Patent Application
20150166423
Embodiments relate to a process for converting lower molecular weight, gaseous alkanes to higher molecular weight hydrocarbons that may be useful as fuels or chemical feedstocks. More particularly, embodiments relate to a process wherein a gas containing lower molecular weight alkanes is reacted with bromine vapor, to form alkyl bromides which in turn are reacted over a crystalline alumino-silicate catalyst to form paraffins, olefins, naphthenic and aromatic compounds or mixtures thereof and hydrobromic acid, and that may use in-line hydrogen bromide oxidation for capture of hydrogen bromide produced in the process and re-conversion to bromine.
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. 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 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 from methane. Among these are steam-methane reforming (SMR), partial oxidation (PDX), 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. PDX 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 and 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. Thus, a need exists for an economic process for the conversion of methane and other alkanes found in natural gas to olefins, higher molecular weight hydrocarbons including valuable aromatic compounds or mixtures thereof which, due to their higher density and value, are more economically transported thereby significantly aiding development of remote natural gas reserves. Further, a need exists for such a process that is relatively inexpensive, safe and simple.
An embodiment may provide a process for producing elemental halogen comprising: providing a first stream comprising a hydrogen halide; contacting the first stream with a metal oxide to form water, elemental halogen, and at least some metal halide, wherein the metal oxide comprises a metal capable of forming a plurality of stable oxidation states; and contacting the metal halide with an oxygen source to produce a regenerated metal oxide, wherein the oxygen source contacts the metal halide under conditions sufficient to avoid release of elemental halogen.
An embodiment may provide a process for producing higher hydrocarbons comprising: providing a first stream comprising lower molecular weight alkanes and a hydrogen halide; contacting the first stream with a metal oxide to form oxidation products comprising water, elemental halogen, and a metal halide; reacting at least a portion the methane and at least a portion of the elemental halogen to form halogenation products comprising alkyl halides and a first portion of produced hydrogen halide; and contacting at least a portion of the alkyl halides with a catalyst to produce synthesis products comprising higher hydrocarbons and a second portion of produced hydrogen halide.
An embodiment may comprise a system for producing higher hydrocarbons comprising: an oxidation unit comprising a metal oxide reactor bed for converting a hydrogen halide to a metal halide and elemental halogen, wherein at least a portion of the metal oxide in the metal oxide reactor bed is converted to a metal halide, and wherein the oxidation unit further comprises an offline metal oxide bed containing converted metal halide in fluid communication with an oxygen source; a bromination unit in fluid communication with the oxidation unit for reacting an alkane with the elemental halogen; and a synthesis unit in fluid communication with the bromination unit, wherein the synthesis unit comprises a catalyst for converting alkyl halides to higher hydrocarbons.
The features and advantages of the present invention will be readily apparent to those skilled in the art. While numerous changes may be made by those skilled in the art, such changes are within the spirit of the invention.
These drawings illustrate certain aspects of some of the embodiments of the present invention, and should not be used to limit or define the invention.
Embodiments relate to a process for converting lower molecular weight, gaseous alkanes to higher molecular weight hydrocarbons that may be useful as fuels or chemical feedstocks. More particularly, embodiments relate to a process for converting lower molecular weight, gaseous alkanes to higher molecular weight hydrocarbons, such as paraffins, olefins, naphthenic and aromatic compounds, or mixtures thereof, that may use in-line hydrogen bromide oxidation for capture of hydrogen bromide produced in the process.
As utilized throughout this description, the term “lower molecular weight alkanes” refers to methane, ethane, propane, butane, pentane or any combination thereof including combinations such as natural gas. As used herein, the use of the term Cn refers to a hydrocarbon with a carbon number corresponding to the subscript “n”. In addition, the use of Cn+ refers to a hydrocarbon or a hydrocarbon group comprising hydrocarbons of “n” carbon atoms and/or any hydrocarbon having a number of carbon atoms greater than “n.” Therefore, the use of the designation C5+ indicates that the fraction may include any proportion of C5 hydrocarbons and may also include any hydrocarbons with 5 or more carbon atoms (e.g., C6 hydrocarbons, C7 hydrocarbons, etc.) As also utilized throughout this description, “alkyl halides” and “alkyl bromides” refer to mono, di, and tri halogenated and brominated alkanes, respectively.
