U.S. patent application number 12/957036 was filed with the patent office on 2011-03-24 for process for converting gaseous alkanes to liquid hydrocarbons.
This patent application is currently assigned to MARATHON GTF TECHNOLOGY, LTD.. Invention is credited to John J. Waycuilis.
Application Number | 20110071326 12/957036 |
Document ID | / |
Family ID | 39940024 |
Filed Date | 2011-03-24 |
United States Patent
Application |
20110071326 |
Kind Code |
A1 |
Waycuilis; John J. |
March 24, 2011 |
PROCESS FOR CONVERTING GASEOUS ALKANES TO LIQUID HYDROCARBONS
Abstract
Embodiments disclose a process for converting gaseous alkanes to
higher molecular weight hydrocarbons, olefins or mixtures thereofs
wherein a gaseous feed containing alkanes may be reacted with a dry
bromine vapor to form alkyl bromides and hydrobromic acid vapor.
The mixture of alkyl bromides and hydrobromic acid then may be
reacted over a synthetic crystalline alumino-silicate catalyst,
such as a ZSM-5 or an X or Y type zeolite, at a temperature of from
about 250.degree. C. to about 500.degree. C. so as to form
hydrobromic acid vapor and higher molecular weight hydrocarbons,
olefins or mixtures thereof. Various methods are disclosed to
remove the hydrobromic acid vapor from the higher molecular weight
hydrocarbons, olefins or mixtures thereof and to generate bromine
from the hydrobromic acid for use in the process.
Inventors: |
Waycuilis; John J.;
(Cypress, TX) |
Assignee: |
MARATHON GTF TECHNOLOGY,
LTD.
Houston
TX
|
Family ID: |
39940024 |
Appl. No.: |
12/957036 |
Filed: |
November 30, 2010 |
Related U.S. Patent Documents
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Application
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Patent Number |
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12123924 |
May 20, 2008 |
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12957036 |
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12112926 |
Apr 30, 2008 |
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12123924 |
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11254438 |
Oct 19, 2005 |
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12112926 |
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11101886 |
Apr 8, 2005 |
7348464 |
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11254438 |
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10826885 |
Apr 16, 2004 |
7244867 |
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11101886 |
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Current U.S.
Class: |
570/252 ;
585/733 |
Current CPC
Class: |
C01B 7/093 20130101;
C07C 2529/40 20130101; C07C 17/10 20130101; C07C 1/321 20130101;
C10G 2400/02 20130101; C01B 7/096 20130101; C07C 1/30 20130101;
C07C 17/10 20130101; C07C 1/26 20130101; C01B 9/04 20130101; C10G
50/00 20130101; C07C 17/35 20130101; C10G 2400/20 20130101; C10G
29/205 20130101; C07C 1/26 20130101; C07C 17/07 20130101; C10G
29/26 20130101; C10G 50/02 20130101; C07C 19/075 20130101; C07C
9/00 20130101 |
Class at
Publication: |
570/252 ;
585/733 |
International
Class: |
C07C 17/10 20060101
C07C017/10; C07C 1/26 20060101 C07C001/26 |
Claims
1. A process comprising: separating hydrobromic acid from a gaseous
stream comprising hydrobromic acid and hydrocarbons; converting
said hydrobromic acid to at least bromine; and contacting said
bromine with gaseous alkanes to form bromination products
comprising alkyl bromides.
2. The process of claim 1 wherein said step of separating said
hydrobromic acid from said hydrocarbons comprises: contacting said
gaseous stream with water.
3. The process of claim 2 wherein said step of contacting said
gaseous stream with said water comprises: neutralizing said
hydrobromic acid to form an aqueous solution comprising said water
and a metal bromide salt, the metal of said metal bromide salt
being selected from Cu, Zn, Fe, Co, Ni, Mn, Ca or Mg bromide.
4. The process of claim 3 wherein said step of converting
comprises: oxidizing said aqueous solution containing said metal
bromide salt to form at least said bromine and a reaction product
selected from the group consisting of a metal hydroxide, a metal
oxy-bromide species and combinations thereof; and separating said
bromine from said reaction product
5. The process of claim 4 wherein said water that contacts said
gaseous stream comprises said reaction product.
6. The process of claim 2 wherein said hydrobromic acid dissolves
into said water forming a hydrobromic acid solution, said process
further comprising: vaporizing said hydrobromic acid solution; and
reacting said vaporized hydrobromic acid solution with a metal
oxide to form a reaction product comprising a metal bromide salt,
the metal of said metal bromide salt being selected from the group
of Cu, Zn, Fe, Co, Ni, Mn, Ca or Mg.
7. The process of claim 6 wherein said step of converting
comprises: oxidizing said metal bromide salt to form oxidation
products comprising said bromine and said metal oxide; and
separating said bromine from said metal oxide.
8. The process of claim 6 wherein said metal bromide salt is
contained on a porous support.
9. The process of claim 1 wherein said step of separating
hydrobromic acid from hydrocarbons comprises reacting said
hydrobromic acid with a metal oxide to form reaction products
comprising a metal bromide and steam.
10. The process of claim 9 wherein the metal of said metal oxide is
magnesium, calcium, vanadium, chromium, manganese, iron, cobalt,
nickel, copper, zinc or tin.
11. The process of claim 10 wherein said metal oxide is supported
on a solid carrier.
12. The process of claim 11 wherein said metal oxide is contained
in a bed in a vessel.
13. The process of claim 9 wherein said step of converting
comprises: reacting said metal bromide with an oxygen containing
gas to obtain reaction products comprising said metal oxide and
said bromine.
14. The process of claim 1 wherein said bromine from the reaction
of said metal bromide with said oxygen containing gas is recycled
to said step of contacting said gaseous alkanes to form alkyl
bromides.
15. The process of claim 1 further comprising: reacting said alkyl
bromides in the presence of said hydrobromic acid and a synthetic
crystalline alumino-silicate catalyst to form reaction products
comprising said hydrocarbons.
16. A process comprising: contacting a gaseous stream comprising
hydrobromic acid and hydrocarbons with an aqueous solution
comprising a base selected from the group consisting of a metal
hydroxide, a metal oxy-bromide species, and combinations thereof
such that the hydrobromic acid is neutralized to form a metal
bromide salt in the aqueous solution; oxidizing said aqueous
solution containing said metal bromide salt to form oxidation
products comprising bromine and said base; separating said bromine
from said aqueous solution comprising said base; and contacting
said bromine with gaseous alkanes to form alkyl bromides.
17. The process of claim 16 further comprising: reacting said alkyl
bromides in the presence of said hydrobromic acid and a synthetic
crystalline alumino-silicate catalyst to form said
hydrocarbons.
18. A process comprising: contacting a gaseous stream comprising
hydrobromic acid and hydrocarbons with water, wherein said
hydrobromic acid dissolves in said water to form an aqueous
solution comprising said water and said hydrobromic acid;
neutralizing said hydrobromic acid to form a metal bromide salt;
oxidizing said metal bromide salt to form an oxidation product
comprising bromine; and contacting said bromine with gaseous
alkanes to form bromination products comprising alkyl bromides.
19. The process of claim 18 further comprising: reacting said alkyl
bromides in the presence of said hydrobromic acid and a synthetic
crystalline alumino-silicate catalyst to form reaction products
comprising said hydrocarbons.
20. A process comprising: reacting hydrobromic acid with a metal
oxide to form reaction products comprising a metal bromide and
steam, wherein said hydrobromic acid is contained in a gaseous
stream comprising said hydrobromic acid and hydrocarbons; reacting
said metal bromide with a gas comprising oxygen to form reaction
products comprising bromine and said metal oxide; and contacting
said bromine with gaseous alkanes to form bromination products
comprising alkyl bromides.
21. The process of claim 19 further comprising: reacting said alkyl
bromides in the presence of said hydrobromic acid and a synthetic
crystalline alumino-silicate catalyst to form reaction products
comprising said hydrocarbons.
Description
REFERENCE TO RELATED PATENT APPLICATION
[0001] This application is a continuation-in-part of copending U.S.
patent application Ser. No. 12/112, 926 flied Apr. 30, 2008 and
entitled "Process for Converting Gaseous Alkanes to Olefins" which
is a continuation of U.S. patent application Ser. No. 11/254,438
filed on Oct. 19, 2005 and entitled "Process for Converting Gaseous
Alkanes to Olefins," which is a continuation-in-part of U.S. Pat.
No. 7,348,464 issued on Mar. 25, 2008 and entitled "Process for
Converting Gaseous Alkanes to Liquid Hydrocarbons," which is a
continuation-in-part of U.S. Pat. No. 7,244,867 issued on Jul. 17,
2007 and entitled "Process for Converting Gaseous Alkanes to Liquid
Hydrocarbons".
