WO2016108731A1

WO2016108731A1 - Method of complex extraction of valuable impurities from helium-rich hydrocarbon natural gas with high nitrogen content

Info

Publication number: WO2016108731A1
Application number: PCT/RU2015/000843
Authority: WO
Inventors: Igor Anatol'evich MNUSHKIN
Original assignee: Mnushkin Igor Anatol Evich
Priority date: 2014-12-29
Filing date: 2015-12-02
Publication date: 2016-07-07
Also published as: RU2597081C2; RU2014153653A

Abstract

Invention refers to technique of valuable components additional extraction from natural carbon gas and can be applied at gas processing factories. Method of complex extraction of valuable impurities from helium-rich natural carbon gas with high nitrogen content includes following stages: first level of purification to clean feed flow of natural hydrocarbon gas from mechanical impurities and dropping liquid, second level to purify first purified hydrocarbon gas flow from hydrogen sulphide, carbon dioxide and methanol, regeneration of saturated adsorbent, water stripping of methanol, hydrogen sulphide, carbon dioxide, compression and drying of low-pressure acid gas, third level of purification and drying of second flow of purified hydrocarbon gas from mercury compounds, low temperature separation of third flow of dried and purified hydrocarbon gas, cooling and expansion of deethanizated gas with its partial condensation in "cold box", cryogen additional treatment of liquid helium subproduct to extract nitrogen, oxygen, argon and neon impurities, cryogen helium separation, BFLH adsorbtion treatment, gas fractionation of purified BFLH, commercial fuel gas treatment, storage of liquid nitrogen and helium in Dewar vessels in product park.

Description

METHOD OF COMPLEX EXTRACTION OF VALUABLE IMPURITIES FROM HELIUM-RICH HYDROCARBON NATURAL GAS

WITH HIGH NITROGEN CONTENT

TECHNICAL FIELD Invention refers to technique of additional extraction of valuable components from natural hydrocarbon gas and can be applied at gas processing factories.

Methane is main component of natural hydrocarbon gas, which is widely used as gas fuel in industry and household. However, along with methane natural hydrocarbon gas includes many impure organic and inorganic components: ethane and heavier hydrocarbons, water, hydrogen sulphide, mercaptan, carbon dioxide, helium, nitrogen. Peculiarity of impure components is that from one side, these components decrease heat value of natural hydrocarbon gas as fuel gas and increase transportation cost as natural hydrocarbon gas transportation ballast and on the other side, all listed components except water are valuable raw materials for chemical and petrochemical industries and can be used for further production of polymer, automobile transport fuel, sulfur, methanol, inert gas, nuclear fuel, ammonia, nitric acid and other products.

PREVIOUS TECHNICAL KNOWLEDGE

One of the main impure components in natural hydrocarbon gas is nitrogen, its content in natural hydrocarbon gas can reach 10 % vol., for example, in large Kovykta field nitrogen content in natural hydrocarbon gas is 7.2 % vol. Along with it gas from Kovykta field has up to 0.4 % vol. of helium, that allows defining this gas as rather rich of helium natural gas. Demand for helium in world technology increases intensively. By "Gazprom VNIIGAZ" LLC estimation world consumption of helium can reach 238-312 mln m³ to 2030, while expected production of helium will be 213-238 mln m³, so expected level of helium production deficit is 25-74 mln m³. 100 % of Russian helium is manufactured at the only one helium plant "Gazprom dobycha Orenburg" LLC in amount of 9 mln.m³/year from gas condensate with volume of helium only 0.055 % (Kontorovich A. E., Korzhubaev A. G., Eder L. V. Feedstock base and helium production development prospects in Russia and worldwide (oil-and-gas geology Institute SB RAS) // 29.10.2007: http://www.geoinform.ru/). With deep processing of natural hydrocarbon gases at Kovykta and other fields in Eastern Siberia we can remove up to 60 mln m³/year of helium and it can be highly profitable article of Russian export as well as allows compensating global deficit of helium. Thus, during completion of huge gas fields it is desirable to make deep processing of hydrocarbon natural gas with simultaneous extraction of wide range of impurities, but analysis of patent and technical literature has shown, that existing technical solutions work, mainly, at extraction of one typical impurity or group of similar impurities.

Known nitrogen de-aeration method from natural gas includes: a) absorption of natural gas hydrocarbon component by straight-run naphtha, essentially consisting of C₅-C₈ paraffin, in device for unabsorbed nitrogen absorption and discharge; b) hydrocarbon component desorption from straight-run naphtha in stripper by bottom temperature in range 150-200 °C; c) recirculation of recovered after desorption straight-run naphtha to stage (a); d) supply of desorbed hydrocarbon component to distribution system (Nitrogen de-aeration method from natural gas: pat. 2185226 Russian Federation. No. 2000103939/12; appl. dated 06.11.2012; publ. on 20.07.2002). This method has the following disadvantages:

• discharge of extracted nitrogen to atmosphere and loss of this raw material perspective for further processing;

· absorption of all hydrogen part of natural gas by naphtha with following gas desorption leads to economically inefficient power consumption for desorption process and big naphtha flow in the system;

• provided only natural gas pretreatment for carbon dioxide impurity removal by membrane method, but other impurities, such as hydrogen, are not removed, thus resulting eventually in environment contamination.

Outstanding method of ethane fraction extraction by low-temperature rectification, which includes hydrocarbon gas supply in rectification tower for deethanization, low temperature condensation of vapors from heat exchanger and their supply to reflux tank, partial drain of condensed fluid to rectification tower for reflux and discharge of commercial products (Method of ethane fraction extraction: pat. 2459160 Rus. Federation No. 2010136352/20; appl. 30.08.2010; publ. on 20.08.2012). This method has the following disadvantages:

• necessity for preliminary extraction of C₂ fraction and higher from source natural gas;

• deethanization of raw material at low pressure (1.01 MPa instead of 2.63 MPa) needs to keep freezing temperature (minus 27.6 °C), provided by expensive propane refrigeration instead of positive temperature (6.8 °C), provided by cheap cooled water, that gives actual power consumption cost increase despite to some decline in calorific content; • methane fraction, produced after preliminary extraction of C₂ fraction and higher from source natural gas and contaminated by nitrogen, H₂S, carbon dioxide, and other impurities, which decline calorific content of methane fraction as fuel.