Feed gas that may be used in the processes described herein may be from any suitable source, for example, any source of gas that provides lower molecular weight alkanes, whether naturally occurring or synthetically produced. Examples of sources of lower molecular weight alkanes for use in the processes of the present invention include, but are not limited to, natural gas, coal-bed methane, regasified liquefied natural gas, gas derived from gas hydrates and/or chlathrates, gas derived from anaerobic decomposition of organic matter or biomass, gas derived in the processing of tar sands, and synthetically produced natural gas or alkanes. Combinations of these may be suitable as well in some embodiments. Also, the feed gas may comprise natural gas which may be treated to remove sulfur compounds and carbon dioxide. In any event, it is important to note that embodiments may tolerate small amounts of carbon dioxide, e.g., less than about 2 mol %, in the feed gas.
As used herein, the term “higher hydrocarbons” when used in the context of a coupling, synthesis, or oligomerization reaction refers to higher molecular weight hydrocarbons having a greater number of carbon atoms than one or more hydrocarbon components of the feed gas, as well as olefinic hydrocarbons having the same or a greater number of carbon atoms as one or more hydrocarbon components of the feed gas stream. For instance, if the feed gas comprises natural gas, which may typically be a mixture of lower molecular weight alkanes, predominately methane, with lesser amounts of ethane, propane, and butane, and even smaller amounts of longer chain hydrocarbons such as pentane, hexane, etc., the “higher hydrocarbon(s)” produced according to the invention may include a C2 or higher hydrocarbon, such as ethane, propane, butane, C5+ hydrocarbons, aromatic hydrocarbons, etc., and optionally ethylene, propylene, and/or longer olefins.
While there are many advantages to the disclosed processes, only a few are discussed herein. Embodiments may allow for the capture of a hydrogen halide produced by the conversion of lower molecular weight alkanes to higher hydrocarbons according the process disclosed herein. The use of a metal oxide that may be converted to a metal halide with a plurality of stable oxidation states may allow for hydrogen halide to be captured while simultaneously producing elemental halogen. This process may allow oxygen to be introduced into the system without directly contacting a hydrocarbon containing stream. Further, complete separation of the hydrogen halide from the synthesis product stream may be avoided, allowing a more cost effective process to be implemented.
A block flow diagram generally depicting an embodiment of a process for producing higher hydrocarbons is illustrated in
The methane in the C1 recycle stream 25 may pass through the oxidation unit 27 along with the elemental halogen and water vapor or steam produced in the oxidation unit 27 to form an oxidation product stream 29. The oxidation product stream 29 may then pass to the condenser 31 which may be operated at a temperature that is below the dewpoint of water and above the dewpoint of the halogen as would be apparent to the skilled practitioner, so as to remove a substantial amount of the water vapor contained in oxidation product stream 29, forming a water condensate stream 33 comprising mostly water but that may also contain some halogen. The vapor effluent 35 exiting condenser 31 may then be reheated in in-line heater 37 to a temperature in the range of about 150° C. to about 300° C. The reheated vapor effluent stream exiting in-line heater 37 may then be divided into two portions, a first portion stream 39 and a second portion stream 41, which may be supplied to the C2+ halogenation unit 23 and the C1+ halogenation unit 41. Within the C2+ halogenation unit 23, ethane, ethylene, propane, propylene and any other heavier hydrocarbons contained in C2-C3 recycle stream 21 may react with the elemental halogen contained in first portion stream 39 to form C2+ alkyl halide stream 43. The C2+ alkyl halide stream 43 may comprise C2 and higher alkyl halides, hydrogen halide, and unreacted C2 and higher hydrocarbons. The balance of reheated vapor effluent leaving in-line heater 37 comprises the second portion stream 41, which may be routed to C1+ halogenation unit 45. Within the C1+ halogenation unit 41, the methane and small amounts of C2 and heavier hydrocarbons may react with the elemental halogen contained in second portion stream 41 to form C1+ alkyl halide stream 47. The C1+ alkyl halide stream 47 may comprise mostly C1 alkyl halides and small amounts of C2 and heavier alkyl halides, hydrogen halide and unreacted methane. The C2+ alkyl halide stream 43 may be combined with C1+ alkyl halide stream 47 which may then pass to the synthesis unit 5 wherein the alkyl halides may react over a suitable catalyst to form the synthesis product stream 3, wherein the whole process may be repeated. Accordingly, the process illustrated in
As generally described above in reference to
In an alternate embodiment the C2-C3 hydrocarbon fraction may be withdrawn as additional products from the process rather than recycled, obviating the need for C2+ halogenation unit 23.
