Patent Application
20170088484
The systems and methods described herein relate to aromatization of non-aromatic hydrocarbon, to equipment and materials useful in such aromatization, and to the use of such aromatization for, e.g., natural gas upgrading.
Improved production methods have led to an increase in the availability of hydrocarbon gasses of geological origin, e.g., natural gas. Natural gas generally comprises a mixture of one or more hydrocarbon compounds having a number of carbon atoms in the range of from C1 to C9, and can further comprise non-hydrocarbon compounds such as CO2 and/or H2S. Natural gas is used for a wide variety of purposes, e.g., electric power generation, transportation fuel, residential and commercial heating and cooking, and as a feed stock for producing commodity chemicals such as ethylene.
Frequently, natural gas is produced from a geological formation that is located some distance away from the place at which it is consumed. This has stimulated the construction of natural gas pipelines, which transport compressed natural gas from producers to consumers. Some natural gas consumers, e.g., operators of electric power generators, are sensitive the caloric content of the natural gas obtained from the pipeline. When the natural gas has too little caloric content, the electric power generators operate inefficiently. Too large a caloric content can lead to equipment damage, e.g., as a result of overheating the generator's combustion and expansion stages.
To avoid these difficulties, pipeline operators have established specifications for natural gas entering the operator's pipeline. Heating value, which is a measure of the caloric content of a natural gas, is one typical pipeline specification. Another is Wobbe Index, which is closely related to natural gas combustion energy. Two different natural gasses obtained from different sources and having different compositions, but having substantially the same Wobbe Index, will have substantially the same combustion heat output when combusted under substantially the same conditions. As an example, a pipeline operator might require that natural gas entering its pipeline have a heating value of from 36.07 MJ/sm3 to 41.40 MJ/sm3) and a Wobbe Index of from 49.01 MJ/sm3 to 52.22 MJ/sm3).
The heating value and Wobbe Index of a natural gas are strongly influenced by the type and relative amounts of the gas's hydrocarbon compounds. Typically, hydrocarbon compounds having a greater number of carbon atoms provide an increased heat of combustion. Consequently, when a natural gas does not meet the minimum heating value and Wobbe Index specifications, propane enrichment can be used to produce a pipeline-able gas. More typically, the natural gas exceeds the heating value and Wobbe Index specifications. In these cases, a natural gas producer might extract a portion of the gas's relatively-high heating content hydrocarbon compounds, e.g., excess ethane and/or excess propane beyond that needed to achieve the minimum specified heating value and Wobbe Index.
Fractionation and dew point control are conventional methods for producing a pipeline-able natural gas from a natural gas source that does not meet the pipeline's heating value and Wobbe Index specifications. For example, a natural gas of excessive Wobbe Index and/or heating value can be fractionated to produce a bottoms fraction comprising primarily C5+ hydrocarbon, a side stream comprising primarily C3 and C4 hydrocarbon, and an overhead stream comprising primarily methane and ethane. The bottoms fraction is typically conducted away, e.g., for blending with other liquid hydrocarbon. In locations having the facilities to do so, the side stream can be compressed and transported as LPG. In remote locations, the side stream is typically re-injected into the reservoir. The fractionation is configured to produce a methane/ethane mixture that meets pipeline specifications. Since the composition of natural gas and the relevant pipeline specifications vary from location to location, fractionation conditions cannot be readily standardized. In part as a result of this difficulty, fractionation is typically utilized when the natural gas is in the form of a raw natural gas produced from a natural gas well (“gas well gas”), particularly those having a lesser C3+ hydrocarbon content.
Dew point control is typically used when the natural gas has a greater C3+ hydrocarbon content, e.g., when the natural gas is associated gas. Associated gas is a form of natural gas that is typically found with petroleum deposits, e.g., dissolved in the oil or as a free “gas cap” above the oil in a reservoir. To carry out dew point control, a mixture of gaseous and liquid hydrocarbon produced from a reservoir is conducted to a vapor-liquid separation stage. A bottoms product comprising crude oil is conducted away. An overhead stream comprising associated gas is also conducted away, at least a portion of which is conducted to the dew point control stage. Dew point control is carried out by cooling the associated gas to a sufficiently low temperature at a sufficiently elevated pressure to condense a sufficient portion thereof such that the remaining vapor-phase portion meets pipeline specifications at that location. The liquid-phase is disengaged from the vapor phase in, e.g., a separator drum. A pipeline-able natural gas comprising primarily methane and ethane is conducted away from the separator drum as an overhead stream. A bottoms stream comprising primarily C2+ hydrocarbon is also conducted away. It is conventional to convert at least a portion of the bottoms stream to aromatic hydrocarbon, which greatly lessens difficulties associated with transporting the bottoms stream away from remote locations. See, e.g., P.C.T. Patent Application Publication No. WO 2015/084518A1, which is incorporated by reference herein in its entirety.
One aromatization process, disclosed in U.S. Pat. No. 4,855,522, involves converting C2, C3, and C4 hydrocarbon with increased selectivity for aromatic hydrocarbon. The process utilizes a dehydrocyclization catalyst including (a) an aluminosilicate having a silica to alumina molar ratio of at least 5 and (b) a compound of (i) Ga and (ii) at least one rare earth metal. The reference discloses carrying out the aromatization conversion at a space velocity (LHSV) in the range of from 0.5 to 8 hr−1, a temperature ≧450° C. (e.g., 475° C. to 650° C.), a pressure of from 1 bar to 20 bar, and a feed contact time of 1 to 50 seconds. One difficulty observed with this process is that the accumulation of catalyst coke gradually decreases aromatic hydrocarbon yield. One way to overcome this difficulty includes operating the reaction cyclically, e.g., by carrying out the reaction in at least two reactors. The first reactor carries out aromatization (reaction mode) while the second reactor undergoes decoking (regeneration mode), and vice versa. Decreasing the amount of time (the “cycle time”) that a fixed bed reactor is operated in reaction mode before switching to regeneration mode is used to lessen the amount of coke accumulation. However, decreasing cycle time decreases the yield of aromatic hydrocarbon produced by the process, particularly when a significant amount of coke has accumulated. In those cases, it can be necessary to carry out regenerating mode for a relatively long time interval compared to that of aromatization mode. Another way to overcome this difficulty involves utilizing fluidized catalyst beds in the reaction stages. This, however, typically requires specialized catalysts and process conditions, and is difficult to implement in practice, particularly in remote locations.
Still other difficulties encountered when operating conventional process for natural gas aromatization arise from the compositional variation observed for natural gas produced in different locations. Some natural gas is relatively rich in ethane; other natural gas contains much less. It is observed when operating conventional aromatization of C2-C4 hydrocarbon that conditions needed for C2 aromatization typically lead to excessive conversion to coke of the C3 and C4 hydrocarbon. Further, conditions optimized for the aromatization of C3 and C4 hydrocarbon typically exhibit less aromatics yield from the natural gas's C2 hydrocarbon. This situation is worsened by the wide variation in ethane content of natural gas produced from different reservoirs.
One way to overcome this is to tailor the aromatization process so that it is compatible with a particular reservoir. For example, U.S. Pat. No. 8,835,706 discloses aromatization of an ethane-propane feed in two reactors operated in a complex series/parallel arrangement. The feed to the process, which is obtained from natural gas by cryogenically separating methane, is reacted in a first stage operated under conditions which maximize the conversion of propane to aromatics. Ethane is reacted in the second stage to produce additional aromatics. The second stage is operated under conditions which maximize the conversion of ethane to aromatic hydrocarbon. The process can be tailored to a particular natural gas source by, e.g., appropriately sizing the first, reactor, second reactor, and associated interconnection and separations stages. Doing so, however, limits the usefulness of the process to natural gas within a relatively narrow compositional range.
There is therefore a need for improved processes for aromatizing C2-C9 non-aromatic hydrocarbon, such as those extracted from natural gas. More particularly, there is a need for aromatization processes that are capable of operating at relatively long cycle time, and which avoid the complexities of conventional process. Even more desired are aromatization processes that do not require tailoring reactors and associated equipment to the composition of a particular source of C2-C9 hydrocarbon, and especially processes that can be carried out using standardized modules that are readily transported to remote locations.
The invention generally relates to upgrading gaseous hydrocarbon of different composition, as may be produced from various sources at various geographic locations. The sources each comprise methane, C2 hydrocarbon and C3+ non-aromatic hydrocarbon but contain significantly different relative amounts of these components. Controlling the volumetric flow rate of such a gaseous hydrocarbon from the source in an inverse relationship with the hydrocarbon source's C2+ hydrocarbon concentration upstream of a separation stage is found to provide a separated C2+ hydrocarbon stream having desirable features. The separated C2+ hydrocarbon stream features both a substantially-constant volumetric flow rate and a substantially constant C2+ hydrocarbon concentration. Advantageously, it has been discovered that these features are achieved for a wide variety of gaseous hydrocarbon, as may be produced at a wide variety of geographic locations by a wide variety of processes.