[0002] This application is related to the following copending
patent applications: U.S. patent application Ser. No. 11/778,479
filed on Jul. 16, 2007 and entitled "Process for Converting Gaseous
Alkanes to Liquid Hydrocarbons" ; and U.S. patent application Ser.
No. 11/957,261 filed on Dec. 14, 2007 and entitled "Process for
Converting Gaseous Alkanes to Liquid Hydrocarbons".
BACKGROUND OF THE INVENTION
[0003] 1. Field of the Invention
[0004] The present invention relates to a process for converting
lower molecular weight, gaseous alkanes to olefins, higher
molecular weight hydrocarbons, or mixtures thereof that may be
useful as fuels or monomers and intermediaries in the production of
fuels or chemicals, such as lubricant and fuel additives, and more
particularly, in one or more embodiments, 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 olefins, higher molecular weight hydrocarbons or mixtures
thereof.
[0005] 2. Description of Related Art
[0006] 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
[0007] 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 H.sub.2) at high
temperatures (approximately 1000.degree. 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.degree. 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.degree. 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.
[0008] Numerous alternatives to the conventional production of
synthesis gas as a route to methanol or synthetic liquid
hydrocarbons have been proposed.
[0009] 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. Nos.
5,243,098 or 5,334,777 to Miller, teach reacting a lower alkane,
such as methane, with a metallic halide to form a metal 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 (CH.sub.3Cl), resulting in unwanted
by-products such as CH.sub.2Cl.sub.2 and CHCl.sub.3 which are
difficult to convert or require severe limitation of conversion per
pass and hence very high recycle rates.
[0010] Other prior art processes propose the catalytic chlorination
or bromination of methane as an alternative to generation of
synthesis gas (CO and H.sub.2). 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. Nos. 6,462,243 B1, U.S. 6,472,572 B1, and U.S.
6,525,230 to Grosso are also capable of attaining higher
selectivity to mono-halogenated intermediates by the use of
bromination. The resulting alkyl bromide 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 bromide/oxide solids that is
operated cyclically in four steps. While certain ethers, such as
dimethyl ether ("DME") are 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. No. 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.
[0011] 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 conversion of methane and other alkanes
found in natural gas to olefins, higher molecular weight
hydrocarbons 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.
SUMMARY OF THE INVENTION
[0012] 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 comprising:
[0013] separating hydrobromic acid from a gaseous stream comprising
hydrobromic acid and hydrocarbons; converting said hydrobromic acid
to at least bromine; and contacting said bromine with gaseous
alkanes to form bromination products comprising alkyl bromides.
[0014] In another characterization of the present invention, a
process is provided that comprises: contacting a gaseous stream
comprising hydrobromic acid and hydrocarbons with an aqueous
solution comprising a base selected from the group consisting of a
metal hydroxide, a metal oxybromide species, and combinations
thereof such that the hydrobromic acid is neutralized to form a
metal bromide salt in the aqueous solution; oxidizing said aqueous
solution containing said metal bromide salt to form oxidation
products comprising bromine and said base; separating said bromine
from said aqueous solution comprising said base; and contacting
said bromine with gaseous alkanes to form alkyl bromides.
[0015] In another characterization of the present invention, a
process is provided that comprises: contacting a gaseous stream
comprising hydrobromic acid and hydrocarbons with water, wherein
said hydrobromic acid dissolves in said water to form an aqueous
solution comprising said water and said hydrobromic acid;
neutralizing said hydrobromic acid to form a metal bromide salt;
oxidizing said metal bromide salt to form an oxidation product
comprising bromine; and contacting said bromine with gaseous
alkanes to form bromination products comprising alkyl bromides.
[0016] In another characterization of the present invention, a
process is provided that comprises: reacting hydrobromic acid with
a metal oxide to form reaction products comprising a metal bromide
and steam, wherein said hydrobromic acid is contained in a gaseous
stream comprising said hydrobromic acid and hydrocarbons; reacting
said metal bromide with a gas comprising oxygen to form reaction
products comprising bromine and said metal oxide; and contacting
said bromine with gaseous alkanes to form bromination products
comprising alkyl bromides.
BRIEF DESCRIPTION OF THE DRAWINGS
[0017] The accompanying drawings, which are incorporated in and
form a part of the specification, illustrate the embodiments of the
present invention and, together with the description, serve to
explain the principles of the invention.
[0018] In the drawings:
[0019] FIG. 1 is a simplified block flow diagram of an embodiment
of the process of the present invention;
[0020] FIG. 2 is a schematic view of one embodiment of the process
of the present invention;
[0021] FIG. 3 is a schematic view of another embodiment of process
of the present invention;
[0022] FIG. 4A is schematic view of another embodiment of the
process of the present invention;
[0023] FIG. 4B is a schematic view of the embodiment of the process
of the present invention illustrated in FIG. 4A depicting an
alternative processing scheme which may be employed when oxygen is
used in lieu of air in the oxidation stage;
[0024] FIG. 5A is a schematic view of the embodiment of the process
of the present invention illustrated in FIG. 4A with the flow
through the metal oxide beds being reversed; FIG. 5B is a schematic
view of the embodiment of the process of the present invention
illustrated in FIG. 5A depicting an alternative processing scheme
which may be employed when oxygen is used in lieu of air in the
oxidation stage;
[0025] FIG. 6A is a schematic view of another embodiment of the
process of the present invention;
[0026] FIG. 6B is a schematic view of the embodiment of the process
of the present invention illustrated in FIG. 6A depicting an
alternative processing scheme which may be employed when oxygen is
used in lieu of air in the oxidation stage;
[0027] FIG. 7 is a schematic view of another embodiment of the
process of the present invention;
[0028] FIG. 8 is a schematic view of the embodiment of the process
of the present invention illustrated in FIG. 7 with the flow
through the metal oxide beds being reversed; and
[0029] FIG. 9 is a schematic view of another embodiment of the
process of the present invention.
[0030] FIG. 10 is a graph of methyl bromide conversion and product
selectivity for the oligimerization reaction of the process of the
present invention as a function of temperature;
[0031] FIG. 11 is a graph comparing conversion and selectivity for
the example of methyl bromide, dry hydrobromic acid and methane
versus only methyl bromide plus methane;
[0032] FIG. 12 is a graph of product selectivity from reaction of
methyl bromide and dibromomethane vs. product selectivity from
reaction of methyl bromide only;
[0033] FIG. 13 is a graph of a Paraffinic Olefinic Napthenic and
Aromatic (PONA) analysis of a typical condensed product sample of
the process of the present invention; and
[0034] FIG. 14 is a graph of a PONA analysis of another typical
condensed product sample of the present invention
DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS
[0035] 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 lines 11 and 111 in the
embodiments of the process of the present invention as illustrated
in FIGS. 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.
[0036] A block flow diagram generally depicting an embodiment of a
process of the present invention is illustrated in FIG. 1, while
specific embodiments of the process illustrated in FIG. 1 are
illustrated in FIGS. 2 and 3. Referring to
[0037] FIG. 1, a gas stream comprising recycle gas and a natural
feed gas is combined with dry bromine vapor and fed to an alkane
bromination reactor. The recycle gas and the natural gas feed may
comprise lower molecular weight hydrocarbons. In the alkane
bromination reactor, the gas stream and the dry bromine vapor are
reacted to produce gaseous alkyl bromides and hydrobromic acid
vapors. As illustrated, gaseous alkyl bromides and hydrobromic acid
vapors are fed to the alkyl bromide conversion reactor. In the
alkyl bromide conversion reactor, the gaseous alkyl bromides are
reacted to form higher molecular weight hydrocarbons and additional
hydrobromic acid vapors. In the illustrated embodiment, the
hydrobromic acid vapors are then removed from the higher molecular
hydrocarbons in the hydrobromic acid removal unit by a recirculated
aqueous solution. As illustrated in FIG. 1, the recirculated
aqueous solution carries the hydrobromic acid (or metal bromide
salt if the acid is neutralized by the aqueous solution) to the
bromide oxidation unit. As will be discussed in more detail below,
the hydrobromic acid may be neutralized in the bromide oxidation
unit to form a metal bromide salt. Oxygen or air is supplied to the
bromide oxidation unit to oxidize the metal bromide salt to form
the bromine, which is then recycled to the alkane bromination
reactor.
[0038] In the illustrated embodiment, a natural gas feed is also
introduced into the hydrobromic acid removal unit. From the
hydrobromic acid removal unit, the natural gas feed and the higher
molecular hydrocarbons are fed to the dehydration and product
recovery unit. In the dehydration and product recovery unit, water
is removed from the higher molecular weight hydrocarbons and a
hydrocarbon liquid product is produced. In addition, a gas stream
of recycle gas and the natural gas feed are conveyed to the alkane
bromination reactor. Accordingly, the process illustrated in FIG. 1
may be used to produce a liquid hydrocarbon product from lower
molecular hydrocarbons.