In another embodiment, if the feed gas 1 comprises essentially pure methane, the C1 hydrocarbon fraction may include the relatively small amount of C2 hydrocarbon produced in the synthesis unit 5, for example, comprising less than 2 mol % of C2 of the fraction, and also include the hydrogen halide entering the separation unit 9 and may leave the separation unit 9 in methane stream 15; in this case C2-C3 recycle stream 21 may comprise mostly C3 hydrocarbons and may include only small amounts of C2 hydrocarbons and hydrogen halide.
The amount of C2-C3 hydrocarbons in the methane stream 15 may be varied by changing the conditions in the separation unit 9 but, in some embodiments, may be maintained at low concentration levels of less than about 2 mol % C2 of the fraction, and less than about 0.1 mol % C3 of the fraction. As shown in
As illustrated in
Without wishing to be limited by theory, it is believed that when a gaseous stream (e.g., C1 recycle stream 25) comprising a hydrogen halide (e.g., hydrogen bromide) contacts a metal oxide (e.g., metal oxide reactor bed 49) with a plurality of oxidation states, the hydrogen halide may be removed from the stream with the production of the corresponding molecular halogen. The rate of reaction and the production of the corresponding elemental halogen may depend on the operating conditions, and is particularly dependent on the temperature at which the reaction occurs. In an embodiment, the capture of a hydrogen halide with a metal oxide/metal halide with multiple oxidation states may occur at a temperature ranging from about 150° C. to about 350° C. and more alternatively at a temperature within a range about 200° C. to about 300° C. The reactions may be exemplified by the following equations using copper oxide and hydrogen bromide as examples:
2 HBr(g)+CuO(s)→CuBr2(s)+H2O(g) (Keq on the order of 7.8×108 @ 250° C.)
2 CuBr2(s)→2 CuBr(s)+Br2(g) (Keq on the order of 3×10−1 @ 250° C.)
Upon contact of the hydrogen halide with the metal oxide, the reaction may continue until a large portion of the metal oxide on the metal oxide reactor bed 49 may be converted into a metal halide. However, before all, or substantially all, of the metal oxide has been converted to metal halide, the metal oxide reactor bed 49 may be taken offline and regenerated. As illustrated on