This discovery has led to the development of a system and process which includes at least two main functionalities: an aromatization functionality and a product recovery functionality. The aromatization functionality is configured to receive a gaseous aromatization feed having (i) a preselected volumetric flow rate and (ii) a preselected concentration of C2+ non-aromatic hydrocarbon. The aromatization functionality produces a reaction effluent comprising aromatic hydrocarbon produced by dehydrocyclization of at least a portion of the aromatization feed's C2+ non-aromatic hydrocarbon. The product recovery functionality separates at least aromatic hydrocarbon from the reaction effluent. The process is advantageous in that it can be operated at relatively long cycle times without the complexities of conventional process. Since the aromatization functionality is configured for aromatizing a preselected volumetric flow rate of an aromatization feed having a predetermined concentration of C2+ hydrocarbon, the process can be carried out in modules configured for the specified functionalities, e.g., using individual feed pre-treatment, aromatization, and product recovery modules configured for performing those functions. The modules, which can be of a preselected size and capacity, can be readily transported to a wide variety of remote locations, e.g., for natural gas processing.
The advantages of the present techniques are better understood by referring to the following detailed description and the attached drawings, in which:
Although certain aspects of the invention are described in greater detail, the invention is not limited thereto, and this description is not meant to foreclose other aspects within the within the true spirit and scope of the appended claims.
Certain aspects of the systems and techniques described herein relate to hydrocarbon conversion utilizing an aromatizer, a recoverer, and an optional pretreater. The hydrocarbon conversion utilizes a gaseous feed comprising C2 hydrocarbon and C3+ non-aromatic hydrocarbon. The gaseous feed is treated in the feed pretreater, to provide an aromatization feed for aromatization, the gaseous aromatization feed having (i) a preselected volumetric flow rate and (ii) a preselected concentration of C2+ non-aromatic hydrocarbon. The feed pretreatment can be carried out in stages, e.g., by transferring heat away from the gaseous feed in a heat transfer stage in order to condense a portion of the gaseous feed. A separation stage can be used for separating condensate from the remainder of the gaseous feed which comprises gaseous hydrocarbon. The aromatization feed is produced by vaporizing at least a portion of the condensate. Condensate vaporization can be carried out at any convenient location, typically upstream of the dehydrocyclization reaction. The form of heat transfer used is not critical, and can be direct, indirect, or a combination thereof. The form of separation is not critical, and can include one or more fractionators, one or more separator drums, one or more membranes, or a combination thereof. Conventional heat transfer and separation technology can be used, but the invention is not limited thereto. For gaseous feeds available from a wide variety of sources at a wide variety of geographic locations, the amount of gaseous feed condensation can be regulated so that the pretreatment produces a preselected volumetric flow rate of aromatization feed, with the aromatization feed having a preselected concentration of C2+ non-aromatic hydrocarbon. Advantageously, the preselected values are substantially constant, e.g., no significant variation is needed to practice the invention using gaseous feeds available at different geographic locations.
This can be accomplished, e.g., by one or more of (i) decreasing the volumetric flow rate of gaseous feed to the pretreatment at geographic locations where the gaseous feed has a greater concentration of C2+ non-aromatic hydrocarbon, (ii) increasing the volumetric flow rate of gaseous feed at geographic locations where the gaseous feed has a lesser concentration of C2+ non-aromatic hydrocarbon, (iii) decreasing the amount of heat withdrawn from the gaseous feed at geographic locations where the gaseous feed has a greater concentration of C2+ non-aromatic hydrocarbon, (iv) increasing the amount of heat withdrawn from the gaseous feed at geographic locations where the gaseous feed has a lesser concentration of C2+ non-aromatic hydrocarbon, (v) increasing the gaseous feed's pressure at geographic locations where the gaseous feed has a lesser concentration of C2+ non-aromatic hydrocarbon, and (vi) decreasing the gaseous feed's pressure at geographic locations where the gaseous feed has a greater concentration of C2+ non-aromatic hydrocarbon.
Since the preselected volumetric flow rate and the preselected concentration of C2+ non-aromatic hydrocarbon are substantially independent of the gaseous feed utilized, the capacities (e.g., hydraulic capacity) of the pretreatment, reaction, and recovery functionalities can be preselected, in the sense that these capacities can be substantially the same for any geographic location where a suitable gaseous feed may be available. Advantageously, this permits the standardization of the pretreatment, reaction, recovery, and auxiliary functionalities. For example, one or more of these functionalities can be modularized, e.g., a functionality can be carried out in one or more modules of standardized size, shape, and/or capacity. Conventional separators, aromatizers, and recoverers are customized for conversion of a gaseous feed in a particular geographic location. Unlike the conventional case, the pretreater, aromatizer, and recoverer of the invention are standardized. Where a greater conversion capacity is needed in a particular geographic location, fewer standardized pretreaters, aromatizers, and recoverers can be used (e.g., fewer standard modules for one or more of these functionalities). In locations where a greater conversion capacity is needed, a greater number of standardized aromatizers, pretreaters, aromatizers, and recoverers can be used (e.g., more standardized modules for one or more of these functionalities). In other words, one advantage of the invention is that it compensates for variations in needed conversion capacity by increasing or decreasing capacity in discreet, standardized amounts, instead of requiring customization by continuous capacity variation as in the prior art.
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 aromatization feed's non-aromatic hydrocarbon aromatic hydrocarbon and optionally other cyclic hydrocarbon. The number of reactors utilized for carrying 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 plurality of reactors, a portion of the aromatic hydrocarbon produced in a first reactor may be removed before the remainder of the first reactor's effluent is reacted in one or more additional reactors. In other aspects, the first reactor's reaction effluent may be reacted in one or more additional reactors with little or no separation of aromatic hydrocarbon from the first reaction effluent. The presence of the first product's aromatic hydrocarbon and methane in the second stage is less detrimental to additional aromatic hydrocarbon production than may be expected from the teachings of the prior art.
The pretreatment, reaction, recovery, and auxiliary functionalities will now be described in more detail. To avoid doubt about the meaning of certain terms used in this description and appended claims, the following definitions are provided.
Two or more items are, “proximate” to one another when they are spatially close, without regard to whether the spatial relationship places one item underneath, over, or beside another item. Items of definite size and/or shape (e.g., physical components) can be proximate to one another and/or proximate to items that might be of indefinite size and/or shape (e.g., certain chemical reactions).
As used herein, “substantially” or other words of degree are relative modifiers intended to indicate permissible variation from the characteristic so modified. It is not intended to be limited to the absolute value or characteristic which it modifies, but rather possessing more of the physical or functional characteristic than its opposite, and preferably, approaching or approximating such a physical or functional characteristic. Unit conversions to Normal Meters' per Day (“NM3D”) of a gaseous composition, and to m3 per day (“M3PD of a liquid composition, use conversion factors from the Petroleum Engineering Handbook, Vol. VII, Indexes and Standards, 169, Society of Petroleum Engineers, 2007.
The term “Cn” hydrocarbon means hydrocarbon having n carbon atom(s) per molecule, wherein n is a positive integer. The term “Cn+” hydrocarbon means hydrocarbon having at least n carbon atom(s) per molecule. The term “Cn−” hydrocarbon means hydrocarbon having no more than n carbon atom(s) per molecule. The term “hydrocarbon” means a class of compounds containing hydrogen bound to carbon, and encompasses (i) saturated hydrocarbon, (ii) unsaturated hydrocarbon, and (iii) mixtures of hydrocarbons, and including mixtures of hydrocarbon compounds (saturated and/or unsaturated), including mixtures of hydrocarbon compounds having different values of n. A “Gaseous” composition is one that is primarily in the vapor phase.
The terms “alkane” and “paraffinic hydrocarbon” mean substantially-saturated compounds containing hydrogen and carbon only, e.g., those containing ≦1% (molar basis) of unsaturated carbon atoms. As an example, the term alkane encompasses C2 to C20 linear, iso, and cyclo-alkanes. Aliphatic hydrocarbon means hydrocarbon that is substantially free of hydrocarbon compounds having carbon atoms arranged in one or more rings.
The term “unsaturate” and “unsaturated hydrocarbon” refer to one or more C2+ hydrocarbon compounds which contain at least one carbon atom directly bound to another carbon atom by a double or triple bond. The term “olefin” refers to one or more unsaturated hydrocarbon compound containing at least one carbon atom directly bound to another carbon atom by a double bond. In other words, an olefin is a compound which 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 term “aromatics” and aromatic hydrocarbon mean hydrocarbon compounds containing at least one aromatic core.
The term “Periodic Table” means the Periodic Chart of the Elements, as it appears on the inside cover of The Merck Index, Twelfth Edition, Merck & Co., Inc., 1996.
The term “reaction zone” or “reactor zone” mean a location within a reactor, e.g., a specific volume within a reactor, for carrying out a specified reaction. A reactor or reaction stage can encompass 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 stage can include a first zone for carrying out first and second reactions and a second zone for carrying out a third reaction, where the first reaction (e.g., dehydrocyclization) can be the same as or different from the second reaction, and the third reaction (e.g., selective oxidation) can be the same as or different from the second reaction.
