CA2203115C - Single stage fixed bed oxychlorination of ethylene - Google Patents

Single stage fixed bed oxychlorination of ethylene Download PDF

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Publication number
CA2203115C
CA2203115C CA002203115A CA2203115A CA2203115C CA 2203115 C CA2203115 C CA 2203115C CA 002203115 A CA002203115 A CA 002203115A CA 2203115 A CA2203115 A CA 2203115A CA 2203115 C CA2203115 C CA 2203115C
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Prior art keywords
ethylene
reactor
catalyst
oxygen
method according
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Expired - Fee Related
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CA002203115A
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French (fr)
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CA2203115A1 (en
Inventor
Pierluigi Fatutto
Andrea Marsella
Dario Vio
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EVC Technology AG
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EVC Technology AG
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Priority to GB9421136A priority Critical patent/GB2294262B/en
Priority to GB9421136.4 priority
Application filed by EVC Technology AG filed Critical EVC Technology AG
Priority to PCT/IB1995/000872 priority patent/WO1996012693A1/en
Publication of CA2203115A1 publication Critical patent/CA2203115A1/en
Application granted granted Critical
Publication of CA2203115C publication Critical patent/CA2203115C/en
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Classifications

    • YGENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02PCLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00Technologies relating to chemical industry
    • Y02P20/50Improvements relating to the production of products other than chlorine, adipic acid, caprolactam, or chlorodifluoromethane, e.g. bulk or fine chemicals or pharmaceuticals
    • Y02P20/58Recycling
    • Y02P20/582Recycling of unreacted starting or intermediate materials

Abstract

A method for the oxychlorination of ethylene to produce 1,2-dichloroethane comprises reacting ethylene with a chlorine source and an oxygen source in a fixed-bed oxychlorination reactor in the presence of a catalyst, and is characterized in that a single reactor is used and ethylene is present in a large molar excess with respect to chlorine. The chlorine source is suitably HC1, the catalyst is suitably a cupric chloride catalyst, and the molar excess of ethylene is preferably between 200 and 700 %.

Description

WO 96/12693 PCTlIB9510~872 SINGLE STAGE FIXED BED OXYCHLORINATION OF ETHYLENE
The present invention relates to the oxychlorination of ethylene in a fixed bed reactor system which consists of a single reactor, to produce chlorinated hydrocarbons, ., particularly 1, 2-dichloroethane (EDC).
It is well known that hydrocarbons such as ethylene may be chlorinated by reacting them with hydrogen chloride and gases containing elemental oxygen, particularly air or oxygen enriched air, in the presence of a catalyst at elevated temperatures and pressures in order to produce chlorinated hydrocarbons such as EDC. The reaction may be carried out with two different reactor technologies. The 15- first is fluid bed reactor technology wherein a gaseous mixture of reactants is contacted with a fluidizable catalyst powder. The second is fixed bed reactor technology, in which the gaseous reactants flow over a fixed catalyst inside the reactor.
Fluid bed reactors have a number of drawbacks, such as potential stickiness of the catalyst powder, unsteady operation, poor selectivity owing to the gas and catalyst solids back mixing in the reactor, loss of heat transfer owing to fouling of the cooler bundle and limits in reagent velocity imposed by the need to avoid catalyst loss by elutriation from the reactor.
Fixed bed reactor technology has been developed in order to overcome these problems (see US patent 3,892,816 and US
patent 4,123,467).
Although the fixed bed reactor overcomes many of the problems incurred with the fluid bed reactor system, a number of new problems have been encountered. A major problem is the difficulty, in the fixed bed reactor, of transferring the heat developed by the exothermic oxychlorination reaction away from the reactor to prevent CA 02203115 1997-04-18 ' overheating. For this reason, all the necessary reagents may not be fed in the correct stoichiomentric ratio to the reactor. Moreover, because it can be unsafe to have an oxygen concentration of above 8% in the mixture feeding the reactor, for flammability reasons, the reaction is carried out in two or more subsequent stages (usually three) such that the ethylene is introduced into the first reactor while the HC1 and oxygen feeds are split between the reactors.
Unreacted ethylene plus some inert gases are recycled back to the first reactor.
In a further attempt to reduce the incidence of hot spots and the like, it is known to alter the activity profile of the catalyst within a fixed bed reactor such that the activity increases in the direction of flow. For example, see European patent application 0146925. However, in the prior art, even when a profiled catalyst is used it has been deemed necessary to use a multi-reactor system.
We have now developed a new process for the catalytic oxychlorination of ethylene which makes use of a single fixed-bed reactor. Nevertheless, hot-spots are avoided and good selectivity to ethylene is achieved, as well as over 99% utilisation of HC1.
According to a first aspect of the invention, we provide a method for the oxychlorination of ethylene to produce 1,2-dichloroethane (EDC), comprising reacting ethylene, a chlorine source and an oxygen source in a fixed-bed oxychlorination reactor in the presence of a catalyst, characterised in that a single reactor is used and ethylene is present in a large molar excess with respect to chlorine.
Preferably, the chlorine source is HCl. ' Preferably, the ethylene is introduced in a 200 - 700 % molar equivalent excess with respect to stoichiometric HC1, in order to produce a high partial pressure of ethylene.