2 CuBr(s)+O2(g)→2 CuO(s)+Br2(g) (Keq on the order of 2.6×106 @ 100° C.)
2 CuBr(s)+Br2(g)→2 CuBr2(s) (Keq on the order of 2.7×104 @ 100° C.)
Based on these assumed reaction mechanisms, a metal halide with multiple oxidation states may be used to shuttle elemental oxygen from an oxygen source 51 into the process for producing higher hydrocarbons without directly contacting a hydrocarbon stream (e.g., C1 recycle stream 25) with elemental oxygen. The oxygen source 51 may include, but is not limited to, air, oxygen enriched air, pure oxygen, any other source of elemental oxygen, or any combination thereof. Further, the halogen used in the process may be contained in a metal halide salt and be retained within the process. Other metal oxide/metal halide systems in which the metal may exist in multiple oxidation states, such as chromium (Cr), tin (Sn) or Vandium (V), can in principle also be utilized except that these metal halides are more stable to decomposition, requiring substantially higher temperatures before significant elemental halogen is evolved. However, without wishing to be limited by theory, some metal halides may be directly reactive with alkanes to yield alkyl halides without free elemental halogen being formed. Further, mixtures of one or more metal halides, for example copper bromide and vanadium bromide, or iron bromide and copper bromide may be employed in certain embodiments to implement the invention. Regeneration off-gas 55 comprising mostly inert components contained in the oxygen source (e.g. nitrogen from air) and excess oxygen and smaller amounts other gases from the regeneration may be withdrawn from the offline metal oxide reactor bed 53. The regeneration off-gas 55 may then be treated in scrubbing unit 57, for example, to remove any trace amounts of halogen that might be normally present, or larger amounts of halogen that might be intermittently present due to process upsets or mechanical failures, etc., and then vented to the atmosphere or utilized in another process.
The cyclic nature of this process may require a plurality of reactor vessels to allow at least one metal oxide reactor bed (e.g., metal oxide reactor bed 49) to contact a stream containing a hydrogen halide while another reactor (offline metal oxide reactor bed 53) containing a metal halide may be allowed to contact an oxygen source (e.g., oxygen source 51). Multiple metal oxide/metal halide reactor beds may be used with any number of reactors cycling between oxidation and regeneration to allow for the system to be run continuously. Such reactors may comprise fixed bed reactors, radial bed reactors, fixed fluidized bed reactors (e.g., a fluidized bed that remains substantially confined within a single vessel), or any other suitable reactor type. Alternatively, a circulating fluidized bed reactor or moving bed reactor design may be used. In this embodiment, the metal oxide may contact a stream containing a hydrogen halide in one reactor or reaction zone and then be transported to another reactor or reaction zone to contact an oxygen source. Exemplary moving bed reactors may include vertical moving bed reactors, radial moving bed reactors, circulating fluidized bed reactors (e.g., with multiple reactors or reaction zones), or any other reactor configuration allowing for physical transport of a solid reactant or catalyst between reaction zones.
As shown in
The resulting aqueous condensate stream 33 may be passed through a stripping vessel 59, for example, to remove any elemental halogen, hydrocarbons, and possibly any hydrogen halide absorbed in the aqueous phase. The stripping vessel 59 may use a stripping gas 61, such as air, along with suitable process conditions so that any elemental halogen absorbed in the aqueous phase may be stripped into the vapor phase and carried along with the stripping gas 61, which may leave the stripping vessel 59 as oxygen source 51. In an embodiment, the oxygen source 51 may then pass to the offline metal oxide bed 53 undergoing regeneration. Such an arrangement may allow for any elemental halogen to be recaptured by the metal halide for further use in the system. The water effluent 63 exiting the stripping vessel 59 may contain only trace amounts of a halogen and may be treated in an additional process to remove any traces of halogen and halides in the stream for disposal, or may be used in other processes requiring an aqueous stream.
With continued reference to
As illustrated on
CH4(g)+Br2(g)→CH3Br(g)+HBr(g)
This reaction may occur 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 3:1, a temperature of about 500° C. and a residence time of about 60 seconds, about 90% selectivity to the mono-halogenated methyl bromide may occur. Small amounts of dibromomethane and trace amounts of tribromomethane may also form in the bromination reaction. Small amounts of higher alkanes, such as ethane, propane contained in the C1 recycle stream 25, may also be readily brominated resulting in mono and multiple brominated species such as ethyl bromides and propyl bromides. If a methane to bromine ratio of significantly less than 2.5 to 1 is utilized, substantially lower selectivity to methyl bromide may occur and significant formation of undesirable carbon soot may be observed. Further, the amount of water vapor contained in the feed into the C1+ halogenation unit 45 should be minimized. It has been discovered that minimization of water vapor in the feed to the C1+ halogenation unit 45 may minimize 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.
