Method of producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons, and installation for implementing same

Information

  • Patent Grant
  • 10550331
  • Patent Number
    10,550,331
  • Date Filed
    Wednesday, March 25, 2015
    9 years ago
  • Date Issued
    Tuesday, February 4, 2020
    4 years ago
Abstract
A method and an installation for producing a concentrate of aromatic hydrocarbons from light aliphatic hydrocarbons and from mixtures thereof with oxygenates. Initial raw material is fed into two in-series-connected reaction units, with zeolite catalysts a mixture obtained following the reaction units is separated into a liquid fraction and a gas fraction, and the gas fraction is fed to the inlet of the first and second reaction unit. The method is characterized in that the gas fraction obtained following the reaction units is separated into a hydrogen-containing gas and into a broad fraction of light hydrocarbons, containing olefins, and in that the hydrogen-containing gas is fed into an oxygenate synthesis unit, in that the resultant oxygenates are fed to the inlet of the first and second reaction unit, and in that the broad fraction of light hydrocarbons, containing olefins, is fed to the inlet of the first reaction unit.
Description

This application is a US nationalization pursuant to 35 U.S.C. § 371 of PCT/RU2015/000171, now WO 2015/147700, filed Mar. 25, 2015, which claims priority to RU Patent Application Serial No. 2014111985, filed Mar. 28, 2014, the entire disclosure of each of which is incorporated herein by reference.


AREA OF TECHNOLOGY TO WHICH THE INVENTION BELONGS

The invention belongs to the area of gas chemistry and gas refining, more specifically to methods and devices for producing aromatic hydrocarbon concentrate (AHCC) from light aliphatic hydrocarbons and oxygenated compounds (C1-C4 aliphatic alcohols), in which the feedstock is fed to two serially connected reactors, a first and second with pentasil-based zeolite catalysts, which differ in the conditions of conversion of aliphatic to aromatic hydrocarbons, the mixture obtained from the reactors is separated into liquid and gaseous fractions, and the gaseous fraction is fed to the inputs of the first and second reactors and can be used to produce aromatic hydrocarbons.


As hydrocarbon feedstock containing aliphatic hydrocarbons, the method can use C2-C4 hydrocarbon fractions, including those containing C5+ components: propane-butane fractions (PBF), natural gas liquids (NGL), light fractions of natural gas liquids, and straight-run gasolines, as well as light naphtha from the Fischer-Tropsch process, etc. PBF containing up to 60-80% propane by weight are preferable. C1-C4 aliphatic alcohols, including bioethanol and biopropanol, can also be used as feedstock.


Technological Level of the Method


In most prior-art methods of producing aromatic hydrocarbons from light aliphatic feedstock, the stream consists mainly of C3-C4 hydrocarbons, and the feedstock is converted without separation into components in a single reaction zone, under identical conditions, while propane is converted to aromatic hydrocarbons with a high yield at higher temperature than butane or propylene, ethylene, while the conversion of ethane requires still higher temperatures. Therefore, the mixed feedstock is brought into contact with the catalyst either at high temperature, to address the less reactive feedstock components, thereby increasing the coking rate and reducing the catalyst's life, or at lower temperature, with a relatively low degree of propane and ethane conversion, with recycling of the unconverted feedstock.


An investigation of light aliphatic feedstock conversion processes noted that propane conversion proceeds more efficiently at a temperature at least 30° C. higher than that needed for butane conversion. This has been noted in U.S. Pat. No. 5,171,912, which describes a process for producing gasoline from n-butane with conversion of the propane produced in n-butane conversion in a separate reactor, and proposes to separate the butane and propane fractions in a rectification column for recycling.


Patent WO2007053061A2 states that the addition of ˜25% olefins to the conversion feedstock permits compensation of the endothermic effect in conversion of saturated aliphatic hydrocarbons (propane and butanes) which permits conversion of aliphatic hydrocarbons to be performed in an adiabatic reactor under near-isothermal conditions.


RF Patent 2477656 describes a method of two-stage methanol conversion in which thermostabilization in both reaction zones is effected by circulation of aliphatic hydrocarbons formed in the synthesis of aromatic hydrocarbons. Circulation is effected at a ratio of seven moles of paraffins (C3-C5) per mole of conversion methanol, which ensures adiabatic conditions in both reaction zones with a 20-30° C. temperature gradient across the catalyst bed.


U.S. Pat. No. 4,642,402 proposes to recycle benzene and/or toluene extracted from the catalyzate at the propane aromatization stage to increase the alkylbenzene yield in the synthesized gasoline.


U.S. Pat. No. 5,043,502 proposes to increase the alkylbenzene yield by using an additional reaction zone where benzene methylation occurs.


These technical solutions have several disadvantages. For example, the tubular reaction furnaces that deliver heat for conversion of C3-C4 light paraffins (Aroforming process described in Thermal and Catalytic Processes in Petroleum Refining by Serge Raseev, New York, 2003) are difficult to operate (catalyst loading/unloading) and expensive. Solutions with thermostabilization through feedstock recycling increase the volumes of reactor equipment (the volume of the reaction zones) and increase energy costs through feedstock circulation. In addition, some of the feedstock will inevitably be carried off with the discharge gases from the feedstock fractionation unit, and the quantity of lost feedstock increases with the number of circulation cycles.


Solutions involving the delivery of heat through superheating of the feedstock fed to the reactor cause soot formation in the heating furnace coils, especially when the feedstock contains olefins. In addition, feedstock superheating causes local superheating of the front bed of the catalyst, causing premature aging or deactivation of the catalyst and uneven catalyst activity across the catalyst bed, since the catalyst's activity rises with temperature.


Solutions involving the feeding of paraffin and olefin mixtures (˜25%) to the reactor permit use of a simple reactor for conversion and avoidance of local catalyst superheating, and stabilization of catalyst activity across its bed, reducing the volume of the reaction zone, but also have disadvantages related to soot production on olefin heating and require additional outlays for olefin production.


Solutions involving increased propane conversion by increasing the temperature in the second reaction zone simultaneously cause a reduction in arene yield.


The prior art also knows a method of producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons in which the feedstock is fed to two serially connected reactors, a first and second with pentasil-based zeolite catalysts, where the reactors differ in the conditions of conversion of hydrocarbons to aromatics, the mixture obtained from the reactors is separated into liquid and gaseous fractions, and the gaseous fraction is fed to the inputs of the first and second reactors, as described in RF Patent 2277524, published in 2006.


RF Patent 2277524 proposes to perform conversion of the propane-butane fraction preferably of 80% butanes sequentially in two reactors, the first of which converts predominantly butane, and the second high-temperature reactor converts propane unconverted in the first reactor and other mainly saturated aliphatic hydrocarbons formed during feedstock conversion in the first reactor. This method is most similar in technical essence and achieved technical result and has been selected as the prototype for the claimed invention.


