Patent ID: 12234416

The invention is susceptible to various modifications and alternative forms, specific embodiments thereof are shown by way of example in the drawings. The drawings may not be to scale. It should be understood that the drawings are not intended to limit the scope of the invention to the particular embodiment illustrated.

DETAILED DESCRIPTION

The present disclosure provides processes to convert a mixture of light hydrocarbons comprising isobutane to liquid transportation fuels. The process and systems described herein relate primarily to the conversion of a feed stream predominantly comprising isobutane to generate upgraded products that meet specifications for a blend component of a liquid transportation fuel.

The feed stream is catalytically activated in a two-step activation that occurs in two separate catalyst beds arranged in series. The feed stream is received in a first catalyst bed and catalytically activated in the first catalyst bed by contacting a first catalyst to produce a first effluent comprising C2-C4 olefins and aromatics. The entire first effluent is then immediately conveyed to a second catalyst bed that is contained within a second reactor and is maintained at a temperature that is approximately 50 degrees cooler than the first catalyst bed (measured at the start or inlet of the catalyst bed). The first effluent is catalytically activated in the second catalyst bed by contacting a second catalyst that is compositionally distinct from the first catalyst to produce a second effluent that comprises aromatics and olefins in an approximate 1:1 molar ratio. The second effluent is then further upgraded by in a third reaction zone by oligomerization and/or alkylation to produce value-added chemicals and/or products including larger hydrocarbons and alkylated aromatics that meet specifications as a liquid transportation fuel blend component. An advantage of the inventive processes and systems is a significant decrease in production of light alkanes containing from 1-4 carbon atoms that do not meet government specifications for use as a transportation fuel and are typically referred to as fuel gas. Another advantage is that the feed stream is enriched in isobutane and contains only a minor amount of n-butane. Isobutane possesses increased chemical reactivity and is more easily catalytically activated than n-butane, which allows the two stage activation of the feed stream to be performed at lower temperatures. Additional advantages will become evident from the detailed disclosure provided below.

In some embodiments, the feed stream is generally comprised of light alkanes and is enriched in isobutane (iC4), In certain embodiments, the feed stream comprises at least 40 wt. % isobutane; optionally, at least 50 wt. % isobutane; %; optionally, at least 60 wt. % isobutane; %; optionally, at least 70 wt. % isobutane; optionally, at least 80 wt. % isobutane; optionally, at least wt. % isobutane.

In certain embodiments, the feed stream may be obtained by processing a stream of natural gas liquids to remove lighter hydrocarbon components (i.e., C1-C3) by way of conventional natural gas separation technologies that are well-characterized, such as, for example, de-methanizer, de-ethanizer and de-propanizer fractionation columns. A typical product of such separation technologies is commonly characterized as natural gasoline, comprising about 72 wt. % pentanes, with a majority of the remainder comprising C6 hydrocarbons.

A first embodiment of the inventive processes and systems is illustrated by the process flow-diagram ofFIG.1. A feed stream101that is enriched in isobutane is converted in a system50. The feed stream101is received by a first reactor110comprising a first catalyst bed112that contains a first catalyst114. The first reactor110is operable to receive the feed stream101and to maintain a temperature and a pressure in the first catalyst bed112that facilitates catalytic conversion of the feed stream101by the first catalyst114to produce a first effluent116comprising predominantly C2-C4 olefins, along with some monocyclic aromatics and alkanes (including unreacted iC4).

The first effluent116is conveyed to and received by a second reactor120that comprises a second catalyst bed122containing a second catalyst124. The second reactor120is operable to receive maintain a temperature and a pressure in the second catalyst bed122that facilitates conversion of the first effluent116by the second catalyst124to produce a second effluent126that is enriched in monocyclic aromatic content (on a molar basis relative to the first effluent) and additionally comprises olefins and some residual alkanes, including unreacted isobutane from the feed stream.

The second effluent126is next conveyed to a compressor128that compresses the second effluent to a pressure that is higher than the pressure that is maintained in the first reactor110and second reactor120to produce a compressed second effluent129that is immediately conveyed to alkylation reactor130that comprises an alkylation catalyst bed132containing an alkylation catalyst134. The alkylation reactor130is designed to maintain a temperature and a pressure that facilitates the contacting of the compressed second effluent129with the alkylation catalyst134at a temperature and a pressure that facilitate the alkylation of monocyclic aromatics in the compressed second effluent129with C1-C4 olefins present in the compressed second effluent to produce an alkylation effluent134that is enriched in mono-alkylated aromatics.