There is also method of helium production where helium is extracted from helium-, nitrogen- and methane gas fraction, thus helium-rich fraction at least is condensed and divide into helium-rich gas fraction (a) and helium depleted liquid fraction (b). Then, helium-rich gas fraction (a) is condensed till helium content in resulted gas fraction reaches 90 %, preferably, or at least 95 %, and particularly, at least 98 %. Then, helium depleted liquid fraction (b) is extended, vaporized till, at least, 70 % of helium content, preferably, at least 85 % of helium content passes into gaseous state and then it is divided into helium-rich gas fraction and helium depleted liquid fraction. After that helium-rich gas fraction is heated and added into source helium content gas (Method of helium production: pat. 2486131 Russian Federation. No. 2009104099/05; appl. dated 06.1 1.2012; publ. on 27.06.2013). This method has the following disadvantages:

· need for preliminary production of concentrated source helium-rich fraction, which contains up to 10 % of helium;

• helium-rich fraction return in source helium-rich fraction is unreasonable, because this fraction was formed from fraction (b) and was condensed before, so it is more reasonable to direct it in gas fraction (a).

There is also method to extract helium from natural gas with helium concentrate production and further low temperature separation and impurities absorption purification, it is defined by conversion of natural gas flow and getting synthetic gas, catalytic synthesis of products, which then condensed with helium concentrate separation (Method of helium extraction from natural gas: pat. 2478569 Russian Federation. No. 2011 146306/05; appl. dated 16.11.201 1 ; publ. on 10.04.2013). This method has the following disadvantages:

• total conversion of source natural hydrocarbon gas into synthesis gas with further production of methanol and loss of hydrocarbon gas as fuel;

• to produce helium concentrate off-gas flow after CO conversion supply to carbon dioxide absorption cleaning in absorber with solutions of potassium and monoethanol amine; and/or pressure swing adsorption in adsorber and obtain rich hydrogen gas, and then finally extract hydrogen in short circuit activated carbon adsorption or zeolite adsorbent, however during hydrogen adsorption activity of examined adsorbents is very low, at the level 100 nsnvVgr (0.001 gr/gr or 0.1 % wt.)over zeolites (Big Oil and Gas Encyclopedia. Hydrogen adsorption isotherms over synthetic zeolites, http://www.ngpedia.ru/) and 1-2 kg/m³ (0.002 gr/gr or 0.2 % wt.) activated carbon adsorption (Tarasov B. P., Goldshleger N. F. Carbon nanostructure hydrogen sorption // International Scientific Journal for Alternative Energy and Ecology. 2002.

No. 3. P. 20-38). It makes hydrogen adsorption from helium concentrate with high content of hydrogen economically unviable (common adsorbents activity (adsorption capacity) of other, quite easily, sorbing agents 10 times higher and equal to 10-15 % vol. in technological flows cleaning process, in industrial conditions).

There is also known method of natural gas flow treatment including following stages: (a) supply of partially condensed feed stream in first unit for gas and liquid separation at pressure higher 20 bars and lower than 40 bars;

(b) feed stream separation in first unit to divide gas and liquid into gaseous flow and liquid flow;

(c) expansion of liquid flow obtained at stage (b), and supply it in second unit for gas and liquid separation;

(d) split of following gaseous flow, at least, in to two subflows, and what is more, straight after separation subflows have the same composition and phase;

(e) expansion of first subflow, from (d) stage, resulting, at least into partially condensed first subflow and supply this partially condensed first subflow in second unit for gas and liquid separation;

(f) cooling of second subflow collected at (d) stage by cold flow, and get at least, partially condensed second subflow and supply this at least partially condensed second subflow to second unit to separate gas and liquid;

(g) drain gaseous flow from second unit to separate gas and liquid;

(h) drain liquid flow from second unit to separate gas and liquid (Method and device for hydrocarbons flow treatmen pat. 2460022 Russian Federation. No. 2009119469/06; appl. dated 23.10.2007, publ. 27.08.2012). This method has the following disadvantages:

• low production effectiveness due to quite complicated scheme of combined operation of first and second units, which is limited to feed stream separation after first unit into four flows, three of which are returned into second unit as a feed;

• in this example it is not absolutely clear how second unit should operate, because flow 100 coming from bottom of second unit can not be in vapour-liquid state and flow 110 in vapour phase (table 1);

• given calculations do not convince in opportunity of invention practical realization, because parameters characterizing process realization do not comply with process indicators. DISCLOSURE OF INVENTION

At invention development the aim was to extract valuable impurities from natural helium- rich hydrocarbon gas, which has high nitrogen content and preliminary gas drying and purification and extract, in sequential order (by boiling temperatures): pentan-hexane hydrocarbon fraction, butane fraction, propane fraction, ethane, commercial fuel gas, liquid nitrogen and high purity liquid helium by using effective techniques of low-temperature, rectification, absorption, gravity separation and filtration and minimization of power consumption for process realization due to effective usage of heat recuperation of heat exchange between hot and cold flows and application of turbodetanders and heat pumps.

The set task can be solved on the account that complex method of valuable impurities extraction from natural helium-rich hydrocarbon gas, which has high content of nitrogen includes following stages:

(a) first level purification treats natural hydrocarbon gas feed stream from mechanical impurities and dropping liquid by gravity separation resulting in first flow of purified hydrocarbon gas;

(b) second level purification treats first purified hydrocarbon flow from H₂S, carbon dioxide and methanol by absorption in high pressure absorbing column with regenerated absorbent and obtains purified hydrocarbon gas second flow and rich absorbent flow;

(c) rich absorbent flow regeneration in expansion units with separators and gradually rising absorbent' s temperature and pressure decrease in expansion unit and then in low pressure regenerating tower with production of regenerated absorbent returned after recuperating heaters to stage (b), and acid expansion gas. Acid gas is purified from methanol and partially from H₂S and carbon dioxide in washing tower by process water with production of acid water and low- pressure acid gas and expansion gas is supplied into fuel network;

(d) acid water from stage (c) is stripped of methanol, H₂S and carbon dioxide with production of stripped process water returned to stage (c) and stripped gases, further subject to cooling, partial condensation and separation with production of technological methanol and low- pressure stripped gases; and then technological methanol subject to additional treatment with production of commercial high-purity methanol:

(e) compression and drying of low-pressure acid gases from stage (c) with net part of low-pressure stripped gases from stage (d) in first adsorbers system with adsorbent stationary layer with usage of adsorbent on the basis of zeolite, with production of compressed and dried ballast gas and its further pumping in high-purity methanol and technological methanol mixes from stage (d) in commercial fuel gas providing necessary commercial fuel gas calorific heat, remain part of low-pressure stripped gases from stage (d) is supplied to the flare, or low-pressure acid gas from stage (c) in mix with total volume of low-pressure stripped gases from stage (d) are directed as feed stream to gas chemical facility;

(f) third level of drying and purification of second purified hydrocarbon flow from stage (b) from mercury compounds in second adsorbers system with stationary adsorber layer uses adsorbent on the basis of aluminum oxide and zeolite KA with pore size 3A in gas purifying adsorbers from mercury - adsorbent on the basis of aluminum oxide, copper promoted with production of third flow of dried and purified hydrocarbon gas, which partially used as desorbed agent in adsorbent regeneration;

(g) separation of low-temperature dried and purified third hydrocarbon gas flow from stage (f) in first rectification column with production of deethanizated gas from the column top and first residue, containing C₂ hydrocarbons and higher in column bottom;

(h) deethanized gas expansion and cooling from stage (g) with its partial condensation in "cold box" in system of recuperative heat exchangers with use of refrigerants of different temperatures.

(i) low-temperature denitration, of partially condensed deethanizated gas from stage (h) in second rectification column and separation of nitrogen and helium with production of nitrogen- helium mix from column top and second residual from second rectification column bottom, being methane fraction;

(j) hydrogen de-aeration from nitrogen-helium mix from stage (i) by catalytic gas phase method, hydrogen oxidation by ambient oxygen to water with following drying of wet nitrogen- helium mix in forth adsorbent system with stationary layer of zeolite adsorbent, dewatering of nitrogen-helium mix after hydrogen oxidation, low-temperature helium extraction with production of liquid nitrogen and liquid helium preproduct as commercial product;

(k) low-temperature additional treatment of liquid helium preproduct with purity 99.97 % to remove nitrogen, oxygen, argon and neon in the system of sequentially working adsorbers with stationary layer of adsorbent with use activated charcoal as adsorbent and getting at this stage liquid helium commercial product with purity up to 99.9999 %;

(1) separation of residue from stage (g) in fourth rectification column with production of ethane from column top and fourth residue from column bottom - broad fraction of light hydrocarbons (BFLH), which mainly consists of propane and heavier hydrocarbons;

(m) BFLH adsorption cleaning from sulphur compounds, including mercaptans, in third system of adsorbers with stationary layer of adsorbent on the basis of zeolite with production of purified BFLH and zeolite regeneration in commercial methane fraction after booster compressor station without mercaptans impurities and discharge of regeneration gas, saturated with methanol and mercaptans to booster compressor stations;

(n) gas fractionation of purified BFLH directed from stage (m), at least, in two rectification columns to separate hydrocarbons with production of propane, butane, pentane- hexane fractions for their further pumping into storage tanks for commercial products storage;

(o) commercial fuel gas treatment by compounding of methane fraction from stage (i) and compressed and dried ballast gas from stage (e) with addition of technological methanol from stage (d), compressing of gas with 1.9 MPa to 5.6-7.5 MPa pressure for pumping it in main pipeline, besides gas regeneration supply from stage (m) and booster compressor stations is separate;

(p) forming of optimum flow set of cooling agents of different temperatures and flow from "heat pump" for recuperation heat exchangers of "cold box" taken from stages (g), (i), (j), (k), providing power consumption minimization for realization of deep processing of natural gas in general;

(q) storage of liquid nitrogen and helium in Dewar vessels in product park and their loading in liquefied gases transportation system;

(r) production of methane purified from nitrogen in liquid form and its storage and transportation at stage (i) or by capacity increase at heat pump compressors at stages (j) and (q) and by keeping permanent capacity with decrease of nitrogen amount in processed gas.

Suggested multiple stage diagram of complex extraction of valuable impurities from helium-rich hydrocarbon gas with high amount of nitrogen provides:

1) preliminary deep drying and purification of feed hydrocarbon natural gas from carbon dioxide, hydrogen sulphide, mercaptan and mercury compounds;

2) extraction, in sequential order (by boiling temperatures), of pentane-hexane hydrocarbon fraction, butane fraction, propane fraction, ethane, commercial fuel gas, liquid nitrogen and liquid helium from purified hydrocarbon gas;

3) additional helium treatment of nitrogen, hydrogen, oxygen, carbon dioxide, argon and neon purification with production of high purity liquid helium (over 99.9999 %) and get helium of competitive quality in comparison with other international producers;

4) build system of optimum use of recuperative heat during heat exchange between hot and cold flows due to conjunction of exchange heat equipment in "cold box", thus minimizes cold losses to atmosphere and reduces power consumption for cooling of recycled nitrogen- methane fraction in "heat pump"; 5) minimization of environmental discharge of acid waste waters and components of feed hydrocarbon natural gas due to acid water stripping and usage of process gas flows for adsorbents regeneration;

6) expand industrial base of undeveloped Russian eastern region due to construction of range of gas and chemical facilities, using worked-out hydrocarbon fractions as raw material for further processing to polyethylene, polypropylene, ethylene oxide, ethanol, methanol, glycol, components of automotive fuel and etc.

It is efficient for first level purification of feed stream natural hydrocarbon gas from mechanical impurities and dropping liquid at stage (a) to place built-in coalesce filter in gravity separator providing junction and integration of liquid drops, so it increases deposition speed and gas purification intensity.

It is also efficient at second level of joint treatment of purified hydrocarbon gas from H₂S, carbon dioxide and methanol by absorption use at stage (b) as regenerated absorbent aqua solution of amine, bearing 20-40 % amine, therewith as amine we can use monoethanolamine, diethanolamine, triethanolamine, methyldethanolamine and their mixtures and in case of high concentration of hydrogen sulphide in purified gas we should add in amine aqua solution, containing triethanolamine 5-12 % piperazine to intensify activity to C0₂. Usage of contact devices in the form of cross flow nozzles at stages (b) and (c) in absorbing, regenerating and washing columns, which have separate throughput areas for gas and liquid phases, it allows us to optimize production process at these stages, decrease operational costs for realization of second treatment level of first purified hydrocarbon gas flow and increase quality of gas purification from extracted impurities.