“Dehydrocyclization” means removing hydrogen from and cyclizing a non-cyclic hydrocarbon to produce, e.g., one or more of cyclo-paraffin, cyclo-olefin, and aromatic hydrocarbon. The reaction can be carried out in one or more of (i) one step, which includes both dehydrogenation and cyclization; (ii) two steps, e.g., dehydrogenation followed by cyclization of the dehydrogenated intermediate; and (iii) three or more steps, e.g., normal paraffin dehydrogenation, cyclization of the olefinic intermediate, and additional dehydrogenation (aromatization) of the cyclo-olefin intermediate. The dehydrocyclization (including any dehydrogenation carried out in connection with dehydrocyclization) is “non-oxidative” meaning that the reaction is carried out with little if any oxidative coupling of feed hydrocarbon, intermediate hydrocarbon (if any), or dehydrocyclization product.
The term “selectivity” refers to the production (on a weight basis) of a specified compound in a catalytic reaction. As an example, the phrase “a light hydrocarbon conversion reaction has a 100% selectivity for aromatic hydrocarbon” means that 100% of the light hydrocarbon (weight basis) that is converted in the reaction is converted to aromatic hydrocarbon. When used in connection with a specified reactant, the term “conversion” means the amount of the reactant (weight basis) consumed in the reaction. For example, when the specified reactant is C4 paraffinic hydrocarbon, 100% conversion means 100% of the C4 paraffinic hydrocarbon is consumed in the reaction. The yield, on a weight basis, is the product of the conversion and the selectivity.
The term “modularization” means to arrange process equipment and associated piping, valving, instrumentation, controls and support systems into compact and transportable units. Such a unit or plurality thereof can be supported on one or more moveable skids. Modularization desirably provides operational flexibility. Processing capacity is readily increased, e.g., by adding additional modules.
Certain aspects of the invention, which include processing hydrocarbon of geological origin, will now be described in more detail with reference to
At least two streams are conducted away from pretreater 114. The first stream is the aromatization feed, which is produced at a preselected volumetric flow rate, and which has a preselected concentration of C2+ non-aromatic hydrocarbon. The second stream, which comprises mainly methane and ethane, has been found to meet heating value and Wobbe Index specifications of a variety of natural gas pipelines. The first stream is conducted away via conduit 120 (e.g., as condensed liquid and/or vapor) to aromatizer 124. The second stream is conducted away via conduit 115, e.g., for one or more of storage; transportation, e.g., via pipeline; further processing, e.g., cooling/compression to produce a compressed gas and/or LNG; re-injection into hydrocarbon source 102; etc. In addition to the first and second streams, a primarily liquid third stream comprising hydrocarbon is optionally produced and conducted away from the pretreater via conduit 122. The third stream can be generated in aspects in which the pretreater includes at least one fractionator (e.g., one or more distillation columns), where the third stream can be, e.g., a fractionator bottoms stream.
The aromatization feed is conducted via conduit 120 to aromatizer 124, for conversion to aromatic hydrocarbon of at least a portion of the feed's C2+ non-aromatic hydrocarbon. A reaction effluent is conducted away via conduit 126. The reaction effluent comprises (i) non-aromatic hydrocarbon and (ii) aromatic hydrocarbon and molecular hydrogen, at least a portion of the aromatic hydrocarbon and the molecular hydrogen (and optionally a portion of the non-aromatic hydrocarbon) being formed during the dehydrocyclization.
The reaction effluent is conducted via conduit 126 to recoverer 128, for recovering at least a first product and a tail gas from the reaction effluent. The first product comprises at least a portion of the reaction effluent's aromatic hydrocarbon. The tail gas comprises (i) at least a portion of the reaction effluent's molecular hydrogen and (ii) at least a portion of the reaction effluent's non-aromatic hydrocarbon. The tail gas, which typically comprises primarily molecular hydrogen, methane, and ethane, is conducted away via conduit 130, e.g., for one or more of storage; transportation, such as via pipeline; further processing, e.g., cooling/compression to produce a compressed gas and/or LNG; re-injection into hydrocarbon source 102; etc. Since the tail gas has been found to meet heating value and Wobbe Index specifications of a variety of natural gas pipelines, in certain aspects (not shown) at least a portion of the tail gas is combined with at least a portion of the pretreater's second stream, e.g., for pipeline transportation of the combined stream. In the aspects illustrated schematically in
As shown in
Tail gas, in addition to molecular hydrogen, methane, and ethane, can further comprise sulfur compounds such as mercaptan and/or hydrogen sulfide. Typically, this is the case when the aromatization feed includes organosulfur compounds, H2S, or both. If desired, the tail gas can be treated (not shown in
All or a portion of the tail gas can be utilized for on-site power generation, e.g., by combusting tail gas in one or more gas turbines, the gas turbine being used to power one or more electric generators of power generator 140. The electricity can be used in the process, e.g., for heating aromatizer 124, e.g., to provide heat for the (typically endothermic) dehydrocyclization reaction. Alternatively, or in addition, the electricity generated can be introduced into one or more power grids. Alternatively or in addition to the gas turbine, one or more combustion engines can be used for combusting tail gas, e.g., to power the electric generator.
In particular aspects illustrated in
Alternatively or in addition to any electric power as may be derived from tail gas combustion (or the combustion of a tail gas component such as molecular hydrogen) in power generator 140, thermal energy produced by such combustion or a portion thereof can be transferred to aromatizer 124 via thermal conduit 146. Doing so can, e.g., provide heat for the dehydrocyclization. One or more heat recovery steam generators, including those commonly adapted for gas turbine service, can be used for this purpose. The generated steam can be used as a heat transfer fluid in thermal conduit 146, for example, for providing heat to aromatizer 124. Alternatively or additionally, a portion of the separated molecular hydrogen of conduit 144, or a portion of the tail gas of conduit 130, or both, may be used to provide heat to aromatizer 124. For example, one or more of these streams may be combusted in a heater or chemically reacted in an exothermic reaction that releases heat for the aromatization reaction (or reactor(s)).
Continuing with reference to
The block diagram of
Part or all of the system 100 may be constructed as a modular system that may be quickly moved and set up at a location, used, then broken down and moved to a new location when no longer needed.
Certain gaseous feeds that are suitable for producing the specified aromatization feed will now be described in more detail. The invention is not limited to these gaseous feeds, and this description is not meant to foreclose the use of other gaseous feeds within the broader scope of the invention.
The gaseous feed, which in aspects illustrated in
The gaseous feed typically contains C3 and/or C4 hydrocarbon e.g., (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 can contain C5+ hydrocarbon, the amount of C5+ hydrocarbon when present is typically small, e.g., ≦30 wt. %, such as ≦10 wt. %, or ≦01 wt. %. Typically, the gaseous feed contains ≦15 wt. % of C6+ saturated hydrocarbon, e.g., ≦5 wt. %.
The gaseous feed can contain methane, e.g., ≧1 wt. % methane, such as ≧10 wt. %, or ≧20 wt. %, or ≧60 wt. %. Even though methane is a diluent, i.e., it typically does not react to produce aromatic hydrocarbon or catalyst coke in the presence of the specified dehydrocyclization catalyst under the specified reaction conditions, its presence is beneficial. It is believed that this benefit results at least in part from a decrease in the partial pressure of the C2-C9 hydrocarbon compounds that is achieved when the gaseous feed further includes methane. Decreasing the partial pressure of the C2-C9 hydrocarbons, particularly the partial pressure of the C2-C5 hydrocarbon, has been found to lessen the amount of catalyst coke formed under the specified dehydrocyclization process conditions. Typically, the gaseous feed includes a total of ≦10 wt. % of impurities such as CO, CO2, H2S, and total mercaptan; e.g., ≦1 wt. %, or ≦0.1 wt. %. Although not typically found in gaseous feeds separated from raw natural gas, certain gaseous feeds include molecular hydrogen, e.g., ≧1 wt. % molecular hydrogen based on the weight of the gaseous feed 118, such as ≧5 wt. %.
The gaseous feed includes ethane, typically in an ≧1 wt. %, based on the weight of the gaseous feed, e.g., ≧5 wt. %, or ≧10 wt. %, such as in the range of from 10 wt. % to 40 wt. %. Suitable gaseous feed include those containing a major amount of ethane that is >50 wt. %, such as ≧75 wt. %, or ≧90 wt. %, or ≧95 wt. %. One representative gaseous feed includes (i) ≧10 wt. % ethane, such as in the range of from 10 wt. % to 40 wt. %; and further includes (ii) 1 wt. % to 40 wt. % methane, (iii) 20 wt. % to 50 wt. % propane, and (iv) 20 wt. % to 50 wt. % butanes. In other aspects, the amount of ethane in the gaseous feed is <1 wt. %, e.g., ≦0.1 wt. %, or ≦0.1 wt. %.