[0039] Referring to FIG. 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 conduit
62, mixed with dry bromine liquid transported via line 25 and pump
24, and passed to heat exchanger 26 wherein the liquid bromine is
vaporized. The mixture of lower molecular weight alkanes and dry
bromine vapor is fed to reactor 30. Preferably, the molar ratio of
lower molecular weight alkanes to dry bromine vapor in the mixture
introduced into reactor 30 is in excess of 2.5:1. Reactor 30 has an
inlet pre-heater zone 28 which heats the mixture to a reaction
initiation temperature in the range of about 250.degree. C. to
about 400.degree. C.
[0040] In first reactor 30, the lower molecular weight alkanes are
reacted exothermically with dry bromine vapor at a relatively low
temperature in the range of about 250.degree. C. to about
600.degree. 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:
CH.sub.4 (g)+Br.sub.2 (g).fwdarw.CH.sub.3Br (g)+HBr (g)
[0041] This reaction occurs with a significantly high degree of
selectivity to methyl bromide. For example, in the case of
bromination of methane, a methane to bromine ratio of about 4.5:1
increases the selectivity to the mono-halogenated methyl bromide to
that obtained using smaller methane to bromine ratios. 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 such as ethyl bromides, propyl bromides
and butyl bromides. If an alkane to bromine ratio of significantly
less than about 2.5 to 1 is utilized, a lower selectivity to methyl
bromide occurs and significant formation of undesirable carbon soot
is observed. Further, the dry bromine vapor that is feed into first
reactor 30 is substantially water-free. Applicant has discovered
that elimination of substantially all water vapor from the
bromination step in first reactor 30 substantially 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.
[0042] The effluent that contains alkyl bromides and hydrobromic
acid is withdrawn from the first reactor via line 31 and is
partially cooled in heat exchanger 32 before flowing to a second
reactor 34. The temperature to which the effluent is partially
cooled in heat exchanger 34 is in the range of about 150.degree. C.
to about 350.degree. C. when it is desired to convert the alkyl
bromides to higher molecular weight hydrocarbons in second reactor
34, or to range of about 150.degree. C. to about 450.degree. C.
when it is desired to convert the alkyl bromides to olefins a
second reactor 34. In second reactor 34, the alkyl bromides are
reacted exothermically over a fixed bed 33 of crystalline
alumino-silicate catalyst, preferably a zeolite catalyst. The
temperature and pressure employed in second reactor, as well as the
zeolite catalyst, will determine the product(s) that is formed from
the reaction of alkyl bromides occurring in second reactor 34.
[0043] The crystalline alumino-silicate catalyst employed in second
reactor 34 is preferably a zeolite catalyst and most preferably a
ZSM-5 zeolite catalyst when it is desired to form higher molecular
weight hydrocarbons, 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 reactor 34 as
will be evident to a skilled artisan.
[0044] When it is desired to form olefins from the reaction of
alkyl bromides in reactor 34, the crystalline alumino-silicate
catalyst employed in second reactor 34 is preferably a zeolite
catalyst, and most preferably an X type or Y type zeolite catalyst.
A preferred zeolite is 10 X or Y type zeolite, although other
zeolites with differing pore sizes and acidities, which are
synthesized by varying the alumina-to-silica ratio may be used in
the process of the present invention as will be evident to a
skilled artisan. Although the zeolite catalyst is preferably used
in a protonic form, a sodium form or a mixed protonic/sodium form,
the zeolite 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 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 second reactor 34 as will be evident to a
skilled artisan.
[0045] The temperature at which the second reactor 34 is operated
is an important parameter in determining the selectivity of the
reaction to higher molecular hydrocarbons or to olefins.
[0046] Where a catalyst is selected to form higher molecular weight
hydrocarbons in reactor 34, it is preferred to operate second
reactor 34 at a temperature within the range of about 150.degree.
to 450.degree.. Temperatures above about 300.degree. 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.degree. C. significant
methyl bromide conversion on, the order of 20% is noted, with a
high selectivity towards C.sub.5+ products. 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.degree. C., methyl bromide
conversion increases towards 90% or greater, however selectivity
towards C.sub.5+ products decreases and selectivity towards lighter
products, particularly undesirable methane, increases.
Surprisingly, very little ethane or C.sub.2,-C.sub.3 olefin
components are formed. At temperatures approaching 450.degree. C.,
almost complete conversion of methyl bromide to methane occurs. In
the optimum operating temperature range of between about
300.degree. C. and 400.degree. 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.degree. 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.degree.
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.degree. C. to about
450.degree. C., but preferably in the range of about 300.degree. C.
to about 400.degree. C. in the second reactor 34 balance increased
selectivity of the desired C.sub.5+ 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.
[0047] Where a catalyst is selected to form olefins in reactor 34,
it is preferred to operate second reactor 34 at a temperature
within the range of about 250.degree. C. to 500.degree. C.
Temperatures above about 450.degree. C. in the second reactor can
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 the preferred 10 X zeolite
catalyst, it is believed that cyclization reactions also occur such
that the C7+ fractions contain substantial substituted aromatics.
At increasing temperatures approaching 400.degree. C., it is
believed that methyl bromide conversion increases towards 90% or
greater, however selectivity towards C.sub.5+ products decreases
and selectivity towards lighter products, particularly olefins
increases. At temperatures exceeding 550.degree. C., it is believed
that a high conversion of methyl bromide to methane and
carbonaceous, coke occurs. In the preferred operating temperature
range of between about 300.degree. C. and 450.degree. C., as a
byproduct of the reaction, a lesser amount of coke probably 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.degree. 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.degree. 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.degree. C. to about 500.degree. C., but preferably in
the range of about 300.degree. C. to about 450.degree. C. in the
second reactor 34 balance increased selectivity of the desired
olefins and C.sub.5+ 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.
[0048] The catalyst may be periodically regenerated in situ, by
isolating reactor 34 from the normal process flow, purging with an
inert gas via line 70 at a pressure in a range from about 1 to
about 5 bar at an elevated temperature in the range of about
400.degree. C. to about 650.degree. C. to remove unreacted material
adsorbed on the catalyst insofar as is practical, and then
subsequently oxidizing the deposited carbon to CO.sub.2 by addition
of air or inert gas-diluted oxygen to reactor 34 via line 70 at a
pressure in the range of about 1 bar to about 5 bar at an elevated
temperature in the range of about 400.degree. C. to about
650.degree. C. Carbon dioxide and residual air or inert gas is
vented from reactor 34 via line 75 during the regeneration
period.
[0049] The effluent which comprises hydrobromic acid and higher
molecular weight hydrocarbons, olefins or mixtures thereof is
withdrawn from the second reactor 34 via line 35 and is cooled to a
temperature in the range of 0.degree. C. to about 100.degree. C. in
exchanger 36 and combined with vapor effluent in line 12 from
hydrocarbon stripper 47, which contains feed gas and residual
higher molecular weight hydrocarbons stripped-out by contact with
the feed gas in hydrocarbon stripper 47. The combined vapor mixture
is passed to a scrubber 38 and contacted with a concentrated
aqueous partially-oxidized metal bromide salt solution containing
metal hydroxide, metal oxide, metal oxy-bromide or mixtures of
these species, which is transported to scrubber 38 via line 41. 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.degree. C.
to about 180.degree. C., allowing the use of glass-lined or
fluorpolymer-lined equipment; although Co(II), Ni(II), Mn(ll),
V(II), Cr(ll) 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
hydrocarbons condensed in scrubber 38 may be skimmed and withdrawn
in line 37 and added to liquid hydrocarbons exiting the product
recovery unit 52 in line 54. Hydrobromic acid is dissolved in the
aqueous solution and neutralized by the metal hydroxide, metal
oxide, metal oxy-bromide or mixtures of these species to yield
metal bromide salt in solution and water which is removed from the
scrubber 38 via line 44.
