Process for converting gaseous alkanes to liquid hydrocarbons

A process for converting gaseous alkanes to liquid hydrocarbons wherein a gaseous feed containing alkanes is reacted with a dry bromine vapor to form alkyl bromides and hydrobromic acid vapor. The mixture of alkyl bromides and hydrobromic acid are then reacted over a synthetic crystalline alumino-silicate catalyst, such as a ZSM-5 zeolite, at a temperature of from about 150° C. to about 450° C. so as to form higher molecular weight hydrocarbons and hydrobromic acid vapor. Propane and butane which comprise a portion of the products may be recovered or recycled back through the process to form additional C5+ hydrocarbons. Various methods are disclosed to remove the hydrobromic acid vapor from the higher molecular weight hydrocarbons and to generate bromine from the hydrobromic acid for use in the process.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates to a process for converting lower molecular weight, gaseous alkanes to liquid hydrocarbons useful for the production of fuels, and more particularly, to a process wherein a gas containing lower molecular weight alkanes is reacted with a dry bromine vapor to form alkyl bromides and hydrobromic acid which in turn are reacted over a crystalline alumino-silicate catalyst to form liquid hydrocarbons.

2. Description of Related Art

Natural gas which is primarily composed of methane and other light alkanes has been discovered in large quantities throughout the world. Many of the locales in which natural gas has been discovered are far from populated regions which have significant gas pipeline infrastructure or market demand for natural gas. Due to the low density of natural gas, transportation thereof in gaseous form by pipeline or as compressed gas in vessels is expensive. Accordingly, practical and economic limits exist to the distance over which natural gas may be transported in gaseous form exist. Cryogenic liquefaction of natural gas (LNG) is often used to more economically transport natural gas over large distances. However, this LNG process is expensive and there are limited regasification facilities in only a few countries that are equipped to import LNG.

Another use of methane found in natural gas is as feed to processes for the production of methanol. Methanol is made commercially via conversion of methane to synthesis gas (CO and H2) at high temperatures (approximately 1000° C.) followed by synthesis at high pressures (approximately 100 atmospheres). There are several types of technologies for the production of synthesis gas (CO and H2) from methane. Among these are steam-methane reforming (SMR), partial oxidation (POX), autothermal reforming (ATR), gas-heated reforming (GHR), and various combinations thereof. SMR and GHR operate at high pressures and temperatures, generally in excess of 600° C., and require expensive furnaces or reactors containing special heat and corrosion-resistant alloy tubes filled with expensive reforming catalyst. POX and ATR processes operate at high pressures and even higher temperatures, generally in excess of 1000° C. As there are no known practical metals or alloys that can operate at these temperatures, complex and costly refractory-lined reactors and high-pressure waste-heat boilers to quench & cool the synthesis gas effluent are required. Also, significant capital cost and large amounts of power are required for compression of oxygen or air to these high-pressure processes. Thus, due to the high temperatures and pressures involved, synthesis gas technology is expensive, resulting in a high cost methanol product which limits higher-value uses thereof, such as for chemical feedstocks and solvents. Furthermore production of synthesis gas is thermodynamically and chemically inefficient, producing large excesses of waste heat and unwanted carbon dioxide, which tends to lower the conversion efficiency of the overall process. Fischer-Tropsch Gas-to-Liquids (GTL) technology can also be used to convert synthesis gas to heavier liquid hydrocarbons, however investment cost for this process is even higher. In each case, the production of synthesis gas represents a large fraction of the capital costs for these methane conversion processes.

Numerous alternatives to the conventional production of synthesis gas as a route to methanol or synthetic liquid hydrocarbons have been proposed. However, to date, none of these alternatives has attained commercial status for various reasons. Some of the previous alternative prior-art methods, such as disclosed in U.S. Pat. No. 5,243,098 or 5,334,777 to Miller, teach reacting a lower alkane, such as methane, with a metallic halide to form a metalous halide and hydrohalic acid which are in turn reduced with magnesium oxide to form the corresponding alkanol. However, halogenation of methane using chlorine as the preferred halogen results in poor selectivity to the monomethyl halide (CH3Cl), resulting in unwanted by-products such as CH2Cl2and CHCl3which are difficult to convert or require severe limitation of conversion per pass and hence very high recycle rates.

Other prior art processes propose the catalytic chlorination or bromination of methane as an alternative to generation of synthesis gas (CO and H2). To improve the selectivity of a methane halogenation step in an overall process for the production of methanol, U.S. Pat. No. 5,998,679 to Miller teaches the use of bromine, generated by thermal decomposition of a metal bromide, to brominate alkanes in the presence of excess alkanes, which results in improved selectivity to mono-halogenated intermediates such as methyl bromide. To avoid the drawbacks of utilizing fluidized beds of moving solids, the process utilizes a circulating liquid mixture of metal chloride hydrates and metal bromides. Processes described in U.S. Pat. No. 6,462,243 B1, U.S. Pat. No. 6,472,572 B1, and U.S. Pat. No. 6,525,230 to Grosso are also capable of attaining higher selectivity to mono-halogenated intermediates by the use of bromination. The resulting alkyl bromides intermediates such as methyl bromide, are further converted to the corresponding alcohols and ethers, by reaction with metal oxides in circulating beds of moving solids. Another embodiment of U.S. Pat. No. 6,525,230 avoids the drawbacks of moving beds by utilizing a zoned reactor vessel containing a fixed bed of metal oxide/metal bromide that is operated cyclically in four steps. These processes also tend to produce substantial quantities of dimethylether (DME) along with any alcohol. While DME is a promising potential diesel engine fuel substitute, as of yet, there currently exists no substantial market for DME, and hence an expensive additional catalytic process conversion step would be required to convert DME into a currently marketable product. Other processes have been proposed which circumvent the need for production of synthesis gas, such as U.S. Pat. Nos. 4,655,893 and 4,467,130 to Olah in which methane is catalytically condensed into gasoline-range hydrocarbons via catalytic condensation using superacid catalysts. However, none of these earlier alternative approaches have resulted in commercial processes.

It is known that substituted alkanes, in particular methanol, can be converted to olefins and gasoline boiling-range hydrocarbons over various forms of crystalline alumino-silicates also known as zeolites. In the Methanol to Gasoline (MTG) process, a shape selective zeolite catalyst, ZSM-5, is used to convert methanol to gasoline. Coal or methane gas can thus be converted to methanol using conventional technology and subsequently converted to gasoline. However due to the high cost of methanol production, and at current or projected prices for gasoline, the MTG process is not considered economically viable. Thus, a need exists for an economic process for the for the conversion of methane and other alkanes found in natural gas to useful liquid hydrocarbon products which, due to their higher density and value, are more economically transported thereby significantly aiding development of remote natural gas reserves. A further need exists for a process for converting alkanes present in natural gas which is relatively inexpensive, safe and simple in operation.

SUMMARY OF THE INVENTION

To achieve the foregoing and other objects, and in accordance with the purposes of the present invention, as embodied and broadly described herein, one characterization of the present invention is a process for converting gaseous alkanes to liquid hydrocarbons. The process comprises reacting a gaseous feed having lower molecular weight alkanes with bromine vapor to form alkyl bromides and hydrobromic acid. The alkyl bromides and hydrobromic acid are reacted in the presence of a synthetic crystalline alumino-silicate catalyst and at a temperature sufficient to form higher molecular weight hydrocarbons and hydrobromic acid vapor. The hydrobromic acid vapor is removed from the higher molecular weight hydrocarbons by reacting the hydrobromic acid vapor with a metal oxide to form a metal bromide and steam.

In another characterization of the present invention, a process is provided for converting gaseous alkanes to liquid hydrocarbons wherein a gaseous feed having lower molecular weight alkanes is reacted with bromine vapor to form alkyl bromides and hydrobromic acid. The alkyl bromides and hydrobromic acid are reacted in the presence of a synthetic crystalline alumino-silicate catalyst and at a temperature sufficient to form higher molecular weight hydrocarbons and hydrobromic acid vapor. The hydrobromic acid vapor and the higher molecular weight hydrocarbons are transported to a first vessel having a bed of metal oxide particles, the hydrobromic acid vapor reacting with the bed of metal oxide particles to form metal bromide particles and steam.