The disadvantage of the method claimed in the prototype invention is its low aromatic hydrocarbon yield and its poor selectivity with respect to alkylbenzenes, in particular xylenes.


Disclosure of the Invention as a Method


The present invention has the principal goal of offering a method of producing AHCC from light aliphatic hydrocarbons and mixtures thereof with oxygenated compounds (C1-C4 aliphatic alcohols) that increases the efficiency of production of aromatic hydrocarbon concentrates and the selectivity with respect to alkylbenzenes, in particular xylenes.


The stated objective is accomplished by separating the gaseous fraction separated from the reaction product, aromatic hydrocarbons, into hydrogen-containing gas and NGL containing olefins. The hydrogen-containing gas is fed to a unit for synthesizing oxygenates (methanol or mixtures thereof with ethers and/or C2-C4 aliphatic alcohols), and the resulting oxygenates are fed to the inputs of the first and second reactors. The olefin-containing NGL are fed to the input of the first reactor. This advantageous characteristic makes it possible to increase the useful yield of aromatic hydrocarbons as a whole and alkylbenzenes produced in aromatic hydrocarbons in particular.


A variant of the invention exists in which oxygenates are synthesized by producing synthesis gas using autothermal reforming technology, with subsequent oxygenate synthesis in a circulating or flow-through scheme. This makes it possible to increase the efficiency of arene concentrate production by involving the oxygenates in the conversion process, increase the alkylbenzene selectivity, and solve the problem of heat delivery to the endothermal reaction of light saturated hydrocarbon aromatization (dehydrocyclization) by parallel conduct of exothermal reactions of oxygenate conversion to aromatic hydrocarbons and alkylation of aromatic hydrocarbons by alcohols in the reaction zone, which permits the conversion of hydrocarbon-oxygenate mixtures to be performed using reactor equipment that does not use elements of heat exchange with the reaction zone in its design.


A variant of the invention exists in which oxygenate synthesis is performed with simultaneous fine purification of discharge hydrogen-containing gas to remove sulfur compounds. This makes it possible to increase the useful product yield through additional purification to remove undesirable impurities containing sulfur compounds, which are catalytic poisons for catalysts of hydrocarbon conversion to synthesis gas and oxygenate synthesis.


A variant of the invention exists in which a temperature of 400-500° C. is maintained in the first reactor and a temperature of 450-520° C. is maintained in the second reactor. This makes it possible to maintain a higher temperature in the second reactor, since the conversion of propane with oxygenates proceeds more efficiently at a temperature at least 15° C. higher than the temperature required for conversion of butanes or 40° C. higher than the temperature required for conversion of butanes with oxygenates.


A variant of the invention exists in which the temperature in the first and second reactors is controlled by the oxygenate flow rate. This makes it possible to control the required temperature by adjusting the oxygenate feed from the oxygenate synthesis unit.


A variant of the invention exists in which the reactor catalyst is a catalyst that contains a mechanical mixture of two zeolites, the first of which is characterized by a silicate modulus SiO2/Al2O3=20 and is pretreated with an aqueous alkali solution and modified by rare-earth oxides in quantities of 0.5-2.0% by wt. of the weight of the first zeolite, while the second is characterized by a silicate modulus SiO2/Al2O3=82, contains residual quantities of sodium oxide 0.04% by wt. of the weight of the second zeolite, and is modified by magnesium oxide in a quantity of 0.5-5.0% by wt. of the weight of the second zeolite, where the zeolites are used in a mass ratio of 1.7/1 to 2.8/1, and the binder contains at least silica and is used in a quantity of 20-25% by wt. of the weight of the catalyst. This makes it possible, in the co-conversion of hydrocarbons and oxygenates, to achieve a higher yield of aromatic hydrocarbons with practically full conversion of the hydrocarbon feedstock, better selectivity with respect to the formation of methylbenzenes in the aromatic hydrocarbon concentrate, in particular xylenes, makes it possible to increase the catalyst's regeneration period by at least double (to 800 h), reduce the hydrocarbon feedstock conversion temperature by at least 15° C., and extend the catalyst's life cycle to 30-50 regenerations.


A variant of the invention exists in which a liquid hydrocarbon and water condenser is installed after the second reactor, serially connected to a three-phase product separator into reaction water, liquid hydrocarbons, and discharge gases. This makes it possible to increase the useful product yield through the possibility of recirculating the discharge gases.


A variant of the invention exists in which a liquid hydrocarbon and water condenser is installed after the first reactor, serially connected to a separator for separating liquid products fed to the three-phase separator. This makes it possible to perform intermediate removal of the liquid fraction and increase the aromatic hydrocarbon yield. Through intermediate separation of aromatic hydrocarbons, favorable conditions are created for their synthesis, which also reduces the recycling of olefin-containing fractions to reactor 1.


A variant of the invention exists in which discharge gases from the three-phase separator undergo stripping with removal of the natural gas liquids containing olefins. This makes it possible to remove 90% of the propane from the discharge gases, as well as ethylene. Deethanization of the NGL (distillation of some of the dissolved methane and ethane) is desirable, since it permits reduction of the circulation of ethane and methane that do not participate in the aromatization process.


A variant of the invention exists in which the benzene and/or benzene-toluene fraction is removed from the aromatic hydrocarbon concentrate and fed to the input of the first and/or second reactor. This makes it possible to increase alkylbenzene production further through recirculation of lower aromatic hydrocarbons, including benzene.


The combination of essential features of the claimed invention is unknown from the technological level for methods with similar purposes, which supports the conclusion that the invention as a method meets the novelty criterion.


Technological Level of the Installation


In another respect, the present invention relates to an installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons, including two serially connected reactors, a first and second with pentasil-based zeolite catalysts, where the reactors differ in the conditions of conversion of hydrocarbons to aromatics, a unit for separating the mixture produced after the reaction zones into liquid and gaseous fractions, and the gaseous fraction output is connected to the inputs of the first and second reactors. Such an installation is described in RF Patent 2277524 of 2006. Said installation is the most similar in technical essence and has been selected as the prototype for the claimed invention as a device.


The disadvantage of the prototype invention is its low efficiency in producing aromatic hydrocarbon concentrate using the claimed installation, and its high concentration of benzene and naphthalenes and low concentration of alkylbenzenes in the resulting aromatic hydrocarbons. In addition, we can note the short period between regenerations of the catalysts used in the reactors, as well as the need to use more complex isothermal reactor equipment (reactors affording nearly isothermal conversion through delivery of heat to the reactor).


Disclosure of the Invention as an Installation


The present invention also has the goal of offering an installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons and mixtures thereof with aliphatic alcohols, including two serially connected reactors, a first and second with pentasil-based zeolite catalysts, where the reactors differ in the conditions of conversion of hydrocarbons to aromatics, a unit for separating the mixture produced after the reaction zones into liquid and gaseous fractions, and the gaseous fraction output is connected to the inputs of the first and second reactors, permitting at least reducing of the aforementioned disadvantage.