Although the gaseous feed can contain unsaturated C2+ hydrocarbon, such as C2-C5 unsaturated hydrocarbon, the amount of these unsaturated compounds is typically ≦20 wt. %, e.g., ≦10 wt. %, such as ≦1 wt. %, or ≦0.1 wt. %, or in the range of from 0.1 wt. % to 10 wt. %. The gaseous feed 118 can be substantially-free of non-aliphatic hydrocarbon. More particularly, the gaseous feed 118 can be substantially-free of aromatic hydrocarbon, where substantially-free in this context means <1 wt. % based on the weight of the gaseous feed 118, such as ≦0.1 wt. %, or ≦0.01 wt. %, or ≦0.001 wt. %.
In certain aspects, the source of the gaseous feed includes natural gas, e.g., raw natural gas (“raw gas”). Natural gas is (i) a mixture including hydrocarbon, (ii) primarily in the vapor phase at a temperature of 15.6° C. and a pressure of 1.016 bar (absolute), and (iii) withdrawn from a geologic formation. Natural gas can be obtained, e.g., from one or more of petroleum deposits, coal deposits, and shale deposits. One suitable raw natural gas includes 3 mole % to 70 mole % methane, 10 mole % to 50 mole % ethane, 10 mole % to 40 mole % propane, and 5 mole % to 40 mole % butanes and 1 mole % to 10 mole % of total C5 to C9 hydrocarbon. In certain aspects, ≧50 wt. % of the gaseous feed includes natural gas, such as raw natural gas, e.g., ≧75 wt. %, or ≧90 wt. %, or ≧95 wt. %.
As described with respect to
In certain aspects, the gaseous feed includes one or more of (i) gas obtained from a natural gas well (“Gas Well”, Non-associated”, or “Dry” gas), (ii) natural gas obtained from a condensate well (“Condensate Well Gas”), and (iii) casing head gas (“Wet” or “Associated” gas). For example, Associated Gas can comprise 0 mole % to 95 mole % of methane, 5 mole % to 50 mole % of ethane, 2 mole % to 40 mole % of propane, 0.1 mole % to 30 mole % of i-butane, 1 mole % to 30 mole % of n-butane, and 0.05 mole % to 25 mole % of i-pentane. Table 1 includes typical raw gas compositional ranges (mole %) and, parenthetically, typical average composition (mole %) of certain raw gases that may be found in these streams.
One suitable gaseous feed includes ≧75 wt. % Associated Gas, based on the weight of the gaseous feed, e.g., ≧90 wt. %, or ≧95 wt. %. In conventional petroleum production, the lack of effective natural transportation facilities, e.g., the lack of natural gas liquefaction and/or pipeline facilities, can result in Associated Gas being stranded at or near the reservoir. The stranded Associated Gas may be uneconomical to transport, and, thus, may be flared, which is undesirable. Moreover, even in locations where pipeline facilities are available, Associated Gas may be excluded from the pipeline because it typically exceeds one or more pipeline specifications, e.g., ≦12 wt. % ethane, ≦5 wt. % propane, ≦2 wt. % butanes, a Wobbe Index of from 49.01 MJ/sm3 to 52.22 MJ/sm3), and a heating value of from 36.07 MJ/sm3 to 41.40 MJ/sm3).
Since methane is not detrimental to the process, and is in at least some aspects beneficial, the techniques described herein can at least partially obviate the need for costly and inefficient cryogenic methane separation facilities. Typically, obtaining the gaseous feed from the hydrocarbon source does not include (i) exposing the gaseous feed, the hydrocarbon source, or any intermediate thereof to a temperature ≦−37° C., e.g., ≦−46° C., such as ≦−60° C. Certain aspects of the techniques do not include cryogenic processing, e.g., cryogenic methane separation is not used.
The techniques are therefore particularly advantageous in remote or under-developed locations, where (i) the lack of cryogenic methane separation facilities limits the utility of conventional natural gas aromatization processes, (ii) the lack of a pipeline or natural gas production infrastructure, may result in significant quantities of light hydrocarbon being flared or burned as fuel, and (iii) Associated Gas remains stranded at a remote location for lack of pipeline facilities or a failure to meet one or more specifications of an available pipeline.
The gaseous feed is pretreated to produce the specified aromatization feed. Certain forms of pretreatment, illustrated schematically in
Pretreater 114 utilizes gaseous feed provided via conduit 118 to produce a aromatization feed for aromatizer 124 at a preselected amount ≧1 MSCFD (29,000 NM3D), e.g., ≧10 MSCFD (290,000 NM3D), such as in the range of from 2 MSCFD (57,000 NM3D) to 80 MSCFD (2.3·106 NM3D), or in the range of from 10 MSCFD (290,000 NM3D) to 30 MSCFD (860,000 NM3D). The aromatization feed has a preselected C2+ non-aromatic hydrocarbon concentration in the range of from 15 mole % to 90 mole % per mole of aromatization feed, typically 55 mole % to 80 mole %. Typically, the gaseous feed is provided to the pretreater in an amount in the range of from 10 MSCFD (290,000 NM3D) to 150 MSCFD (4.3·106 NM3D).
In the aspects illustrated in
The aromatization feed is conducted away from separator drum 203 as a bottoms stream via conduit 120. In order to provide an aromatization feed to aromatizer 124 having the specified preselected C2+ non-aromatic hydrocarbon concentration and the specified preselected volumetric flow rate, the separation in drum 203 is typically carried out at a predetermined dew point in the range of from −40° F. (−40° C.) to −20° F. (−28.9° C.), typically at a pressure in the range of from about 300 psia (2068 kPa) to about 2000 psia (13,800 kPa), e.g., 400 psia (2760 kPa) to 700 psia (4830 kPa).
In the aspects shown in
In certain aspects, not shown in
When operating in accordance with these aspects, the quantity of heat transferred from the gaseous feed to the refrigerant in the first heat transfer stage typically varies by no more than +/−50% (substantially equivalent to a compressor power variation of no more than +/−50%) over a very wide variety of gaseous feeds obtained from a very wide variety of hydrocarbon sources. Typically, this variation is achieved when the condensed portion of the gaseous feed has a volume variation of ≦+/−25%, e.g., ≦+/−10%, or +/−5%. Further decreases in this variation can be achieved, e.g., no more than +/−40%, such as no more than +/−30%, provided (A) the preselected amount of aromatization feed is in the range of from 5 MSCFD (140,000 NM3D) to 35 MSCFD (1·106 NM3D), (B) the amount gaseous feed is in the range of from 30 MSCFD (860,000 NM3D) to 140 MSCFD (4·106 NM3D), and (C) the gaseous feed has a molar ratio of C2 hydrocarbon to C3+ non-aromatic hydrocarbon in the range of from 0.5 to 1.5. This effect is a significant advantage in standardization of pretreater refrigeration capacity, e.g., since modularization is less difficult to achieve when pretreater refrigeration capacity is substantially independent of the choice of gaseous feed.
In certain aspects, the pretreatment is configured to advantageously decrease variations recoverer capacity as might otherwise be needed when changing from the first gaseous feed to the second gaseous feed. Fewer aromatic hydrocarbon compounds are produced (molar basis) by dehydrocyclization of the aromatization feed's C2 hydrocarbon than are produced by dehydrocyclization of the aromatization feed's C3+ non-aromatic hydrocarbon. Consequently, gaseous feeds having a greater value of M will need a recoverer of comparatively lesser capacity than for those having a lesser value of M. It has been found that this effect can be at least partially overcome by operating the pretreatment to produce an aromatization feed having a preselected concentration of C3+ hydrocarbon within a relatively narrow range. For example, it has been observed that when the aromatization feed comprises C2 hydrocarbon in an amount in the range of from 10 mole % to 35 mole % and C3+ hydrocarbon in an amount in the range of from 40 mole % to 50 mole %, that the first product is recovered in an amount that varies by no more than +/−25%. Further decreases in this variation can be achieved, e.g., no more than +/−20%, such as no more than +/−15%, when (A) the preselected amount of aromatization feed is in the range of from 5 MSCFD (140,000 NM3D) to 35 MSCFD (1·106 NM3D), (B) the amount of gaseous feed is in the range of from 30 MSCFD (860,000 NM3D) to 140 MSCFD (4·106 NM3D), and (C) the gaseous feed has a molar ratio of C2 hydrocarbon to C3+ non-aromatic hydrocarbon in the range of from 0.5 to 1.5. This effect is a significant advantage in standardization of recoverer capacity, e.g., since modularization is less difficult to achieve when the recoverer capacity is substantially independent of the choice of gaseous feed.
The feed pretreater 114 may be constructed onto a first set of modules, wherein the division of the units into the separate modules is based on the amount of flow. For example, the separator drum 203, and supporting units, may be constructed into a single skid for small scale applications, e.g., less than about 10 million standard cubic feet per day (MSCFD) (290,000 NM3D), less than about 5 MSCFD (140,000 NM3D), or less than about 1 MSCFD (29,000 NM3D).