In still another characterization of the present invention, a process is provided for converting gaseous alkanes to liquid hydrocarbons wherein a gaseous feed having lower molecular weight alkanes is reacted with bromine vapor to form alkyl bromides and hydrobromic acid. The alkyl bromides and hydrobromic acid are reacted in the presence of a synthetic crystalline alumino-silicate catalyst and at a temperature sufficient to form higher molecular weight hydrocarbons and hydrobromic acid vapor. The hydrobromic acid vapor is removed from said higher molecular weight hydrocarbons by reaction with a metal oxide to form a first metal bromide and steam. The first metal bromide is oxidized with an oxygen containing gas to form bromine vapor. The bromine vapor is reacted a reduced metal bromide to form a second metal bromide.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

As utilized throughout this description, the term “lower molecular weight alkanes” refers to methane, ethane, propane, butane, pentane or mixtures thereof. As also utilized throughout this description, “alkyl bromides” refers to mono, di, and tri brominated alkanes. Also, the feed gas in lines11and111in the embodiments of the process of the present invention as illustrated inFIGS. 2 and 3, respectively, is preferably natural gas which may be treated to remove sulfur compounds and carbon dioxide. In any event, it is important to note that small amounts of carbon dioxide, e.g. less than about 2 mol %, can be tolerated in the feed gas to the process of the present invention.

A block flow diagram generally depicting the process of the present invention is illustrated inFIG. 1, while specific embodiments of the process of the present invention are illustrated inFIGS. 2 and 3. Referring toFIG. 2, a gas stream containing lower molecular weight alkanes, comprised of a mixture of a feed gas plus a recycled gas stream at a pressure in the range of about 1 bar to about 30 bar, is transported or conveyed via line, pipe or conduit62, mixed with dry bromine liquid transported via line25and pump24, and passed to heat exchanger26wherein the liquid bromine is vaporized. The mixture of lower molecular weight alkanes and dry bromine vapor is fed to reactor30. Preferably, the molar ratio of lower molecular weight alkanes to dry bromine vapor in the mixture introduced into reactor30is in excess of 2.5:1. Reactor30has an inlet pre-heater zone28which heats the mixture to a reaction initiation temperature in the range of about 250° C. to about 400° C.

In first reactor30, the lower molecular weight alkanes are reacted exothermically with dry bromine vapor at a relatively low temperature in the range of about 250° C. to about 600° C., and at a pressure in the range of about 1 bar to about 30 bar to produce gaseous alkyl bromides and hydrobromic acid vapors. The upper limit of the operating temperature range is greater than the upper limit of the reaction initiation temperature range to which the feed mixture is heated due to the exothermic nature of the bromination reaction. In the case of methane, the formation of methyl bromide occurs in accordance with the following general reaction:
CH4(g)+Br2(g)→CH3Br(g)+HBr(g)

This reaction occurs with a significantly high degree of selectivity to methyl bromide. For example, in the case of bromination of methane with a methane to bromine ratio of about 4.5:1 selectivity to the mono-halogenated methyl bromide is in the range of 90 to 95%. Small amounts of dibromomethane and tribromomethane are also formed in the bromination reaction. Higher alkanes, such as ethane, propane and butane, are also readily bromoninated resulting in mono and multiple brominated species. If an alkane to bromine ratio of significantly less than about 2.5 to 1 is utilized, selectivity to methyl bromide substantially lower than 90% occurs and significant formation of undesirable carbon soot is observed. It has also been shown that other alkanes such as ethane, propane and butane which may be present in the feed gas to the bromination reactor are readily brominated to form ethyl bromides, propyl bromides and butyl bromides. Further, the dry bromine vapor that is feed into first reactor30is substantially water-free. Applicant has discovered that elimination of substantially all water vapor from the bromination step in first reactor30substantially eliminates the formation of unwanted carbon dioxide thereby increasing the selectivity of alkane bromination to alkyl bromides and eliminating the large amount of waste heat generated in the formation of carbon dioxide from alkanes.

The effluent that contains alkyl bromides and hydrobromic acid is withdrawn from the first reactor via line31and is partially cooled to a temperature in the range of about 150° C. to about 350° C. in heat exchanger32before flowing to a second reactor34. In second reactor34, the alkyl bromides are reacted exothermically at a temperature range of from about 150° C. to about 450° C., and a pressure in the range of about 1 to 30 bar, over a fixed bed33of crystalline alumino-silicate catalyst, preferably a zeolite catalyst, and most preferably a ZSM-5 zeolite catalyst. Although the zeolite catalyst is preferably used in the hydrogen, sodium or magnesium form, the zeolite may also be modified by ion exchange with other alkali metal cations, such as Li, Na, K or Cs, with alkali-earth metal cations, such as Mg, Ca, Sr or Ba, or with transition metal cations, such as Ni, Mn, V, W, or to the hydrogen form. Other zeolite catalysts having varying pore sizes and acidities, which are synthesized by varying the alumina-to-silica ratio may be used in the second reactor34as will be evident to a skilled artisan. In this reactor, the alkyl bromides are oligimerized to produce a mixture of higher molecular weight hydrocarbon products, primarily C3, C4 and C5+ gasoline-range and heavier hydrocarbon fractions, and additional hydrobromic acid vapor.

The temperature at which the second reactor34is operated is an important parameter in determining the selectivity of the oligimerization reaction to various higher molecular weight liquid hydrocarbon products. It is preferred to operated second reactor34at a temperature within the range of about 150° to 450°. Temperatures above about 300° C. in the second reactor result in increased yields of light hydrocarbons, such as undesirable methane, whereas lower temperatures increase yields of heavier molecular weight hydrocarbon products. At the low end of the temperature range, with methyl bromide reacting over ZSM-5 zeolite at temperatures as low as 150° C. significant methyl bromide conversion on the order of 20% is noted, with a high selectivity towards C5+ products. Also it is noted that methyl bromide appears to be more reactive over a lower temperature range relative to methyl chloride or other substituted methyl compounds such as methanol. Notably, in the case of the alkyl bromide reaction over the preferred zeolite ZSM-5 catalyst, cyclization reactions also occur such that the C7+ fractions are composed primarily of substituted aromatics. At increasing temperatures approaching 300° C., methyl bromide conversion increases towards 90% or greater, however selectivity towards C5+ products decreases and selectivity towards lighter products, particularly undesirable methane, increases. Surprisingly, very little ethane or C2, —C3olefin components are formed. At temperatures approaching 450° C., almost complete conversion of methyl bromide to methane occurs. In the optimum operating temperature range of between about 300° C. and 400° C., as a byproduct of the reaction, a small amount of carbon will build up on the catalyst over time during operation, causing a decline in catalyst activity over a range of hours, up to hundreds of hours, depending on the reaction conditions and the composition of the feed gas. It is believed that higher reaction temperatures above about 400° C., associated with the formation of methane favor the thermal cracking of alkyl bromides and formation of carbon or coke and hence an increase in the rate of deactivation of the catalyst. Conversely, temperatures at the lower end of the range, particularly below about 300° C. may also contribute to coking due to a reduced rate of desorption of heavier products from the catalyst. Hence, operating temperatures within the range of about 150° C. to about 450° C., but preferably in the range of about 300° C. to about 400° C. in the second reactor34balance increased selectivity of the desired C5+ products and lower rates of deactivation due to carbon formation, against higher conversion per pass, which minimizes the quantity of catalyst, recycle rates and equipment size required.

The catalyst may be periodically regenerated in situ, by isolating reactor34from the normal process flow, purging with an inert gas via line70at a pressure in a range from about 1 to about 5 bar at an elevated temperature in the range of about 400° C. to about 650° C. to remove unreacted material adsorbed on the catalyst insofar as is practical, and then subsequently oxidizing the deposited carbon to CO2by addition of air or inert gas-diluted oxygen to reactor34via line70at a pressure in the range of about 1 bar to about 5 bar at an elevated temperature in the range of about 400° C. to about 650° C. Carbon dioxide and residual air or inert gas is vented from reactor34via line75during the regeneration period.