[0050] The residual vapor phase containing olefins, higher
molecular weight hydrocarbons or mixtures thereof that is removed
as effluent from the scrubber 38 is forwarded via line 39 to
dehydrator 50 to remove substantially all water via line 53 from
the vapor stream. The water is then removed from the dehydrator 50
via line 53. The dried vapor stream containing olefins, higher
molecular weight hydrocarbons or mixtures thereof is further passed
via line 51 to product recovery unit 52 to recover olefins, the
C.sub.5+ gasoline-range hydrocarbon fraction or mixtures thereof as
a liquid product in line 54. Any conventional method of dehydration
and liquids recovery, such as solid-bed desiccant adsorption
followed by refrigerated condensation, cryogenic expansion, or
circulating absorption oil or other solvent, as used to process
natural gas or refinery gas streams, and/or to recover olefinic
hydrocarbons, 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 unit 52 is then split into a
purge stream 57 which may be utilized as fuel for the process and a
recycled residual vapor which is compressed via compressor 58. The
recycled residual vapor discharged from compressor 58 is split into
two fractions. A first fraction that is equal to at least 2.5 times
the feed gas molar volume is transported via line 62 and is
combined with dry liquid bromine conveyed by pump 24, heated in
exchanger 26 to vaporize the bromine and fed into first reactor 30.
The second fraction is drawn off of line 62 via line 63 and is
regulated by control valve 60, at a rate sufficient to dilute the
alkyl bromide concentration to reactor 34 and absorb the heat of
reaction such that reactor 34 is maintained at the selected
operating temperature, preferably in the range of about 300.degree.
C. to about 450.degree. C. in order to maximize 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 reactor 30 to be controlled in addition to moderating the
temperature in second reactor 34.
[0051] Water containing metal bromide salt in solution which is
removed from scrubber 38 via line 44 is passed to hydrocarbon
stripper 47 wherein residual dissolved hydrocarbons are stripped
from the aqueous phase by contact with incoming feed gas
transported via line 11. The stripped aqueous solution is
transported from hydrocarbon stripper 47 via line 65 and is cooled
to a temperature in the range of about 0.degree. C. to about
70.degree. C. in heat exchanger 46 and then passed to absorber 48
in which residual bromine is recovered from vent stream in line 67.
The aqueous solution ,effluent from adsorber 48 is transported via
line 49 to a heat exchanger 40 to be preheated to a temperature in
the range of about 100.degree. C. to about 600.degree. C., and most
preferably in the range of about 120.degree. C. to about
180.degree. C. and passed to third reactor 16. Oxygen or air is
delivered via line 10 by blower or compressor 13 at a pressure in
the range of about ambient to about 5 bar to bromine stripper 14 to
strip residual bromine from water. Water is removed from stripper
14 in line 64 and combined with water stream 53 from dehydrator 50
to form water effluent stream in line 56 which is removed from the
process. The oxygen or air leaving bromine stripper 14 is fed via
line 15 to reactor 16 which operates at a pressure in the range of
about ambient to about 5 bar and at a temperature in the range of
about 100.degree. C. to about 600.degree. C., but most preferably
in the range of about 120.degree. C. to about 180.degree. C. so as
to oxidize an aqueous metal bromide salt solution to yield
elemental bromine and metal hydroxide, metal oxide, metal
oxy-bromide or mixtures of these 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.degree. C. to about
180.degree. C., which should allow the use of glass-lined or
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,
metal oxide, metal oxy-bromide or mixtures of these species so
formed to once again yield the metal bromide salt and water. Heat
exchanger 18 in reactor 16 supplies heat to vaporize water and
bromine. Thus, it is believed that the overall reactions result in
the net oxidation of hydrobromic acid produced in first reactor 30
and second reactor 34 to 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)+3/2O.sub.2(g)=3Br.sub.2(g)+Fe(OH).sub.3
1)
3HBr(g)+H.sub.2O=3H(+a)+3Br(-a)+H.sub.2O 2)
3H(+a)+3Br(-a)+Fe(OH).sub.3=Fe(+3a)+3Br(-a)+3H.sub.2O 3)
In the case of the metal bromide being CU(II)Br2, the reactions are
believed to be:
4Cu(+2a)+8Br(-a)+3H.sub.2O+3/2O.sub.2(g)=3Br.sub.2(g)+CuBr.sub.2.3Cu(OH)-
.sub.2 1)
6HBr(g)+H.sub.2O=6H(+a)+6Br(-a)+H.sub.2O 2)
6H(+a)+6Br(-a)+CuBr.sub.2.3Cu(OH).sub.2=4Cu(+2a)+8Br(-a)+6H.sub.2O
3)
[0052] 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 reactor 16 via line 19, are cooled in condenser
20 at a temperature in the range of about 0.degree. C. to about
70.degree. C. and a pressure in the range of about ambient to 5 bar
to condense the bromine and water and passed to three-phase
separator 22. In three-phase separator 22, 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.
[0053] Liquid bromine is pumped in line 25 from three-phase
separator 22 via pump 24 to a pressure sufficient to mix with vapor
stream 62. 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 separator 22 and is
passed via line 23 to bromine scrubber 48, wherein residual bromine
is recovered by solution into and by reaction with reduced metal
bromides in the aqueous metal bromide solution stream 65. Water is
removed from separator 22 via line 27 and introduced into stripper
14.
[0054] In another embodiment of the invention, referring to FIG. 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 conduit 162, mixed with dry bromine
liquid transported via pump 124 and passed to heat exchanger 126
wherein the liquid bromine is vaporized. The mixture of lower
molecular weight alkanes and dry bromine vapor is fed to reactor
130. Preferably, the molar ratio of lower molecular weight alkanes
to dry bromine vapor in the mixture introduced into reactor 130 is
in excess of 2.5:1. Reactor 130 has an inlet pre-heater zone 128
which heats the mixture to a reaction initiation temperature in the
range of about 250.degree. C. to about 400.degree. C.. In first
reactor 130, the lower molecular weight alkanes are reacted
exothermically with dry bromine vapor at a relatively low
temperature in the range of about 250.degree. C. to about
600.degree. 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:
CH.sub.4 (g)+Br.sub.2 (g).fwdarw.CH.sub.3Br (g)+HBr (g)
[0055] This reaction occurs with a significantly high degree of
selectivity to methyl bromide. For example, in the case of
bromination of methane, a methane to bromine ratio of about 4.5:1
increases the selectivity to the mono-halogenated methyl bromide.
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 such as ethyl bromides, propyl
bromides and butyl bromides. If an alkane to bromine ratio of
significantly less than about 2.5 to 1 is utilized, a lower
selectivity to methyl bromide occurs and significant formation of
undesirable carbon soot is observed. Further, the dry bromine vapor
that is feed into first reactor 30 is preferably substantially
water-free. Applicant has discovered that elimination of
substantially all water vapor from the bromination step in first
reactor 30 substantially 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.
[0056] The effluent that contains alkyl bromides and hydrobromic
acid is withdrawn from the first reactor via line 131 and is
partially cooled in heat exchanger 132 before flowing to a second
reactor 134. The temperature to which the effluent is partially
cooled in heat exchanger 134 is in the range of about 150.degree.
C. to about 350.degree. C. where it is desired to convert the alkyl
bromides to higher molecular weight hydrocarbons in second reactor
134, or to range of about 150.degree. C. to about 450.degree. C.
where it is desired to convert the alkyl bromides to olefins in
second reactor 134. In second reactor 134, the alkyl bromides are
reacted exothermically over a fixed bed 133 of crystalline
alumino-silicate catalyst, preferably a zeolite catalyst. The
temperature and pressure employed in second reactor 134, as well as
the zeolite catalyst, will determine the product that is formed
from the reaction of alkyl bromides occurring in second reactor
134.
[0057] The crystalline alumino-silicate catalyst employed in second
reactor 134 is preferably a zeolite catalyst and most preferably a
ZSM-5 zeolite catalyst when it is desired to form higher molecular
weight hydrocarbons, 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 reactor 134
as will be evident to a skilled artisan.
[0058] When it is desired to form olefins from the reaction of
alkyl bromides in reactor 134, the crystalline alumina-silicate
catalyst employed in second reactor 134 is preferably a zeolite
catalyst and most preferably an X type or Y type zeolite catalyst.
A preferred zeolite is 10 X or Y type zeolite, although other
zeolites with differing pore sizes and acidities, which are
synthesized by varying the alumina-to-silica ratio may be used in
the process of the present invention as will be evident to a
skilled artisan. Although the zeolite catalyst is preferably used
in a protonic form, a sodium form or a mixed protonic/sodium form,
the zeolite 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 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 second reactor 134 as will be evident to a
skilled artisan.
[0059] The temperature at which the second reactor 134 is operated
is an important parameter in determining the selectivity of the
reaction to higher molecular weight hydrocarbons or to olefins.
[0060] When a catalyst is selected to form higher molecular weight
hydrocarbons in reactor 134, it is preferred to operate second
reactor 134 at a temperature within the range of about 150.degree.
to 450.degree.. Temperatures above about 300.degree. 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.degree. C. significant
methyl bromide conversion on the order of 20% is noted, with a high
selectivity towards C.sub.530 products. 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.degree. C., methyl bromide conversion
increases towards 90% or greater, however selectivity towards
C.sub.5+ products decreases and selectivity towards lighter
products, particularly undesirable methane, increases.