For larger applications, e.g., greater than about 20 MSCFD (570,000 NM3D), greater than about than about 35 MSCFD (1·106 NM3D), or greater than about 50 MSCFD (1.4·106 NM3D), first heat transfer stage 201, second heat transfer stage 202, and separator drum 203 may be separated among two or more modules. A typical skid has a gross weight ≦200,000 pounds (91,000 kg), e.g., ≦80,000 pounds (36,300 kg), such as in the range of from 10,000 pounds (4500 kg) to 200,000 pounds (91,000 kg), or 50,000 pounds (22,700 kg) to 150,000 pounds (68,000 kg); a width ≦15 feet (4.6 m), e.g., ≦102 inches (2.6 m); and a height ≦15 feet (4.6 m), e.g., ≦14.6 feet (4.45 m). Those skilled in the art will appreciate that these arrangements are only one example, and any number of different arrangements and systems as may be used for the feed pretreater 114.
Certain forms of aromatizer 124 will now be described in more detail with reference to
The reactor 224 uses a dehydrocyclization process to convert to aromatic hydrocarbon at least a portion of the aromatization feed's C2+ non-aromatic hydrocarbon. In this form of dehydrocyclization, the aromatization feed obtained from conduit 120 is heated in heat transfer stages 226 and 228 by transferring heat from the reaction effluent of 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 a portion of the aromatization feed's C2-C9 non-aromatic hydrocarbon to aromatic hydrocarbon and molecular hydrogen. Typically, the dehydrocyclization catalyst includes ≧10 wt. % of a molecular sieve component and ≧0.005 wt. % of a dehydrogenation component.
When the molecular sieve component and dehydrogenation component together include less than 100 wt. % of the catalyst, ≧90 wt. % of the remainder of the catalyst can include a matrix component, such as ≧99 wt. % of the remainder. The catalyst typically includes the molecular sieve component in an amount ≧20 wt. %, based on the weight of the catalyst, e.g., ≧25 wt. %, such as in the range of from 30 wt. % to 99.9 wt. %. In certain aspects, the molecular sieve component includes aluminosilicate, e.g., ≧90 wt. % of at least one aluminosilicate. The aluminosilicate can be an un-substituted aluminosilicate, a substituted aluminosilicate, or a combination thereof. For example, the aluminosilicate can be in a form where at least a portion of its original metal has been replaced, e.g., by ion exchange, with other suitable metal (typically metal cation) of Groups 1-13 of the Periodic Table. Typically, the aluminosilicate includes zeolite aluminosilicate, e.g., ≧90 wt. % of at least one zeolite based on the weight of the aluminosilicate. The term zeolite includes those in which at least part of the aluminum is replaced by a different trivalent metal, such as gallium or indium.
The molecular sieve component typically includes ≧90 wt. % of one or more of the specified molecular sieves, e.g., ≧95 wt. %. In certain aspects, the molecular sieve component includes at least one zeolite molecular sieve, e.g., ≧90 wt. % zeolite, such as ≧95 wt. %, based on the weight of the molecular sieve component. Although, the molecular sieve component can consist essentially of or even consist of zeolite, in alternative aspects the zeolite(s) is present in the molecular sieve component in combination with other (e.g., non-zeolitic) molecular sieve. The zeolite can be one that is in hydrogen form, e.g., one that has been synthesized in the alkali metal form, but is then converted from the alkali to the hydrogen form. Typically the zeolite is one having a medium pore size and a Constraint Index of 2-12 (as defined in U.S. Pat. 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 and mixtures and intermediates thereof such as ZSM-5/ZSM-11 admixture. Optionally, the zeolite is one including at least one set of pores of substantially uniform size extending through the molecular sieve, wherein geometric mean of the cross-sectional dimensions of each of the sets of pores is >5 Å, or >5.3 Å, e.g., ≧5.4 Å such as ≧5.5 Å, or in the range of 5 Å to 7 Å, or 5.4 Å to 7 Å. ZSM-5 and/or ZSM-12 are suitable, particularly H-ZSM-5. For example, the molecular sieve component can include ≧90 wt. % of (A) ZSM-5 and/or (B) ZSM-12, based on the weight of the molecular sieve component, e.g., ≧95 wt. % of H-ZSM-5. In certain aspects, the molecular sieve has a relatively small crystal size, e.g., small crystal ZSM-5, meaning ZSM-5 having a crystal size <0.05 micrometers (μm), such as in the range of 0.02 μm to 0.05 μm. Small crystal ZSM-5 and the method for determining molecular sieve crystal size are disclosed in U.S. Pat. No. 6,670,517, which is incorporated by reference herein in its entirety.
In other aspects, the molecular sieve component includes at least one molecular sieve of the MCM-22 family, e.g., MCM-22 alone or in combination with other molecular sieve such as one or more of the specified zeolites. The MCM-22 family includes those molecular sieves having an X-ray diffraction pattern including d-spacing maxima 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 are obtained by standard techniques using the K-alpha doublet of copper as incident radiation and a diffractometer equipped with a scintillation counter and associated computer as the collection system. Examples of suitable MCM-22-family molecular sieve is described in U.S. Pat. Nos. 4,954,325; 4,439,409 (PSH-3); U.S. Pat. No. 4,826,667 (SSZ-25); U.S. Pat. No. 6,077,498 (ITQ-1); U.S. Pat. No. 5,250,277 (MCM-36); U.S. Pat. No. 5,236,575 (MCM-49); and U.S. Pat. No. 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 MCM-22-family molecular sieve can be used as well.
When the molecular sieve component includes at least one aluminosilicate, the aluminosilicate's silica:alumina ratio (substantially the same as the aluminosilicate's Si:Al2 atomic ratio) is typically ≧2, e.g., in the range of from 5 to 100. The silica:alumina ratio is meant to represent the Si:Al2 atomic ratio in the rigid anionic framework of the crystalline aluminosilicate. Alternatively or in addition, the catalyst can be made more resistant to deactivation (and increase aromatic hydrocarbon yield) by including phosphorous with the molecular sieve component. When used, the amount of phosphorous is typically ≧1 wt. % based on the weight of the molecular sieve component. For example, when the molecular sieve component includes aluminosilicate, the phosphorous:aluminum atomic ratio can be in the range of from 0.01 to 1. Zeolite having a higher silica:alumina ratio can be utilized when a lower catalyst acidity is desired, e.g., in the range of from 44 to 100, such as from 50 to 80, or 55 to 75.
In addition to the molecular sieve component, the catalyst includes ≧0.005 wt. %, based on the weight of the catalyst, of a dehydrogenation component, e.g., at least one dehydrogenation metal. The dehydrogenation component can include one or more neutral metals selected from Groups 3 to 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 oxides, sulfides and/or carbides of these metals. For example, the dehydrogenation component can be Ga, Zn, or a combination thereof, optionally supported on a catalyst including ZSM-5 as the molecular sieve component.
Typically, the dehydrogenation component includes ≧90 wt. % of the one or more of the specified dehydrogenation metals and/or oxide thereof, e.g., ≧95 wt. %, or ≧99 wt. %. For example, the dehydrogenation component can include ≧90 wt. % of (A) Ga and/or (B) Zn, including oxides thereof. Typically, the catalyst includes ≧0.01 wt. % of the dehydrogenation component, based on the weight of the catalyst, e.g., ≧0.1 wt. % of the dehydrogenation component, such as ≧0.5 wt. %, or ≧1 wt. %.
Those skilled in the art will appreciate that when the dehydrogenation component includes one or more metals of greater catalytic dehydrogenation activity, e.g., Pt, and/or Pd, a lesser amount of dehydrogenation component is needed, e.g., in the range of 0.005 wt. % to 0.1 wt. %, based on the weight of the catalyst, such as 0.01 wt. % to 0.6 wt. %, or 0.01 wt. % to 0.05 wt. %. When the dehydrogenation component includes one or more metals of lesser dehydrogenation activity, e.g., one or more of Ga, In, Zn, Cu, Re, Mo, and W, a greater amount of dehydrogenation component is needed, e.g., in the range of 0.05 wt. % to 10 wt. %, based on the weight of the catalyst, such as 0.1 wt. % to 5 wt. %, or 0.5 wt. % to 2 wt. %.
The dehydrogenation component can be provided on, in, or proximate to the catalyst in any manner, for example by conventional methods such as impregnation or ion exchange. At least part of the dehydrogenation metal may also be present in the crystalline framework of the molecular sieve. For one representative catalyst, (i) the dehydrogenation component includes ≧95 wt. % of (A) Ga and/or (B) Zn, and (ii) the first molecular sieve component includes ≧95 wt. % of H-ZSM-5.
In certain aspects, the dehydrogenation component includes ≧99 wt. % of one or more of Ga, Zn, and In, and the molecular sieve component includes ≧99 wt. % of ZSM-5-type zeolite that has been impregnated with the dehydrogenation metal component and/or ion exchanged with the dehydrogenation metal component. For example, the catalyst can include Ga-impregnated and/or In-impregnated H-ZSM-5, Ga-exchanged and/or In-exchanged H-ZSM-5, H-gallosilicate of ZSM-5 type structure and H-galloaluminosilicate of ZSM-5 type structure. 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 present in the zeolite channels in close vicinity to the zeolitic protonic acid sites. While not wishing to be bound by any theory or model, the tetrahedral or framework Al and/or Ga is believed to contribute to acid function of the catalyst and octahedral or non-framework Ga and/or In is believed to contribute to the dehydrogenation function of the catalyst. Although typically the zeolite is impregnated or ion-exchanged with the dehydrogenation metal, other forms of zeolite can be used, such as H-galloaluminosilicate of ZSM-5 type structure having framework (tetrahedral) Si/Al and Si/Ga atomic ratios of about 10:1 to 100:1 and 15:1 to 150:1, respectively, and non-framework (octahedral) Ga of about 0.5 wt. % to 0 wt. %.