It is also efficient to perform at stage (d) acid water stripping from stage (c) with production of stripped process water, returned to stage (c) and stripped gases in first rectification column, thus more effective for power consumption than conventional stripping in evaporator, where low concentration of methanol in stripped process water should be kept at all mass of fluid phase, balanced to vapor phase with high water vapor concentration, where at stripping diluted methanol water solution is produced, while in rectification column this balance is kept by low amount of liquid phase in column bottom and perform additional process methanol treatment with production of commercial high purity methanol in auxiliary unit of rectification columns or, at least, in second auxiliary rectification column with possibility to produce methanol of 99 % purity. Additionally, suggested technology practically excludes acid waste waters discharge into environment, because all stripped water is returned to stage (c) and is made up with fresh water if needed. From environmental point of view there is also a problem to utilize combination of acid gases from stage (c) and low-pressure stripping gases from stage (d), as former mainly consists of hydrogen sulphide and carbon oxide with impurity of methanol, and as second mainly consists of carbon dioxide, and direct discharge of these gases into atmosphere as well as supply to the flare leads to considerable local air pollution. However, problem of these gases utilization can be solved because fuel gas produced by this method contains not more than 2 % vol. of nitrogen and has heating value rather higher than standard. Substantial reserve in heating value of produced fuel gas allows us to pump full volume of acid gases from stage (c) into worked-out methane fraction with low nitrogen content and in balance quantity providing necessary heating value of commercial fuel gas, low-pressured stripping gases from stage (d) subjected to additional compression to 2.0 MPa (abs.) pressure and drying with 20-40 °C up to dew point not more than minus 20 °C. At stage (e) in drying process of low-pressured acid gases in absorbers to remove moisture from gas is advisable to use zeolite of KA type, having highest water selectivity during gases drying process in comparison with other adsorbents, though it has relatively low moisture retention capacity. Alternative solution for drying of low-pressure acid gases in adsorbers for moisture removal is usage of two layers of adsorbent in flow direction, and in first layer granulated aluminum oxide can be used, it has high moisture retention capacity, thus provides quite low depth of gas dewatering, but in second layer use zeolite of KA type, providing deep additional gas drying. Adsorbent regeneration at stage (e) is advisable to perform by low-pressure gaseous nitrogen flow from stage (j) heated to 300-350 °C, and perform adsorbent cooling after regeneration by low-pressure cold gaseous nitrogen flow from stage (j) and then low-pressure gaseous nitrogen after regeneration and cooling shall be discharged to the flare.

It is efficient at stage (f) - third level of drying and purification from mercury compounds of purified hydrocarbon gas second flow from stage (b) and its additional treatment from methanol vapors in second system of adsorbers with stationary layer of adsorbent with gas drying and purification from methanol vapors with adsorbent on the basis of aluminum oxide and zeolite KA to provide regeneration of adsorbents in adsorbers of gas purification from mercury - adsorbent on the basis of aluminum oxide, copper promoted and unload adsorbent from adsorbers as worked out. Adsorbent regeneration of aluminum oxide and zeolite KA can be performed by agent gaseous flow of low or middle pressure, heated up to 300-350 °C and adsorbent cooling after regeneration lead by cold gaseous flow of low or middle pressure; in addition as gaseous agent of low or middle pressure methane fraction from denitration stage (i) followed by compression or part of purified hydrocarbon gas second flow from stage (b) followed by cooling, condensing and separation of desorbed liquid and return it to the beginning of stage (f);as for gas drying adsorbent cooling after regeneration as agent gaseous flow of low or middle pressure part of third dried and purified hydrocarbon gas from stage (f), at that rate of agent gaseous flow of low or middle pressure is 5-10 % of total amount of second flow of purified hydrocarbon gas supplied to stage (f). As a result of deep drying and treatment of purified hydrocarbon gas at stage (f) moisture content in dried third flow of dried and purified natural gas is not more than 0.1 ppmv (or 0.00001 % vol.), carbon dioxide content not more than 2 ppmv, methanol is not more than 50 ppmv, thus, provide normal operation of following cryogenic equipment. To protect equipment from possible ingress of microscopic adsorbent specks collected third flow of dried and purified natural gas is subjected to filtration, at that in third flow of dried and purified natural gas adsorbent specks size should not be more than 50 micron.

Presents of recuperation heat exchangers in "cold box" at stage (h) with use of flows of different temperature makes it feasible to place in "cold box" also cryogenic equipment from stages (i) and (j), these actions reduce significantly cooling losses and decrease power consumption for method realization in general.

At stage (k) liquid helium subproduct final treatment to remove nitrogen, hydrogen, oxygen, argon and neon impurities is provided, it is performed in adsorbers with activated carbon at 70-100 K temperature and 4 MPa pressure because argon, neon, nitrogen, hydrogen, adsorption with activated carbon can be provided only in cryogenic conditions at high pressure in adsorbers; thus regeneration of activated carbon adsorbent can be performed by low-pressure gaseous nitrogen flow from stage (j) heated up to 200-250 °C and discharge strippings to the flare and adsorbent cooling after regeneration is provided by cold flow of low-pressure gaseous nitrogen from stage (j). Liquid helium purification quality at stage (k) can be measured by neon slippage in liquid helium flow at output of last adsorber in sequential system of adsorbers. Alternative for cost-ineffective final treatment of liquid helium subproduct can be replacement of adsorbing liquid-phase helium purification from hydrogen in absence of neon and xenon in favor of catalytic gas-phase hydrogen oxidation by ambient oxygen with production of water with following drying of undried nitrogen-helium mixture at stage (j) in fourth adsorbers system with stationary layer of zeolite adsorbent, where zeolite of KA type can be used; thus adsorbent regeneration is provided by low-pressure gaseous nitrogen flow from stage (j) heated to 300-350 °C, as for adsorbent cooling after regeneration it can be conducted by cold flow of low-pressure gaseous nitrogen, after regeneration and cooling adsorbent is directed to the flare.