Surprisingly, very little ethane or C.sub.2,-C.sub.3 olefin
components are formed. At temperatures approaching 450.degree. C.,
almost complete conversion of methyl bromide to methane occurs. In
the optimum operating temperature range of between about
300.degree. C. and 400.degree. 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.degree. 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.degree.
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.degree. C. to about
450.degree. C., but preferably in the range of about 300.degree. C.
to about 400.degree. C. in the second reactor 134 balance increased
selectivity of the desired C.sub.5+ 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.
[0061] When a catalyst is selected to form olefins in reactor 134,
it is preferred to operate second reactor 134 at a temperature
within the range of about 250.degree. to 500.degree. C.
Temperatures above about 450.degree. C. in the second reactor can
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 the preferred 10 X zeolite
catalyst, it is believed that cyclization reactions also occur such
that the C7+ fractions contain substantial substituted aromatics.
At increasing temperatures approaching 400.degree. C., it is
believed that methyl bromide conversion increases towards 90% or
greater, however selectivity towards C.sub.5+ products decreases
and selectivity towards lighter products, particularly olefins
increases. At temperatures exceeding 550.degree. C., it is believed
that a high conversion of methyl bromide to methane and
carbonaceous, coke occurs. In the preferred operating temperature
range of between about 300.degree. C. and 450.degree. C., as a
byproduct of the reaction, a lesser amount of coke probably 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.degree. 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.degree. 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.degree. C. to about 500.degree. C., but preferably in
the range of about 300.degree. C. to about 450.degree. C. in the
second reactor 134 balance increased selectivity of the desired
olefins and C.sub.5+ 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.
[0062] The catalyst may be periodically regenerated in situ, by
isolating reactor 134 from the normal process flow, purging with an
inert gas via line 170 at a pressure in the range of about 1 bar to
about 5 bar and an elevated temperature in the range of 400.degree.
C. to 650.degree. C. to remove unreacted material adsorbed on the
catalyst insofar as is practical, and then subsequently oxidizing
the deposited carbon to CO.sub.2 by addition of air or inert
gas-diluted oxygen via line 170 to reactor 134 at a pressure in the
range of about 1 bar to about 5 bar and an elevated temperature in
the range of 400.degree. C. to 650.degree. C. Carbon dioxide and
residual air or inert gas are vented from reactor 134 via line 175
during the regeneration period.
[0063] The effluent which comprises hydrobromic acid and higher
molecular weight hydrocarbons, olefins or mixtures thereof is
withdrawn from the second reactor 134 via line 135, cooled to a
temperature in the range of about 0.degree. C. to about 100.degree.
C. in exchanger 136, and combined with vapor effluent in line 112
from hydrocarbon stripper 147. The mixture is then passed to a
scrubber 138 and contacted with a stripped, recirculated water that
is transported to scrubber 138 in line 164 by any suitable means,
such as pump 143, and is cooled to a temperature in the range of
about 0.degree. C. to about 50.degree. C. in heat exchanger 155.
Any liquid hydrocarbon product condensed in scrubber 138 may be
skimmed and withdrawn as stream 137 and added to liquid hydrocarbon
product 154. Hydrobromic acid is dissolved in scrubber 138 in the
aqueous solution which is removed from the scrubber 138 via line
144, and passed to hydrocarbon stripper 147 wherein residual
hydrocarbons dissolved in the aqueous solution are stripped-out by
contact with feed gas 111. The stripped aqueous phase effluent from
hydrocarbon stripper 147 is cooled to a temperature in the range of
about 0.degree. C. to about 50.degree. C. in heat exchanger 146 and
then passed via line 165 to absorber 148 in which residual bromine
is recovered from vent stream 167.
[0064] The residual vapor phase containing olefins, higher
molecular weight hydrocarbons or mixtures thereof is removed as
effluent from the scrubber 138 and forwarded to dehydrator 150 to
remove substantially all water from the gas stream. The water is
then removed from the dehydrator 150 via line 153. The dried gas
stream containing olefins, higher molecular weight hydrocarbons or
mixtures thereof is further passed via line 151 to product recovery
unit 152 to recover olefins, the C.sub.5+ gasoline range
hydrocarbon fraction or mixtures thereof as a liquid product in
line 154. 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, or other solvents as used to process
natural gas or refinery gas streams and recover olefinic
hydrocarbons, as known to a skilled artisan, may be employed in the
implementation of this invention. The residual vapor effluent from
product recovery unit 152 is then split into a purge stream 157
that may be utilized, as fuel for the process and a recycled
residual vapor which is compressed via compressor 158. The recycled
residual vapor discharged from compressor 158 is split into two
fractions. A first fraction that is equal to at least 2.5 times the
feed gas volume is transported via line 162, combined with the
liquid bromine conveyed in line 125 and passed to heat exchanger
126 wherein the liquid bromine is vaporized and fed into first
reactor 130. The second fraction which is drawn off line 162 via
line 163 and is regulated by control valve 160, at a rate
sufficient to dilute the alkyl bromide concentration to reactor 134
and absorb the heat of reaction such that reactor 134 is maintained
at the selected operating temperature, preferably in the range of
about 300.degree. C. to about 450.degree. C. in order to maximize
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 reactor 130 to be controlled in addition
to moderating the temperature in second reactor 134.
[0065] Oxygen, oxygen enriched air or air 110 is delivered via
blower or compressor 113 at a pressure in the range of about
ambient to about 5 bar to bromine stripper 114 to strip residual
bromine from water which leaves stripper 114 via line 164 and is
divided into two portions. The first portion of the stripped water
is recycled via line 164, cooled in heat exchanger 155 to a
temperature in the range of about 20.degree. C. to about 50.degree.
C., and maintained at a pressure sufficient to enter scrubber 138
by any suitable means, such as pump 143. The portion of water that
is recycled is selected such that the hydrobromic acid solution
effluent removed from scrubber 138 via line 144 has 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
exchanger 141 and preheater 119 and to minimize the vapor pressure
of HBr over the resulting acid. A second portion of water from
stripper 114 is removed from line 164 and the process via line
156.
[0066] The dissolved hydrobromic acid that is contained in the
aqueous solution effluent from adsorber 148 is transported via line
149 and is combined with the oxygen, oxygen enriched air or air
leaving bromine stripper 114 in line 115. The combined aqueous
solution effluent and oxygen, oxygen enriched air or air is passed
to a first side of heat exchanger 141 and through preheater 119
wherein the mixture is preheated to a temperature in the range of
about 100.degree. C. to about 600.degree. C. and most preferably in
the range of about 120.degree. C. to about 250.degree. C. and
passed to third reactor 117 that contains a metal bromide salt or
metal oxide. The preferred metal of the bromide salt or metal oxide
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 reactor 117
can 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. More preferably, the oxide form of the metal in a range
of 10 to 20% by weight is deposited on an inert support such as
alumina with a specific surface area in the range of 50 to 200
m2/g. The oxidation reactor 117 operates at a pressure in the range
of about ambient to about 5 bar and at a temperature in the range
of about 100.degree. C. to 600.degree. C., but most preferably in
the range of about 130.degree. C. to 350.degree. C.; therein, the
metal bromide is oxidized by oxygen, yielding elemental bromine and
metal hydroxide, metal oxide or metal oxybromide species or, metal
oxides in the case of the supported metal bromide salt or metal
oxide 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, it is believed that the
overall reaction results in the net oxidation of hydrobromic acid
produced in first reactor 130 and second reactor 134 to 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)+3/2O.sub.2(g)=3Br.sub.2(g)+Fe(OH)3 1)
3HBr(g)+H.sub.2O=3H(+a)+3Br(-a)+H.sub.2O 2)
3H(+a)+3Br(-a)+Fe(OH)3=Fe(+3a)+3Br(-a)+3H.sub.2O 3)
In the case of the metal bromide being CU(II)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:
4Cu(+2a)+8Br(-a)+3H.sub.2O+3/2O.sub.2(g)=3Br.sub.2(g)+CuBr.sub.2.3Cu(OH)-
.sub.2 1)
6HBr(g)+H.sub.2O=6H(+a)+6Br(-a)+H.sub.2O 2)
6H(+a)+6Br(-a)+CuBr.sub.2.3Cu(OH).sub.2=4Cu(+2a)+8Br(-a)+6H.sub.2O
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+O.sub.2(g)=Br2(g)+2Cu(II)O 2)
2HBr(g)+Cu(II)O=Cu(II)Br.sub.2+H.sub.2O(g) 3)
[0067] 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 reactor 117, are cooled
in the second side of exchanger 141 and condenser 120 to a
temperature in the range of about 0.degree. C. to about 70.degree.