The alkylation effluent134is conveyed to and received by first separator135, which is designed and operable to separate hydrocarbons according to boiling point to produce a light fraction138comprising C1-C4 hydrocarbons and hydrogen and a heavy fraction137comprising C5+hydrocarbons and predominantly comprises monocyclic aromatics and some benzene. In certain embodiments, the first separator135is a two-phase splitter and separation of the alkylation effluent134is achieved by partial condensation.

The heavy fraction137is conveyed to and received by a second separator140that may be a naphtha stabilizer. Second separator140is designed to separate the heavy fraction137to produce 1) an aromatics fraction142that predominantly comprises hydrocarbons containing six or more carbon atoms, including monocyclic alkylated aromatics, benzene and 2) a C5 hydrocarbons fraction144that predominantly comprises C5 alkanes and olefins formed during the process. In some embodiments, the second separator140is a naphtha stabilizer. The C5 recycle fraction144may optionally be recycled and combined with feed stream101upstream from the first reactor110. Aromatics fraction142may optionally be conveyed to and received by a third separator147that separates benzene from the alkylated aromatics in the aromatics fraction to produce an alkylated aromatics product150that meets or exceeds government specifications for a liquid transportation fuel (or a blend component thereof) and a benzene stream148, at least a portion of which is optionally conveyed via conduit to be combined with the second effluent126upstream from compressor128.

The light fraction138can optionally be utilized as a fuel gas. Alternatively, hydrogen stream152may be separated from the light fraction138in a fourth separator139to produce recycle stream150. At least a portion of recycle stream150can be redirected via conduit and combined with feed stream101upstream from the first reactor110, thereby serving as a reaction diluent for the feed stream101in the first reactor110and a mechanism for recycling unreacted isobutane.

A second embodiment of the inventive processes and systems is illustrated by the process flow-diagram ofFIG.2. A feed stream201that is enriched in isobutane is converted in a system60. The feed stream201is received by a first reactor210comprising a first catalyst bed212that contains a first catalyst214and a second catalyst bed216comprising a second catalyst218, with the two catalyst beds arranged in series within the first reactor210. The first reactor210is operable to receive the feed stream201and to maintain a temperature and a pressure in the first catalyst bed212that facilitates catalytic conversion of the feed stream201by the first catalyst214to produce a first effluent (not depicted) comprising predominantly C2-C4 olefins, along with some monocyclic aromatics and alkanes (including unreacted isobutane).

The first effluent (not depicted) is next received by the second catalyst bed216containing a second catalyst218. The first reactor210is operable to receive and maintain a temperature and a pressure in the second catalyst bed216that facilitates conversion of the first effluent by the second catalyst218to produce a second effluent220that is enriched in monocyclic aromatic content (on a molar basis relative to the first effluent) and additionally comprises olefins and some residual alkanes, including unreacted isobutanes from the feed stream201.

Speaking generally, embodiments that combine the first and second catalyst beds in the first reactor are designed such that the first reactor is operable to maintain the second catalyst bed at a temperature (typically measured at the inlet to the second catalyst bed) that is at least 50° F. lower, (optionally, at least 50° F. lower) than the temperature that is maintained in the first catalyst bed (typically measured at the inlet to the first catalyst bed). In these embodiments, the first and second catalyst beds are arranged in series such that the feed stream contacts the first catalyst bed prior to contacting the second catalyst bed.

Again referring to the embodiment diagram depicted inFIG.2, the second effluent220is next conveyed to a compressor228that compresses the second effluent to a pressure that is higher than the pressure that is maintained in the first reactor210to produce a compressed second effluent229that is immediately conveyed to alkylation reactor230that comprises an alkylation catalyst bed232containing an alkylation catalyst234. The alkylation reactor230is designed to maintain a temperature and a pressure that facilitates the contacting of the compressed second effluent229with the alkylation catalyst234at a temperature and a pressure that facilitate the alkylation of monocyclic aromatics in the compressed second effluent229with C1-C4 olefins present in the compressed second effluent to produce an alkylation effluent234that is enriched in mono-alkylated aromatics.

The alkylation effluent234is conveyed to and received by first separator235, which is designed and operable to separate hydrocarbons according to boiling point to produce a light fraction238comprising C1-C4 hydrocarbons and hydrogen and a heavy fraction237comprising C5+hydrocarbons and predominantly comprises monocyclic aromatics and some benzene. In certain embodiments, the first separator235is a two-phase splitter and separation of the alkylation effluent234is achieved by partial condensation.