Besides the molecular sieve component and dehydrogenation component, the catalyst can further include an optional matrix component, e.g., one or more inorganic binders. The amount of matrix component is not critical. When present, the amount of matrix component is typically in the range of 0.01 times the weight of the molecular sieve component to about 0.9 times the weight of the molecular sieve component, e.g., in the range of 0.02 to 0.8. The matrix component can include active materials, such as synthetic or naturally occurring zeolites. Alternatively, or in addition, the matrix component can include clays and/or oxides such as alumina, silica, silica-alumina, zirconia, titania, magnesia or mixtures of these and other oxides. The matrix component can include naturally occurring materials and/or materials in the form of gelatinous precipitates or gels including mixtures of silica and metal oxides. Clays may also be included with the oxide type binders to modify the mechanical properties of the catalyst or to assist in its manufacture.
Alternatively or in addition, the matrix component can include one or more substantially inactive materials. Inactive materials suitably serve as diluents to control the amount of conversion so that products may be obtained economically and orderly without employing other means for controlling the rate of reaction. Alternatively or in addition to any phosphorous added to or impregnated into the molecular sieve component, the matrix component can optionally include phosphorous, e.g., to lessen catalyst acidity. Those skilled in the art will appreciate that lessening catalyst acidity decreases the amount of catalyst coke produced during the catalytic conversion of the feed's light hydrocarbon to aromatic hydrocarbon. Suitable phosphorous-containing matrices are disclosed in U.S. Pat. No. 5,026,937, which is incorporated by reference herein in its entirety. The matrix component is optional. In certain aspects, the catalyst is substantially-free of matrix, e.g., contains ≦1 wt. % of matrix, such as ≦0.1 wt. %. In particular, the catalyst can be substantially free of binder, e.g., contains ≦1 wt. % of binder, such as ≦0.1 wt. %. For example, the catalyst's molecular sieve component can includes ≧95 wt. % of self-bound bound molecular sieve, e.g., ≧95 wt. % of self-bound ZSM-5, and in particular small crystal H-ZSM-5.
The catalyst can be one that has been subjected to one or more treatments, e.g., a selectivation treatment to increase selectivity for producing desired aromatic hydrocarbon compounds such as para-xylene. For example, the selectivation can be carried out before introduction of the catalyst into the reactor and/or in-situ in the reactor, e.g., by contacting the catalyst with a selectivating agent, such as at least one organosilicon in a liquid carrier and subsequently calcining the catalyst at a temperature of 350° C. to 550° C. This selectivation procedure can be repeated two or more times and alters the diffusion characteristics of the catalyst such that the formation of para-xylene over other xylene isomers is favored. Such a selectivation process is described in detail in U.S. Pat. Nos. 5,633,417 and 5,675,047.
Typically, the catalyst has a surface area as measured by nitrogen physisorption in the range of from 100 m2/g to 600 m2/g, e.g., in the range of from 200 m2/g to 500 m2/g. When the catalyst includes aluminosilicate which includes phosphorous, the phosphorous:aluminum atomic ratio is typically in the range of from 0.01 to 0.5. For example, the catalyst can contain ≧10 wt. % of phosphorous-modified alumina, such as ≧15 wt. %, or in the range of from 10 wt. % to 20 wt. %.
During dehydrocyclization, at least a portion of the C2+ non-aromatic hydrocarbon of the specified aromatization feed of conduit 120 is converted to aromatic hydrocarbon, molecular hydrogen, and optionally additional non-aromatic hydrocarbon. The catalytic dehydrocyclization conditions can include exposing the feed to a temperature in the range of from 400° C. to 650° C., a pressure in the range of from 100 kPa to 2200 kPa. Typically, the catalytic dehydrocyclization conditions further include a weight hourly space velocity (WHSV) ≧0.1 hr−1. More typically, the catalytic dehydrocyclization conditions include a temperature in the range of from 500° C. to 625° C., a pressure in the range of from 30 psia (207 kPa) to 80 psia (522 kPa). WHSV can be in the range of from 0.1 hr−1 to 20 hr−1. Typically, the WHSV of C2+ hydrocarbon (the “C2+ WHSV”) in the specified aromatization feed with respect to the dehydrocyclization catalyst is in the range of from 0.1 hr−1 to 20 hr−1, e.g. 0.2 hr−1 5 hr−1, or 0.3 hr−1 to 1 hr−1. The C2+ WHSV is the hourly rate of the C2+ hydrocarbon (in grams per hour) exposed to the second catalyst per gram of the dehydrocyclization catalyst. The reaction is typically endothermic. Generally, the average temperature drop across the reaction zone is ≦600° C., more typically in the range of from 20° C. to 200° C., e.g., in the range of from 50° C. to 150° C.
The reaction effluent of conduit 126 from reactor 224 may be cooled by exchanging heat with the reactor feed stream of conduit 120 through the heat transfer stages 226 and 228. Further cooling may be provided by flowing the reaction effluent of conduit 126 through other coolers, e.g., air cooler 236. The reaction effluent is conducted away from aromatizer 124 via conduit 126 to recoverer 128, e.g., for recovery of at least a portion of the reaction effluent's aromatics.
When the aromatizer is operated using the specified aromatization feed under the specified conditions, the dehydrocyclization typically results in (i) ≧25 wt. % conversion of the aromatization feed's C2+, e.g., ≧25%, such as ≧50%, or ≧75%; and (ii) a C3+ conversion ≧50%, e.g., ≧75%, or ≧90%, or ≧95%, or ≧98%. The reaction effluent typically includes (i) unreacted feed, including unconverted diluent, and (ii) vapor-phase and/or liquid-phase products of the dehydrocyclization reaction and other reaction pathways, product phase being that subsisting at the outlet of the downstream-most dehydrocyclization reaction zone of reactor 224. Typically, the reaction effluent does not include solid (e.g., catalyst coke) and/or semi-solid (e.g., catalyst coke precursors) products of dehydrocyclization and other reaction pathways, as these generally remain in the first stage until they are removed during regeneration mode.
The reaction effluent typically comprises ≧10 wt. % aromatic hydrocarbon, molecular hydrogen, ≧1 wt. % ethane, 1 wt. % to 40 wt. % methane, ≦2 wt. % propane, and ≦1 wt. % butanes. The amount of ethane in the reaction product can be, e.g., ≧5 wt. %, such as ≧10 wt. %, particularly when operating the dehydrocyclization at a relatively high temperature within the specified temperature range. Since ethane is typically produced by hydrogenolysis of the aromatization feed's C3+ non-aromatic hydrocarbon during the dehydrocyclization, the reaction effluent generally contains ethane and methane even when the aromatization feeds contains ≦1 wt. % of ethane and/or ≦1 wt. % of methane.
Certain forms of recoverer 128 are illustrated schematically in
In the aspects illustrated in
The first product typically comprises ≧50 wt. % of the reaction effluent's aromatic hydrocarbon, e.g., ≧75 wt. %, such as ≧90 wt. %, or ≧95 wt. %. The recovery of the first product from the reaction effluent's aromatic hydrocarbon is typically carried out at a recovery rate in the range of from 200 barrels per day (“BPD”) (3.2 M3PD) to 10,000 BPD (160 M3PD), e.g., 500 BPD (7.9 M3PD) to 5000 BPD (79 M3PD), such as 1000 BPD (16 M3PD) to 4000 BPD (64 M3PD). The tail gas typically comprises ethane and molecular hydrogen, and typically further comprises methane. For example, the tail gas can comprise 1 wt. % ethane, e.g., ≧5 wt. %, such as ≧10 wt. %. In a particular aspect, the tail gas comprises (A) ≧75 wt. % of the reaction effluent's ethane, e.g., ≧90 wt. %, such as ≧95 wt. %; and (B) ≦10 wt. % of the reaction product's aromatic hydrocarbon, e.g., ≦5 wt. %, such as ≦1 wt. %. Typically, the tail gas further comprises (a) ≧10 wt. % of the reaction effluent's molecular hydrogen, e.g., ≧25 wt. %, such as ≧50 wt. %, or ≧95 wt. %, and/or ≧10 wt. % of the reaction effluent's methane, e.g., ≧25 wt. %, such as ≧50 wt. %, or ≧95 wt. %.
The aromatization of the aromatization feed is not limited to a single reactor zone, or a single reactor, or a single reaction stage. Dividing the reaction among multiple zones and/or reactors may allow for more efficient production of the cyclic products, for example, with higher yields and longer periods before coking. Further, dividing the reactor among multiple zones and multiple vessels lessens the difficulties associated with construction of a higher throughput module system, for example, having a throughput of 50 MSCFD (1.4·106 NM3D) or greater.