To achieve this goal, the unit for separating the mixture into liquid and gaseous fractions contains a module for separating the gaseous fraction into hydrogen-containing gas and the natural gas liquids containing olefins. The installation additionally includes an oxygenate synthesis unit whose input is connected to the hydrogen-containing gas output of the gaseous fraction separation module, while the output of the oxygenate synthesis unit is connected to the inputs of the first and second reactors. This makes it possible to increase the yield of aromatic hydrocarbons as a whole and alkylbenzenes therein.


A variant of the invention exists in which the oxygenate synthesis unit includes a unit for producing synthesis gas, adapted to the production of synthesis gas using autothermal reforming technology. This makes it possible to increase the efficiency of oxygenate production in a circulating or flow-through scheme through intermediate production of synthesis gas with optimal H2/CO and H2/CO2 stoichiometric ratios for subsequent oxygenate synthesis.


A variant of the invention exists in which the oxygenate synthesis unit includes a unit for fine purification of discharge hydrogen-containing gas to remove sulfur compounds. This makes it possible to increase the useful yield of aromatic hydrocarbon concentrate through additional purification to remove undesirable impurities containing sulfur compounds, which are catalytic poisons for catalysts of hydrocarbon conversion to synthesis gas and oxygenate synthesis.


A variant of the invention exists in which the first and second reactors include a catalyst that contains a mechanical mixture of two zeolites, the first of which is characterized by a silicate modulus SiO2/Al2O3=20 and is pretreated with an aqueous alkali solution and modified by rare-earth oxides in quantities of 0.5-2.0% by wt. of the weight of the first zeolite, while the second is characterized by a silicate modulus SiO2/Al2O3=82, contains residual quantities of sodium oxide 0.04% by wt. of the weight of the second zeolite, and is modified by magnesium oxide in a quantity of 0.5-5.0% by wt. of the weight of the second zeolite, where the zeolites are used in a mass ratio of 1.7/1 to 2.8/1, and the binder contains at least silica and is used in a quantity of 20-25% by wt. of the weight of the catalyst. This advantageous characteristic makes it possible to achieve a higher yield of aromatic hydrocarbons with practically full conversion of the hydrocarbons and better selectivity with respect to the formation of alkylbenzenes in the aromatic hydrocarbon concentrate.


A variant of the invention exists in which the installation additionally includes a liquid hydrocarbon and water condenser installed after the first reactor and before the second reactor and a separator for separating the liquid fraction. This makes it possible to perform intermediate removal of the liquid fraction and increase arene production. (Through separation of aromatic hydrocarbons, more favorable conditions are created for their synthesis, which also reduces the recycling of the olefin-containing fraction to reactor 1.)


A variant of the invention exists in which the installation additionally includes a liquid hydrocarbon and water condenser installed after the second reactor, serially connected to a three-phase conversion product separator into reaction water, liquid hydrocarbons, and discharge gases. This makes it possible to extract the liquid hydrocarbon fraction and remove reaction water.


A variant of the invention exists in which the installation additionally includes a module for stripping discharge gases for removal of the natural gas liquids containing olefins, installed after the three-phase separator. This advantageous characteristic makes it possible to remove 90% of the propane from the discharge gases, as well as ethylene. Deethanization of the natural gas liquids (distillation of at least some of the dissolved methane and ethane from the natural gas liquids) is desirable, since it permits reduction of the circulation of ethane and methane contained in the natural gas liquids that do not participate in the process.


A variant of the invention exists in which the installation additionally includes a circulation compressor installed after the three-phase separator. This makes it possible to reduce the costs of separating the discharge gases from the three-phase separator into hydrogen-containing gas and the natural gas liquids.


A variant of the invention exists in which the installation additionally includes a unit for removing the benzene and/or benzene-toluene fraction from the aromatic hydrocarbon concentrate, whose output is connected to the input of the first and/or second reactor. This advantageous characteristic makes it possible to increase alkylbenzene production further through recirculation of lower aromatic hydrocarbons.


The combination of essential features of the claimed invention is unknown from the technological level for devices with similar purposes, which supports the conclusion that the invention as an installation meets the novelty criterion.





BRIEF DESCRIPTION OF DRAWINGS

Other distinguishing features and advantages of the invention clearly follow from the specification which is presented below for illustration purposes and is not restrictive, with references to the attached figures, in which:



FIG. 1 schematically depicts the overall view of the installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons and mixtures thereof with oxygenates according to the invention;



FIG. 2 schematically depicts the steps of the method of producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons and mixtures thereof with oxygenates according to the invention.





Pursuant to FIG. 1, the installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons and mixtures thereof with C1-C4 aliphatic alcohols, including two serially connected reactors, a first reactor 1 and a second reactor 2 with pentasil-based zeolite catalysts, where reactors 1 and 2 differ in the conditions of conversion of hydrocarbons to aromatics, and unit 3 for separating the mixture obtained after the reaction zones into a liquid fraction containing C5+ and water and a gaseous fraction containing H2, C1-C2 and C2-C5, i.e., olefin-containing NGL. The NGL output is connected to the inputs of the first and second reactors.


Unit 3 for separating the mixture obtained at the output of the second reactor into liquid and gaseous fractions contains module 4 for separating the gaseous fraction into hydrogen-containing gas containing mainly hydrogen, methane, and ethane, and NGL containing C2-C5 olefins and paraffins.


The installation additionally includes oxygenate synthesis unit 5, whose input is connected to the hydrogen-containing gas output of gaseous fraction separation module 4, and the output of oxygenate synthesis unit 5 is connected to the input of the first and second reactors. Oxygenate synthesis unit 5 includes synthesis gas unit 6, adapted to the production of synthesis gas by autothermal reforming technology.


Oxygenate synthesis unit 5 also includes unit 7 for fine purification of discharge hydrogen-containing gas to remove sulfur compounds.


The purpose of unit 7 for fine purification to remove sulfur compounds is the chemosorptive or adsorptive purification of hydrogen-containing gas to remove sulfur compounds in order to meet requirements for sulfur content of crude hydrocarbons defined by requirements for prereforming, reforming, and oxygenate synthesis catalysts. In addition, oxygenate synthesis unit 5 includes unit 8 for oxygenate synthesis from synthesis gas by a flow-through and/or circulating scheme.


The preferable method of synthesizing oxygenates in oxygenate synthesis unit 5 is to obtain synthesis gas by autothermal refining technology, with subsequent oxygenate synthesis by a flow-through and/or circulating scheme.


Unit 6 consists of steam-oxygen (autothermal) conversion, prereforming, and heat recovery sections (not shown in FIG. 1). The purpose of the steam-oxygen (autothermal) conversion, pre-reforming, and heat recovery sections is to obtain synthesis gas by heating feedstock, mixing it with superheated steam, stabilizing the composition of the feedstock by adiabatic prereforming (adiabatic steam conversion and destructive hydrogenation of hydrocarbon feedstock), steam-oxygen or steam-air conversion of hydrocarbon feedstock, heat recovery, and steam condensation and dewatering.