The heavy fraction237is conveyed to and received by a second separator240that may be a naphtha stabilizer. Second separator240is designed to separate the heavy fraction237to produce 1) an aromatics fraction242that predominantly comprises hydrocarbons containing six or more carbon atoms, including monocyclic alkylated aromatics, benzene and 2) a C5 hydrocarbons fraction244that predominantly comprises C5 alkanes and olefins formed during the process. In some embodiments, the second separator240is a naphtha stabilizer. The C5 recycle fraction244may optionally be recycled and combined with feed stream201upstream from the first reactor210. Aromatics fraction242may optionally be conveyed to and received by a third separator247that separates benzene from the alkylated aromatics in the aromatics fraction to produce an alkylated aromatics product250that meets or exceeds government specifications for a liquid transportation fuel (or a blend component thereof) and a benzene stream248, at least a portion of which is optionally conveyed via conduit to be combined with the second effluent226upstream from compressor228.

The light fraction238can optionally be utilized as a fuel gas. Alternatively, hydrogen stream252may be separated from the light fraction238in a fourth separator239to produce recycle stream250. At least a portion of recycle stream250can be redirected via conduit and combined with feed stream201upstream from the first reactor210, thereby serving as a reaction diluent for the feed stream201in the first reactor210and a mechanism for recycling unreacted isobutane.

Speaking generally, the first catalyst comprises a zeolite that contacts alkanes present in the feed stream and facilitates C—H bond activation as an initial step toward producing products comprising olefins and aromatics. Speaking generally, the combination of first catalyst and conditions of temperature, pressure and space velocity in the first reaction zone (that are maintained by the first reactor) are selected to facilitate catalytic conversion of the feed stream by the first catalyst with high selectivity toward the production of C2-C4 olefins while minimizing selectivity toward the catalytic conversion of the feed stream to produce aromatics.

The second catalyst comprises a zeolite that contacts the first effluent and facilitates catalytic conversion of C2-C4 olefins present in the first effluent to produce aromatics and larger olefins (having a greater molecular weight) from the olefins in the first effluent. To a lesser extent, the second activation catalyst may also convert alkanes that remained unconverted in the first activation zone to produce olefins and aromatics. Speaking generally, the combination of second activation catalyst and conditions of temperature, pressure and space velocity in the second activation zone (that are maintained by the second activation reactor) are selective to facilitate C—H bond activation by the second catalyst as an initial step toward producing products comprising predominately monocyclic aromatics and a smaller amount of olefins. Restated, the second catalyst catalytically convert the first effluent by increased selectivity to production of aromatics while decreasing selectivity toward the production of olefins.

Each of the first and second catalysts comprise zeolites capable of activating hydrocarbons to produce olefins and/or aromatic hydrocarbons. In some embodiments the first activation catalyst is a zeolite that is selective toward the conversion of isobutane to produce olefins via protonation of a C—H bond leading to a carbocation intermediate, and less selective for the production of monocyclic aromatics. In some embodiments, the second activation catalyst is a zeolite that is selective for the conversion of C2-C4 light olefins to monocyclic aromatics. In some embodiments, the first catalyst is compositionally-distinct from the second catalyst. In some embodiments, the first and second catalyst beds are fixed beds. In some embodiments the first and second catalyst beds are fixed beds that are arranged in series within a single reactor. Alternatively, the first and second catalyst beds may alternatively comprise a variety of known moving catalyst bed configurations, but an advantage of the present process is that it enables a longer catalytic life span due to the omission of dehydrogenating metals from the activation catalysts, enabling the use of a fixed bed configuration for the first and second catalysts. In addition, the present two step activation is advantaged over a single step activation as this dehydrogenation reaction is equilibrium-limited for conversion to olefins, and a single step activation therefore results in lower overall percent conversion of the light alkanes feed stream (less than 50 mole %). Generation of additional aromatics in the two step activation process relieves these equilibrium constraints, allowing additional conversion of the feed stream to olefins.