Certain aspects of the invention which include an aromatizer having first and second reaction stages are illustrated schematically in
As shown in
Heat is transferred away from the second reaction effluent in heat transfer stages 326, 336, and 328, and is conducted via conduit 325 to recoverer 128 for recovery of a second tail gas and a second product. As shown in
The aromatizer's second stage typically converts ≧5 wt. % of the tail gas's ethane, e.g., ≧10 wt. %, such as ≧20 wt. %. Since the first product typically contains ≦5 wt. % of C3+ hydrocarbon, e.g., ≦2 wt. % propane, and ≦1 wt. % butanes, little if any conversion of C3+ hydrocarbon occurs in reactor 334. This effect desirably decreases the accumulation of catalyst coke in stage 2.
Surprisingly, it has been found that appreciable conversion of the first product's ethane to aromatic hydrocarbon can be achieved with less or substantially no removal of the first reaction effluent's aromatic hydrocarbon upstream of the second stage, even when the reaction temperature of the second stage exceeds that of the first stage. Accordingly, in certain aspects that are not shown in
As an example of such alternative aspects, the feed to reactor 334 can include ≧90 wt. % of the reaction effluent of reactor 224, e.g., ≧95 wt. %, such as ≧99 wt. %. Although little if any compositional separation between reactors 224 and 334 is used in these aspects, it is within their scope to divide the entirety of first reaction effluent into two or more streams, with one stream including ≧50 wt. % of the entirety of the first reaction effluent being conducted to reactor 334, e.g., ≧75 wt. %, such as ≧90 wt. %, or ≧95 wt. %; and one or more additional streams being conducted away from the process. It is also within the scope of these aspects to separate at least a portion of the first reaction effluent's aromatic hydrocarbon upstream of 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. % of the first reaction effluent's aromatic hydrocarbon, e.g., ≧75 wt. %, such as ≧90 wt. %, or ≧95 wt. %; and ≧50 wt. % of the first reaction effluent's molecular hydrogen, such as ≧90 wt. %, or ≧95 wt. %. Also typical of these aspects is that the portion of the first reaction effluent that is reacted under dehydrocyclization conditions in reactor 334 includes ≧50 wt. % of the first reaction effluent's ethane, e.g., ≧75 wt. %, such as ≧90 wt. %, or ≧95 wt. %; and ≧50 wt. % of any other (i.e., besides ethane) non-aromatic hydrocarbon in the first reaction effluent, e.g., ≧75 wt. %, such as ≧90 wt. %, or ≧95 wt. %. The inter-stage separations of these aspects, when used, are typically carried out in separation stage 129 of recoverer 128. Alternatively or in addition to the aspects illustrated in
Although dehydrocyclization process conditions in reactor 334 can be the same as those of reactor 224, typically they are different. Similarly, although the same dehydrocyclization catalyst can be utilized in reactors 224 and 334, typically they are different. Although
Those skilled in the art will appreciate that within the ranges of process parameters specified for the reactor 224, there are process conditions which if selected would result in a maximum propane conversion to aromatic hydrocarbon “XMP”. Unlike conventional multi-stage processes, reactor 224 is typically operated at a propane conversion to aromatic hydrocarbon that is less than XMP. Instead, when the aromatization feed includes propane and/or when propane is produced in reactor 224, process conditions are generally selected so that the initial (start of run) propane conversion to aromatic hydrocarbon in reactor 224 “X1P” is ≦0.95·XMP. Typically, X1P≦0.90·XMP, e.g., ≦0.85·XMP, or ≦0.80·XMP, or ≦0.75·XMP. It has been found that operating reactor 224 under conditions which provide X1P≧XMP lead to excessive catalyst coking, typically resulting in a shortened cycle time in fixed bed operation. The same effect is observed for conversion of butanes to aromatic hydrocarbon, but with less sensitivity to changes the process conditions of reactor 224. Instead of operating at reactor 224 at maximum conversion of butanes to aromatic hydrocarbon XMB, the specified process conditions typically result in an initial (start of run) conversion of butenes to aromatic hydrocarbon (“X1B”) that is less than XMB, e.g., X1B≦0.995·XMB, such as ≦0.99·XMB, or ≦0.985·XMB.
Dehydrocyclization conditions in reactor 334 generally include a temperature T2 in the range of from 450° C. to 700° C., and a pressure P2≦35 psia (241.3 kPa). Typically, T1≦0.9·T2, e.g., T1≦0.85·T2, such as T1≦0.8·T2. The pressure in reactor 334 is typically less than the pressure in reactor 224, e.g., P2≦0.95·P1, such as P2≦0.90·P1, or P2≦0.85·P1, or P2≦0.8·P1. Typically, reaction conditions in reactor 334 include T2 in the range of from 500° C. to 675° C. and P2≦34 psia (234.4 kPa), e.g., ≦32 psia (220.6 kPa), such as ≦30 psia (207 kPa), or in the range of from 10 psia (68.9 kPa) to 35 psia (241.3 kPa) or from 12 psia (82.8 kPa) to 34 psia (234.4 kPa). Generally, the reaction in reactor 334 is carried out at a C2+ hydrocarbon WHSV of the specified first product with respect to the second catalyst in the range of from 0.1 hr−1 to 20 hr−1, e.g. 0.2 hr−1 to 5 hr−1, or 0.3hr−1 to 1 hr−1. In congruence with the WHSV of reactor 224, the WHSV of reactor 334 is based on C2+ hydrocarbon content, and is the hourly rate of C2+ hydrocarbon (in grams) introduced into reactor 334 per gram of second catalyst. Those skilled in the art will appreciate that T1 and T2 represent average temperatures across a reaction zone, or more particularly, across a catalyst bed located within a reaction zone. Average temperature is calculated by adding the zone's inlet temperature to the zone's outlet temperature, and then dividing the sum by 2. P1 and P2 are not average pressures. Instead, they correspond to the inlet pressure at the specified reactor, e.g., reactor 224 for P1 and reactor 334 for P2. Unlike reactor 224, it has been found that increased reaction pressure reactor 334 (i) decreases the yield of aromatic hydrocarbon and (ii) increases the rate of catalyst coke accumulation. A pressure of P2≦35 psia (241.3 kPa) is generally needed to achieve a reactor 334 cycle time ≧50 hours.
Contrary to expectations, it is detrimental to operate reactor 334 at a temperature sufficient for maximum conversion of ethane, typically T2>700° C. Doing so is observed to result in a decrease in selectivity to the desired aromatic hydrocarbon product. Typically, the average temperature across any reaction zone within reactor 334 (and across any catalyst bed located within a reaction zone reactor 334) is ≦700° C. Typically, the aromatization feed is not exposed to a temperature ≧700° C. at the inlet of reactor 334.
It also has been found that operating reactor 334 at a temperature >700° C. can lead to a chemical conversion of the catalyst's dehydrogenation component and a loss of catalytic dehydrocyclization activity, particularly when the dehydrogenation component includes one or more oxide of Zn. While not wishing to be bound by any theory or model, it is believed that utilizing a temperature >700° C. results in a conversion from the oxide form to a metallic form of Zn, which has a greater vapor pressure than does the oxide form. The loss of catalytic dehydrocyclization activity is thus attributed at least partially to the evaporation of Zn from the catalyst. Typically, total ethane conversion in the second stage is ≦60%, e.g., ≦50, such as ≦40%, or in the range of from 25% to 60%, or 30% to 55%, or 30% to 50%, or 30% to 40%.
At least to lessen the rate of catalyst coke accumulation, typically the dehydrocyclization conditions of reactor 224 are selected to convert (i) ≧50 wt. %, e.g., ≧75 wt. %, such as ≧90 wt. % of the aromatization feed's C3+ non-aromatic hydrocarbon, but (ii) ≦50 wt. %, e.g., ≦25%, such as ≦10%, or ≦1% of the aromatization feed's C2 hydrocarbon. Since the feed to reactor 334 typically contains fewer C3+ non-aromatic hydrocarbon compounds than does the aromatization feed, reactor 334 is typically operated under more severe conditions in order to achieve at least some ethane conversion to aromatic hydrocarbon. Those skilled in the art will appreciate that within the ranges of process parameters specified for reactor 334, there are process conditions which if used would result in a maximum ethane conversion to aromatic hydrocarbon “XME”. Unlike conventional multi-stage processes, reactor 334 typically is not operated at XME. Instead, process conditions are generally selected so that the initial (start of run) conversion of ethane to aromatic hydrocarbon in reactor 334 “X2E” is less than XME, e.g., ≦0.9·XME. Typically, X2E≦0.85·XME, e.g., ≦0.8·XME, or ≦0.75·XME. It has been found that operating stage two under conditions which provide X2E≧XME lead to excessive catalyst coking, which typically necessitates a shortened cycle time in fixed bed operation. Since there is a net conversion of ethane in reactor 334, X2E is greater than zero. Since ethane can be (and typically is) produced in reactor 224, the amount of ethane conversion in reactor 224 (X1E) can be less than zero. Generally, when operating under the specified conditions using the specified feeds, X2E is greater than X1E. Typically, [1−(X1E/X2E)] is ≦5, e.g., ≦2, such as ≦1, or in the range of 0.5 to 5, or 0.6 to 2, or 0.6 to 1.