C. wherein the bromine and water are condensed and passed to
three-phase separator 122. In three-phase separator 122, 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
separator 122 via pump 124 to a pressure sufficient to mix with
vapor stream 162. 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
separator 122 and is passed via line 123 to bromine scrubber 148,
wherein residual bromine is recovered by dissolution into
hydrobromic acid solution stream conveyed to scrubber 148 via line
165. Water is removed from the three-phase separator 122 via line
129 and passed to stripper 114.
[0068] 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 CO.sub.2 and the subsequent efficient and selective reactions of
alkyl bromides to primarily C.sub.2 to C.sub.4 olefins, heavier
products, the C.sub.5+ fraction of which contains substantial
branched alkanes and substituted aromatics, or mixtures thereof.
Byproduct hydrobromic acid vapor from the bromination reaction and
subsequent reaction in reactor 134 are readily dissolved into an
aqueous phase and neutralized by the metal hydroxide or metal oxide
species resulting from oxidation of the metal bromide.
[0069] In accordance with another embodiment of the process of the
present invention illustrated in FIG. 4A, the alkyl bromination and
alkyl bromide conversion stages are operated in a substantially
similar manner to those corresponding stages described with respect
to FIGS. 2 and 3 above. 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
conduits 262 and 211, respectively, and mixed with dry bromine
liquid in line 225. The resultant mixture is transported via pump
224 and passed to heat exchanger 226 wherein the liquid bromine is
vaporized. The mixture of lower molecular weight alkanes and dry
bromine vapor is fed to reactor 230. Preferably, the molar ratio of
lower molecular weight alkanes to dry bromine vapor in the mixture
introduced into reactor 230 is in excess of 2.5:1. Reactor 230 has
an inlet pre-heater zone 228 which heats the mixture to a reaction
initiation temperature in the range of 250.degree. C. to
400.degree. C. In first reactor 230, the lower molecular weight
alkanes are reacted exothermically with dry bromine vapor at a
relatively low temperature in the range of about 250.degree. C. to
about 600.degree. 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:
CH.sub.4 (g)+Br.sub.2 (g).fwdarw.CH.sub.3Br (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, a high selectivity to the mono-halogenated methyl
bromide occurs. 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 such as ethyl
bromides, propyl bromides and butyl bromides. If an alkane to
bromine ratio of significantly less than 2.5 to 1 is utilized,
substantially lower selectivity to methyl bromide occurs and
significant formation of undesirable carbon soot is observed.
Further, the dry bromine vapor that is feed into first reactor 230
is substantially water-free. Applicant has discovered that
elimination of substantially all water vapor from the bromination
step in first reactor 230 substantially 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.
[0070] The effluent that contains alkyl bromides and hydrobromic
acid is withdrawn from the first reactor via line 231 and is
partially cooled in heat exchanger 232 before flowing to a second
reactor 234. The temperature to which the effluent is partially
cooled in heat exchanger 234 is in the range of about 150.degree.
C. to about 350.degree. C. when it is desired to convert the alkyl
bromides to higher molecular weight hydrocarbons in second reactor
234, or to range of about 150.degree. C. to about 450.degree. C.
when it is desired to convert the alkyl bromides to olefins a
second reactor 234. In second reactor 234, the alkyl bromides are
reacted exothermically over a fixed bed 233 of crystalline
alumino-silicate catalyst, preferably a zeolite catalyst. The
temperature and pressure employed in second reactor, as well as the
zeolite catalyst, will determine the product that is formed from
the reaction of alkyl bromides occurring in second reactor 234.
[0071] The crystalline alumino-silicate catalyst employed in second
reactor 234 is preferably a zeolite catalyst and most preferably a
ZSM-5 zeolite catalyst when it is desired to form higher molecular
weight hydrocarbons, 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 reactor 234
as will be evident to a skilled artisan.
[0072] When it is desired to form olefins from the reaction of
alkyl bromides in reactor 234, the crystalline alumino-silicate
catalyst employed in second reactor 234 is preferably a zeolite
catalyst, and most preferably an X type or Y type zeolite catalyst.
A preferred zeolite is 10 X or Y type zeolite, although other
zeolites with differing pore sizes and acidities, which are,
synthesized by varying the alumina-to-silica ratio may be used in
the process of the present invention as will be evident to a
skilled artisan. Although the zeolite catalyst is preferably used
in a protonic form, a sodium form or a mixed protonic/sodium form,
the zeolite 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 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 second reactor 234 as will be evident to a
skilled artisan.
[0073] The temperature at which the second reactor 234 is operated
is an important parameter in determining the selectivity of the
reaction to higher molecular hydrocarbons, or to olefins.
[0074] Where a catalyst is selected to form higher molecular weight
hydrocarbons in reactor 234, it is preferred to operate second
reactor 234 at a temperature within the range of about 150.degree.
to 450.degree.. Temperatures above about 300.degree. 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.degree. C. significant
methyl bromide conversion on the order of 20% is noted, with a high
selectivity towards C.sub.5+ products. 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.degree. C., methyl bromide conversion
increases towards 90% or greater, however selectivity towards
C.sub.5+ products decreases and selectivity towards lighter
products, particularly undesirable methane, increases.
Surprisingly, very little ethane or C.sub.2,-C.sub.3 olefin
components are formed. At temperatures approaching 450.degree. C.,
almost complete conversion of methyl bromide to methane occurs. In
the optimum operating temperature range of between about
300.degree. C. and 400.degree. 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.degree. 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.degree.
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.degree. C. to about
450.degree. C., but preferably in the range of about 300.degree. C.
to about 400.degree. C. in the second reactor 234 balance increased
selectivity of the desired C.sub.5+ 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.
[0075] Where a catalyst is selected to form olefins in reactor 234,
it is preferred to operated second reactor 234 at a temperature
within the range of about 250.degree. to 500.degree. C.
Temperatures above about 450.degree. C. in the second reactor can
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 the preferred 10 X zeolite
catalyst, it is believed that cyclization reactions also occur such
that the C7+ fractions contain substantial substituted aromatics.
At increasing temperatures approaching 400.degree. C., it is
believed that methyl bromide conversion increases towards 90% or
greater, however selectivity towards C.sub.5+ products decreases
and selectivity towards lighter products, particularly olefins
increases. At temperatures exceeding 550.degree. C., it is believed
that a high conversion of methyl bromide to methane and
carbonaceous, coke occurs. In the preferred operating temperature
range of between about 300.degree. C. and 450.degree. C., as a
byproduct of the reaction, a lesser amount of coke probably 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.degree. 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.degree. 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.degree. C. to about 500.degree. C., but preferably in
the range of about 300.degree. C. to about 450.degree. C. in the
second reactor 234 balance increased selectivity of the desired
olefins and C.sub.5+ 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.
[0076] The catalyst may be periodically regenerated in situ, by
isolating reactor 234 from the normal process flow, purging with an
inert gas via line 270 at a pressure in the range of about 1 bar to
about 5 bar and an elevated temperature in the range of about
400.degree. C. to about 650.degree. C. to remove unreacted material
adsorbed on the catalyst insofar as is practical, and then
subsequently oxidizing the deposited carbon to CO.sub.2 by addition
of air or inert gas-diluted oxygen via line 270 to reactor 234 at a
pressure in the range of about 1 bar to about 5 bar and an elevated
temperature in the range of about 400.degree. C. to about
650.degree. C. Carbon dioxide and residual air or inert gas are
vented from reactor 234 via line 275 during the regeneration
period. The effluent which comprises hydrobromic acid and higher
molecular weight hydrocarbons, olefins or mixtures thereof is
withdrawn from the second reactor 234 via line 235 and cooled to a
temperature in the range of about 100.degree. C. to about
600.degree. C. in exchanger 236. As illustrated in FIG. 4A, the
cooled effluent is transported via lines 235 and 241 with valve 238
in the opened position and valves 239 and 243 in the closed
position and introduced into a vessel or reactor 240 containing a
bed 298 of 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.degree. 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. Or more preferably, an alumina support with a
specific surface area of about 50 to 200 m2/g. In reactor 240,
hydrobromic acid is reacted with the metal oxide at temperatures
below about 600.degree. C. and preferably between about 100.degree.