Favored catalysts include supported or unsupported structured silica-aluminas that do not include additional impregnated metals. In some embodiments, ZSM-5 zeolite catalysts are utilized that are characterized by Si/Al ratios ranging from 12-80. In some embodiments, the activation catalyst utilized in the first catalyst bed is different from the activation catalyst utilized in the second catalyst bed. In some embodiments, the first catalyst is selective for the production of olefins (versus aromatics) when contacted with the feed stream and the second catalyst is selective for the production of aromatics (versus olefins) when contacted with the first effluent. In some embodiments, the first catalyst has a higher Bronsted Acidity than the second catalyst. In some embodiments, the first catalyst has lower Lewis Acidity than to the second catalyst. In some embodiments, the first catalyst has a higher Bronsted acidity and a lower Lewis acidity than the second catalyst. In some embodiments, the first activation catalyst comprises a ZSM-5 catalyst with a Si/Al ratio in the range from 12-30, optionally in the range from 20-30. In some embodiments, the second activation catalyst comprises a ZSM-5 catalyst with a Si/Al ratio in the range from 31-80, optionally in the range from 31 to 50. In some embodiments, all catalysts utilized in the inventive process and system lack impregnated metals (other than Al) that promote dehydrogenation of alkanes, The lack of impregnated metals makes the zeolite catalyst utilized in the inventive processes and systems less expensive (i.e., decreased OPEX) and more resistant to deactivation by high sulfur containing feed stream (i.e., increased catalytic lifespan), therefore more useful with contaminated feeds.

In the present inventive process a sufficient concentration of light olefins (C2-C4) is generated through the sequential, two-step catalytic activation such that activation catalysts can be utilized that do not incorporate typical dehydrogenation promoting metals (such as platinum, zinc, molybdenum, or gallium) without significantly decreasing product yield. Catalysts comprising dehydrogenation-promoting metals are not only more expensive to utilize, but also prone to poisoning by contaminants (including but not limited to, sulfur, nitrogen, arsenic and lead) that are often present in hydrocarbon feed streams derived from petroleum. Thus, the ability of the present process to proceed effectively in the absence of poison-sensitive catalysts is highly advantageous to the present process. Further, dehydration is an equilibrium-limited reaction which limits yields production yields of favored olefins and aromatics. Utilizing activation catalysts without impregnated metals (other than Aluminum) allows increased yields of favored products without the upgrading reactions being equilibrium limited. This alleviates the need to recycle the feed one or more times to increase the yield of olefins and aromatics, as is done in many conventional processes.

The first catalyst bed contained within the first reactor is maintained at a temperature in the range from 500° C. to 650° C. (typically measured at the inlet of the first activation reactor). In some embodiments, the temperature in the first catalyst bed is maintained within the range from 525° C. to 625° C. In some embodiments, the temperature in the first catalyst bed is maintained within the range from 525° C. to 600° C. In some embodiments, the temperature in the first catalyst bed is maintained within the range from 550° C. to 600° C. In some embodiments, the temperature in the first catalyst bed is maintained within the range from 575° C. to 600° C.

The second catalyst bed contained within the second reactor is maintained at a temperature in the range from 475° C. to 625° C. (typically measured at the start of the catalyst bed) and at a temperature that is in the range from 25° C. to 75° C. cooler than the temperature that is maintained in the first catalyst bed. In some embodiments, the temperature in the second catalyst bed is maintained at a temperature in the range from 450° C. to 620° C. (typically measured at the inlet of the first catalyst bed) and at a temperature that is in the range from 30° C. to 70° C. cooler than the temperature that is maintained in the first catalyst bed. In some embodiments, the temperature in the second catalyst bed is maintained at a temperature in the range from 500° C. to 575° C. (typically measured at the inlet of the first catalyst bed) and at a temperature that is in the range from 40° C. to 60° C. cooler than the temperature that is maintained in the first catalyst bed. In some embodiments, the temperature in the second catalyst bed is maintained at a temperature that is 50° C. cooler than the temperature that is maintained in the first catalyst bed. In some embodiments, the temperature in the second catalyst bed is maintained within the range from 500° C. to 600° C. In some embodiments, the temperature in the second catalyst bed is maintained within the range from 500° C. to 575° C. In some embodiments, the temperature in the second catalyst bed is maintained within the range from 525° C. to 575° C.

The first catalyst bed contained within the first reactor is maintained at a pressure that favors the catalytic conversion of the feed stream to olefins (versus aromatics) by the first catalyst in the first catalytic bed. The conversion of light alkanes to light olefins (as opposed to aromatics) is favored at lower pressure. Thus, the first reactor maintains the first catalyst bed at a pressure in the range from 0 psig to 150 psig; Optionally, at a pressure in the range from 0 psig to 100 psig; Optionally, at a pressure in the range from 0 psig to 50 psig.