Due to the presence of hydrogen in the initial hydrocarbon feedstock (stripped discharge gases), the H/C molar ratio for the initial hydrocarbon feedstock will be ≈4.5 (for methane, H/C=4), which ensures the production of synthesis gas using autothermal reforming technology with a stoichiometric ratio f=(MFH2−MFCO2)/(MFCO+MFCO2)≥2 (MF=“mole fraction”) at a low ratio MFCO2/MFCO≤0.17, which permits production of methyl alcohol with a concentration no less than 94% suitable for conversion to aromatic hydrocarbons without a concentration (distillation) stage. At lower ratios f<2, mixtures of methanol and C2-C3 aliphatic alcohols, as well as mixtures of alcohols with ethers, can be synthesized.


The purpose of unit 8 for oxygenate synthesis from synthesis gas is to produce oxygenates suitable for co-conversion with aliphatic hydrocarbons by a circulating or flow-through scheme. The most suitable method of oxygenate synthesis is oxygenate synthesis by a circulating scheme.


In addition, the ratio 2 will be met either if the mass fraction of carbon in the discharge gases is increased, which makes it possible to obtain methanol, or if the composition of discharge gases is altered while the activity of conversion catalysts declines during their service.


First and second reactors 1 and 2 include the catalyst claimed in the present invention, whose composition is described above.


The installation additionally includes unit 9 installed after first reactor 1 and before second reactor 2, consisting of liquid hydrocarbon and water condenser 10 and liquid fraction separator 11, which is connected to three-phase separator 12.


The installation additionally includes liquid hydrocarbon and water condenser 13 installed after second reactor 2, serially connected to three-phase separator 12 of conversion product to reaction water, liquid hydrocarbons and discharge gases.


The installation additionally includes module 14 installed after three-phase separator 12, designed to stabilize liquid hydrocarbons leaving 12, in which the light aliphatic hydrocarbon fraction (NGL) is distilled from the hydrocarbon condensate in addition to the fraction obtained in unit 4.


The installation additionally includes circulation compressor 14 installed after three-phase separator 12 and before gaseous fraction separation module 4.


The installation additionally includes unit 15 for extracting the benzene and/or benzene-toluene fraction from the aromatic hydrocarbon concentrate, whose output is connected to the input of first reactor 1 and/or second reactor 2.


Reactor 1 is designed for aromatization of a mixture of saturated and unsaturated aliphatic hydrocarbons and oxygenates. It contains at least one hydrocarbon feedstock heater 16, at least one hydrocarbon feedstock mixer 17, and at least one reaction zone 18.


By “reaction zone” here, we mean the entire reactor space in which hydrocarbon conversion occurs, including that which is divided into separate segments. The reactor may be a multi-bed type, for example, with mixing of streams within the reactor. It may have several mixing and feedstock feed zones. The reactor may also be tubular with catalyst contained in the reaction tubes, etc. The conversion feedstock is chosen so that exo- and endothermal reactions proceed efficiently, which affords several aforementioned advantages.


During conversion of hydrocarbons to aromatic hydrocarbon concentrate, fixed-bed reactors with periodic catalyst regeneration or fluidized-bed catalytic reactors with continuous catalyst regeneration are used.


Reactor 2 is designed for aromatization of a mixture of saturated and unsaturated aliphatic hydrocarbons and oxygenates and contains at least one hydrocarbon feedstock heater 19, at least one hydrocarbon feedstock mixer 20, and at least one reaction zone 21.


Unit 3 for separating the conversion products into reaction water, hydrogen-containing gas, stable aromatic hydrocarbon concentrate, and the natural gas liquids contains a three-phase conversion product separator for reaction water, liquid hydrocarbons, and discharge gases, as well as a module for stripping discharge gases from the three-phase separator, which permits extraction of the natural gas liquids containing olefins from the discharge gases. Unit 3 may also contain circulation compressor 14.


To maximize the AHCC yield, the discharge gas stripping module must afford extraction of 90% of the propane from the discharge gases. Deethanization of the natural gas liquids (distillation of at least part of the dissolved methane and ethane from the natural gas liquids) is desirable, since it permits reduction of the circulation of ethane and methane which are contained in the natural gas liquids and do not participate in the process.


EMBODIMENT OF THE INVENTION

Aromatic hydrocarbon production according to the invention proceeds as follows.


Step A1. The natural gas liquids or mixtures thereof with C1-C4 aliphatic alcohols are fed to reactor 1 of the installation. The preferred feedstock is a propane-butane fraction containing 70-80% propane, as well as circulating aliphatic saturated and unsaturated hydrocarbons from unit 3 and oxygenates from unit 5. The hydrocarbons are evaporated and thoroughly mixed. To increase the alkylbenzene content of the produced aromatic hydrocarbon concentrate, the benzene or benzene-toluene fraction, including that containing aliphatic hydrocarbons, may also be fed to reactor 1 (without an extractive distillation stage to remove aliphatic hydrocarbons).


Step A2. A mixture consisting of PBF or NGL hydrocarbon feedstock, recirculating olefin-containing NGL, and oxygenates is converted in reaction zone 1 in the gaseous phase. The unsaturated aliphatic hydrocarbons recycled from unit 3 are nearly completely dehydrocyclized, oxygenate vapors are totally converted, and part of the saturated aliphatic hydrocarbons, both circulating and arriving with the feedstock stream, are converted.


Step A3. The conversion product from reactor 1 is fed to reactor 2, to which oxygenates from unit 5 are also fed. In reactor 2, the incoming mixture from the output of reactor 1, which is thoroughly mixed with oxygenate vapors arriving from unit 5, undergoes gas-phase conversion. To increase the concentration of alkylbenzenes in the produced aromatic hydrocarbon concentrate, the benzene or benzene-toluene fraction, including that containing aliphatic hydrocarbons, may also be fed to reactor 2 (without an extractive distillation stage to remove aliphatic hydrocarbons).


Step A4. Additionally, with the aid of hydrocarbon and liquid condenser 10, where the liquid part of the conversion product (C5+ and reaction water) condenses, and with the aid of separator 11, the gaseous part of the product is fed to reactor 2, while the liquid part of the product is removed and fed immediately to unit 3. The presence of condenser 10 and separator 11 permits reduction of the hydrocarbon circulation because in reactor 2, the extraction of aromatic hydrocarbons creates more favorable conditions for synthesis of aromatic hydrocarbons, which reduces recycling of the olefin-containing fraction to reactor 1.