In the C2+ halogenation unit 23, the C2+ hydrocarbons contained in the C2-C3 recycle stream 21 may react exothermically with halogen vapor from the first portion stream 39 at a relatively low temperature, for example, in the range of about 250° C. to about 400° C., and at a pressure in the range of about 1 bar to about 40 bar to produce gaseous higher alkyl halides and hydrogen halide. In particular embodiments, the C2+ halogenation unit 23 may be operated at a temperature in the range of 300° C. to about 375° C. In some embodiments, the C2+ halogenation unit 23 may have an inlet pre-heater zone for heating its feed to a reaction initiation temperature in the range of about 200° C. to about 350° C. The C2 alkyl halide stream 43 may reach a temperature in the range of about 350° C. to about 375° C. As previously described, halogen may be supplied to C2+ halogenation unit 23 by way of first portion stream 39. The portion of reheated vapor effluent stream from in-line heater 37 going to first portion stream 39 may be adjusted such that the molar ratio of halogen to the sum of C2 plus C3 hydrocarbons in C2-C3 recycle stream 21 may be in the range of about 0.75 to 0.5 in the combined feed to C2+ halogenation unit 23. While
As will be appreciated by those of ordinary skill in the art, with the benefit of this disclosure, the reactions in C1+ halogenation unit 45 and in C2+ halogenation unit 23 may be a homogeneous gas phase reaction or a heterogeneous (catalytic) reaction. Examples of suitable catalysts that may be utilized in C1+ halogenation unit 45 or C2+ halogenation unit 23 may include, but are not limited to, platinum, palladium, or supported non-stoichiometric metal oxy-halides such as FeOxBry or FeOxCly or supported stoichiometric metal oxy-halides such as TaOF3, NbOF3, ZrOF2, SbOF3 as described in Olah, et al, J. Am. Chem. Soc. 1985, 107, 7097-7105, which is incorporated herein in its entirety. In some embodiments, a multi-zone halogenation reactor may be used in the C1+ halogenation unit 45. In this embodiment, the vapor effluent stream 35 comprising halogen and alkanes may be heated to a temperature sufficient to initiate thermal bromination of the alkanes. A downstream zone may then comprise one or more catalytic materials to complete the halogenation reaction and increase the selectivity of the reaction products to mono-halogenated methane (e.g., methyl bromide). The C2-C3 recycle stream 21 may be recycled into a downstream zone of the C1+ halogenation unit 45 in some embodiments, or, as illustrated, fed to a separate, parallel C2+ halogenation unit 23 reactor.
As shown in
In an embodiment, the crystalline alumino-silicate catalyst employed in the synthesis unit 5 may be a zeolite catalyst, such as a ZSM-5 zeolite catalyst, when it is desired to form higher molecular weight hydrocarbons. The zeolite catalyst may be used in the hydrogen, sodium, magnesium form, or any combination thereof. 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 or that have been modified by chemical de-alumination may be used in the synthesis reactor as will be evident to a skilled artisan.
When it is desired to form olefins from the reaction of alkyl halides in the synthesis unit 5, the crystalline alumino-silicate catalyst employed in the synthesis reactor may be a zeolite catalyst. For example, the zeolite catalyst may be an X type, Y type or SAPO zeolite catalyst, although other zeolites with differing pore sizes and acidities, which are synthesized by varying the alumina-to-silica ratio may be used in certain embodiments as will be evident to a skilled artisan. The zeolite catalyst may be used in a protonic form, a sodium form, or a mixed protonic/sodium form. The zeolite catalyst may also be modified by ion exchange with other alkali metal cations, such as Li, 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. These various alternative cations may have an effect of shifting reaction selectivity. Other zeolite catalysts having varying pore sizes and acidities, which are synthesized by varying the alumina-to-silica ratio may be used in the synthesis reactor as will be evident to a skilled artisan.