WO 96112693 PCTlIB951008~2 The oxygen source may be pure oxygen, or an oxygen-enriched gas. Oxygen is preferably supplied in a molar excess of up to 15%, more preferably between 2 and 8%, with respect to HC1. -The large excess of ethylene present functions to increase the selectivity of the reaction, as well as acting as a heat sink, exploiting its high specific heat capacity. Unreacted ethylene is preferably recovered and recycled back to the l0 reactor, or to other processes requiring ethylene such as in direct chlorination reactions.
The composition of the recycle gas reaches an equilibrium depending mainly on combustion rate, the amount of inert 15 gases in the raw materials and the purge rate. Depending on these factors, ethylene concentration can vary between 10 and 90%. As a consequence, the actual ethylene excess used will depend on its concentration in the recycled vent gas and on the recycle flow rate.
In general, the ethylene excess with respect to its stoichiometric requirement as determined by the amount of HC1 can be expressed as a percentage calculated according to the formula:
(Q1 + Q2 - Q6) where 100 is the stoichiometirc requirement, and wherein Q2 = Q4 Q1 -1/2 Q3 + Q5 and 100Q7 - Q8% 02 in 3 5 Q4 =
%02 in - %02 rec CA 02203115 1997-04-18 . , WO 96/12693 . PCT/1B95/00872 where %02in - oxygen at inlet of reactor 02 rec ~ oxygen in recycle stream.
The symbols are defined as follows:
Q1 - mol/h fresh ethylene Q2 - mol/h recycled ethylene Q3 - mol/h HC1 Q4 - mol/h total recycle Q5 - mol/h burned ethylene Q6 - mol/h fed inert gases Q7 - mol/h fed oxygen Q8 - mol/h total fresh reagents Control of the recycled gas flow rate may be used to adjust the oxygen concentration at the inlet of the reactor and thereby the hotspot temperature. In the conditions of temperature and pressure existing in the inlet of the reactor, the lower flammability limit of the mixture occurs when the oxygen concentration is around 8%. For safety and operational reasons, the concentration is advantageously between 5 and 6% v/v, as use of a higher concentration can result in an elevated hotspot temperature in the catalytic bed.
Typically, the hotspot temperature would be about 230 280~C, depending on a number of factors, including reactor diameter.
The reactor employed in the method of the invention is a tubular reactor. Advantageously, it consists of a plurality of tubes stacked together within a single coolant jacket.
The internal diameter of each tube is preferably between 20 and 40 millimetres. Diameters of less than 20 millimetres are disadvantageous as an excessive number of tubes is required in an industrial reactor in order to obtain a satisfactory throughput of materials, while diameters larger than 40 millimetres result in excessively high hotspot temperatures inside the catalytic bed.