Step A5. The conversion product from reactor 2 is fed through hydrocarbon and water condenser 13 in a mixture with hydrocarbon condensate from unit 11 (or without mixing) to unit 3. There, it is separated into reaction water to be recycled and unstable hydrocarbon condensate entering separation unit 22, where the latter is separated into a C5+ or C6+ hydrocarbon fraction and C2-C5 natural gas liquids. The discharge gases, with the aid of module 4, are separated into hydrogen-containing gas, which is fed to unit 5 for conversion to oxygenates, and the natural gas liquids containing olefins, which together with the natural gas liquids from separation unit 22 are recycled to reactor 1. The ratio of the circulating natural gas liquids from reactor 2 to the feedstock ranges from 0.3:1 to 1:1, depending on the composition of the feedstock. The ratio of oxygenates to the hydrocarbon feedstock is 1:1-1:4.


Step A6. The temperature at the outlet of each reaction zone in reactors 1 and 2 is controlled by the oxygenate flow rate. The pressure in the reaction zones is 0.5-2.5 MPa. Temperatures are from 400° C. to 520° C. Thermostabilization of the reaction zone of reactor 2 is fully or partially effected by the heat capacity of the conversion feedstock and the presence in the conversion mixture of oxygenates, whose conversion to aromatic hydrocarbons and methylbenzenes or alkylbenzenes releases heat, and paraffins, whose conversion to aromatic hydrocarbons consumes heat.


The oxygenates are distributed between reaction zones 1 and 2 so as to ensure adiabatic heating of the feedstock in the reaction zone of reactor 1 to 400-500° C. and in the second to 450-520° C., respectively, and excess oxygenates from unit 5 can be discarded.


In reactors 1 and 2, endothermal reactions of saturated aliphatic hydrocarbon conversion to aromatic hydrocarbon concentrate, exothermal reactions of oxygenate aromatization, and exothermal reactions of aromatic compound alkylation occur. As a result, adiabatic conditions can be maintained in each reaction zone of reactors 1 and 2, which permits simple reactor equipment to be used, even without the use of additional heat supply/removal from the reaction zone.


The figure additionally designates the following channels:

    • 23: feedstock feed to reactor 1;
    • 24: conversion output from reactor 1;
    • 25: connection of the output of oxygenate synthesis unit 5 to the input of first reactor 1 for oxygenate feed;
    • 26: connection of the output of unit 4 to the input of first reactor 1 for NGL feed;
    • 27: connection of unit 11 to the input of separation unit 3 for water-hydrocarbon condensate feed;
    • 28: connection of the output of unit 9 to the input of second reactor 2 for discharge gas feed;
    • 29: connection of the output of oxygenate synthesis unit 5 to the input of second reactor 2 for oxygenate feed;
    • 30: connection of the output of unit 13 to the input of separation unit 3 for conversion product and condensate feed;
    • 31: removal of reaction water from separation unit 3;
    • 32: boiler-quality water feed to the input of oxygenate synthesis unit 5;
    • 33: oxygen feed to the input of oxygenate synthesis unit 5;
    • 34: output from unit 4 to the connection unit from streams 35 for NGL feed from discharge gases;
    • 35: connection of the output from separation unit 22 to channel 34-26 for arene concentrate stabilization gas feed;
    • 36: output from oxygenate synthesis unit 5 for hydrogen containing gas (HCG) purge feed;
    • 37: condensation water feed to the input of oxygenate synthesis unit 5;
    • 38: HCG feed from unit 4 to unit 5;
    • 39: arene concentrate stabilization gas feed from unit 3 to unit 15.


The sequence of steps is illustrative and permits some operations to be reordered, added, or performed simultaneously without loss of the capability of producing aromatic hydrocarbon concentrate from natural gas liquids.


INDUSTRIAL APPLICABILITY

The claimed installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons may be embodied in practice, and when embodied it affords realization of the claimed purpose, which supports the conclusion that the invention meets the industrial applicability criterion.


In accordance with the claimed invention, calculations of the method of operation of the installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons have been performed with the following process parameters: pressure 0.5-1.5 MPa; temperature according to specification; rate of oxygenate feed to reaction zones W=1 to 2 h−1 (in liquid); rate of aliphatic hydrocarbon feed to reaction zones W=200 to 1500 h−1 (in gas).


According to the technological process modeling data, the process claimed in this invention is highly efficient, permitting production of up to 820 kg of aromatic hydrocarbons from a metric ton of liquefied hydrocarbon gases containing 80% propane, and up to 900 kg of aromatic hydrocarbons from a metric ton of butanes, exceeding the stated parameters for processes for producing aromatic hydrocarbons by catalytic reforming of naphtha (accounting for the recycling of aliphatic hydrocarbons separated from the reformate to the reforming stage, the aromatic hydrocarbon yield is no more than 75% of the feedstock) and by Cyclar® technology (a joint development of BP and UOP), the aromatic hydrocarbon yield is up to 66% from n-butane, and no more than 60% from propane, and the AHCC yield is 53% according to data presented in the prototype specification.


Another distinguishing feature of the process is the increased yield of alkylbenzenes, in particular xylenes, which permits the use of the resulting aromatic hydrocarbon concentrate to produce xylenes, in particular paraxylene. The product made using Cyclar® technology contains 20-23% xylenes, and a similar concentration is claimed in the specification of the invention prototype.


The Advantages of the Technological Solution are:

    • high aromatic hydrocarbon yield, 82-90%;
    • elevated content of alkylbenzenes, including xylenes, in the aromatic hydrocarbon concentrate, ˜40%;
    • simplified reactor equipment design due to the offsetting thermal effect in conversion of hydrocarbon mixtures. This solution permits use of a simple reactor for conversion, avoidance of local superheating of the catalyst, and stabilization of the catalyst's productivity across its bed, which reduces the volume of the reaction zone;
    • possibility of separating high-added-value byproducts, for example paraxylene, from the synthesis products;
    • low aliphatic hydrocarbon content in the aromatic hydrocarbon concentrate, ˜1% for conversion of propane-butane fractions;
    • possibility of involving benzene fractions (including mixtures containing aliphatics) in the process for further conversion to alkylbenzenes;
    • possibility of recycling discharge hydrogen-containing gas from other processes, including processes of refining aromatic hydrocarbon concentrate to marketable aromatic hydrocarbons;
    • possibility of adjusting the oxygenate feed to reactors 1 and 2 to maintain the required process temperature;
    • possibility of using oxygenates for co-conversion;
    • possibility of reducing the propane conversion temperature by at least 15° C.;
    • reduction in naphthalene content in the conversion product compared to the option of converting C3-C4 paraffins alone;
    • two- to three-fold increase in time between regenerations compared to conversion of C3-C5 paraffins alone.


Additional technical results are:

    • recycling of discharge gases from synthesis of arenes;
    • stabilization of the yield of aromatic hydrocarbon concentrate in case of variation in feedstock composition and catalyst deactivation;
    • increased aromatic hydrocarbon concentrate yield;
    • simplification of reactor equipment design;
    • reduction in energy costs due to reduction in circulation of aliphatic hydrocarbons from feedstock and conversion product;
    • improved efficiency in the use of discharge gases from neighboring and main processes to increase the marketable product yield;
    • diversification of feedstock: transition from expensive feedstock (naphtha) to cheaper feedstock (NGL).