The effluent which comprises the higher molecular weight hydrocarbon products and hydrobromic acid is withdrawn from the second reactor34via line35and is cooled to a temperature in the range of 0° C. to about 100° C. in exchanger36and combined with vapor effluent in line12from hydrocarbon stripper47, which contains feed gas and residual hydrocarbon products stripped-out by contact with the feed gas in hydrocarbon stripper47. The combined vapor mixture is passed to a scrubber38and contacted with a concentrated aqueous partially-oxidized metal bromide salt solution containing metal hydroxide and/or metal oxide and/or metal oxy-bromide species, which is transported to scrubber38via line41by any suitable means, such as by pump42. The preferred metal of the bromide salt is Fe(III), Cu(II) or Zn(II), or mixtures thereof, as these are less expensive and readily oxidize at lower temperatures in the range of about 120° C. to about 180° C., allowing the use of fluorpolymer-lined equipment; although Co(II), Ni(II), Mn(II), V(II), Cr(II) or other transition-metals which form oxidizable bromide salts may be used in the process of the present invention. Alternatively, alkaline-earth metals which also form oxidizable bromide salts, such as Ca (II) or Mg(II) may be used. Any liquid hydrocarbon product condensed in scrubber38may be skimmed and withdrawn in line37and added to liquid hydrocarbon product exiting the product recovery unit52in line54. Hydrobromic acid is dissolved in the aqueous solution and neutralized by the metal hydroxide and/or metal oxide and/or metal oxy-bromide species to yield metal bromide salt in solution and water which is removed from the scrubber38via line44.

The residual vapor phase containing the higher molecular weight hydrocarbon products that is removed as effluent from the scrubber38is forwarded via line39to dehydrator50to remove substantially all water via line53from the vapor stream. The water is then removed from the dehydrator50via line53. The dried vapor stream containing the higher molecular weight hydrocarbon products is further passed via line51to product recovery unit52to recover propane and butane as desired, but primarily the C5+ fraction as a liquid product in line54. Any conventional method of dehydration and liquids recovery, such as solid-bed dessicant adsorption followed by refrigerated condensation, cryogenic expansion, or circulating absorption oil, as used to process natural gas or refinery gas streams, as will be evident to a skilled artisan, may be employed in the process of the present invention. The residual vapor effluent from product recovery unit52is then split into a purge stream57which may be utilized as fuel for the process and a recycled residual vapor which is compressed via compressor58. The recycled residual vapor discharged from compressor58is split into two fractions. A first fraction that is equal to at least 2.5 times the feed gas molar volume is transported via line62and is combined with dry liquid bromine conveyed by pump24, heated in exchanger26to vaporize the bromine and fed into first reactor30. The second fraction is drawn off of line62via line63and is regulated by control valve60, at a rate sufficient to dilute the alkyl bromide concentration to reactor34and absorb the heat of reaction such that reactor34is maintained at the selected operating temperature, preferably in the range of about 300° C. to about 400° C. in order to optimize conversion versus selectivity and to minimize the rate of catalyst deactivation due to the deposition of carbon. Thus, the dilution provided by the recycled vapor effluent permits selectivity of bromination in the first reactor30to be controlled in addition to moderating the temperature in second reactor34.

Water containing metal bromide salt in solution which is removed from scrubber38via line44is passed to hydrocarbon stripper47wherein residual dissolved hydrocarbons are stripped from the aqueous phase by contact with incoming feed gas transported via line11. The stripped aqueous solution is transported from hydrocarbon stripper47via line65and is cooled to a temperature in the range of about 0° C. to about 70° C. in heat exchanger46and then passed to absorber48in which residual bromine is recovered from vent stream in line67. The aqueous solution effluent from scrubber48is transported via line49to a heat exchanger40to be preheated to a temperature in the range of about 100° C. to about 600° C., and most preferably in the range of about 120° C. to about 180° C. and passed to third reactor16. Oxygen or air is delivered via line10by blower or compressor13at a pressure in the range of about ambient to about 5 bar to bromine stripper14to strip residual bromine from water which is removed from stripper14in line64and is combined with water stream53from dehydrator50to form water effluent stream in line56which is removed from the process. The oxygen or air leaving bromine stripper14is fed via line15to reactor16which operates at a pressure in the range of about ambient to about 5 bar and at a temperature in the range of about 100° C. to about 600° C., but most preferably in the range of about 120° C. to about 180° C. so as to oxidize an aqueous metal bromide salt solution to yield elemental bromine and metal hydroxide and/or metal oxide and or metal oxy-bromide species. As stated above, although Co(II), Ni(II), Mn(II), V(II), Cr(II) or other transition-metals which form oxidizable bromide salts can be used, the preferred metal of the bromide salt is Fe(III), Cu(II), or Zn(II), or mixtures thereof, as these are less expensive and readily oxidize at lower temperatures in the range of about 120° C. to about 180° C., allowing the use of fluorpolymer-lined equipment. Alternatively, alkaline-earth metals which also form oxidizable bromide salts, such as Ca (II) or Mg(II) could be used.

Hydrobromic acid reacts with the metal hydroxide and/or metal oxide and/or metal oxy-bromide species so formed to once again yield the metal bromide salt and water. Heat exchanger18in reactor16supplies heat to vaporize water and bromine. Thus, the overall reactions result in the net oxidation of hydrobromic acid produced in first reactor30and second reactor34to elemental bromine and steam in the liquid phase catalyzed by the metal bromide/metal oxide or metal hydroxide operating in a catalytic cycle. In the case of the metal bromide being Fe(III)Br3, the reactions are believed to be:
Fe(+3a)+6Br(−a)+3H(+a)+ 3/2O2(g)=3Br2(g)+Fe(OH)3  1)
3HBr(g)+H2O=3H(+a)+3Br(−a)+H2O  2)
3H(+a)+3Br(−a)+Fe(OH)3=Fe(+3a)+3Br(−a)+3H2O  3)

The elemental bromine and water and any residual oxygen or nitrogen (if air is utilized as the oxidant) leaving as vapor from the outlet of third reactor16via line19, are cooled in condenser20at a temperature in the range of about 0° C. to about 70° C. and a pressure in the range of about ambient to 5 bar to condense the bromine and water and passed to three-phase separator22. In three-phase separator22, since liquid water has a limited solubility for bromine, on the order of about 3% by weight, any additional bromine which is condensed forms a separate, denser liquid bromine phase. The liquid bromine phase, however, has a notably lower solubility for water, on the order of less than 0.1%. Thus a substantially dry bromine vapor can be easily obtained by condensing liquid bromine and water, decanting water by simple physical separation and subsequently re-vaporizing liquid bromine.

Liquid bromine is pumped in line25from three-phase separator22via pump24to a pressure sufficient to mix with vapor stream62. Thus bromine is recovered and recycled within the process. The residual oxygen or nitrogen and any residual bromine vapor which is not condensed exits three-phase separator22and is passed via line23to bromine scrubber48, wherein residual bromine is recovered by solution into and by reaction with reduced metal bromides in the aqueous metal bromide solution stream65. Water is removed from separator22via line27and introduced into stripper14.

In another embodiment of the invention, referring toFIG. 3, a gas stream containing lower molecular weight alkanes, comprised of mixture of a feed gas plus a recycled gas stream at a pressure in the range of about 1 bar to about 30 bar, is transported or conveyed via line, pipe or conduit162, mixed with dry bromine liquid transported via pump124and passed to heat exchanger126wherein the liquid bromine is vaporized. The mixture of lower molecular weight alkanes and dry bromine vapor is fed to reactor130. Preferably, the molar ratio of lower molecular weight alkanes to dry bromine vapor in the mixture introduced into reactor130is in excess of 2.5:1. Reactor130has an inlet pre-heater zone128which heats the mixture to a reaction initiation temperature in the range of about 250° C. to about 400° C. In first reactor130, the lower molecular weight alkanes are reacted exothermically with dry bromine vapor at a relatively low temperature in the range of about 250° C. to about 600° C., and at a pressure in the range of about 1 bar to about 30 bar to produce gaseous alkyl bromides and hydrobromic acid vapors. The upper limit of the operating temperature range is greater than the upper limit of the reaction initiation temperature range to which the feed mixture is heated due to the exothermic nature of the bromination reaction. In the case of methane, the formation of methyl bromide occurs in accordance with the following general reaction:
CH4(g)+Br2(g)→CH3Br(g)+HBr(g)
This reaction occurs with a significantly high degree of selectivity to methyl bromide. For example, in the case of bromine reacting with a molar excess of methane at a methane to bromine ratio of 4.5:1, selectivity to the mono-halogenated methyl bromide is in the range of 90 to 95%. Small amounts of dibromomethane and tribromomethane are also formed in the bromination reaction. Higher alkanes, such as ethane, propane and butane, are also readily brominated resulting in mono and multiple brominated species. If an alkane to bromine ratio of significantly less than 2.5 to 1 is utilized, selectivity to methyl bromide substantially lower than 90% occurs and significant formation of undesirable carbon soot is observed. It has also been shown that other alkanes such as ethane, propane and butane which may be present in the feed gas to the bromination are readily brominated to form ethyl bromides, propyl bromides and butyl bromides. Further, the dry bromine vapor that is feed into first reactor130is substantially water-free. Applicant has discovered that elimination of substantially all water vapor from the bromination step in first reactor130substantially eliminates the formation of unwanted carbon dioxide thereby increasing the selectivity of alkane bromination to alkyl bromides and eliminating the large amount of waste heat generated in the formation of carbon dioxide from alkanes.