The preferred length of the reactor is between 3.5 and 8 metres. A length of less than 3.5 metres results in too short a residence time and therefore either low reactant conversion or low specific throughput; a length of more than 5 8 metres is not necessary in order to achieve both high HCL
and oxygen conversion and large specific throughput.
Catalyst layers within the reactor can be arranged in a number of ways. For example, the reactor may simply be filled with catalyst in the normal manner, not employing a profiled catalyst distribution. Alternatively, a simple loading pattern maybe employed whereby the catalyst is loaded in two layers, a first of low activity catalyst or diluted catalyst (See USP 4,123,467) in order to avoid hotspots and a second of a more active or more concentrated catalyst, in order to increase the rate of reaction. A
further, more complex loading pattern consists of a succession of several layers of catalyst with increasing activity (or concentration) from the first to the last layer. The choice of suitable catalyst loading pattern will depend on the maximum temperature of the hotspot, as well as the inside diameter and length of the tubular reactor and on the projected throughput.
Invariably, it is advantageous to fill the last part of the reactor with a high activity catalyst as used in the third reactor of a three stage oxychlorination process.
Catalysts for use in the invention are known in the art and are supported catalysts in which cupric chloride is the major active component and alumina, silica gel, alumino silicate and the like form the supports. The support material may be present in the form of spheres, cubes, cones, hollow cylinders, cylindrical pellets, multilobate pellets and other shapes.
In addition to cupric chloride, the catalyst may also comprise promoters such as potassium, magnesium, sodium, lithium, calcium, cerium and cesium chlorides for improving the selectivity to EDC. The activity prolfile of the catalyst in the catalytic bed would be arranged in such a way as to have the HC1 conversion above 98% at a point 70 to 80~ of the distance along the catalytic bed. The last 20 to 30~ of the catalytic bed will perform as a finisher, so that the whole reaction is assured of a high conversion even if the first part of the catalytic bed loses activity over time.
Preferably, the reactants are preheated to between 100 and 200~C. The reaction pressure can range up to 20 barg, the preferred range being between 4 and 7 barg.
The invention will now be described, for the purposes of exemplification only, in the following examples with reference to the appended figure, which is a systematic diagram of a single reactor oxychlorination apparatus.
Example 1 The reactor is a unit composed of 1 inch external diameter (e. d) nickel tube 14 BWG 25 feet long; inside on the axis there is a thermowell of 6 mm e.d. containing 8 thermocouples with which it is possible to record the thermal profile of the reactor. The reactor is surrounded with an external jacket in which steam at 210~C and 18 barg is used to control the temperature of the reaction. The reactor pressure is controlled with a pneumatic valve on an effluent line.
The reagents were preheated in 18 barg steam heated exchangers; subsequently, ethylene, HC1 and nitrogen were mixed together and oxygen was added in a mixer where the velocity of the gases is higher than the speed of ethylene flame propagation. The catalyst used was a normal industrial catalyst for oxygen three stage fixed bed process consisting of hollow cylinders containing copper and potassium chloride arranged such that the amount of the WO 96/12693 PCTlIB95/00872 copper varies from 24 to 60gr/litre from inlet to outlet of the reactor. The catalytic bed was 2.6 litres. In this reactor, a mixture of 223.1 moles/h of ethylene, 66.8 '~ moles/h of HC1, 17.5 moles/h of oxygen and 18 moles/h of nitrogen was introduced. The oxygen excess was 4.8%.
Nitrogen was used to stimulate the inert gases of the recycle gas and its amount was a function of the recycle ratio. In this case the recycle composition was 90%
ethylene and 10~ inert gases. The pressure at inlet was 5.1 barg and at outlet 3.5 barg. The temperature of cooling jacket was held at 210~C. The outlet stream, consisting of a mixture of ethylene, oxygen, HC1, nitrogen, EDC, water, COx and byproducts, was analyzed and the results were:
-Oxygen conversion to crude EDC 94.7%
-HC1 conversion to crude EDC 99.40 -EDC production 32.8 mol/h.
-Selectivity of ethylene to COx 0.6% mol -Selectivity of ethylene to EC1 0.24% mol -Selectivity of ethylene to EDC 98,66% mol -Selectivity of ethylene to chloral 0,15% mol -Selectivity of ethylene to impurities 0,35% mol -Hotspot 233~C
Example 2 This example was carried out with the same reactor and catalytic scheme as in example 1. It was fed a mixture of ethylene 169 moles/h, HC1 71.5 moles/h, oxygen 19.5 moles/h and nitrogen 65 moles/h. The oxygen excess was 9%. Recycle composition was ethylene 70%, inerts 30% v/v. The outlet pressure was 7.2 barg. The cooling jacket temperature was 2IO~C.
The results were:
-Oxygen conversion to crude EDC 90.8%
-HC1 conversion to crude EDC 99.50 -EDC production 35.6 moles/h -Selectivity of ethylene to COx 0,34% mol -Selectivity of ethylene to EC1 0,17% mol -Selectivity of ethylene to chloral 0,20% mol '' -Selectivity of ethylene to EDC 98,92% mol -Selectivity of ethylene to impurities 0.37% mol -Hotspot 2'17 ~C