In some embodiments, the second reactor generally maintains the second catalyst bed at a pressure that is in the range between 0 psig and 500 psig. However, the conversion of light olefins and light alkanes (present in the first effluent) to aromatics is favored by higher pressure. Therefore, in some embodiments the second reactor maintains the second catalyst bed at a pressure in the range between 35 psig and 500 psig; optionally in the range between 50 psig and 300 psig; optionally, in the range between 50 psig and 200 psig; optionally, in the range between psig and 150 psig.

In certain embodiments, the temperature maintained within the first catalyst bed (typically measured at the inlet of the first reactor) is lower than the temperature maintained in the second catalyst bed (typically measured at the inlet of the second activation reactor). Optionally, the first temperature may be at least 10° F. lower, at least 20° F. lower, at least 25° F. lower, at least 30° F. lower, at least 35° F. lower, at least 40° F. lower, at least 45° F. lower or at least lower than the second temperature. Utilizing a lower temperature for the conversion of the first effluent in the second catalyst bed favors conversion of the first effluent to aromatics (versus cracking), thereby producing a second effluent with an approximate 1:1 molar ratio of C2-C4 light olefins to monocyclic aromatics. This makes the second effluent an excellent feed stream for a downstream alkylation process that is selective for the production of mono-alkylated aromatics having a Reid vapor pressure and octane rating that exceed specifications for a liquid transportation fuel blend stock. The final product of the process may be blended into liquid transportation fuel, separated and sold as one or more chemical products (i.e., Benzene, Xylene, Toluene, etc.) or utilized as feed stock for one or more additional process to produce value-added chemicals while minimizing the production of undesirable light hydrocarbons containing four or fewer carbons.

Speaking generally, the alkylation reactor is maintained at a feed inlet temperature and pressure suitable to facilitate the catalytic alkylation of aromatics present in the second effluent by the alkylation catalyst. The aromatics that are alkylated may be produced by aromatization that takes place in the first or second reactor, may be present in the feed stream for the process, or a combination of these possibilities. These aromatics are alkylated by olefins that are largely produced by the activation of alkanes in the first and, to a lesser extent, second reactors. Alkylation of aromatics by the alkylation catalyst produces larger aromatic hydrocarbons comprising at least seven carbon atoms that are preferably mono-alkylated and have a boiling point that is in the boiling point range of a liquid transportation fuel (e.g., gasoline or diesel). Typically, the alkylation effluent comprises an increased percentage of alkylated aromatic compounds comprising from seven to nine carbon atoms. Optionally, the larger hydrocarbons also are characterized by a lower Reid vapor pressure and an increased octane rating.

The alkylation reactor contains a bed of alkylation catalyst that may be in any known configuration including fixed bed, moving bed or ebulliated bed. The alkylation reactor is generally maintained at a pressure in a range from 0 psig to 800 psig, optionally in the range from 35 psig to 600 psig. The alkylation reactor is typically maintained at a temperature (measured within the alkylation reactor inlet or at the start of the alkylation catalyst bed) that is in the range from 150° C. to 350° C.; optionally between 200° C. to 350° C.; optionally, 250° C. to 350° C. Typically, flow thorough the alkylation reactor is maintained at a weighted hourly space velocity (WHSV) in the range from 0.5 hr−1to 10 hr−1on an olefin basis. Optionally, the WHSV is in the range from 0.5 hr−1to 6.0 hr−1. While higher overall throughput is desirable, ideally the chosen WHSV allows for conversion of at least 85% of hydrocarbons present in the second effluent at the selected operating temperature and pressure.

In some embodiments, the catalytic conversion occurring in the alkylation reactor produces an alkylation effluent wherein the product comprises at least 30 wt. % (optionally, at least 40 wt %) of hydrocarbon molecules having a boiling-point that is within the boiling point range of a liquid transportation fuel. In some embodiments, the catalytic conversion occurring in the alkylation reactor produces an alkylation effluent wherein the product comprises less than 10 wt. % paraffins. In some embodiments, the catalytic conversion occurring in the alkylation reactor produces an alkylation effluent wherein the product comprises at least 50 wt. % aromatics; optionally at least 60 wt. % aromatics; optionally, at least 65 wt. % aromatics. In some embodiments, the catalytic conversion occurring in the alkylation reactor produces an alkylation effluent wherein the product comprises at least 40 wt. % mono-alkylated aromatics; optionally at least 50 wt. % mono-alkylated aromatics; optionally at least 60 wt. % mono-alkylated aromatics.