It is efficient to use wide-pore zeolite of NaX type, having high selectivity to sulfur compounds for BFLH adsorption treatment from sulfur compounds, including mercaptans in third system of adsorbers with stationary layer of adsorbent at stage (m). At transfer from adsorption phase to regeneration phase it is also advisable to use nitrogen for adsorber drainage and adsorbent layer blowdown before regeneration, and to perform adsorbent regeneration by methane fraction at 0.08 MPa pressure, heated up to 300-350 °C, for adsorbent cooling after regeneration use cold methane fraction flow from stage (i). It allows creating nitrogen blanket above displaced liquid BFLH and prevent its contamination of methan by equitable transition of methane from desorption gases to BFLH, which may decrease quality of propane fraction extracted from BFLH. Methane fraction after regeneration and adsorbent cooling is compressed and pumped into commercial fuel gas.

For power consumption decrease and maximum usage of inner energy resources it is advisable to provide heat in equipment of cryogen stages (h), (i), (j), (k) from recuperative heat exchangers combined in "cold box" by hotter technological flows or by condensation heat of gas flows of higher temperature or "heat pumps" and heat removal in equipment of cryogen stages (h), (i), (j) (k), combined in "cold box" provided by cooler technological flows or by vaporization heat of liquid flows with lower temperature or cooling cycles.

It is efficient at stage (n) as heat source for boilers heating in two rectification columns for hydrocarbons separation (debutanizer and depropanizer) use oil AMT-300 or other similar heat carrier with warming temperature up to 260 °C that allows excluding necessity to provide facility with expensive and deficient heat carrier, which is high pressure water steam. Heat carrier can be warmed up to 260 °C in tube furnace, it can use gas from fuel network where blowdowns from technological stages are discharged as a fuel.

At stage (n), gas fractionation of purified BFLH, it is advisable to split hydrocarbons in rectification columns at technological processes, providing high clarity of separation and high quality of produced products, thus obtained propane fraction contains at least 97 % vol. of propane, butane fraction - min. 98 % vol. sums of butanes, pentane-hexane fraction - min. 98 % vol. of pentane and higher hydrocarbons.

It is practical for power consumption decrease to use at stage (j) as working medium for "heat pump" methane-nitrogen mixture totally absent of ethane, produced in additional side column, which is connected to second rectification column and where cross flow of methane- nitrogen fraction is made periodically for fraction makeup and removal of cumulated impurities.

To provide flexibility in facility operation in case of changes of quantity and quality of feed stream and market situation it is advisable at helium-rich natural gas processing facility to have, at least, two parallel process lines; and thus, there should be more lines including stages for ethane, BFLH production and denitration than lines for helium production, it allows collecting nitrogen-helium mixture from all lines in large header, where from mixture can be distributed to helium production stages in accordance of helium content in nitrogen-helium mixture and capacity of liquid-helium generator and supply excess to the flare or store it till next production or change of production target.

LIST OF DRAWINGS

Shown in figure 1 concept drawing of complex extraction of valuable impurities from helium-rich hydrocarbon natural gas with high nitrogen content performed at station, including units, stages and pipelines.

100 - feed stream separation unit;

1 10 - unit of hydrocarbon gases absorption from acid gases;

120 - regeneration of saturated absorbent;

130 - process condensate stripping unit (sour water);

140 -low-pressure acid gases compression and drying unit;

150 - hydrocarbon gas drying and mercury extraction unit;

160 - low-temperature rectification unit;

170 - cooling and expansion of deethanizated gas ("cold box");

180 - low-temperature denitration unit;

190 - helium production unit;

200 - ethane and BFLH extraction unit;

210 - BFLH adsorption unit;

220 - gas fractionation unit;

1-39 - pipelines.

"Cold box" in general is marked by dotted line, which combines units 160, 170 and 180.

BRIEF DESCRIPTION OF DRAWINGS

Feed stream of hydrocarbon natural gas is supplied by 1 pipeline to unit 100, feed stream separation unit, consisting of 4 separators which realize stage (a) of first level of hydrocarbon natural gas purification from mechanical impurities and dropping liquid. In three-phase separators feed stream phases gravity separation method is used, based on density difference of gas, dropped liquid and solid mechanical substances and phases separation inertial method due to centrifugal force action and flow direction change in separator itself. After hydrocarbon natural gas feed stream purification, first separated purified hydrocarbon natural gas flow consisting of hydrogen, helium, nitrogen, carbon dioxide, methane, ethane, propane, sums of butanes, pentane and higher, methanol, hydrogen sulphide and mercaptans is supplied from separators by common flow by pipeline 2 to unit of hydrocarbon gases adsorption of acid gases 110, and liquid from separators' bottom goes by gravity flow to drainage tank.

In acid gases absorption unit 110 stage (b) second level of purification of purified hydrocarbon gas from hydrogen sulphide, carbon dioxide and methanol impurities is realized. First flow of purified hydrocarbon gas is supplied to bottom section of absorber packing, working at 7 MPa pressure and cooled in water additional cooler at 30-60 °C regenerated water solution of amine adsorbent with concentration of methyldiethanolamine (MDEA) 40 % vol. to the top section of absorber packing thus provides deepness and selectivity of gas purification. Purified gas from top part of absorbers with temperature 30-40 °C and carbon dioxide content, max. 5 ppm passes recuperative heat exchanger and then cooled purified gas at temperature not higher than 30 °C and not lower than 20 °C goes to purified gas separator where gas cleaning from trapped absorber drops is taken place. Worked-out absorber water solution, saturated by acid components (carbon dioxide (C0₂) and hydrogen sulfide (H₂S)) and methanol from the bottom of absorber, is supplied by pipeline 4 in regeneration of saturated absorbent unit 120. Separated, purified second flow of purified hydrocarbon gas comes from the top part of the separator by pipeline 3 into unit 150 drying and mercury removal.