The effluent that contains alkyl bromides and hydrobromic acid is withdrawn from the first reactor130via line131and is partially cooled to a temperature in the range of about 150° C. to 350° C. in heat exchanger132before flowing to a second reactor134. In second reactor134, the alkyl bromides are reacted exothermically at a temperature range of from about 150° C. to about 450° C., and a pressure in the range of about 1 bar to 30 bar, over a fixed bed133of crystalline alumino-silicate catalyst, preferably a zeolite catalyst, and most preferably a ZSM-5 zeolite catalyst. Although the zeolite catalyst is preferably used in the hydrogen, sodium or magnesium form, the zeolite may also be modified by ion exchange with other alkali metal cations, such as Li, Na, K or Cs, with alkali-earth metal cations, such as Mg, Ca, Sr or Ba, or with transition metal cations, such as Ni, Mn, V, W, or to the hydrogen form. Other zeolite catalysts having varying pore sizes and acidities, which are synthesized by varying the alumina-to-silica ratio may be used in the second reactor134as will be evident to a skilled artisan. In this reactor, the alkyl bromides are oligimerized to produce a mixture of higher molecular weight hydrocarbon products and additional hydrobromic acid vapor.

The temperature at which the second reactor134is operated is an important parameter in determining the selectivity of the oligimerization reaction to various higher molecular weight liquid hydrocarbon products. It is preferred to operate second reactor134at a temperature within the range of about 150° to 450°, but more preferably within the range of about 300 C. to 400 C. Temperatures above about 300° C. in the second reactor result in increased yields of light hydrocarbons, such as undesirable methane, whereas lower temperatures increase yields of heavier molecular weight hydrocarbon products. At the low end of the temperature range, methyl bromide reacting over ZSM-5 zeolite at temperatures as low as 150° C. significant methyl bromide conversion on the order of 20% is noted, with a high selectivity towards C5+ products. Notably, in the case of alkyl bromides reacting over the preferred ZSM-5 zeolite catalyst, cyclization reactions occur such that the C7+ fractions produced contain a high percentage of substituted aromatics. At increasing temperatures approaching 300° C., methyl bromide conversion increases towards 90% or greater, however selectivity towards C5+ products decreases and selectivity towards lighter products, particularly undesirable methane, increases. Surprisingly, very little ethane or C2-C4olefin compounds are produced. At temperatures approaching 450° C. almost complete conversion of methyl bromide to methane occurs. In the optimum range of operating temperatures of about 300° C. to 400° C., as a byproduct of the reaction, a small amount of carbon will build up on the catalyst over time during operation, causing a decline in catalyst activity over a range of several hundred hours, depending on the reaction conditions and feed gas composition. It is observed that higher reaction temperatures above about 400° C. favor the thermal cracking of alkyl bromides with formation of carbon and hence increases the rate of deactivation of the catalyst. Conversely, operation at the lower end of the temperature range, particularly below about 300° C. may also promote coking, likely to the reduced rate of desorption of hydrocarbon products. Hence, operating temperatures within the range of about 150° C. to 450° C. but more preferably in the range of about 300° C. to 400° C. in the second reactor134balance increased selectivity towards the desired products and lower rates of deactivation due to carbon formation, against higher conversion per pass, which minimizes the quantity of catalyst, recycle rates and equipment size required.

The catalyst may be periodically regenerated in situ, by isolating reactor134from the normal process flow, purging with an inert gas via line170at a pressure in the range of about 1 bar to about 5 bar and an elevated temperature in the range of 400° C. to 650° C. to remove unreacted material adsorbed on the catalyst insofar as is practical, and then subsequently oxidizing the deposited carbon to CO2by addition of air or inert gas-diluted oxygen via line170to reactor134at a pressure in the range of about 1 bar to about 5 bar and an elevated temperature in the range of 400° C. to 650° C. Carbon dioxide and residual air or inert gas are vented from reactor134via line175during the regeneration period.

The effluent which comprises the higher molecular weight hydrocarbon products and hydrobromic acid is withdrawn from the second reactor134via line135, cooled to a temperature in the range of about 0° C. to about 100° C. in exchanger136, and combined with vapor effluent in line112from hydrocarbon stripper147. The mixture is then passed to a scrubber138and contacted with a stripped, recirculated water that is transported to scrubber138in line164by any suitable means, such as pump143, and is cooled to a temperature in the range of about 0° C. to about 50° C. in heat exchanger155. Any liquid hydrocarbon product condensed in scrubber138may be skimmed and withdrawn as stream137and added to liquid hydrocarbon product154. Hydrobromic acid is dissolved in scrubber138in the aqueous solution which is removed from the scrubber138via line144, and passed to hydrocarbon stripper147wherein residual hydrocarbons dissolved in the aqueous solution are stripped-out by contact with feed gas111. The stripped aqueous phase effluent from hydrocarbon stripper147is cooled to a temperature in the range of about 0° C. to about 50° C. in heat exchanger146and then passed via line165to absorber148in which residual bromine is recovered from vent stream167.

The residual vapor phase containing the higher molecular weight hydrocarbon products is removed as effluent from the scrubber138and forwarded via line139to dehydrator150to remove substantially all water from the gas stream. The water is then removed from the dehydrator150via line153. The dried gas stream containing the higher molecular weight hydrocarbon products is further passed via line151to product recovery unit152to recover C3and C4as desired, but primarily the C5+ fraction as a liquid product in line154. Any conventional method of dehydration and liquids recovery such as solid-bed dessicant adsorption followed by, for example, refrigerated condensation, cryogenic expansion, or circulating absorption oil, as used to process natural gas or refinery gas streams, as known to a skilled artisan, may be employed in the implementation of this invention. The residual vapor effluent from product recovery unit152is then split into a purge stream157that may be utilized as fuel for the process and a recycled residual vapor which is compressed via compressor158. The recycled residual vapor discharged from compressor158is split into two fractions. A first fraction that is equal to at least 2.5 times the feed gas volume is transported via line162, combined with the liquid bromine conveyed in line125and passed to heat exchanger126wherein the liquid bromine is vaporized and fed into first reactor130. The second fraction which is drawn off line162via line163and is regulated by control valve160, at a rate sufficient to dilute the alkyl bromide concentration to reactor134and absorb the heat of reaction such that reactor134is maintained at the selected operating temperature, preferably in the range of about 300° C. to about 400° C. in order to optimize conversion vs. selectivity and to minimize the rate of catalyst deactivation due to the deposition of carbon. Thus, the dilution provided by the recycled vapor effluent permits selectivity of bromination in the first reactor130to be controlled in addition to moderating the temperature in second reactor134.

Oxygen, oxygen enriched air or air110is delivered via blower or compressor113at a pressure in the range of about ambient to about 5 bar to bromine stripper114to strip residual bromine from water which leaves stripper114via line164and is divided into two portions. The first portion of the stripped water is recycled via line164, cooled in heat exchanger155to a temperature in the range of about 20° C. to about 50° C., and maintained at a pressure sufficient to enter scrubber138by any suitable means, such as pump143. The portion of water that is recycled is selected such that the hydrobromic acid solution effluent removed from scrubber138via line144has a concentration in the range from about 10% to about 50% by weight hydrobromic acid, but more preferably in the range of about 30% to about 48% by weight to minimize the amount of water which must be vaporized in exchanger141and preheater119and to minimize the vapor pressure of HBr over the resulting acid. A second portion of water from stripper114is removed from line164and the process via line156.