Thus, this invention achieves its stated objective of improving the efficiency of aromatic hydrocarbon concentrate production and increasing selectivity with respect to alkylbenzenes, in particular xylenes.


The distinguishing features and advantages of the invention also follow from the tables, which are presented below for illustration and are not restrictive, in which:

    • Table 1 tabulates a comparison of product yields;
    • Table 2 tabulates the material and component balance of the claimed method;
    • Table 3 tabulates data on co-aromatization of propane, n-butane, and a mixture of propane, propene, butanes, butenes, and oxygenates (methanol and isopropanol).









TABLE 1







Product Yield Comparison















Present





Prototype,
Patent,




Propane
Conversion
Conversion




Conversion,
of
of PBF



Naphtha
Cyclar ®
Butane
(80%


Technology
Reforming
Process
Fraction
Propane)














Benzene, %
6
27
14
3


Toluene, %
21
43
45
35


Xylenes and
20
20
23
40


ethylbenzene, %


Higher aromatics,
20
9
11
21


%


Nonaromatic
33
1
7
1


compounds, %


AHCC yield, %
78
60
53
82
















TABLE 2







Material and Component Balance


Without Use of Additional Unit 11




















Stream/Unit


Total,
CO







Ar



Designation
Name of
From-To/
×1000 tons/
and





C5+

and



on Diagram
Stream
Remark
yr
CO2
H2O
Methanol
H2
C1-C2
C3-C4
Aliphatics
Aromatics
N2
O2










Reactor 1




















Reactor 1
Entered reactor

1025
1
20
279
0
38
683
0
3
0
0



1, total:














23
PBF
Feed-
500
0
0
0
0
10
490
0
0
0
0




stock













25
Oxygenates
From
299
1
19
279
0
0
0
0
0
0
0




unit 5













26
NGL
From
225
0
1
0
0
28
193
0
3
0
0




unit 3


















Including:




















34
NGL from
From
173
0
1
0
0
21
148
0
3
0
0



discharge gases
unit 3













35
Arene
From
53
0
0
0
0
7
45
0
0
0
0



concentrate
unit 3














stabilization















gases














Reactor 1
Received from

1025
1
177
0
15
120
497
3
213
0
0



reactor 1, total














24
Conversion
To unit
1025
1
177
0
15
120
497
3
213
0
0



product from
9














reactor 1



















Unit 9




















Unit 9
Entered unit 5,

1025
1
177
0
15
120
497
3
213
0
0



total:














24
Conversion
To unit
1025
1
177
0
15
120
497
3
213
0
0



product from
9














reactor 1














Unit 9
Received from

1025
1
177
0
15
120
497
3
213
0
0



unit 5, total:














27
Water-hydrocarbon
To unit
455
0
175
0
0
3
68
2
207
0
0



(HC)
3














condensate














28
Discharge gases
To unit
570
1
2
0
15
117
428
1
6
0
0




2


















Reactor 2




















Reactor 2
Entered reactor

814
2
17
229
15
117
428
1
6
0
0



2, total:














29
Oxygenates
From
245
1
15
229
0
0
0
0
0
0
0




unit 5













28
Discharge gases
From
570
1
2
0
15
117
428
1
6
0
0




unit 9













Reactor 2
Received from

814
2
145
0
22
251
219
2
173
0
0



reactor 2, total:















Conversion
To unit
814
2
145
0
22
251
219
2
173
0
0



product from
13














reactor 2



















Unit 3




















Unit 3
Entered unit 3,

1270
2
320
0
22
254
288
4
380
0
0



total














27
Water-hydrocarbon
From
455
0
175
0
0
3
68
2
207
0
0



condensate
unit 9













30
Conversion
From
814
2
145
0
22
251
219
2
173
0
0



product and
unit 13














condensate















thereof














Unit 3
Received from

1270
2
320
0
22
254
288
4
380
0
0



unit 3, total














96
Hydrogen-
To unit
313
1
0
0
22
226
63
0
0
0
0



containing gas
5













34 + 35
NGL
To
225
0
1
0
0
28
193
0
3
0
0


total

reactor 1













31
Reaction water
Product
320
0
320
0
0
0
0
0
0
0
0


39
Stabilized arene
Product
411
0
0
0
0
0
30
3
377
0
0



concentrate
or to















unit 15


















Unit 5




















Unit 5
Entered unit 5,

786
1
210
0
22
226
63
0
0
2
261



total














32
Boiler-quality
Feed-
210
0
210
0
0
0
0
0
0
0
0



water
stock













33
Oxygen
Feed-
263
0
0
0
0
0
0
0
0
2
261




stock













38
HCG
From
313
1
0
0
22
226
63
0
0
0
0




unit 3













Unit 5
Received from

787
29
197
510
11
39
0
0
0
2
0



unit 4, total














37
Condensation
Product
163
0
163
0
0
0
0
0
0
0
0



water














25 + 29 total
Oxygenates
To
544
1
34
508
0
1
0
0
0
0
0




reactors















1 and 2













36
HCG discharge
Product
80
28
0
2
11
38
0
0
0
2
0







Installation As a Whole


















Total feedstock
973
0
210
0
0
10
490
0
0
2
261




















23
PBF
Feed-
500
0
0
0
0
10
490
0
0
0
0




stock













32
Boiler-quality
Feed-
210
0
210
0
0
0
0
0
0
0
0



water
stock













33
Oxygen 99%
Feed-
263
0
0
0
0
0
0
0
0
2
261




stock





























Total products
973
28
482
2
11
38
30
3
377
2
0




















31
Reaction water
Product
320
0
320
0
0
0
0
0
0
0
0


37
Condensation
Product
163
0
163
0
0
0
0
0
0
0
0



water














39
Stabilized arene
Product
411
0
0
0
0
0
30
3
377
0
0



concentrate














36
HCG discharge
Product
80
28
0
2
11
38
0
0
0
2
0
















TABLE 3







Data on Co-Aromatization of Propane, n-Butane,


and a Mixture of Saturated and Unsaturated C3-C4 Hydrocarbons


and Oxygenates









Example Number











1
2
3














Temperature, ° C.
490
515   
450


Pressure, atm.
8
8  
6


Volumetric gas flow rate, hr−1
300
300   
1500







TAKEN (ratio, % by wt.)










Propane

76.9 



n-Butane
83.3


Methyl alcohol
16.7
23.1 


Isopropyl alcohol


26.5


NGL (C3-C4 fraction, 50% olefins by wt.)


73.5


TOTAL:
100
100   
100


Oxygenate conversion
100
100   
100


AHCC yield per feedstock pass (per HC part
44.2*
36*  
78.2*


of feedstock)*, % by wt.







C1-C4 Hydrocarbon Composition of Gas, % by wt.