In unit 120, which is regeneration of saturated absorbent, stage (c) is realized: worked-out flow of MDEA solution, saturated by acid components and methanol, is supplied by pipeline 4, throttled to 3^1 MPa pressure and heated in recuperative heat exchangers due to MDEA regenerated heat. Worked-out flow of MDEA solution, saturated by acid components and heated to 70-80 °C temperature then throttled further to 1 MPa pressure and supplied to expansion station (expanser) of worked-out MDEA solution. There is separation of expansion gas from worked-out MDEA solution due to pressure drop to 0.8 MPa (abs.) at expansion station (expanser). MDEA solution, saturated by acid components, is supplied from the bottom of expansion station (expanser) to tube side of recuperative heat exchangers, where it is heated to 95-105 °C and directed to the top section of desorber packing working at 0.18 MPa pressure. Absorber is equipped with thermal siphon to heat MDEA saturated solution up to 123-125 °C temperature. The resulted gas phase with trapped absorbent drops after heating of saturated MDEA solution is discharged from the top of desorber. Next, this mixture consisting, mainly, of acid components and methanol is pumped into acid gas separator, where is divided in gas and liquid phases. Acid gas separator is equipped with packed column for acid gas washing to remove methanol by process water, coming to unit 120 by pipeline 5. Washed from methanol acid gas is supplied to compression and drying of low-pressure gases unit 140 as low-pressure acid gas, and condensate (acid water) consisting of water, methane and carbon dioxide from the bottom of acid gas separator is directed to condensate (acid water) stripping unit 130. Regenerated MDEA mixture is returned by pipeline 6 to unit 110, which is absorption cleaning of hydrocarbon gases from acid gases.

Stage (d) is realized in system of additional rectification columns with reboilers, cooling condensers and separators in unit 130, which is process condensate (sour water) stripping. Condensate (sour water) from regeneration of saturated absorbent, unit 120, by pipeline 7 is heated in recuperative heat exchanger and supplied to middle part of first additional rectification column where temperature in the bottom of column is 118 °C. Condensate (sour water) is stripped from methanol and carbon dioxide in the column. Stripped condensate from the column bottom follows to recuperative heat exchanger, air cooler, cooler and goes to process water header. Mixed vapors of methanol, carbon dioxide and water is discharged from the top of first additional rectification column, then it is cooled and pumped in separator, where water solution of methanol is divided from stripping gases. Next, water solution of methanol passes to recuperative heat exchanger, then after heating it follows to second additional rectification column for methanol concentration, where bottom temperature is 188 °C. Vapors from the top of second rectification column are cooled in air cooler, condensed and supplied to reflux tank, where methanol is extracted from stripping gases. Stripped condensate is returned to regeneration unit of worked-out amine solvent 120 by pipeline 9 for MDEA regenerated solution makeup and preservation of its needed concentration. Stripping gases escaped in first and second additional rectification columns are collected in one flow and directed to low-pressure compression and drying unit 140. High-purity methanol is discharged from unit 130 by pipeline 8.

Stage (e) - stage of low-pressure gases drying and compression is done in unit 140. Low- pressure condensate stripping gases and low-pressure acid gases by pipelines 10 from unit 130 and pipeline 11 from unit 120, respectively, are supplied to separator, where separation of condensed moisture is taken place. Next, gases undergo two-stage compression. After first stage of two-stage compression, gases are directed to cooling at 0.4 MPa pressure, after second stage of compression at pressure 1.99 MPa compressed gases are cooled and supplied to separator, where condensed moisture is separated and then, they are flowed to first system of adsorbers with stationary layer of adsorbent for drying to provide required water dew-point of commercial gas. Drying is performed in three adsorbers working in cyclone with combined layer of adsorbent on the basis of grain aluminum oxide (top layer) and zeolite KA (bottom layer), which doesn't absorb sulphur compounds and methanol, thus allows releasing regeneration gas to atmosphere. At adsorption stage flow of dried gas goes through layers of adsorbent from the top to the bottom to exclude adsorbent layer fluidization and chafing. Ballast, compressed and dried gas with carbon dioxide 88-89 % vol., nitrogen 2 % vol., methane 7 % vol., methanol 1.6 % vol. from adsorbers is gone through filter, where it is cleaned from trapped zeolite and discharged through pipeline 14 to be mixed with calculated amount of high-purity methanol to get standard fuel gas calorific capacity, next, this gas mixture is supplied to booster compressors and discharged after them as commercial fuel gas by pipeline 15. There is also a split to "pure" methane fraction and "impure" methane fraction, which is used for zeolite regeneration in BFLH adsorption unit 210. Adsorber is switched from adsorption stage to regeneration after adsorbents are saturated with moisture. At regeneration and cooling stages gas flow (desorbing agent) goes through adsorbent layers from bottom to top. For regeneration stage nitrogen is heated in electric furnace in recuperative heat exchanger to 240-300 °C, directed to adsorber, where it increases adsorbent temperature to 200-250 °C and keeps adsorbent temperature at level of 230-290 °C during all regeneration process. Saturated regenerated gas is supplied from top of adsorber, cooled in recuperative heat exchanger and air coolers to temperature not more than 50 °C and is discharged by pipeline 13 to the flare. For adsorbent cooling after regeneration cold nitrogen is used, supplied by pipeline 12 from bottom to top in adsorber and decreasing temperature of complex layer of adsorbent to 20-30 °C. Next, in accordance with cyclogram of drying process apparatus switches to adsorption stage.

Wet second flow of purified hydrocarbon gas from unit 1 10 by pipeline 3 comes to stage (f) to provide third level of drying, purification from methanol and mercury compounds in second adsorbers system with stationary adsorbent layer, consisting of four adsorbers of methanol and moisture, working in pairs, in parallel and two adsorbers of gas purification from mercury compounds in unit 150. At the beginning gas to be purified is directed to top of two adsorbers in parallel, working at drying stage with stationary layer of adsorbent on the basis of aluminum oxide and zeolite KA, thus gas is purified from methanol and moisture to water vapor concentration in dried gas not more than 0.1 ppmv. Coming out dry gas from bottom of adsorbers is divided to two flows: main flow is supplied to gas purification section from mercury compounds, and relatively low second flow by pipeline 16 is directed as desorbing agent to adsorbent regeneration of gas drying. As an alternative of adsorbent regeneration usage of methane fraction as desorbing agent is provided, supplied by pipeline 17 or low-pressure nitrogen - by pipeline 18. After saturation of adsorbents adsorbers are switched from adsorption stage to regeneration. At regeneration stage and cooling gas flow (desorbing agent) passes through adsorbent layers from bottom to top. For regeneration desorbing agent is heated in recuperative heat exchanger and then in process furnace to 300 °C temperature and directed to adsorber and then, first adsorbents' temperature increases to 300 °C and then lasts during adsorbent regeneration process at 290 °C level. Moisture-saturated regeneration gas is discharged from top of adsorber and passed through separator, where condensed water is separated. After regeneration adsorbent layer is cooled to temperature max. 35 °C by corresponding flow of cold desorbing agent with recuperation of heat removal. Thus, heat recuperation of hot gases improves unit performance indicators due to rational use of power resources. After adsorbents regeneration and cooling apparatus are switched to adsorption stage in accordance with cyclogram of drying process. By usage of methane fraction as desorbing agent, desorption products are discharged to pipeline 19 by pipeline 20, and if low-pressured nitrogen is used, desorption products are flushed into pipeline 21 by pipeline 22.