C. to about 500.degree. C. in accordance with the following general
formula wherein M represents the metal:
2HBr+MO.fwdarw.MBr.sub.2+H.sub.2O
The steam resulting from this reaction is transported together with
olefins and/or the high molecular hydrocarbons in line 244, 218 and
216 via opened valve 219 to heat exchanger 220 wherein the mixture
is cooled to a temperature in the range of about 0.degree. C. to
about 70.degree. C. This cooled mixture is forwarded to dehydrator
250 to remove substantially all water from the gas stream. The
water is then removed from the dehydrator 250 via line 253. The
dried gas stream containing olefins, higher molecular weight
hydrocarbons or mixtures thereof is further passed via line 251 to
product recovery unit 252 to recover olefins, the C.sub.5+
fraction, or mixtures thereof as a liquid product in line 254. 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 or other solvent, as used to process natural gas or
refinery gas streams and recover olefinic hydrocarbons, as known to
a skilled artisan, may be employed in the implementation of this
invention. The residual vapor effluent from product recovery unit
252 is then split into a purge stream 257 that may be utilized as
fuel for the process and a recycled residual vapor which is
compressed via compressor 258. The recycled residual vapor
discharged from compressor 258 is split into two fractions. A first
fraction that is equal to at least 1.5 times the feed gas volume is
transported via line 262, combined with the liquid bromine and feed
gas conveyed in line 225 and passed to heat exchanger 226 wherein
the liquid bromine is vaporized and fed into first reactor 230 in a
manner as described above. The second fraction which is drawn off
line 262 via line 263 and is regulated by control valve 260, at a
rate sufficient to dilute the alkyl bromide concentration to
reactor 234 and absorb the heat of reaction such that reactor 234
is maintained at the selected operating temperature, preferably in
the range of about 300.degree. C. to about 450.degree. C. in order
to maximize 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 reactor 230 to be
controlled in addition to moderating the temperature in second
reactor 234. Oxygen, oxygen enriched air or air 210 is delivered
via blower or compressor 213 at a pressure in the range of about
ambient to about 10 bar to bromine via line 214, line 215 and valve
249 through heat exchanger 215, wherein oxygen, oxygen enriched air
or air is preheated to a temperature in the range of about
100.degree. C. to about 500.degree. C. to a second vessel or
reactor 246 containing a bed 299 of a solid phase metal bromide.
Oxygen reacts with the metal bromide in accordance with the
following general reaction wherein M represents the metal:
MBr.sub.2+1/2O.sub.2.fwdarw.MO+Br.sub.2
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. Reactor 246 is
operated below 600.degree. C., and more preferably between about
300.degree. C. to about 500.degree. C. The resultant bromine vapor
is transported from reactor 246 via line 247, valve 248 and line
242 to heat exchanger or condenser 221 where the bromine is
condensed into a liquid. The liquid bromine is transported via line
242 to separator 222 wherein liquid bromine is removed via line 225
and transported via line 225 to heat exchanger 226 and first
reactor 230 by any suitable means, such as by pump 224. The
residual air or unreacted oxygen is transported from separator 222
via line 227 to a bromine scrubbing unit 223, 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 line
212 to line 225. The scrubbed air or oxygen is vented via line 229.
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.
[0077] 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.
[0078] Reactors 240 and 246 may be operated in a cyclic fashion. As
illustrated in FIG. 4A, valves 238 and 219 are operated in the open
mode to permit hydrobromic acid to be removed from the effluent
that is withdrawn from the second reactor 234, while valves 248 and
249 are operated in the open mode to permit air, oxygen enriched
air or oxygen to flow through reactor 246 to oxidize the solid
metal bromide contained therein. Once significant conversion of the
metal oxide and metal bromide in reactors 240 and 246,
respectively, has occurred, these valves are closed. At this point,
bed 299 in reactor 246 is a bed of substantially solid metal
bromide, while bed 298 in reactor 240 is substantially solid metal
oxide. As illustrated in FIG. 5A, valves 245 and 243 are then
opened to permit oxygen, oxygen enriched air or air to flow through
reactor 240 to oxidize the solid metal bromide contained therein,
while valves 239 and 217 are opened to permit effluent which
comprises olefins, the higher molecular weight hydrocarbons and/or
hydrobromic acid that is withdrawn from the second reactor 234 to
be introduced into reactor 246. The reactors are operated in this
manner until significant conversion of the metal oxide and metal
bromide in reactors 246 and 240, respectively, has occurred and
then the reactors are cycled back to the flow schematic illustrated
in FIG. 4A by opening and closing valves as previously
discussed.
[0079] When oxygen is utilized as the oxidizing gas transported in
via line 210 to the reactor being used to oxidize the solid metal
bromide contained therein, the embodiment of the process of the
present invention illustrated in FIGS. 4A and 5A can be modified
such that the bromine vapor produced from either reactor 246 (FIG.
4B) or 240 (FIG. 5B) is transported via lines 242 and 225 directly
to first reactor 230. 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. Compressor 213 is not illustrated in FIGS. 4B and 5B as
substantially all commercial sources of oxygen, such as a
commercial air separator unit, will provide oxygen to line 210 at
the required pressure. If not, a compressor 213 could be utilized
to achieve such pressure as will be evident to a skilled
artisan.
[0080] In the embodiment of the present invention illustrated in
FIG. 6A, the beds of solid metal oxide particles and solid metal
bromide particles contained in reactors 240 and 246, 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 olefins, the higher molecular weight
hydrocarbons and/or hydrobromic acid is withdrawn from the second
reactor 234 via line 235, cooled to a temperature in the range of
about 100.degree. C. to about 500.degree. C. in exchanger 236, and
introduced into the bottom of reactor 240 which contains a bed 298
of solid metal oxide particles. The flow of this introduced fluid
should induce the particles in bed 298 to move upwardly within
reactor 240 as the hydrobromic acid is reacted with the metal oxide
in the manner as described above with respect to FIG. 4A. At or
near the top of the bed 298, 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 reactor 240 are withdrawn via
a weir or cyclone or other conventional means of solid/gas
separation, flow by gravity down line 259 and are introduced at or
near the bottom of the bed 299 of solid metal bromide particles in
reactor 246. In the embodiment illustrated in FIG. 6A, oxygen,
oxygen enriched air or air 210 is delivered via blower or
compressor 213 at a pressure in the range of about ambient to about
10 bar, transported via line 214 through heat exchanger 215,
wherein the oxygen, oxygen enriched air or air is preheated to a
temperature in the range of about 100.degree. C. to about
500.degree. C. and introduced into second vessel or reactor 246
below bed 299 of a solid phase metal bromide. Oxygen reacts with
the metal bromide in the manner described above with respect to
FIG. 4A to produce a dry, substantially HBr free bromine vapor. The
flow of this introduced gas should induce the particles in bed 299
to flow upwardly within reactor 246 as oxygen is reacted with the
metal bromide. At or near the top of the bed 298, 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 reactor 246 are
withdrawn via a weir or cyclone or other conventional means of
solid/gas separation, flow by gravity down line 264 and are
introduced at or near the bottom of the bed 298 of solid metal
oxide particles in reactor 240. In this manner, reactors 240 and
246 may be operated continuously without changing the parameters of
operation.
[0081] In the embodiment illustrated in FIG. 6B, oxygen is utilized
as the oxidizing gas and is transported in via line 210 to reactor
246. Accordingly, the embodiment of the process of the present
invention illustrated in FIG. 6A is modified such that the bromine
vapor produced from reactor 246 is transported via lines 242 and
225 directly to first reactor 230. Since oxygen is reactive and
will not build up in the system, it is believed that the need to
condense the bromine vapor to a liquid to remove unreactive
components, such as nitrogen, should be obviated. Compressor 213 is
not illustrated in FIG. 6B as substantially all commercial sources
of oxygen, such as a commercial air separator unit, will provide
oxygen to line 210 at the required pressure. If not, a compressor
213 could be utilized to achieve such pressure as will be evident
to a skilled artisan.
[0082] In accordance with another embodiment of the process of the
present invention that is illustrated in FIG. 7, the alkyl
bromination and alkyl bromide conversion stages are operated in a
substantially similar manner to those corresponding stages
described in detail with respect to FIG. 4A except as discussed
below. Residual air or oxygen and bromine vapor emanating from
reactor 246 is transported via line 247, valve 248 and line 242 and
valve 300 to heat exchanger or condenser 221 wherein the
bromine-containing gas is cooled to a temperature in the range of
about 30.degree. C. to about 300.degree. C. The bromine-containing
vapor is then transported via line 242 to vessel or reactor 320
containing a bed 322 of 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, Maryland. More
preferably the metal is deposited in oxide form in a range of about
10 to 20 wt % on an alumina support with a specific surface area in
the range of about 50 to 200 m2/g, In reactor 320, bromine vapor is
reacted with the solid phase metal bromide, preferably retained on
a suitable attrition-resistant support at temperatures below about
300.degree. C. and preferably between about 30.degree. C. to about
200.degree. C. in accordance with the following general formula
wherein M.sup.2 represents the metal:
2M.sup.2Br.sub.n+Br.sub.2.fwdarw.2M.sup.2Br.sub.n+1
In this manner, bromine is stored as a second metal bromide, i.e.