Speaking generally, the alkylation catalyst may comprise any catalyst characterized as either Bronsted or Lewis acidic. In some embodiments, the alkylation catalyst is supported on an acidic support which is well understood by to those having experience in the field. The Si/Al ratio of the alkylation catalyst may range from 12-80. In some embodiments, the catalyst may comprise a ZSM-5 catalyst with a Si/Al ratio of 23. A wide variety of catalysts have been found to promote aromatic alkylation including, but not limited to, aluminum chloride, phosphoric acid, sulfuric acid, hydrofluoric acid, silica, alumina, sulfated zirconia, zeolites (including, for example, ZSM-5, ZSM-3, ZSM-4, ZSM-18, ZSM-20, zeolite-beta, H-Y, MCM-22, MCM-36 and MCM-49).

Certain embodiments comprise mixing a diluent with the first fraction and/or the second fraction prior to contacting with the activation catalysts. The diluent may be added to the feed stream in a molar ratio ranging from 10:1 to 1:10 relative to the quantity of feed stream that is fed to the first activation zone. The diluent may be added at any point that is upstream from the first activation zone in the first activation reactor.

The diluent may comprise any substance that is less likely to be catalytically converted by one or more of the first activation catalyst and the second activation catalyst than the C2-C5 alkanes, olefins and aromatics present in the feed stream or first effluent at the conditions of temperature and pressure that are maintained within the first and second activation zones. This prevents the diluent from reacting with the first and second activation catalysts, which slows the overall catalytic conversion rate of compounds within the feed stream and/or first effluent by the activation catalysts. A large number of chemical compounds may serve as the diluent, the identity of which is fully within the grasp of one having experience in the art. In some embodiments the diluent comprises one or more light paraffins having from one to four carbon atoms, including C1-C4 light paraffins produced by the processes and systems described herein and recycled to be combined with the feed stream. In some embodiments the diluent may comprise any of methane, ethane, propane, butanes, benzene, toluene, xylenes, alkyl- or dialkyl-benzenes, naphthenes, C2-C5 olefins and combinations thereof.

The presence of diluent during catalytic activation (i.e., activation) provides several advantages. First, it effectively decreases the concentration of catalytically-convertible hydrocarbons in the feed stream within the first reactor, decreases the concentration of catalytically-convertible hydrocarbons within the second reactor, or both. This results in a small increase in the total conversion of pentanes (an increase of approximately 5-6 wt. %, typically) to produce olefins or aromatics within each activation reactor. However, it increases the selectivity toward the production of olefins in both the first and second effluent, while slightly decreasing the selectivity toward aromatics. Adjusting the ratio of diluent to feed stream changes the ratio of olefins to aromatics exiting the reactor, thereby providing a valuable point of operational control for downstream processes.

Addition of a diluent also advantageously favors the production of value-added olefins relative to C1-C4 light paraffins and decreases the dimerization rate of C5 hydrocarbons to form durene (1,2,4,5-tetramethylbenzene), a detrimental byproduct that can precipitate as a solid from liquid transportation fuels.

Typically, the composition of the feed stream, the first catalyst, the second catalyst, and the conditions maintained in the first reactor and the second reactor together produce a second effluent having a molar ratio of olefins to aromatics that is in the range from 0.5:1 to 1.5:1. In some embodiments, the composition of the feed stream, the first catalyst, the second catalyst, and the conditions maintained in the first reactor and the second reactor together produce a second effluent having a molar ratio of olefins to aromatics that is in the range from 0.8:1 to about 1.2:1 In some embodiments, the composition of the feed stream, the first catalyst, the second catalyst, and the conditions maintained in the first reactor and the second reactor together produce a second effluent having a molar ratio of olefins to aromatics that is in the range from 0.9:1 to about 1.1:1, which maximizes the conversion of aromatics and olefins in the second effluent to high octane (rating) mono-alkylated aromatics in the alkylation reactor, while minimizing selectivity to the production of di-alkylated and tri-alkylated aromatics. Mono-alkylated aromatics exhibit increased octane rating and decreased Reid vapor pressure properties and are favored gasoline blending components. In contrast, di-alkyl and tri-alkyl aromatics comprising more than nine carbon atoms are not well-suited for blending into gasoline and exhibit non-optimal cetane number for blending into diesel. The present inventive process maximizes conversion of the light alkanes feed stream to olefins and aromatics, while simultaneously producing a second effluent (that serves as feed for the alkylation reactor) that possesses an optimal ratio of olefins to aromatics to maximize production of high octane rating mono-alkylated aromatics that meet specifications for a gasoline blend stock.