In section of gas purification from mercury compounds two adsorbers are loaded with adsorbent layer of the basis of aluminum oxide, cuprum promoted, thus one of them works in adsorption mode (adsorbent regeneration is not provided due to low concentration of extracted component, because purified gas has only mercury amount not more than 2 ppb) as second is in stand-by mode and switches on if adsorbent in first adsorber becomes saturated. Worked-out adsorbent is directed to special facility for mercury recovery and replaced by fresh adsorbent. Dried to moisture amount max. 0.1 ppmv and purified from mercury third flow of dried and purified natural gas is discharged from unit 150 by pipeline 23.

Next, dried and purified from mercury third flow of dried and purified natural gas from drying unit of hydrocarbon gas and mercury removal 150 by pipeline 23 is supplied to low- temperature rectification unit 160, where, in first rectification column separation to deethanizated gas, releasing from column top and first residue, containing hydrocarbons C₂ and higher for stage (g) realization. To decrease cold losses condenser cooler and boiler with column where heat of relatively high temperature flow is recuperated are placed in general "cold box" of cryogen equipment.

Deethanizated gas is directed by pipeline 24 to deethanizated gas expansion and cooling unit 170 in "cold box" in the system of recuperative heat exchangers with usage of refrigerants of different temperature for stage (h) realization. "Cold box" has powerful insulation of environment and in space between "cold box" equipment insulation material is placed, thus excludes convective streams inside "cold box" and thermostats each of heat exchangers or mass exchange apparatus. Since refrigerants of different temperature have different pressure, there is optimum mover in relation to cold flows, used as refrigerants and, relatively activated carbon "warm" flows, used as heat carriers, which allows denying use of pumps for various cold sources pumping. Besides, effective heat transfer and use of nitrogen-methane mixture recycle, formed in denization columns through "heat pump" allow eliminating expensive external cold sources, e. g. liquid nitrogen from outside.

Partially condensed, deethanized gas, consisting of mainly methane and nitrogen with helium impurities is directed by pipeline 25 to cryogen denitration unit 180, where stage (i) is realized in second rectification column, working at 2.6 MPa pressure. Partially condensed in "cold box" deethanized gas is supplied to middle part of second rectification column, equipped with reboiler in bottom part. Second residue, which is methane fraction, is extracted at bottom of second rectification column. Extracted from top of second rectification column gas at temperature near minus 110 °C is supplied to "cold box", where it is partially condensed. Partially condensed gas from "cold box" at temperature near minus 120 °C is pumped to bottom section of column for nitrogen and helium extraction, working at 2.6 MPa pressure. Top product of nitrogen and helium extraction column with temperature near minus 115 °C is supplied to "cold box", where it is heated and directed to unit 190 as nitrogen-helium mixture by pipeline 26. Bottom product of nitrogen and helium extraction column is fed as reflux to second rectification column. Methane fraction from bottom of second rectification column with temperature near minus 101 °C is returned to "cold box", where it is vaporized and extracted as product gas and also can be supplied to hydrogen gas drying and mercury removal unit 150 by pipeline 27 for adsorbents regeneration and by pipeline 27 is directed to booster compression systems to get commercial fuel gas.

Nitrogen-helium gas with partial content of hydrogen and methane is fed by pipeline 26 to helium production unit 190, where stages (j) and (k) is realized, which includes several sections, such as catalytic gas cleaning from hydrogen, cryogenic helium extraction, its fluidization and cryogen helium adsorption from firstly inert gases. In hydrogen extraction unit, in-coming nitrogen-helium mixture at 2.4 MPa pressure and temperature near 44 °C is mixed with recycle stream from helium compression unit, which has oxygen needed for hydrogen and methane oxidizing reaction. Mixed flow is preliminary heated in hydrogen reactor heater to 370 °C and supplied to hydrogen reactor, where hydrogen and methane are oxidated to water and carbon dioxide by catalytic reaction. Temperature of out coming flow at reactor discharge increases to 450 °C due to exothermal reaction. Therefore, moist nitrogen-helium mixture is discharged from reactor with 450 °C temperature without hydrogen trace, but with oxygen content 1000 ppm and some nitrogen content increase due to atmospheric nitrogen. After reactor undried, hot nitrogen-helium mixture is gone through gas regeneration heater, where it heats regeneration gas for drying adsorbers of nitrogen-helium mixture to 230 °C. Next, preliminary cooled gas with 400 °C temperatures is cooled in hydrogen reactor heater to 74 °C temperature and then cooled by air in hydrogen reactor cooler outlet to 47 °C. Undried, cooled nitrogen- helium mixture is cooled to 25 °C in cooler of undried nitrogen-helium mixture, where low- pressure nitrogen is used from helium extraction unit.

Next, undried nitrogen-helium mixture is fed to adsorber of nitrogen-helium mixture. Nitrogen-helium mixture is supplied to one of two adsorbers from bottom to top. Water and carbon dioxide contained in nitrogen-helium mixture are adsorbed by molecular gate and their content in gas is decreased practically to zero level to prevent system freezing of helium extraction and fluidization. Next, dried and helium enriched gas is gone through filter for nitrogen-helium mix to eliminate dust from molecular gates, thus can affect operation of cryogen process in helium extraction unit. To remove water and carbon dioxide from supersaturated molecular gates, adsorber of nitrogen-helium mixture is heated in regeneration mode by regeneration gas approximately 5 hours and then cooled in near 4 hours. As regeneration gas low-pressure nitrogen, extracted from nitrogen-helium mixture in helium extraction unit, is used.