Example 3 This example was carried out with the same reactor and catalytic scheme of examples 1 and 2. It was a fed a mixture of ethylene 267.55 moles/h, HC1 93.45 moles/h, oxygen 25 moles/h, nitrogen 50moles/h. Oxygen excess was 7%.
Recycle composition: ethylene 80%, inerts 20% v/v. Outlet pressure was 3.2 barg and the inlet pressure 6.6 barg. The cooling jacket temperature was 210~C. The results were:
-Oxygen conversion EDC 90.9%
to crude -HC1 conversion to crude EDC 99%

-EDC production 46.2moles/h -Selectivity of ethylene to COx 0.13% mol -Selectivity of ethylene to EC1 0,25% mol -Selectivity of ethylene to chloral 0,10% mol -Selectivity of ethylene to EDC 99,12% mol -Selectivity of ethylene to impurities 0,40% mol -Hotspot 233~C

Example 4 A reactor consisting of one 1.25 inch e.d. nickel tube 14 BWG 12 feet long was set up, the same as that used in a normal three stage industrial reactor, inside of which there was a thermowell of 6 mm e.d. containing 4 sliding thermocouples. The temperature, pressure control and reagent feeding systems were as in Example 1. The catalytic WO 96/12693 PCTlIB95/00872 bed was 1.8 litres. The catalyst used was the same type as in other examples. The copper content ranged from 24 to 46gr/litre.
It was fed a mixture of 272 moles/h of ethylene, HC1 97.5 ~r moles/h, oxygen 26.8 moles/h and nitrogen 24.6 moles/h. The oxygen excess was 10 % . Recycle composition was 90 % ethylene, 10% inerts. The outlet pressure was 5.5 barg and the pressure drop 0.55 barg. The results were:
-Oxygen conversion 89%
to crude.EDC

-HC1 conversion to crude EDC 98%

-EDC production 47.6moles%h -Selectivity of ethylene COx 0.13% mol -Selectivity of ethylene to EC1 0.20% mol -Selectivity of ethylene to EDC 99.10% mol -Selectivity of ethylene to chloral 0.15% mol -Selectivity of ethylene to impurities o.42% mol -Hotspot 258~C

Example 5 The same reactor and catalytic scheme as in Example 4 was fed a mixture of 214.2 moles/h ethylene, 84.7 moles/h of HC1, 24.1 moles/h of oxygen and 21.6 moles/h of nitrogen.
Oxygen excess was 13.4%. Recycle composition was 90%
ethylene, 10% inerts. The outlet pressure was 5.5 barg and the pressure drop was 0.5 barg. The results were:
-Oxygen conversion 87.1%

-HC1 conversion 99.1%

-EDC production 42 moles/h -Selectivity of ethylene to COx 0.30% mol -Selectivity of ethylene to EC1 0.20% mol -Selectivity of ethylene to EDC 98.50% mol -Selectivity of ethylene to chloral 0.20% mol -Selectivity of ethylene to impurities 0.80% mol -Hotspot 270~C

Example 6 Using the same reactor as in Examples 4 and 5 but with the 5 same loading pattern as in the first stage of a three stage process consisting of a catalyst containing 26 gr/litre of copper and l2gr/litre of potassium in the first 3/5 of the reactor and 40 gr/litre of copper and l8gr/litre of potassium in the remainder part, a mixture of ethylene 231 10 moles/h, oxygen 18 moles/h, HC1 64.8 moles/h and nitrogen 42 moles/h was reacted. The outlet pressure was 5.5 berg and the pressure drop was 0.45 berg.
-Oxygen conversion EDC 88.5%
to crude -HCl conversion to crude EDC 97.80 -EDC production 31.70 moles/h -Selectivity of ethylene to COx 0.10% mol -Selectivity of ethylene to ECl 0.20a mol -Selectivity of ethylene to EDC 99.300 mol -Selectivity of ethylene to chloral 0.10% mol -Selectivity of ethylene impur 0.30% mol -Hotspot 255~C

Claims (9)