The dissolved hydrobromic acid that is contained in the aqueous solution effluent from scrubber148is transported via line149and is combined with the oxygen, oxygen enriched air or air leaving bromine stripper114in line115. The combined aqueous solution effluent and oxygen, oxygen enriched air or air is passed to a first side of heat exchanger141and through preheater119wherein the mixture is preheated to a temperature in the range of about 100° C. to about 600° C. and most preferably in the range of about 120° C. to about 180° C. and passed to third reactor117that contains a metal bromide salt. The preferred metal of the bromide salt is Fe(III), Cu(II) or Zn(II) although Co(II), Ni(II), Mn(II), V(II), Cr(II) or other transition-metals which form oxidizable bromide salts can be used. Alternatively, alkaline-earth metals which also form oxidizable bromide salts, such as Ca (II) or Mg(II) could be used. The metal bromide salt in the oxidation reactor117can be utilized as a concentrated aqueous solution or preferably, the concentrated aqueous salt solution may be imbibed into a porous, high surface area, acid resistant inert support such as a silica gel. The oxidation reactor117operates at a pressure in the range of about ambient to about 5 bar and at a temperature in the range of about 100° C. to 600° C., but most preferably in the range of about 120° C. to 180° C.; therein, the metal bromide is oxidized by oxygen, yielding elemental bromine and metal hydroxide, metal oxide or metal oxy-bromide species or, metal oxides in the case of the supported metal bromide salt operated at higher temperatures and lower pressures at which water may primarily exist as a vapor. In either case, the hydrobromic acid reacts with the metal hydroxide, metal oxy-bromide or metal oxide species and is neutralized, restoring the metal bromide salt and yielding water. Thus, the overall reaction results in the net oxidation of hydrobromic acid produced in first reactor130and second reactor134to elemental bromine and steam, catalyzed by the metal bromide/metal hydroxide or metal oxide operating in a catalytic cycle. In the case of the metal bromide being Fe(III)Br2 in an aqueous solution and operated in a pressure and temperature range in which water may exist as a liquid the reactions are believed to be:
Fe(+3a)+6Br(−a)+3H(+a)+ 3/2O2(g)=3Br2(g)+Fe(OH)3  1)
3HBr(g)+H2O=3H(+a)+3Br(−a)+H2O  2)
3H(+a)+3Br(−a)+Fe(OH)3=Fe(+3a)+3Br(−a)+3H2O  3)
In the case of the metal bromide being Cu(II)Br2 supported on an inert support and operated at higher temperature and lower pressure conditions at which water primarily exists as a vapor, the reactions are believed to be:
2Cu(II)Br2=2Cu(I)Br+Br2(g)  1)
2Cu(I)Br+O2(g)=Br2(g)+2Cu(II)O  2)
2HBr(g)+Cu(II)O=Cu(II)Br2+H2O(g)  3)

The elemental bromine and water and any residual oxygen or nitrogen (if air or oxygen enriched air is utilized as the oxidant) leaving as vapor from the outlet of third reactor117, are transported via line127and are cooled in the second side of exchanger141and condenser120to a temperature in the range of about 0° C. to about 70° C. wherein the bromine and water are condensed and passed to three-phase separator122. In three-phase separator122, since liquid water has a limited solubility for bromine, on the order of about 3% by weight, any additional bromine which is condensed forms a separate, denser liquid bromine phase. The liquid bromine phase, however, has a notably lower solubility for water, on the order of less than 0.1%. Thus, a substantially dry bromine vapor can be easily obtained by condensing liquid bromine and water, decanting water by simple physical separation and subsequently re-vaporizing liquid bromine. It is important to operate at conditions that result in the near complete reaction of HBr so as to avoid significant residual HBr in the condensed liquid bromine and water, as HBr increases the miscibility of bromine in the aqueous phase, and at sufficiently high concentrations, results in a single ternary liquid phase.

Liquid bromine is pumped from three-phase separator122via pump124to a pressure sufficient to mix with vapor stream162. Thus the bromine is recovered and recycled within the process. The residual air, oxygen enriched air or oxygen and any bromine vapor which is not condensed exits three-phase separator122and is passed via line123to bromine scrubber148, wherein residual bromine is recovered by dissolution into hydrobromic acid solution stream conveyed to scrubber148via line165. Water is removed from the three-phase separator122via line129and passed to stripper114.

The following examples demonstrate the practice and utility of the present invention, but are not to be construed as limiting the scope thereof.

Various mixtures of dry bromine and methane are reacted homogeneously at temperatures in the range of 459° C. to 491° C. at a Gas Hourly Space Velocity (GHSV which is defined as the gas flow rate in standard liters per hour divided by the gross reactor catalyst-bed volume, including catalyst-bed porosity, in liters) of approximately 7200 hr−1. The results of this example indicate that for molar ratios of methane to bromine greater than 4.5:1 selectivity to methyl bromide is in the range of 90 to 95%, with near-complete conversion of bromine.

FIG. 7andFIG. 8illustrate two exemplary PONA analyses of two C6+ liquid product samples that are recovered during two test runs with methyl bromide and methane reacting over ZSM-5 zeolite catalyst. These analyses show the substantially aromatic content of the C6+ fractions produced.

Methyl bromide is reacted over a ZSM-5 zeolite catalyst at a Gas Hourly Space Velocity (GHSV) of approximately 94 hr−1over a range of temperatures from about 100° C. to about 460° C. at approximately 2 bar pressure. As illustrated inFIG. 4, which is a graph of methyl bromide conversion and product selectivity for the oligimerization reaction as a function of temperature, methyl bromide conversion increases rapidly in the range of about 200° C. to about 350° C. Lower temperatures in the range of about 100° C. to about 250° C. favor selectivity towards higher molecular weight products however conversion is low. Higher temperatures in the range of about 250° C. to about 350° C. show higher conversions in the range of 50% to near 100%, however increasing selectivity to lower molecular weight products, in particular undesirable methane is observed. At higher temperatures above 350° C. selectivity to methane rapidly increases. At about 450° C., almost complete conversion to methane occurs.

Methyl bromide, hydrogen bromide and methane are reacted over a ZSM-5 zeolite catalyst at approximately 2 bar pressure at about 250° C. and also at about 260° C. at a GHSV of approximately 76 hr−1. Comparison tests utilizing a mixture of only methyl bromide and methane without hydrogen bromide over the same ZSM-5 catalyst at approximately the same pressure at about 250° C. and at about 260° C. at a GHSV of approximately 73 hr−1were also run.FIG. 5, which is a graph that illustrates the comparative conversions and selectivities of several example test runs, shows only a very minor effect due to the presence of HBr on product selectivities. Because hydrobromic acid has only a minor effect on conversion and selectivity, it is not necessary to remove the hydrobromic acid generated in the bromination reaction step prior to the conversion reaction of the alkyl bromides, in which additional hydrobromic acid is formed in any case. Thus, the process can be substantially simplified.

Methyl bromide is reacted over a ZSM-5 zeolite catalyst at 230° C. Dibromomethane is added to the reactor.FIG. 6, which is a graph of product selectivity, indicates that reaction of methyl bromide and dibromomethane results in a shift in selectivity towards C5+ products versus. methyl bromide alone. Thus, these results demonstrate that dibromomethane is also reactive and therefore very high selectivity to bromomethane in the bromination step is not required in the process of the present invention. It has been observed, however, that the presence of dibromomethane increases the rate of catalyst deactivation, requiring a higher operating temperature to optimize the tradeoff between selectivity and deactivation rate, as compared to pure methyl bromide.

A mixture of 12.1 mol % methyl bromide and 2.8 mol % propyl bromide in methane are reacted over a ZSM-5 zeolite catalyst at 295 C and a GHSV of approximately 260 hr−1. A methyl bromide conversion of approximately 86% and a propyl bromide conversion of approximately 98% is observed.

Thus, in accordance with all embodiments of the present invention set forth above, the metal bromide/metal hydroxide, metal oxy-bromide or metal oxide operates in a catalytic cycle allowing bromine to be easily recycled within the process. The metal bromide is readily oxidized by oxygen, oxygen enriched air or air either in the aqueous phase or the vapor phase at temperatures in the range of about 100° C. to about 600° C. and most preferably in the range of about 120° C. to about 180° C. to yield elemental bromine vapor and metal hydroxide, metal oxy-bromide or metal oxide. Operation at temperatures below about 180° C. is advantageous, thereby allowing the use of low-cost corrosion-resistant fluoropolymer-lined equipment. Hydrobromic acid is neutralized by reaction with the metal hydroxide or metal oxide yielding steam and the metal bromide.