The temperature at which the synthesis unit 5 is operated may be an important parameter in determining the selectivity of the reaction to higher molecular hydrocarbons or to olefins. Where a catalyst is selected to form higher molecular weight hydrocarbons in the synthesis unit 5, the synthesis unit 5 may be operated at a temperature ranging from about 150° C. to about 450° C. At the low end of the temperature range, with methyl bromide reacting over ZSM-5 zeolite at temperatures of about 150° C., low methyl bromide conversion on the order of 20% is noted, but with a high selectivity towards C5+ products. At increasing temperatures approaching 400° C., methyl bromide conversion increases towards 95% or greater, however selectivity towards C5+ products decreases slightly and selectivity towards lighter products increases. At temperatures approaching about 450° C., almost complete conversion of methyl bromide may occur but may result in further increased yields of light hydrocarbons such as undesirable methane and may significantly increase the rate of coke formation. Notably, in the case of the alkyl bromide reaction over the zeolite ZSM-5 catalyst, cyclization reactions may also occur such that the C7+ fractions are composed primarily of substituted aromatics. Surprisingly, very little ethane or C2-C3 olefin components may be formed. In the optimum operating temperature range of between about 350° C. and about 400° C., as a byproduct of the reaction, a small amount of carbon may build up on the catalyst over time during operation, causing a decline in catalyst activity over a range of hours, up to about 24 hours, depending on space velocity, the reaction conditions and the composition of the feed gas. Without wishing to be limited by theory, it is believed that higher reaction temperatures above about 400° C., associated with the formation of methane may 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 in the range of about 350° C. to about 400° C. in the synthesis reactor may balance increased selectivity of the desired C5+ products and lower rates of deactivation due to carbon formation, against higher conversion per pass, which may minimize the quantity of catalyst, recycle rates and equipment size required.
Where a catalyst is selected to form olefins in the synthesis unit 5, the synthesis unit 5 may be operated at a temperature ranging from about 250° C. to about 500° C. Temperatures above about 450° C. in the synthesis unit 5 may result in increased yields of light hydrocarbons, such as undesirable methane and also deposition of coke, whereas lower temperatures increase yields of ethylene, propylene, butylene and heavier molecular weight hydrocarbon products. Notably, in the case of the alkyl bromide reaction over a 10× zeolite catalyst, it is believed that cyclization reactions may also occur such that the C7+ fractions contain substantial substituted aromatics. At increasing temperatures approaching 400° C., it is believed that methyl bromide conversion increases towards 90% or greater, however selectivity towards C5+ products may decrease and selectivity towards lighter products, particularly olefins may increase. At temperatures exceeding 550° C., it is believed that a high conversion of methyl bromide to methane and carbonaceous coke may occur. In an operating temperature range of between about 300° C. and about 450° C., as a byproduct of the reaction, a lesser amount of coke may build up on the catalyst over time during operation, causing a decline in catalyst activity over a range of hours, depending on the space velocity, reaction conditions and the composition of the feed gas. Without wishing to be limited by theory, it is believed that 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 250° C. to about 500° C., and alternatively in the range of about 300° C. to about 450° C., in the synthesis unit 5 balance increased selectivity of the desired olefins and 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 the synthesis unit 5 from the normal process flow, purging with an inert gas at a pressure in the range of about 1 bar to about 40 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 CO2 by addition of air or inert gas-diluted oxygen to synthesis reactor at a pressure in the range of about 1 bar to about 40 bar and an elevated temperature in the range of about 400° C. to about 650° C. Carbon dioxide and residual air or inert gases are vented from the synthesis reactor during the regeneration period. The resulting gases may be diverted to the oxidation unit 27 in order to recapture any halogens released during the regeneration cycle, or the gases may be directed to a clean-up unit for removal of any trace halogens prior to being vented to the atmosphere.
The synthesis product stream 3 from the synthesis reactor 5, which may comprise methane, hydrogen halide and higher molecular weight hydrocarbons, olefins or mixtures thereof, may be withdrawn from the synthesis unit 5 and cooled to a temperature in the range of about 100° C. to about 500° C. using any suitable heat exchange device or quenching technique. The synthesis product stream 3 may then be combined with the feed gas 1 and the process may be repeated so as to operate in a continuous fashion.
To facilitate a better understanding of the present invention, the following examples of certain aspects of some embodiments are given. In no way should the following examples be read to limit, or define, the scope of the invention.