1. A method for the oxychlorination of ethylene to produce 1,2-dichloroethane (EDC), comprising reacting ethylene with a chlorine source and an oxygen source in a fixed-bed oxychlorination reactor in the presence of a catalyst, characterised in that a single reactor is used and that ethylene is present in a molar excess with respect to chlorine of between 200 and 700 %.
2. A method according to Claim 1 wherein the chlorine source is HCl.
3. A method according to Claim 1 or Claim 2 wherein the catalyst is a cupric chloride catalyst.
4. A method according to Claim 3 wherein the catalyst further comprises the chlorides of potassium, magnesium, cesium, sodium, lithium, calcium or cerium.
5. A method according to any preceding Claim wherein oxygen is present in a molar excess of up to 15% with respect to the chlorine.
6. A method according to Claim 5 wherein the molar excess of oxygen is between 2 and 8%.
7. A method according to any preceding Claim wherein the activity profile of the catalyst increases in the direction of reagent flow.
8. A method according to any preceding Claim wherein the reactor comprises at least one tube having an internal diameter of between 20 and 40 mm, and a length between 3.5 and 8.5m.
9. A method according to any preceding Claim wherein vent gases are recycled.
CA002203115A 1994-10-20 1995-09-27 Single stage fixed bed oxychlorination of ethylene Expired - Fee Related CA2203115C (en)

Priority Applications (3)

Application Number Priority Date Filing Date Title
GB9421136A GB2294262B (en) 1994-10-20 1994-10-20 Single stage fixed bed oxychlorination of ethylene
GB9421136.4 1994-10-20
PCT/IB1995/000872 WO1996012693A1 (en) 1994-10-20 1995-09-27 Single stage fixed bed oxychlorination of ethylene

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CA2203115C true CA2203115C (en) 2006-09-19

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US7674941B2 (en) 2004-04-16 2010-03-09 Marathon Gtf Technology, Ltd. Processes for converting gaseous alkanes to liquid hydrocarbons
US7838708B2 (en) 2001-06-20 2010-11-23 Grt, Inc. Hydrocarbon conversion process improvements
US7847139B2 (en) 2003-07-15 2010-12-07 Grt, Inc. Hydrocarbon synthesis
US7880041B2 (en) 2004-04-16 2011-02-01 Marathon Gtf Technology, Ltd. Process for converting gaseous alkanes to liquid hydrocarbons
US7883568B2 (en) 2006-02-03 2011-02-08 Grt, Inc. Separation of light gases from halogens
US7964764B2 (en) 2003-07-15 2011-06-21 Grt, Inc. Hydrocarbon synthesis
US7998438B2 (en) 2007-05-24 2011-08-16 Grt, Inc. Zone reactor incorporating reversible hydrogen halide capture and release
US8008535B2 (en) 2004-04-16 2011-08-30 Marathon Gtf Technology, Ltd. Process for converting gaseous alkanes to olefins and liquid hydrocarbons
US8053616B2 (en) 2006-02-03 2011-11-08 Grt, Inc. Continuous process for converting natural gas to liquid hydrocarbons
US8173851B2 (en) 2004-04-16 2012-05-08 Marathon Gtf Technology, Ltd. Processes for converting gaseous alkanes to liquid hydrocarbons
US8198495B2 (en) 2010-03-02 2012-06-12 Marathon Gtf Technology, Ltd. Processes and systems for the staged synthesis of alkyl bromides
US8273929B2 (en) 2008-07-18 2012-09-25 Grt, Inc. Continuous process for converting natural gas to liquid hydrocarbons
US8282810B2 (en) 2008-06-13 2012-10-09 Marathon Gtf Technology, Ltd. Bromine-based method and system for converting gaseous alkanes to liquid hydrocarbons using electrolysis for bromine recovery
US8367884B2 (en) 2010-03-02 2013-02-05 Marathon Gtf Technology, Ltd. Processes and systems for the staged synthesis of alkyl bromides
US8436220B2 (en) 2011-06-10 2013-05-07 Marathon Gtf Technology, Ltd. Processes and systems for demethanization of brominated hydrocarbons
US8642822B2 (en) 2004-04-16 2014-02-04 Marathon Gtf Technology, Ltd. Processes for converting gaseous alkanes to liquid hydrocarbons using microchannel reactor
US8802908B2 (en) 2011-10-21 2014-08-12 Marathon Gtf Technology, Ltd. Processes and systems for separate, parallel methane and higher alkanes' bromination
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US8829256B2 (en) 2011-06-30 2014-09-09 Gtc Technology Us, Llc Processes and systems for fractionation of brominated hydrocarbons in the conversion of natural gas to liquid hydrocarbons
US9193641B2 (en) 2011-12-16 2015-11-24 Gtc Technology Us, Llc Processes and systems for conversion of alkyl bromides to higher molecular weight hydrocarbons in circulating catalyst reactor-regenerator systems
US9206093B2 (en) 2004-04-16 2015-12-08 Gtc Technology Us, Llc Process for converting gaseous alkanes to liquid hydrocarbons