2M.sup.2Br.sub.n+1, in reactor 320 while the resultant vapor
containing residual air or oxygen is vented from reactor 320 via
line 324, valve 326 and line 318.
[0083] The gas stream containing lower molecular weight alkanes,
comprised of mixture of a feed gas (line 211) and a recycled gas
stream, is transported or conveyed via line 262, heat exchanger
352, wherein the gas stream is preheated to a temperature in the
range of about 150.degree. C. to about 600.degree. C., valve 304
and line 302 to a second vessel or reactor 310 containing a bed 312
of 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, Maryland.
More preferably the metal is deposited in an oxide state in a range
of 10 to 20 wt % supported on an alumina support with a specific
surface area of about 50 to 200 m2/g. The temperature of the gas
stream is from about 150.degree. C. to about 600.degree. C., and
preferably from about 200.degree. C. to about 450.degree. C. In
second reactor 310, 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 M.sup.2 represents the metal:
2M.sup.2Br.sub.n+1.fwdarw.2M.sup.2Br.sub.n+Br.sub.2
The resultant bromine vapor is transported with the gas stream
containing lower molecular weight alkanes via lines 314, 315, valve
317, line 330, heat exchanger 226 prior to being introduced into
alkyl bromination reactor 230.
[0084] Reactors 310 and 320 may be operated in a cyclic fashion. As
illustrated in FIG. 7, valve 304 is operated in the open mode to
permit the gas stream containing lower molecular weight alkanes to
be transported to the second reactor 310, while valve 317 is
operated in the open mode to permit this gas stream with bromine
vapor that is generated in reactor 310 to be transported to alkyl
bromination reactor 230. Likewise, valve 306 is operated in the
open mode to permit bromine vapor from reactor 246 to be
transported to reactor 320, while valve 326 is operated in the open
mode to permit residual air or oxygen to be vented from reactor
320. Once significant conversion of the reduced metal bromide and
oxidized metal bromide in reactors 320 and 310, respectively, to
the corresponding oxidized and reduced states has occurred, these
valves are closed as illustrated in FIG. 8. At this point, bed 322
in reactor 320 is a bed of substantially metal bromide in an
oxidized state, while bed 312 in reactor 310 is substantially metal
bromide in a reduced state. As illustrated in FIG. 8, valves 304,
317, 306 and 326 are closed, and then valves 308 and 332 are opened
to permit the gas stream containing lower molecular weight alkanes
to be transported or conveyed via lines 262, heat exchanger 352,
wherein gas stream is heated to a range of about 150.degree. C. to
about 600.degree. C., valve 308 and line, 309 to reactor 320 to
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. Valve 332 is also opened to permit the
resultant bromine vapor to be transported with the gas stream
containing lower molecular weight alkanes via lines 324 and 330 and
heat exchanger 226 prior to being introduced into alkyl bromination
reactor 230. In addition, valve 300 is opened to permit. bromine
vapor emanating from reactor 246 to be transported via line 242
through exchanger 221 into reactor 310 wherein the solid phase
metal bromide in a reduced valence state reacts with bromine to
effectively store bromine as a metal bromide. In addition, valve
316 is opened to permit the resulting gas, which is substantially
devoid of bromine to be vented via lines 314 and 318. The reactors
are operated in this manner until significant conversion of the
beds of reduced metal bromide and oxidized metal bromide in
reactors 310 and 320, respectively, to the corresponding oxidized
and reduced states has occurred and then the reactors are cycled
back to the flow schematic illustrated in FIG. 7 by opening and
closing valves as previously discussed.
[0085] In the embodiment of the present invention illustrated in
FIG. 9, the beds 312 and 322 contained in reactors 310 and 320,
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
reactor 246 via line 242 is cooled to a temperature in the range of
about 30.degree. C. to about 300.degree. C. in exchangers 370 and
372, and introduced into the bottom of reactor 320 which contains a
moving solid bed 322 in a fluidized state. The flow of this
introduced fluid should induce the particles in bed 322 to flow
upwardly within reactor 320 as the bromine vapor is reacted with
the reduced metal bromide entering the bottom of bed 322 in the
manner as described above with respect to FIG. 7. At or near the
top of the bed 322, 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 reactor 320 are withdrawn via a weir, cyclone
or other conventional means of solid/gas separation, flow by
gravity down line 359 and are introduced at or near the bottom of
the bed 312 in reactor 310. In the embodiment illustrated in FIG.
9, the gas stream containing lower molecular weight alkanes,
comprised of mixture of a feed gas (line 211) and a recycled gas
stream, is transported or conveyed via line 262 and heat exchanger
352 wherein the gas stream is heated to a range of about
150.degree. C. to about 600.degree. C. and introduced into reactor
310. The heated gas stream thermally decomposes the solid phase
metal bromide in an oxidized valence state present entering at or
near the bottom of bed 312 to yield elemental bromine vapor and a
solid metal bromide in a reduced state. The flow of this introduced
gas should induce the particles in bed 312 to flow upwardly within
reactor 310 as the oxidized metal bromide is thermally decomposed.
At or near the top of the bed 312, the particles which contain
substantially reduced solid metal bromide on the
attrition-resistant support due to the substantially complete
thermal decomposition in reactor 310 are withdrawn via a weir or
cyclone or other conventional means of gas/solid separation and
flow by gravity down line 364 and introduced at or near the bottom
of the bed 322 of particles in reactor 310. In this manner,
reactors 310 and 320 may be operated continuously with changing the
parameters of operation.
[0086] It is believed that the process of the present invention
should be 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.degree. C. to about 600.degree. C. for the gas phase, and
preferably about 20.degree. C. to about 180.degree. C. for the
liquid phase. It is believed that these operating conditions should
permit the use of less expensive equipment of relatively simple
design that are constructed from readily available metal alloys or
glass-lined equipment for the gas phase and polymer-lined or
glass-lined vessels, piping and pumps for the liquid phase. It is
believed that the process of the present invention also should be
more efficient because less energy should be 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),
olefin and motor gasoline fuels range that have substantial
aromatic content thereby significantly increasing the octane value
of the gasoline-range fuel components.
[0087] The following examples demonstrate the practice and utility
of the present invention, but are not to be construed as limiting
the scope thereof.
EXAMPLE 1
[0088] Various mixtures of dry bromine and methane are reacted
homogeneously at temperatures in the range of 459.degree. C. to
491.degree. 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.sup.-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.
EXAMPLE 2
[0089] FIG. 13 and FIG. 14 illustrate two exemplary PONA analyses
of two C.sub.6+ 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 C.sub.6+ fractions produced.
EXAMPLE 3
[0090] Methyl bromide is reacted over a ZSM-5 zeolite catalyst at a
Gas Hourly Space Velocity (GHSV) of approximately 94 hr.sup.-1 over
a range of temperatures from about 100.degree. C. to about
460.degree. C. at approximately 2 bar pressure. As illustrated in
FIG. 10, 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.degree. C. to about 350.degree. C. Lower
temperatures in the range of about 100.degree. C. to about
250.degree. C. favor selectivity towards higher molecular weight
products however conversion is low. Higher temperatures in the
range of about 250.degree. C. to about 350.degree. 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.degree. C. selectivity to methane rapidly increases. At about
450.degree. C., almost complete conversion to methane occurs.
EXAMPLE 4
[0091] Methyl bromide, hydrogen bromide and methane are reacted
over a ZSM-5 zeolite catalyst at approximately 2 bar pressure at
about 250.degree. C. and also at about 260.degree. C. at a GHSV of
approximately 76 hr.sup.-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.degree. C. and at about 260.degree. C. at a GHSV of
approximately 73 hr.sup.-1 were also run. FIG. 11, 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.
EXAMPLE 5
[0092] Methyl bromide is reacted over a ZSM-5 zeolite catalyst at
230.degree. C. Dibromomethane is added to the reactor. FIG. 12,
which is a graph of product selectivity, indicates that reaction of
methyl bromide and dibromomethane results in a shift in selectivity
towards C.sub.5+ 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.
EXAMPLE 6
[0093] 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.sup.-1. A methyl bromide
conversion of approximately 86% and a propyl bromide conversion of
approximately 98% is observed.
[0094] 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.degree. C. to about 600.degree. C. and most preferably
in the range of about 120.degree. C. to about 180.degree. C. to
yield elemental bromine vapor and metal hydroxide, metal
oxy-bromide or metal oxide. Operation at temperatures below about
180.degree. 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.
[0095] 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.
* * * * *