The elemental bromine vapor and steam are condensed and easily separated in the liquid phase by simple physical separation, yielding substantially dry bromine. The absence of significant water allows selective bromination of alkanes, without production of CO2and the subsequent efficient and selective oligimerization and cyclization reactions of alkyl bromides to primarily propane and heavier products, the C5+ fraction of which contains substantial branched alkanes and substituted aromatics. Byproduct hydrobromic acid vapor from the bromination and oligimerization reaction are readily dissolved into an aqueous phase and neutralized by the metal hydroxide or metal oxide species resulting from oxidation of the metal bromide.

In accordance with another embodiment of the process of the present invention illustrated inFIG. 9A, the alkyl bromination and alkyl bromide conversion stages are operated in a substantially similar manner to those corresponding stages described with respect toFIGS. 2 and 3above. More particularly, a gas stream containing lower molecular weight alkanes, comprised of mixture of a feed gas and a recycled gas stream at a pressure in the range of about 1 bar to about 30 bar, is transported or conveyed via line, pipe or conduits262and211, respectively, and mixed with dry bromine liquid in line225. The resultant mixture is transported via pump224and passed to heat exchanger226wherein the liquid bromine is vaporized. The mixture of lower molecular weight alkanes and dry bromine vapor is fed to reactor230. Preferably, the molar ratio of lower molecular weight alkanes to dry bromine vapor in the mixture introduced into reactor230is in excess of 2.5:1. Reactor230has an inlet pre-heater zone228which heats the mixture to a reaction initiation temperature in the range of 250° C. to 400° C. In first reactor230, the lower molecular weight alkanes are reacted exothermically with dry bromine vapor at a relatively low temperature in the range of about 250° C. to about 600° C., and at a pressure in the range of about 1 bar to about 30 bar to produce gaseous alkyl bromides and hydrobromic acid vapors. The upper limit of the operating temperature range is greater than the upper limit of the reaction initiation temperature range to which the feed mixture is heated due to the exothermic nature of the bromination reaction. In the case of methane, the formation of methyl bromide occurs in accordance with the following general reaction:
CH4(g)+Br2(g)→CH3Br(g)+HBr(g)
This reaction occurs with a significantly high degree of selectivity to methyl bromide. For example, in the case of bromine reacting with a molar excess of methane at a methane to bromine ratio of 4.5:1, selectivity to the mono-halogenated methyl bromide is in the range of 90 to 95%. Small amounts of dibromomethane and tribromomethane are also formed in the bromination reaction. Higher alkanes, such as ethane, propane and butane, are also readily bromoninated resulting in mono and multiple brominated species. If an alkane to bromine ratio of significantly less than 2.5 to 1 is utilized, selectivity to methyl bromide substantially lower than 90% occurs and significant formation of undesirable carbon soot is observed. It has also been shown that other alkanes such as ethane and propane which may be present in the feed gas to the bromination are readily brominated to form ethyl bromides and propyl bromides. Further, the dry bromine vapor that is feed into first reactor230is substantially water-free. Applicant has discovered that elimination of substantially all water vapor from the bromination step in first reactor230substantially eliminates the formation of unwanted carbon dioxide thereby increasing the selectivity of alkane bromination to alkyl bromides and eliminating the large amount of waste heat generated in the formation of carbon dioxide from alkanes.

The effluent that contains alkyl bromides and hydrobromic acid is withdrawn from the first reactor230via line231and is partially cooled to a temperature in the range of about 150° C. to 350° C. in heat exchanger232before flowing to a second reactor234. In second reactor234, the alkyl bromides are reacted exothermically at a temperature range of from about 150° C. to about 450° C., and a pressure in the range of about 1 bar to 30 bar, over a fixed bed233of crystalline alumino-silicate catalyst, preferably a zeolite catalyst, and most preferably a ZSM-5 zeolite catalyst. Although the zeolite catalyst is preferably used in the hydrogen, sodium or magnesium form, the zeolite may also be modified by ion exchange with other alkali metal cations, such as Li, K, Na or Cs, with alkali-earth metal cations, such as Mg, Ca, Sr or Ba, with transition metal cations, such as Ni, Mn, V, W, or to the hydrogen form. Other zeolite catalysts having varying pore sizes and acidities, which are synthesized by varying the alumina-to-silica ratio may be used in the second reactor234as will be evident to a skilled artisan. In this reactor, the alkyl bromides are oligimerized to produce a mixture of higher molecular weight hydrocarbon products and additional hydrobromic acid vapor.

The temperature at which the second reactor234is operated is an important parameter in determining the selectivity of the oligimerization reaction to various higher molecular weight liquid hydrocarbon products. It is preferred to operate second reactor234at a temperature within the range of about 150° C. to about 450° C., but more preferably within the range of about 300° C. to about 400° C. Temperatures above about 300° C. in the second reactor result in increased yields of light hydrocarbons, such as undesirable methane, whereas lower temperatures increase yields of heavier molecular weight hydrocarbon products. At the low end of the temperature range, methyl bromide reacting over ZSM-5 zeolite at temperatures as low as about 150° C. significant methyl bromide conversion on the order of 20% is noted, with a high selectivity towards C5+ products. Notably, in the case of alkyl bromides reacting over the preferred ZSM-5 zeolite catalyst, cyclization reactions occur such that the C7+ fractions produced contain a high percentage of substituted aromatics. At increasing temperatures approaching about 300° C., methyl bromide conversion increases towards 90% or greater, however selectivity towards C5+ products decreases and selectivity towards lighter products, particularly undesirable methane, increases. Surprisingly, very little ethane or C2-C4olefin compounds are produced. At temperatures approaching about 450° C. almost complete conversion of methyl bromide to methane occurs. In the optimum temperature range of about 300° C. to about 400° C., as a byproduct of the reaction, a small amount of carbon will build up on the catalyst over time during operation, causing a decline in catalyst activity over a range of hours to several hundred hours, depending on the reaction conditions and feed gas composition. It is believed that higher reaction temperatures over about 400° C. favor the formation of carbon and hence rate of deactivation of the catalyst. Conversely, operation at the lower end of the temperature range, particularly below about 300° C. may also promote coking, likely to the reduced rate of desorption of hydrocarbon products. Hence, operating temperatures within the range of about 150° C. to about 400° C., but more preferably in the range of about 300° C. to about 400° C., in the second reactor234balance increased selectivity towards the desired products and lower rates of deactivation due to carbon formation, against higher conversion per pass, which minimizes the quantity of catalyst, recycle rates and equipment size required.

The catalyst may be periodically regenerated in situ, by isolating reactor234from the normal process flow, purging with an inert gas via line270at a pressure in the range of about 1 bar to about 5 bar and an elevated temperature in the range of about 400° C. to about 650° C. to remove unreacted material adsorbed on the catalyst insofar as is practical, and then subsequently oxidizing the deposited carbon to CO2by addition of air or inert gas-diluted oxygen via line270to reactor234at a pressure in the range of about 1 bar to about 5 bar and an elevated temperature in the range of about 400° C. to about 650° C. Carbon dioxide and residual air or inert gas are vented from reactor234via line275during the regeneration period.