Cited By (26)

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Publication number Priority date Publication date Assignee Title
US8415512B2 (en) 2001-06-20 2013-04-09 Grt, Inc. Hydrocarbon conversion process improvements
US7838708B2 (en) 2001-06-20 2010-11-23 Grt, Inc. Hydrocarbon conversion process improvements
US7847139B2 (en) 2003-07-15 2010-12-07 Grt, Inc. Hydrocarbon synthesis
US7964764B2 (en) 2003-07-15 2011-06-21 Grt, Inc. Hydrocarbon synthesis
US7880041B2 (en) 2004-04-16 2011-02-01 Marathon Gtf Technology, Ltd. Process for converting gaseous alkanes to liquid hydrocarbons
US8642822B2 (en) 2004-04-16 2014-02-04 Marathon Gtf Technology, Ltd. Processes for converting gaseous alkanes to liquid hydrocarbons using microchannel reactor
US8232441B2 (en) 2004-04-16 2012-07-31 Marathon Gtf Technology, Ltd. Process for converting gaseous alkanes to liquid hydrocarbons
US8008535B2 (en) 2004-04-16 2011-08-30 Marathon Gtf Technology, Ltd. Process for converting gaseous alkanes to olefins and liquid hydrocarbons
US7674941B2 (en) 2004-04-16 2010-03-09 Marathon Gtf Technology, Ltd. Processes for converting gaseous alkanes to liquid hydrocarbons
US8173851B2 (en) 2004-04-16 2012-05-08 Marathon Gtf Technology, Ltd. Processes for converting gaseous alkanes to liquid hydrocarbons
US9206093B2 (en) 2004-04-16 2015-12-08 Gtc Technology Us, Llc Process for converting gaseous alkanes to liquid hydrocarbons
US8053616B2 (en) 2006-02-03 2011-11-08 Grt, Inc. Continuous process for converting natural gas to liquid hydrocarbons
US7883568B2 (en) 2006-02-03 2011-02-08 Grt, Inc. Separation of light gases from halogens
US8921625B2 (en) 2007-02-05 2014-12-30 Reaction35, LLC Continuous process for converting natural gas to liquid hydrocarbons
US7998438B2 (en) 2007-05-24 2011-08-16 Grt, Inc. Zone reactor incorporating reversible hydrogen halide capture and release
US8282810B2 (en) 2008-06-13 2012-10-09 Marathon Gtf Technology, Ltd. Bromine-based method and system for converting gaseous alkanes to liquid hydrocarbons using electrolysis for bromine recovery
US8415517B2 (en) 2008-07-18 2013-04-09 Grt, Inc. Continuous process for converting natural gas to liquid hydrocarbons
US8273929B2 (en) 2008-07-18 2012-09-25 Grt, Inc. Continuous process for converting natural gas to liquid hydrocarbons
US9133078B2 (en) 2010-03-02 2015-09-15 Gtc Technology Us, Llc Processes and systems for the staged synthesis of alkyl bromides
US8367884B2 (en) 2010-03-02 2013-02-05 Marathon Gtf Technology, Ltd. Processes and systems for the staged synthesis of alkyl bromides
US8198495B2 (en) 2010-03-02 2012-06-12 Marathon Gtf Technology, Ltd. Processes and systems for the staged synthesis of alkyl bromides
US8815050B2 (en) 2011-03-22 2014-08-26 Marathon Gtf Technology, Ltd. Processes and systems for drying liquid bromine
US8436220B2 (en) 2011-06-10 2013-05-07 Marathon Gtf Technology, Ltd. Processes and systems for demethanization of brominated hydrocarbons
US8829256B2 (en) 2011-06-30 2014-09-09 Gtc Technology Us, Llc Processes and systems for fractionation of brominated hydrocarbons in the conversion of natural gas to liquid hydrocarbons
US8802908B2 (en) 2011-10-21 2014-08-12 Marathon Gtf Technology, Ltd. Processes and systems for separate, parallel methane and higher alkanes' bromination
US9193641B2 (en) 2011-12-16 2015-11-24 Gtc Technology Us, Llc Processes and systems for conversion of alkyl bromides to higher molecular weight hydrocarbons in circulating catalyst reactor-regenerator systems

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