EXAMPLES

The following examples are representative of one embodiment of the inventive processes and systems disclosed herein, and the scope of the invention is not intended to be limited to the embodiment specifically disclosed. Rather, the scope is intended to be as broad as is supported by the complete disclosure and the appending claims.

Example 1

To demonstrate that a two-step in-series activation of a light alkane can improve both the conversion and aromatics production when utilizing an activation catalyst that is not impregnated with metals, a 100 wt. % mixed pentanes feed stream (1:1 molar ratio of n-C5 to i-C5) was catalytically converted in two activation catalyst beds arranged in-series with the first catalyst bed maintained at a temperature of 575° C. and the second catalyst bed maintained at 500° C. with a WHSV=6.0 hr−1. The activation catalyst for each of the first and second catalyst beds was ZSM-5. Results are time averaged for 16 hours and all reactions were carried out at 30 PSIG. Results were compared with a single temperature activation performed at either of the two temperatures utilized in the inventive two step activation process. Product distribution percentage reflects are normalized by subtracting unreacted feed stream.

TABLE 1Two-step catalytic activation of pentanes comparedto single-pass catalytic activation.Single-PassSingle-PassIn-SeriesCatalytic ProcessActivationActivationActivationTemperature575° C.500° C.575° C.then500° C.Conversion (mol %)534273Recycle Molar Ratio4:14:14:1(Recycle:Feed)Product Selectivity767872(wt. %)Lt. Olefin Selectivity654432(wt. %)BTEX Aromatics Selectivity103237(wt. %)Product Distribution (wt. %)Methane336Ethane8610Ethene211211Propanes11510Propenes302016Butanes464Butenes1185Pentenes241C6+ Paraffins020Benzene379Toluene61517Xylene1910Ethyl-Benzene011
The results in Table 1 clearly demonstrate that when compared to a conventional one-step catalytic activation the two-step activation process described herein increases the overall conversion of the feed stream and significantly improves production of monocylic aromatics benzene, toluene and xylene in the activation product (which is a result of additional conversion to olefin intermediates that are then aromatized in the second step).

Example 2

Extending the above result to isobutane, a 100 wt. % isobutane feed stream was catalytically converted in two activation catalyst beds arranged in-series with the first catalyst bed maintained at a temperature of 575° C. and the second catalyst bed maintained at 525° C. with a WHSV=6.0 hr−1(see Table 5, first column). In a second experiment (Table 5, second column), the first catalyst bed was maintained at a temperature of 600° C. and the second catalyst bed maintained at 550° C. The feed stream utilized for the experiment comprised a recycle stream comprising the activation effluent mixed with a 100 wt % isobutane stream at a 4:1 ratio. The activation catalyst for each of the first and second catalyst beds was ZSM-5. Product analysis results were time averaged for 16 hours and all reactions were carried out at 30 PSIG. Product distribution percentage reflects are normalized by subtracting unreacted feed stream.

TABLE 2Two-step catalytic activation of isobutane under different temperaturescenarios.In SeriesIn-SeriesCatalytic ProcessActivationActivationTemperature575° C.600° C.thenthen525° C.550° C.Conversion (mol %)73.592.7Recycle Molar Ratio (Recycle:Feed)4:14:1Product Selectivity (wt. %)79.478.9Lt. Olefin Selectivity (wt. %)37.832.6BTEX Aromatics Selectivity (wt. %)37.343.2Methane8.39.5Ethane1.73.2Ethene6.79.2Propanes6.68.8Propenes11.912.7Butanes (unreacted feed stream)26.57.3Butenes8.48.0Pentanes1.40.6Pentenes0.90.6C6+ Paraffins1.31.7Benzene6.513.2Toluene14.318.5Xylene5.06.0Ethyl-Benzene0.60.6

The results in Table 2 clearly demonstrate that when compared to a conventional one-step catalytic activation for isobutane, the two-step activation process described herein greatly decreased the quantity of C2-C4 olefins in the activation product, while simultaneously increasing the quantity of the monocylic aromatics benzene, toluene and xylene in the activation product.

Example 3

The alkylation reactor provides a means for upgrading the generated stream enriched in light olefins and aromatics. To demonstrate, the activation effluent produced in Example 2 was contacted with an alkylation catalyst comprising an Si/Al ratio of 23 at a temperature of 300° C. and 30 psig and WHSV=2.0 hr−1. Product results are shown in Table 3 and were time-averaged over a period of 9 hours.