The effluent which comprises the higher molecular weight hydrocarbon products and hydrobromic acid is withdrawn from the second reactor234via line235and cooled to a temperature in the range of about 100° C. to about 600° C. in exchanger236. As illustrated inFIG. 9A, the cooled effluent is transported via lines235and241with valve238in the opened position and valves239and243in the closed position and introduced into a vessel or reactor240containing a bed298of a solid phase metal oxide. The metal of the metal oxide is selected form magnesium (Mg), calcium (Ca), vanadium (V), chromium (Cr), manganese (Mn), iron (Fe), cobalt (Co), nickel (Ni), copper (Cu), zinc (Sn), or tin (Sn). The metal is selected for the impact of its physical and thermodynamic properties relative to the desired temperature of operation, and also for potential environmental and health impacts and cost. Preferably, magnesium, copper and iron are employed as the metal, with magnesium being the most preferred. These metals have the property of not only forming oxides but bromide salts as well, with the reactions being reversible in a temperature range of less than about 500° C. The solid metal oxide is preferably immobilized on a suitable attrition-resistant support, for example a synthetic amorphous silica, such as Davicat Grade 57, manufactured by Davison Catalysts of Columbia, Md. In reactor240, hydrobromic acid is reacted with the metal oxide at temperatures below about 600° C. and preferably between about 300° C. to about 500° C. in accordance with the following general formula wherein M represents the metal:
2HBr+MO→MBr2+H2O
The steam resulting from this reaction is transported together with the high molecular hydrocarbon products in line244,218and216via opened valve219to heat exchanger220wherein the mixture is cooled to a temperature in the range of about 0° C. to about 70° C. This cooled mixture is forwarded to dehydrator250to remove substantially all water from the gas stream. The water is then removed from the dehydrator250via line253. The dried gas stream containing the higher molecular weight hydrocarbon products is further passed via line251to product recovery unit252to recover C3and C4as desired, but primarily the C5+ fraction as a liquid product in line254. Any conventional method of dehydration and liquids recovery such as solid-bed dessicant adsorption followed by, for example, refrigerated condensation, cryogenic expansion, or circulating absorption oil, as used to process natural gas or refinery gas streams, as known to a skilled artisan, may be employed in the implementation of this invention. The residual vapor effluent from product recovery unit252is then split into a purge stream257that may be utilized as fuel for the process and a recycled residual vapor which is compressed via compressor258. The recycled residual vapor discharged from compressor258is split into two fractions. A first fraction that is equal to at least 1.5 times the feed gas volume is transported via line262, combined with the liquid bromine and feed gas conveyed in line225and passed to heat exchanger226wherein the liquid bromine is vaporized and fed into first reactor230in a manner as described above. The second fraction which is drawn off line262via line263and is regulated by control valve260, at a rate sufficient to dilute the alkyl bromide concentration to reactor234and absorb the heat of reaction such that reactor234is maintained at the selected operating temperature, preferably in the range of about 300° C. to about 400° C. in order to optimize conversion vs. selectivity and to minimize the rate of catalyst deactivation due to the deposition of carbon. Thus, the dilution provided by the recycled vapor effluent permits selectivity of bromination in the first reactor230to be controlled in addition to moderating the temperature in second reactor234.

Oxygen, oxygen enriched air or air210is delivered via blower or compressor213at a pressure in the range of about ambient to about 10 bar to bromine via line214, line215and valve249through heat exchanger215, wherein oxygen, oxygen enriched air or air is preheated to a temperature in the range of about 100° C. to about 500° C. to a second vessel or reactor246containing a bed299of a solid phase metal bromide. Oxygen reacts with the metal bromide in accordance with the following general reaction wherein M represents the metal:
MBr2+½O2→MO+Br2
In this manner, a dry, substantially HBr free bromine vapor is produced thereby eliminating the need for subsequent separation of water or hydrobromic acid from the liquid bromine. Reactor246is operated below 600° C., and more preferably between about 300° C. to about 500° C. The resultant bromine vapor is transported from reactor246via line247, valve248and line242to heat exchanger or condenser221where the bromine is condensed into a liquid. The liquid bromine is transported via line242to separator222wherein liquid bromine is removed via line225and transported via line225to heat exchanger226and first reactor230by any suitable means, such as by pump224. The residual air or unreacted oxygen is transported from separator222via line227to a bromine scrubbing unit223, such as venturi scrubbing system containing a suitable solvent, or suitable solid adsorbant medium, as selected by a skilled artisan, wherein the remaining bromine is captured. The captured bromine is desorbed from the scrubbing solvent or adsorbant by heating or other suitable means and the thus recovered bromine transported via line212to line225. The scrubbed air or oxygen is vented via line229. In this manner, nitrogen and any other substantially non-reactive components are removed from the system of the present invention and thereby not permitted to enter the hydrocarbon-containing portion of the process; also loss of bromine to the surrounding environment is avoided.

One advantage of removing the HBr by chemical reaction in accordance with this embodiment, rather than by simple physical solubility, is the substantially complete scavenging of the HBr to low levels at higher process temperatures. Another distinct advantage is the elimination of water from the bromine removed thereby eliminating the need for separation of bromine and water phases and for stripping of residual bromine from the water phase.

Reactors240and246may be operated in a cyclic fashion. As illustrated inFIG. 9A, valves238and219are operated in the open mode to permit hydrobromic acid to be removed from the effluent that is withdrawn from the second reactor234, while valves248and249are operated in the open mode to permit air, oxygen enriched air or oxygen to flow through reactor246to oxidize the solid metal bromide contained therein. Once significant conversion of the metal oxide and metal bromide in reactors240and246, respectively, has occurred, these valves are closed. At this point, bed299in reactor246is a bed of substantially solid metal bromide, while bed298in reactor240is substantially solid metal oxide. As illustrated inFIG. 10A, valves245and243are then opened to permit oxygen, oxygen enriched air or air to flow through reactor240to oxidize the solid metal bromide contained therein, while valves239and217are opened to permit effluent which comprises the higher molecular weight hydrocarbon products and hydrobromic acid that is withdrawn from the second reactor234to be introduced into reactor246via line237. The reactors are operated in this manner until significant conversion of the metal oxide and metal bromide in reactors246and240, respectively, has occurred and then the reactors are cycled back to the flow schematic illustrated inFIG. 9Aby opening and closing valves as previously discussed.

When oxygen is utilized as the oxidizing gas transported in via line210to the reactor being used to oxidize the solid metal bromide contained therein, the embodiment of the process of the present invention illustrated inFIGS. 9A and 10Acan be modified such that the bromine vapor produced from either reactor246(FIG. 9B) or240(FIG. 10B) is transported via lines242and225directly to first reactor230. Since oxygen is reactive and will not build up in the system, the need to condense the bromine vapor to a liquid to remove unreactive components, such as nitrogen, is obviated. Compressor213is not illustrated inFIGS. 9B and 10Bas substantially all commercial sources of oxygen, such as a commercial air separator unit, will provide oxygen to line210at the required pressure. If not, a compressor213could be utilized to achieve such pressure as will be evident to a skilled artisan.

In the embodiment of the present invention illustrated inFIG. 11A, the beds of solid metal oxide particles and solid metal bromide particles contained in reactors240and246, respectively, are fluidized and are connected in the manner described below to provide for continuous operation of the beds without the need to provide for equipment, such as valves, to change flow direction to and from each reactor. In accordance with this embodiment, the effluent which comprises the higher molecular weight hydrocarbon products and hydrobromic acid is withdrawn from the second reactor234via line235, cooled to a temperature in the range of about 100° C. to about 500° C. in exchanger236, and introduced into the bottom of reactor240which contains a bed298of solid metal oxide particles. The flow of this introduced fluid induces the particles in bed298to move upwardly within reactor240as the hydrobromic acid is reacted with the metal oxide in the manner as described above with respect toFIG. 9A. At or near the top of the bed298, the particles which contain substantially solid metal bromide on the attrition-resistant support due to the substantially complete reaction of the solid metal oxide with hydrobromic acid in reactor240are withdrawn via a weir or cyclone or other conventional means of solid/gas separation, flow by gravity down line259and are introduced at or near the bottom of the bed299of solid metal bromide particles in reactor246. In the embodiment illustrated inFIG. 11A, oxygen, oxygen enriched air or air210is delivered via blower or compressor213at a pressure in the range of about ambient to about 10 bar, transported via line214through heat exchanger215, wherein the oxygen, oxygen enriched air or air is preheated to a temperature in the range of about 100° C. to about 500° C. and introduced into second vessel or reactor246below bed299of a solid phase metal bromide. Oxygen reacts with the metal bromide in the manner described above with respect toFIG. 9Ato produce a dry, substantially HBr free bromine vapor. The flow of this introduced gas induces the particles in bed299to flow upwardly within reactor246as oxygen is reacted with the metal bromide. At or near the top of the bed298, the particles which contain substantially solid metal oxide on the attrition-resistant support due to the substantially complete reaction of the solid metal bromide with oxygen in reactor246are withdrawn via a weir or cyclone or other conventional means of solid/gas separation, flow by gravity down line264and are introduced at or near the bottom of the bed298of solid metal oxide particles in reactor240. In this manner, reactors240and246can be operated continuously without changing the parameters of operation.

In the embodiment illustrated inFIG. 11B, oxygen is utilized as the oxidizing gas and is transported in via line210to reactor246. Accordingly, the embodiment of the process of the present invention illustrated inFIG. 11Ais modified such that the bromine vapor produced from reactor246is transported via lines242and225directly to first reactor230. Since oxygen is reactive and will not build up in the system, the need to condense the bromine vapor to a liquid to remove unreactive components, such as nitrogen, is obviated. Compressor213is not illustrated inFIG. 11Bas substantially all commercial sources of oxygen, such as a commercial air separator unit, will provide oxygen to line210at the required pressure. If not, a compressor213could be utilized to achieve such pressure as will be evident to a skilled artisan.

