flavoring is configured to flavor (i) at least a portion of the pre-treated non-aromatic C ₃₊ hydrocarbons of the feed and (ii) at least a portion of the pre-processed C ₂ hydrocarbons _{of the} feed, s obtaining an exit reaction stream including molecular hydrogen, non-aromatic hydrocarbons and aromatic hydrocarbons; and (d) a product recuperator in fluid communication with a flavoring, where the product recuperator is configured to recover at least the first and second products from the effluent reaction stream, where the first product comprises at least a portion of the aromatic hydrocarbons in the effluent reaction stream and the second product comprises a tail gas containing (i) at least a portion of the molecular hydrogen of the effluent reaction stream; and (ii) at least a portion of the non-aromatic hydrocarbons of the effluent reaction stream; where the installation for pre-treatment of raw materials includes: (i) at least one heat transfer step having (A) an input in fluid communication with the feed pipe and (B) an output, the heat transfer step is configured to transfer heat from the gaseous feed in order to condense part of the gaseous feed at a specified dew point, and (ii) at least one vapor-liquid separator, where the vapor-liquid separator has an inlet in fluid communication with the heat transfer step, at least one drum separator and first and second outlets, where: (A) the separator inlet is configured to receive gaseous feed and condensed portions from the heat transfer step, (B) a drum separator is configured to separate gaseous feedstocks and condensed portions to produce a bottoms stream including flavoring feedstocks and an overhead stream including separated gaseous hydrocarbons, (C) a first separator outlet configured to transfer the separated gaseous portion from the pretreatment unit, and (D) the second output of the separator is configured to transfer a predetermined amount of raw material for flavoring in the flavoring. 21. The method of conversion of hydrocarbons, including: (a) providing gaseous feed at a volumetric flow rate of ≥2 million standard cubic feet per day [MSCFD] (58000 NM ³ D), where the gaseous feed includes C ₂ hydrocarbons and non-aromatic C ₃₊ hydrocarbons; (b) monitoring the volumetric flow rate of the gaseous feedstock inversely with the concentration of C ₂ hydrocarbons _{of the} gaseous feed and separating a predetermined amount from the gaseous feed, ≥1 MSCFD (29000 NM ³ D), pretreated feed, where the pretreated feed has a predetermined concentration non-aromatic hydrocarbons C ₂₊ in the range from 55 to 80 mol.% per mole of pre-processed raw materials; (c) reacting the pre-treated feed in the presence of at least one dehydrocyclization catalyst under catalytic dehydrocyclization conditions to produce an exit reaction stream comprising (i) non-aromatic hydrocarbons and (ii) aromatic hydrocarbons, and molecular hydrogen, where aromatic hydrocarbons and molecular hydrogen are formed by dehydrocyclization of at least a portion of the C ₂ hydrocarbons _{of the} pre-treated feed and / or non-aromatic C ₃₊ hydrocarbons of the pre-processed wow raw materials; (d) recovering at least the first and second products from the effluent reaction stream, wherein the first product comprises at least a portion of the aromatic hydrocarbons of the effluent reaction stream and the second product comprises a tail gas comprising (i) at least a portion of the molecular hydrogen of the effluent reaction stream and (ii) at least a portion of non-aromatic hydrocarbons from the effluent; and (e) preparation of aromatization feedstock by pretreatment of ≥2 MSCFD (57000 NM ³ D) gaseous feed, gaseous feed comprises from 0 to 98 mol.% Methane, from 1 to 50 mol.% Hydrocarbons C ₂  and from 1 to 40 mol.% non-aromatic hydrocarbons C _{3 +} where the pre-treatment includes (i) transferring heat from the gaseous feed at least in the first heat transfer step, in order to partially condense the gaseous feed at a specified dew point, and (ii) separating the overhead stream, including gaseous hydrocarbons and bottoms, from the partially condensed feed where the bottoms stream is a raw material for flavoring; in which (iii) the first heat transfer step includes at least one heat exchanger in indirect thermal contact with the refrigerant stream, (iv) the heat exchanger indirectly transfers the first, second and third parts of the heat transferred from the gaseous feed, where the first part is transferred to the refrigerant, the second part is transferred separated gaseous hydrocarbons and the third part is transferred to the aromatization feedstock, and (v) the pre-treatment further includes diverting the refrigerant stream from the first heat transfer step to the second heat transfer step and deriving and some heat from the refrigerant to the second stage heat transfer where the amount of heat does not vary more than +/- 50%, when the condensed portion of the condensed gaseous feedstock having a change in volume of ≤ +/- 10%.

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