In particular aspects, the first stage reactor 224 contains at least two fixed beds of the first catalyst and the second stage reactor 334 contains at least two fixed beds of the second catalyst. Particularly in these aspects, the aromatization feed reacts in the presence of the first catalyst in reactor 224 while exposed to a temperature T1 in the range of from 450° C. to 605° C. and a pressure P1 in the range of from 37 psia (255.1 kPa) to 100 psia (689.5 kPa), at a space velocity (WHSV) in the range of from 0.1 hr−1 to 20 hr−1, such as from 0.25 hr−1 to 2 hr−1. That portion of the first reaction effluent which reacts in second stage reactor 334 is exposed to a temperature T2 in the range of from 500° C. to 675° C. and a pressure P2≦32 psia (220.6 kPa), at a space velocity (WHSV) in the range of from 0.1 hr−1 to 20 hr−1, such as from 0.25 hr−1 to 2 hr−1, and typically less than the WHSV used when reacting the aromatization feed in reactor 224. More particularly, T1≦0.9·T2 and P2≦0.9·P1, with P2 in the range of 1 psia to 32 psia (220.6 kPa), e.g., 5 psia (34.5 kPa) to 30 psia (206.8 kPa), or 5 psia (34.5 kPa) to 29 psia (199.9 kPa).
The first and second catalysts can be selected from among any of the dehydrocyclization catalysts described in connection aspects illustrated in the single-stage aromatizer of
In particular 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 specified for use in stage 1 (the first catalyst), although typically the second catalyst has a greater acidity than does the first catalyst. For example, the molecular sieve component includes ≧90 wt. % of an aluminosilicate in hydrogen form, the aluminosilicate having a constraint index in the range of from 2-12 (e.g., small crystal, H-ZSM-5). Typically, (i) the second catalyst has a silica:alumina ratio in the range of from 3 to 60, e.g., from 10 to 40, such as from 15 to 35, and (ii) the catalyst includes ≦0.01 wt. % phosphorus.
In these aspects, the second catalyst's dehydrogenation component can include ≧90 wt. % of at least one oxide of Zn. The matrix component when used includes ≧90 wt. % of alumina, silica, and combinations thereof. The second catalyst typically includes <0.01 wt. % phosphorus.
A second reaction effluent is produced by reacting that portion of the first reaction effluent entering reactor 334 with the specified second catalyst under the specified conditions. The second reaction effluent is conducted away as a second reaction effluent via conduit 325. Generally, the second reaction effluent includes (i) ≧0.5 wt. % of additional aromatic hydrocarbon, (ii) additional molecular hydrogen, and optionally methane and ethane. The second reaction effluent generally includes ≧1 wt. % of total aromatic hydrocarbon, e.g., the total of (i) the additional aromatics and (ii) any portion of the reaction effluent's aromatics which are conveyed to the second stage but are not converted. Typically, the second reaction effluent has a total aromatic hydrocarbon content of ≧5 wt. %, based on the weight of the second reaction effluent, such as ≧10 wt. %, or in the range of from 1 wt. % to 95 wt. %, or 10 wt. % to 75 wt. %. The second reaction effluent's additional aromatic hydrocarbon content is typically ≧1 wt. %, e.g., ≧2 wt. %, such as ≧10 wt. %, or ≧25 wt. %, or ≧50 wt. %, or ≧75 wt. %. The process produces a desirable BTX product. Even though T2 is greater than T1, it has been found that the second reaction effluent has an unexpected increase in desirable xylene isomers, and an unexpected decrease yield of less desirable C11+ aromatic hydrocarbon.
Light hydrocarbon, e.g., ethane, can be present in the second reaction effluent. Methane, when present in the aromatization feed, does not typically convert to other hydrocarbon compounds under the conditions specified for stage 1 and/or stage 2; consequently when inter-stage methane separation is not used, substantially all of the feed methane typically is present in the second reaction effluent. The second reaction effluent typically includes additional methane, namely that produced in stage 2, e.g., by hydrogenolysis of C2+ hydrocarbon, and optionally that produced in stage 1 when inter-stage methane separation is decreased or not used.
Product Recovery from Second Stage Reaction Effluent
Although the second stage reaction effluent can be recovered in separation stage 129 of recoverer 128, it is typical to utilize a second separation stage, e.g., stage 329, as shown in
In certain aspects, the aromatizer (including, e.g., the first stage and/or second stage; and/or additional stages if used) can include one or more additional reaction zones, e.g., additional reactors (not shown). These additional reactors can be substantially the same as the others, namely of substantially the same bed configuration and contain substantially the same amount of substantially the same catalyst. Typically, the reactors of a stage operate in reaction (dehydrocyclization) mode while the additional reactors are operated in regeneration mode, and vice versa. Continuous or semi-continuous operation can be carried out in each stage, e.g., by alternating reactors and additional reactors in sequence in reaction and regeneration modes. In the aspects illustrated in
Typically, a stage's dehydrocyclization catalyst is regenerated at a temperature ≦700° C. Exceeding this temperature during regeneration has been found to result in catalyst de-alumination and/or loss of structure, leading to an undesirable loss of catalyst acidity. Catalyst regeneration for any of the specified catalysts is typically carried out using procedures which limit the maximum temperature to which the catalyst is exposed during regeneration to about 750° C., more typically to about 650° C. Conventional catalyst regeneration methods can be used, e.g., exposing the catalyst to an oxidant such as air or oxygen in air for a time sufficient to remove at least a portion of the catalyst coke, but the invention is not limited thereto. When dehydrocyclization stages 1 and/or 2 are carried out in a set of fixed-bed, adiabatic reactors in series, a suitable regeneration procedure includes circulating a stream of regeneration gas containing a limited amount of oxygen, which limits the size of the exotherm where coke is burned off the catalyst. Typically, at the location where the regeneration gas enters the first (most upstream, with respect to the flow of regeneration gas) reactor, e.g., at the reactor's inlet, the regeneration gas is exposed to a temperature ≦350° C., ≦325° C., such as ≦300° C. If needed, the oxidant content of the regenerating gas can be decreased to lessen the risk of exceeding the maximum temperature.
Certain forms of power generator will now be described in more detail with reference to
As shown in
The resulting electrical power may be used at the site or exported to an electrical grid. The amount of power produced may depend on the aromatization feed's volumetric flow rate. In examples in which the aromatization feed has a volumetric flow rate of about 50 MSCFD (1.4·106 NM3D), the electrical power generated may be about 150 megawatts (MW).
While the present techniques may be susceptible to various modifications and alternative forms, the examples discussed above have been shown only by way of example. However, it should again be understood that the techniques is not intended to be limited to the particular examples disclosed herein. Indeed, the present techniques include all alternatives, modifications, and equivalents falling within the true spirit and scope of the appended claims.
| Number | Date | Country | Kind |
|---|---|---|---|
| 15195311.4 | Nov 2015 | EP | regional |
This invention claims priority to and the benefit of U.S. Patent Application Ser. No. 62/232,609 filed Sep. 25, 2015 and 62/253,268 filed Nov. 10, 2015 entitled AROMATIZATION OF NON-AROMATIC HYDROCARBONS; and European Patent Application Nos. 15195311.4 filed Nov. 19, 2015, all of which are herein incorporated by reference in their entireties. The following related cases are also incorporated by reference in their entireties: U.S. Patent Application Ser. No. 62/234,262, filed Sep. 29, 2015; European Patent Application No. 15195314.8, filed Nov. 19, 2015; U.S. Patent Application Ser. No. 62/234,240, filed Sep. 29, 2015; European Patent Application No. 15197698.2, filed Dec. 3, 2015; U.S. Patent Application Ser. No. 62/247,795, filed Oct. 29, 2015; European Patent Application No. 15197700.6, filed Dec. 3, 2015; U.S. Patent Application Ser. No. 62/248,374, filed Oct. 30, 2015; European Patent Application No. 15197702.2, filed Dec. 3, 2015; U.S. Patent Application Ser. No. 62/253,268, filed Nov. 10, 2015; U.S. Patent Application Ser. No. 62/326,918, filed Apr. 25, 2016; European Patent Application No. 16175163.1, filed Jun. 20, 2016; U.S. Patent Application Ser. No. 62/299,730, filed Feb. 25, 2016; European Patent Application No. 16167395.9, filed Apr. 28, 2016; U.S. Patent Application Ser. No. 62/313,288, filed Mar. 25, 2016; European Patent Application No. 16173587.3, filed Jun. 8, 2016; U.S. Patent Application Ser. No. 62/313,306, filed Mar. 25, 2016; European Patent Application No. 16173980.0, filed Jun. 10, 2016; U.S. Patent Application Ser. No. 62/298,655, filed Feb. 23, 2016; and European Patent Application No. 16167672.1, filed Apr. 29, 2016.
| Number | Date | Country | |
|---|---|---|---|
| 62232609 | Sep 2015 | US | |
| 62253268 | Nov 2015 | US |