In accordance with another embodiment of the process of the present invention that is illustrated inFIG. 12, the alkyl bromination and alkyl bromide conversion stages are operated in a substantially similar manner to those corresponding stages described in detail with respect toFIG. 9Aexcept as discussed below. Residual air or oxygen and bromine vapor emanating from reactor246is transported via line247, valve248and line242and valve300to heat exchanger or condenser221wherein the bromine-containing gas is cooled to a temperature in the range of about 30° C. to about 300° C. The bromine-containing vapor is then transported via line242to vessel or reactor320containing a bed322of a solid phase metal bromide in a reduced valence state. The metal of the metal bromide in a reduced valence state is selected from copper (Cu), iron (Fe), or molybdenum (Mo). The metal is selected for the impact of its physical and thermodynamic properties relative to the desired temperature of operation, and also for potential environmental and health impacts and cost. Preferably, copper or iron are employed as the metal, with copper being the most preferred. The solid metal bromide is preferably immobilized on a suitable attrition-resistant support, for example a synthetic amorphous silica, such as Davicat Grade 57, manufactured by Davison Catalysts of Columbia, Md. In reactor320, bromine vapor is reacted with the solid phase metal bromide, preferably retained on a suitable attrition-resistant support at temperatures below about 300° C. and preferably between about 30° C. to about 200° C. in accordance with the following general formula wherein M2represents the metal:
2M2Brn+Br2→2M2Brn+1
In this manner, bromine is stored as a second metal bromide, i.e. 2M2Brn+1, in reactor320while the resultant vapor containing residual air or oxygen is vented from reactor320via line324, valve326and line318.

The gas stream containing lower molecular weight alkanes, comprised of mixture of a feed gas (line211) and a recycled gas stream, is transported or conveyed via line262, heat exchanger352, wherein the gas stream is preheated to a temperature in the range of about 150° C. to about 600° C., valve304and line302to a second vessel or reactor310containing a bed312of a solid phase metal bromide in an oxidized valence state. The metal of the metal bromide in an oxidized valence state is selected from copper (Cu), iron (Fe), or molybdenum (Mo). The metal is selected for the impact of its physical and thermodynamic properties relative to the desired temperature of operation, and also for potential environmental and health impacts and cost. Preferably, copper or iron are employed as the metal, with copper being the most preferred. The solid metal bromide in an oxidized state is preferably immobilized on a suitable attrition-resistant support, for example a synthetic amorphous silica such as Davicat Grade 57, manufactured by Davison Catalysts of Columbia, Md. The temperature of the gas stream is from about 150° C. to about 600° C., and preferably from about 200° C. to about—450° C. In second reactor310, the temperature of the gas stream thermally decomposes the solid phase metal bromide in an oxidized valence state to yield elemental bromine vapor and a solid metal bromide in a reduced state in accordance with the following general formula wherein M2represents the metal:
2M2Brn+1→2M2Brn+Br2
The resultant bromine vapor is transported with the gas stream containing lower molecular weight alkanes via lines314,315, valve317, line330, heat exchanger226prior to being introduced into alkyl bromination reactor230.

Reactors310and320may be operated in a cyclic fashion. As illustrated inFIG. 12, valve304is operated in the open mode to permit the gas stream containing lower molecular weight alkanes to be transported to the second reactor310, while valve317is operated in the open mode to permit this gas stream with bromine vapor that is generated in reactor310to be transported to alkyl bromination reactor230. Likewise, valve306is operated in the open mode to permit bromine vapor from reactor246to be transported to reactor320via line307, while valve326is operated in the open mode to permit residual air or oxygen to be vented from reactor320. Once significant conversion of the reduced metal bromide and oxidized metal bromide in reactors320and310, respectively, to the corresponding oxidized and reduced states has occurred, these valves are closed as illustrated inFIG. 13. At this point, bed322in reactor320is a bed of substantially metal bromide in an oxidized state, while bed312in reactor310is substantially metal bromide in a reduced state. As illustrated inFIG. 13, valves304,317,306and326are closed, and then valves308and332are opened to permit the gas stream containing lower molecular weight alkanes to be transported or conveyed via lines262, heat exchanger352, wherein gas stream is heated to a range of about 150° C. to about 600° C., valve308and line,309to reactor320to thermally decompose the solid phase metal bromide in an oxidized valence state to yield elemental bromine vapor and a solid metal bromide in a reduced state. Valve332is also opened to permit the resultant bromine vapor to be transported with the gas stream containing lower molecular weight alkanes via lines324and330and heat exchanger226prior to being introduced into alkyl bromination reactor230. In addition, valve300is opened to permit bromine vapor emanating from reactor246to be transported via line242through exchanger221into reactor310wherein the solid phase metal bromide in a reduced valence state reacts with bromine to effectively store bromine as a metal bromide. In addition, valve316is opened to permit the resulting gas, which is substantially devoid of bromine to be vented via lines314and318. The reactors are operated in this manner until significant conversion of the beds of reduced metal bromide and oxidized metal bromide in reactors310and320, respectively, to the corresponding oxidized and reduced states has occurred and then the reactors are cycled back to the flow schematic illustrated inFIG. 12by opening and closing valves as previously discussed.

In the embodiment of the present invention illustrated inFIG. 14, the beds312and322contained in reactors310and320, respectively, are fluidized and are connected in the manner described below to provide for continuous operation of the beds without the need to provide for equipment, such as valves, to change flow direction to and from each reactor. In accordance with this embodiment, the bromine-containing gas withdrawn from the reactor246via line242is cooled to a temperature in the range of about 30° C. to about 300° C. in exchangers370and372, and introduced into the bottom of reactor320which contains a moving solid bed322in a fluidized state. The flow of this introduced fluid induces the particles in bed322to flow upwardly within reactor320as the bromine vapor is reacted with the reduced metal bromide entering the bottom of bed322in the manner as described above with respect toFIG. 12. At or near the top of the bed322, the particles which contain substantially oxidized metal bromide on the attrition-resistant support due to the substantially complete reaction of the reduced metal bromide with bromine vapor in reactor320are withdrawn via a weir, cyclone or other conventional means of solid/gas separation, flow by gravity down line359and are introduced at or near the bottom of the bed312in reactor310. The resulting gas which is substantially devoid of bromine is vented via line350. In the embodiment illustrated inFIG. 14, the gas stream containing lower molecular weight alkanes, comprised of mixture of a feed gas (line211) and a recycled gas stream, is transported or conveyed via line262and heat exchanger352wherein the gas stream is heated to a range of about 150° C. to about 600° C. and introduced into reactor310. The heated gas stream thermally decomposes the solid phase metal bromide in an oxidized valence state present entering at or near the bottom of bed312to yield elemental bromine vapor and a solid metal bromide in a reduced state. The flow of this introduced gas induces the particles in bed312to flow upwardly within reactor310as the oxidized metal bromide is thermally decomposed. At or near the top of the bed312, the particles which contain substantially reduced solid metal bromide on the attrition-resistant support due to the substantially complete thermal decomposition in reactor310are withdrawn via a weir or cyclone or other conventional means of gas/solid separation and flow by gravity down line364and introduced at or near the bottom of the bed322of particles in reactor310. The resulting bromine vapor is transported with the gas stream containing lower molecular weight alkanes via line354and heat exchanger355and introduced into alkyl bromination reactor230. In this manner, reactors310and320can be operated continuously with changing the parameters of operation.

The process of the present invention is less expensive than conventional process since it operates at low pressures in the range of about 1 bar to about 30 bar and at relatively low temperatures in the range of about 20° C. to about 600° C. for the gas phase, and preferably about 20° C. to about 180° C. for the liquid phase. These operating conditions permit the use of less expensive equipment of relatively simple design that are constructed from readily available metal alloys for the gas phase and polymer-lined vessels, piping and pumps for the liquid phase. The process of the present invention is also more efficient because less energy is required for operation and the production of excessive carbon dioxide as an unwanted byproduct is minimized. The process is capable of directly producing a mixed hydrocarbon product containing various molecular-weight components in the liquefied petroleum gas (LPG) and motor gasoline fuels range that have substantial aromatic content thereby significantly increasing the octane value of the gasoline-range fuel components.

While the foregoing preferred embodiments of the invention have been described and shown, it is understood that the alternatives and modifications, such as those suggested and others, may be made thereto and fall within the scope of the invention.