TABLE 3Product specifications for the non-stabilizedliquid product produced in Example 2.ComponentProduct (wt. %)Paraffins1.2i-Paraffins4.3Aromatics91.8Mono-Aromatics89.4n-Olefins0.6Iso-Olefins0.5RON (calculated)113.8MON (calculated)106.4Avg. Molecular Weight96.8Avg. Specific Gravity0.85Avg API @ 60° F.35.9RVP (psia)6.9C/H ratio9.3Benzene Content (vol. %)4.0

Example 4

One advantage of the present inventive process is that the feed stream comprises isobutane as while minimizing n-butane content. To demonstrate, catalytic cracking of feed streams comprising either n-butane or isobutane were conducted by contacting each feed stream with a ZSM-5 zeolite at several temperatures (and a WHSV of 10 g/g X hr), then examining the overall conversion of each (in wt %) as well as the products produced (in wt %).

TABLE 4Catalytic activation of either iso-butane or n-butane to olefins and aromatics by a ZSM-5 zeolite at several temperatures.Temp.ConversionProductFuel GasOlefinAromaticFeed(° C.)(wt %)SelectivitySelectivitySelectivitySelectivityn-butane65064.560.239.841.718.0i-butane65083.771.328.750.419.9n-butane60042.455.944.143.012.5i-butane60065.769.031.048.516.1n-butane55021.452.247.840.810.6i-butane55034.366.633.448.710.2
At each activation temperature, isobutane conversion was signifigantly higher than n-butane, as was selectivity to olefins and aromatics. Notably, selectivity to fuel gas was significantly decreased at all temperature tested, demonstrating the benefit of utilizing isobutane versus n-butane. Further, utilization of isobutane as feed stream decreases the olefin to aromatics ratio, improving the suitability of the activation effluent for immediate alkylation to produce mono-alkylated single ring aromatics.

Example 5

The effect that a methane diluent has on catalytic activation and conversion of a simulated “natural gasoline” comprising a mixture of pentane isomers was next demonstrated. The feed stream was fed at a WHSV of 1.3 hr−1to a reactor containing an activation catalyst comprising a ⅛″ extrudate consisting of 50 wt. % alumina binder and 50 wt. % ZSM-5 zeolite. The temperature of the reactor (at the inlet for the feed stream) was maintained at 600° C. and 20 psig (2.4 Bar) and results were time-averaged for 16.5 hr. For certain reactions, methane diluent was co-fed along with each feed stream at a methane: feed stream molar ratio of 2:1.

The reaction produced an effluent comprising light olefins, aromatics and light paraffins. Table 4 (below) shows the effect of the methane diluent on the total conversion of the 1:1 and 7:3 feed streams, respectively, as well as the selectivity of each conversion toward light olefins, aromatics, and byproduct C1-C4 fuel gas.

TABLE 4Catalytic activation of a 1:1 i-C5:n-C5 feed streamin both the absence and presence of methane diluent.Feed StreamMixed PentanesMixed Pentanes+/−DiluentNo DiluentCH4DiluentMaterial Balance101%103%Conversion92%80%Fuel gas yield37%22%Product Yield54%58%Coke Yield0%0%Lt. Olefin Yield34%44%Lt. Olefin37%55%SelectivityAromatic Yield20%14%Aromatic21%17%SelectivityFuel Gas Yield37%22%Fuel Gas41%27%Selectivity

The data in Table 4 indicate that adding an inert alkane diluent slightly decreased overall conversion rate, but significantly increased the yield and selectivity to light olefin production for the pentanes feed stream. Adding inert diluent also greatly diminished selectivity to the production of C1-C4 fuel gas. Meanwhile, only a small drop in selectivity to aromatics production was observed in the presence of diluent, which was offset by an equivalent increase in aromatics production. All of these results are advantageous to the process, particularly for embodiments where the second effluent is immediately conveyed to an alkylation process. In certain embodiments that comprise an aromatic alkylation process, diluent can be added to the feed stream or the first effluent at a ratio that helps adjust the molar ratio of produced olefins to aromatics to a value that is between 0.5:1 and 1.5:1 by mole, (optimally 1:1), thereby providing an advantageous feed for a downstream aromatic alkylation process.

Definitions:

In the present disclosure, the term “conversion” is defined as any of the chemical reactions that occur during upgrading of hydrocarbons to liquid transportation fuels. Examples of such reactions include, but are not limited to: oligomerization, aromatization, dehydrogenation, alkylation, hydrogenation and cracking.