CH4
11.1
28.1 
6.9


C2H6
9.7
12.4 
21.0


C2H4
5.6
6.1
3.3


C3H8
38.7
24.1 
32.4


C3H6
5.1
8  
4.5


i-C4H10
3.8
7.9
10.6


n-C4H10
25.1
12.6 
18.5


C4H8
0.9
0.8
2.7


Total, % by wt.
100
100   
100







AHCC Composition, % by wt.










Aliphatics
1.3
0.9
8.2


Benzene
2.1
3.6
6.1


Toluene
32.4
36.3 
32.2


Xylenes + ethylbenzene
42.1
38.7 
36.4


C9+ alkylaromatics
22.1
20.4 
18.1


Total aromatic HC in AHCC
98.7
99  
92.8








Claims
  • 1. An installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons, the installation comprising: a serially connected first reactor and second reactor;a pentasil-based zeolite catalyst;the reactors configured whereby each reactor has different conditions of conversion of hydrocarbons to aromatics;a separation unit connected to receive an output mixture from the second reactor, whereby the separation unit separates the output mixture into a liquid fraction and gaseous fraction, and wherein the separation unit comprises a liquid fraction output and a gaseous fraction output;a line connecting the gaseous fraction output to the first and second reactors;the separation unit comprising a module for separating the gaseous fraction into hydrogen-containing gas and the natural gas liquids containing olefins, and wherein the separation unit comprises a hydrogen-containing gas output;a line connecting the hydrogen-containing output to an oxygenate synthesis unit to the first reactor, the second reactor or both the first and the second reactors;wherein the catalyst in the first reactor comprises a mechanical mixture of two zeolites, the first of which is characterized by a silicate modulus SiO2/Al2O3=20 and is pretreated with an aqueous alkali solution and modified by rare-earth oxides in quantities of 0.5-2.0% by wt. of the weight of the first zeolite, while the second is characterized by a silicate modulus SiO2/Al2O3=82, contains residual quantities of sodium oxide 0.04% by wt. of the weight of the second zeolite, and is modified by magnesium oxide in a quantity of 0.5-5.0% by wt. of the weight of the second zeolite, where the zeolites are used in a mass ratio of 1.7/1 to 2.8/1, and the binder contains at least silica and is used in a quantity of 20-25% by wt. of the weight of the catalyst.
  • 2. An installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons, the installation comprising: a serially connected first reactor and second reactor;a pentasil-based zeolite catalyst;the reactors configured whereby each reactor has different conditions of conversion of hydrocarbons to aromatics;a separation unit connected to receive an output mixture from the second reactor, whereby the separation unit separates the output mixture into a liquid fraction and gaseous fraction, and wherein the separation unit comprises a liquid fraction output and a gaseous fraction output;a line connecting the gaseous fraction output to the first and second reactors;the separation unit comprising a module for separating the gaseous fraction into hydrogen-containing gas and the natural gas liquids containing olefins, and wherein the separation unit comprises a hydrogen-containing gas output;a line connecting the hydrogen-containing output to an oxygenate synthesis unit to the first reactor, the second reactor or both the first and the second reactors; wherein the catalyst in the second reactor comprises a mechanical mixture of two zeolites, the first of which is characterized by a silicate modulus SiO2/Al2O3=20 and is pretreated with an aqueous alkali solution and modified by rare-earth oxides in quantities of 0.5-2.0% by wt. of the weight of the first zeolite, while the second is characterized by a silicate modulus SiO2/Al2O3=82, contains residual quantities of sodium oxide 0.04% by wt. of the weight of the second zeolite, and is modified by magnesium oxide in a quantity of 0.5-5.0% by wt. of the weight of the second zeolite, where the zeolites are used in a mass ratio of 1.7/1 to 2.8/1, and the binder contains at least silica and is used in a quantity of 20-25% by wt. of the weight of the catalyst.
  • 3. An installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons, the installation comprising: a serially connected first reactor and second reactor;a pentasil-based zeolite catalyst;the reactors configured whereby each reactor has different conditions of conversion of hydrocarbons to aromatics;a separation unit connected to receive an output mixture from the second reactor, whereby the separation unit separates the output mixture into a liquid fraction and gaseous fraction, and wherein the separation unit comprises a liquid fraction output and a gaseous fraction output;a line connecting the gaseous fraction output to the first and second reactors;the separation unit comprising a module for separating the gaseous fraction into hydrogen-containing gas and the natural gas liquids containing olefins, and wherein the separation unit comprises a hydrogen-containing gas output;a line connecting the hydrogen-containing output to an oxygenate synthesis unit to the first reactor, the second reactor or both the first and the second reactors; wherein the installation comprises a liquid hydrocarbon and water condenser installed after the first reactor and before the second reactor and a separator for separating the liquid fraction.
  • 4. An installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons, the installation comprising: a serially connected first reactor and second reactor;a pentasil-based zeolite catalyst;the reactors configured whereby each reactor has different conditions of conversion of hydrocarbons to aromatics;a separation unit connected to receive an output mixture from the second reactor, whereby the separation unit separates the output mixture into a liquid fraction and gaseous fraction, and wherein the separation unit comprises a liquid fraction output and a gaseous fraction output;a line connecting the gaseous fraction output to the first and second reactors;the separation unit comprising a module for separating the gaseous fraction into hydrogen-containing gas and the natural gas liquids containing olefins, and wherein the separation unit comprises a hydrogen-containing gas output;a line connecting the hydrogen-containing output to an oxygenate synthesis unit to the first reactor, the second reactor or both the first and the second reactors; wherein the installation comprises a liquid hydrocarbon and water condenser installed after the second reactor, serially connected to a three-phase conversion product separator into reaction water, liquid hydrocarbons, and discharge gases.