A laboratory test setup was used to demonstrate the process described herein. This example was used to demonstrate oxidation of hydrogen bromide with a metal oxide. A mixture of simulated process recycle gas is composed of about 54 mol % methane, about 26 mol % hydrogen bromide gas, and about 20% nitrogen to act as an inert tie-element for calculating a mole percent is fed to a tubular reactor containing approximately 20 cm3 of a “solid reactant” composed of iron oxide dispersed on a porous inert support. The solid reactant contains approximately 20 wt % (as Fe) iron oxide dispersed on 3/16th inch spherical porous beads of gamma-alumina with a specific surface area of approximately 30 m2/g, manufactured by St. Gobain. The externally-insulated reactor tube is maintained at an external wall temperature in the range of approximately 150° C. to 200° C. by electrical resistance heating tape under the insulation to prevent excessive heat loss. Evolution of elemental bromine is observed in the reactor effluent, demonstrating that hydrogen bromide can be oxidized by iron oxide to yield elemental bromine.
This example was used to demonstrate regeneration of a metal bromide with an oxygen containing gas mixture. Fifty (50) g of FeBr3 is dissolved in minimal water and imbibed into ˜75 cm3 of St. Gobain spherical alumina support ( 3/16th spheres). This material is loaded into a reactor tube composed of glass-lined steel pipe that is electrically heat-traced and insulated. N2 gas was flowed at a rate at ˜15 ml/min though the reactor while ramping up heating of the reactor tube to 150° C. After a period of time bromine evolution is observed in the effluent gas. N2 gas flow is continued until Br2 evolution fell off significantly, then heating is reduced to cool reactor tube to ˜50° C. under continued N2. Due to the heating under N2 flow FeBr3 is reduced to FeBr2 according to the following reversible reaction:
2FeBr3(s)=2FeBr2(s)+Br2(g)
With temperature at ˜50° C., a mixture of N2 and O2 at total flow rate of about 15 ml/min is initiated. Periodic samples of the inlet and outlet O2 content are taken as a function of time and recorded. Table 1 below shows that FeBr2 at 50° C. can absorb elemental oxygen from a nitrogen/oxygen mixture according to the following reaction:
6FeBr2(s)+3/2O2(g)=Fe2O3+4FeBr3
Table 2 and
Therefore, the present invention is well adapted to attain the ends and advantages mentioned as well as those that are inherent therein. The particular embodiments disclosed above are illustrative only, as the present invention may be modified and practiced in different but equivalent manners apparent to those skilled in the art having the benefit of the teachings herein. Furthermore, no limitations are intended to the details of construction or design herein shown, other than as described in the claims below. It is therefore evident that the particular illustrative embodiments disclosed above may be altered or modified and all such variations are considered within the scope and spirit of the present invention. While processes and systems are described in terms of “comprising,” “containing,” or “including” various components or steps, the processes and systems can also “consist essentially of” or “consist of” the various components and steps. All numbers and ranges disclosed above may vary by some amount. Whenever a numerical range with a lower limit and an upper limit is disclosed, any number and any included range falling within the range is specifically disclosed. In particular, every range of values (of the form, “from about a to about b,” or, equivalently, “from approximately a to approximately b,” or, equivalently, “from approximately a-b”) disclosed herein is to be understood to set forth every number and range encompassed within the broader range of values. Also, the terms in the claims have their plain, ordinary meaning unless otherwise explicitly and clearly defined by the patentee. Moreover, the indefinite articles “a” or “an”, as used in the claims, are defined herein to mean one or more than one of the element that it introduces. If there is any conflict in the usages of a word or term in this specification and one or more patent or other documents that may be incorporated herein by reference, the definitions that are consistent with this specification should be adopted.
The present application claims priority to U.S. Provisional Application No. 61/915,384, filed Dec. 12, 2013, the entire disclosure of which is incorporated herein by reference.
| Number | Date | Country | |
|---|---|---|---|
| 61915384 | Dec 2013 | US |