  • 5. An installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons, the installation comprising: a serially connected first reactor and second reactor;a pentasil-based zeolite catalyst;the reactors configured whereby each reactor has different conditions of conversion of hydrocarbons to aromatics;a separation unit connected to receive an output mixture from the second reactor, whereby the separation unit separates the output mixture into a liquid fraction and gaseous fraction, and wherein the separation unit comprises a liquid fraction output and a gaseous fraction output;a line connecting the gaseous fraction output to the first and second reactors;the separation unit comprising a module for separating the gaseous fraction into hydrogen-containing gas and the natural gas liquids containing olefins, and wherein the separation unit comprises a hydrogen-containing gas output;a line connecting the hydrogen-containing output to an oxygenate synthesis unit to the first reactor, the second reactor or both the first and the second reactors; wherein the installation comprises a module for stripping discharge gas to recover the natural gas liquids containing olefins installed after the three-phase separator and a module for stabilizing hydrocarbon condensate to remove the natural gas liquids from the condensate.
  • 6. An installation for producing aromatic hydrocarbon concentrate from light aliphatic hydrocarbons, the installation comprising: a serially connected first reactor and second reactor;a pentasil-based zeolite catalyst;the reactors configured whereby each reactor has different conditions of conversion of hydrocarbons to aromatics;a separation unit connected to receive an output mixture from the second reactor, whereby the separation unit separates the output mixture into a liquid fraction and gaseous fraction, and wherein the separation unit comprises a liquid fraction output and a gaseous fraction output;a line connecting the gaseous fraction output to the first and second reactors;the separation unit comprising a module for separating the gaseous fraction into hydrogen-containing gas and the natural gas liquids containing olefins, and wherein the separation unit comprises a hydrogen-containing gas output;a line connecting the hydrogen-containing output to an oxygenate synthesis unit to the first reactor, the second reactor or both the first and the second reactors; wherein the installation comprises a circulation compressor installed after the three-phase separator.
Priority Claims (1)
Number Date Country Kind
2014111985 Mar 2014 RU national
PCT Information
Filing Document Filing Date Country Kind
PCT/RU2015/000171 3/25/2015 WO 00
Publishing Document Publishing Date Country Kind
WO2015/147700 10/1/2015 WO A
US Referenced Citations (80)
Number Name Date Kind
3702886 Argauer Nov 1972 A
3756942 Cattanach Sep 1973 A
3911041 Kaeding et al. Oct 1975 A
3941871 Dwyer et al. Mar 1976 A
3957621 Bonacci May 1976 A
4159282 Olson et al. Jun 1979 A
4211640 Garwood et al. Jul 1980 A
4227992 Garwood et al. Oct 1980 A
4356338 Young Oct 1982 A
4456527 Buss et al. Jun 1984 A
4463204 Liu Jul 1984 A
4465886 Rodewald Aug 1984 A
4499314 Seddon et al. Feb 1985 A
4523049 Jones et al. Jun 1985 A
4554260 Pieters et al. Nov 1985 A
4590321 Chu May 1986 A
4720602 Chu Jan 1988 A
4853202 Kuznicki Aug 1989 A
4899011 Chu et al. Feb 1990 A
4963337 Zones Oct 1990 A
5030782 Harandi Jul 1991 A
5108579 Casci Apr 1992 A
5173461 Absil et al. Dec 1992 A
5178748 Casci et al. Jan 1993 A
5306411 Mazanec et al. Apr 1994 A
5321183 Chang et al. Jun 1994 A
5362697 Fung et al. Nov 1994 A
5365003 Chang et al. Nov 1994 A
5453554 Cheng et al. Sep 1995 A
5498814 Chang et al. Mar 1996 A
5516736 Chang et al. May 1996 A
5536894 Degnan et al. Jul 1996 A
5557024 Cheng et al. Sep 1996 A
5935897 Trubenbach et al. Aug 1999 A
5993642 Mohr et al. Nov 1999 A
6046372 Brown et al. Apr 2000 A
6063724 Resasco et al. May 2000 A
6096193 Resasco et al. Aug 2000 A
6143166 Nacamuli Nov 2000 A
6413898 Faber et al. Jul 2002 B1
6423879 Brown et al. Jul 2002 B1
6504072 Brown et al. Jan 2003 B1
6635792 Choi et al. Oct 2003 B2
6906232 Levin et al. Jun 2005 B2
6995111 Levin et al. Feb 2006 B2
7026263 Le Van Mao Apr 2006 B2
7078578 Janssens et al. Jul 2006 B2
7122492 Ou et al. Oct 2006 B2
7122493 Ou et al. Oct 2006 B2
7164052 Carati et al. Jan 2007 B2
7208442 Xu et al. Apr 2007 B2
7419930 Carati et al. Sep 2008 B2
7700816 Xu et al. Apr 2010 B2
7923399 Long et al. Apr 2011 B2
8226740 Chaumonnot et al. Jul 2012 B2
8338655 Chang Dec 2012 B2
9040003 Andersen et al. May 2015 B2
20040192990 Choudhary Sep 2004 A1
20080027255 Blessing et al. Jan 2008 A1
20080051617 Sangar Feb 2008 A1
20080300434 Cortright Dec 2008 A1
20080300435 Cortright Dec 2008 A1
20090288990 Xie et al. Nov 2009 A1
20100145127 Xie et al. Jun 2010 A1
20100185033 Karim Jul 2010 A1
20110112347 Van Den Berg May 2011 A1
20110118518 Nesterenko May 2011 A1
20130066126 Jana Mar 2013 A1
20130317269 Nesterenko et al. Nov 2013 A1
20140018592 Chen et al. Jan 2014 A1
20140058180 Klingelhofer et al. Feb 2014 A1
20140100404 Narula et al. Apr 2014 A1
20140256010 Narula et al. Sep 2014 A1
20140273146 Narula et al. Sep 2014 A1
20140322781 Narula et al. Oct 2014 A1
20170001922 Lishchiner et al. Jan 2017 A1
20170007992 Lishchiner et al. Jan 2017 A1
20170145317 Lischiner et al. May 2017 A1
20170233311 Vladislavovich et al. Aug 2017 A1
20190100477 Lishchiner et al. Apr 2019 A1
Foreign Referenced Citations (44)
Number Date Country
103117 Jul 1982 EP
2160161 Jun 2005 RU
2284343 Jun 2005 RU
2289477 Dec 2006 RU
2293056 Feb 2007 RU
2294799 Mar 2007 RU
2362760 Feb 2008 RU
2320631 Mar 2008 RU
2323777 May 2008 RU
2333033 Sep 2008 RU
2349567 Mar 2009 RU
2349568 Mar 2009 RU
2350591 Mar 2009 RU
2350592 Mar 2009 RU
2354638 May 2009 RU
2354639 May 2009 RU
2433863 Apr 2010 RU
2391135 Jun 2010 RU
2429910 Jul 2010 RU
2009101606 Jul 2010 RU
2429910 Sep 2011 RU
2440189 Jan 2012 RU
2477656 Feb 2012 RU
2446135 Mar 2012 RU
2010135608 Mar 2012 RU
2454388 Jun 2012 RU
2458898 Aug 2012 RU
2466976 Nov 2012 RU
2477656 Mar 2013 RU
2495017 Oct 2013 RU
2544017 Jan 2014 RU
2544241 Jan 2014 RU
2509759 Mar 2014 RU
2550354 Mar 2014 RU
2518091 Jun 2014 RU
2558955 Aug 2014 RU
2549571 Apr 2015 RU
WO 1996016004 May 1996 WO
WO 2008109877 Sep 2008 WO
WO 2015115932 Aug 2015 WO
WO 2015112056 Sep 2015 WO
WO 2015147700 Oct 2015 WO
WO 2016024883 Feb 2016 WO
WO 2017155424 Sep 2017 WO
Non-Patent Literature Citations (4)
Entry
International Search report PCT/RU2015/000171.
Translation of International Search report PCT/RU 2015/000171.
PCT/RU2015/000171 Opinion.
International Application Status Report (IAS).
Related Publications (1)
Number Date Country
20170145317 A1 May 2017 US