Operation of two adsorbers in nitrogen-helium mixture drying unit performs with periodic switch from adsorption mode to regeneration. After regeneration regenerated adsorbent is several grades warmer than adsorber in adsorption mode. To minimize temperature fluctuations of undried nitrogen-helium mixture at switch from worked-out adsorber to regenerated adsorber, so called parallel adsorption mode is used. During this parallel adsorption mode, ongoing for 3-4 hours, main part of gas is still directed to working adsorber, as part of total gas flow is gone through regenerated, warmer adsorber and by leading it to working temperature and minimizes temperature difference between two dried helium enriched gas flows, discharged from adsorbers. After this parallel mode flow is fully switched to generated adsorber.

Helium extraction unit consists of cryogenic nitrogen extraction and pressure swing adsorption (PSA). Dried nitrogen-helium mixture is fed to "cold box" to produce helium for cryogenic split to 4 flows: helium enriched flow from helium production "cold box", liquid nitrogen flow, which is fed to liquid nitrogen storage by pipeline 28 and further used for preliminary cooling in helium fluidization and storage and packing units; tail gas flow, which is nitrogen with helium at 1.0 MPa pressure coming to compression in helium compression unit to increase helium recovery rate, middle pressure nitrogen flow at 0.8-1.0 MPa pressure, which is directed to section boundary as nitrogen for auxiliary systems.

Middle pressure 2500 kg/h nitrogen is fed to section boundary for adsorbers regeneration of BFLH cleaning unit 210 by pipeline 34. Remain nitrogen is supplied to nitrogen booster compressor, compressed, cooled in nitrogen booster compressor cooler, then cooled in helium production "cold box" to minus 21 °C temperature, followed further to nitrogen turboexpander, 43 where nitrogen is expanded to 0.3 MPa with production of low-pressure nitrogen, providing necessary cooling load for nitrogen export. Nitrogen turboexpander and nitrogen boost compressor are placed on general roll. Low-pressure nitrogen is heated in helium production "cold box" and supplied to hydrogen de-aeration unit as regeneration gas and partially can be used in unit 150 for methanol of second purified hydrocarbon gas drying and purifying.

Out coming from helium production "cold box" helium enriched flow contains over 80 % mol. helium and goes to helium purification stage in PSA section. Pressure swing adsorption is based on physical phenomena when high-boiling components of low-polarity, such as hydrogen or helium, are hard to adsorb in comparison to such molecules as C0₂, CO, N₂ and hydrocarbons. Thus, in pressure swing adsorption main part of impurities in helium-rich flow can be selectively adsorbed with production of high purity helium product.

PSA process consists of two stages:

- impurities adsorption by high pressure, which is needed for partial pressure increase and, consequently, increase quantity of adsorbing impurities;

- desorption (or regeneration) is performed by low pressure to decrease maximum possible quantity of impurities at adsorbent, and respectfully, to reach high helium purity rate.

Pressure swing adsorption process is realized in temperature 100-80 K, produced by liquid nitrogen evaporation in shell side of shell-and-tube adsorber, where tubes are filled with adsorbent

Adsorbers in adsorption process are placed stepwise, which allows high flexibility and excludes impact of composition, temperature and pressure changes in feed gas.

PSA performed at high pressure, allows us, from one hand, to reduce significantly adsorbers' size due to adsorption duration decrease, which is provided by quick adsorbent regeneration due to sharp pressure relief in apparatus, and on the other hand, to realize practically adsorber isothermal process at all process stages. Activated carbons are used as adsorbent, e. g. SKT-6 type, which do not, practically, sorb helium, even at very low temperature, and nitrogen, argon, neon are sorbed in sufficient amount.

PSA efficiently separates helium saturated out coming flow to pure helium flow with helium amount 99.97 % mole and desorbs tail gas flow, consisting, mainly, of nitrogen and residual impurities. Tail gas flow is directed to helium compression unit. Fluid pure helium is fed to helium liquidizing unit, where from helium flow as one of the main product flows is directed by pipeline 29 to liquid helium storage and packing.

Compression unit of tail gas is used for collection of tail gas and blowdowns of process equipment, containing helium, after production process and their compression for reinjection to 5 000843 feed gas. This process significantly increases helium extraction rate in the station. Undried nitrogen-helium mixture, after pressure relief in adsorbers for nitrogen-helium mixture drying, is mixed with tail gas flow from PSA and supplied to tail gas compression unit at 0.13 MPa pressure. Residual gas from helium production "cold box" is combined with I&C air from section boundary and supplied at 0.5 MPa pressure to tail gas compression module. Pressurized tail gases are cooled in tail gas compressor cooler to 47 °C and separated from the rest lubrication oil in de-oiling block. Pressurized tail gas is supplied to hydrogen extraction unit for mixing with feed gas.

Liquid helium flow from PSA is fed to helium fluidizing unit where is subjected to fluidizing by usage of modified Brayton cycle with three pressure levels, including preliminary stage of cooling by liquid nitrogen from liquid nitrogen storage and packing block, stages of expansion and preliminary cooling by Claude cycle, and also adiabatic expansion in turbine and isenthalpic expansion by Joule - Tomson cycle for liquid helium production. Liquid helium is discharged from helium fluidizing unit as main product flow and supplied by pipeline 28 to liquid helium storage and packing block. Gaseous helium, boil off gas and endflash gas are directed to liquid helium storage and packing zone for reliquefaction.

Low-pressure helium from helium liquefaction unit, first, compressed in low pressure helium compression block to middle pressure and cooled to 47 °C in cooler of low-pressure helium compressor. Pressurized middle pressure helium is combined with middle pressure helium, coming from helium liquefaction unit and supplied to middle pressure helium compressor for final compression. Pressurized high pressure helium is cooled in cooler of middle pressure helium compressor to 47 °C and separated from remain lubrication oil in de-oiling unit.

First residue, bearing hydrocarbons and higher from unit 160 is directed by pipeline 30 to unit 200, where extraction of ethane and broad fraction of light hydrocarbons (BFLH) from first residue is taken place, thus, realizing stage (1). First residue is heated up to 20 °C in BFLH recuperative heat exchanger and supplied by two-phase flow to middle section of fourth rectification column, equipped with reboiler, condenser-cooler, reflux tank, where is subjected to rectification. Ethane fraction, which is top column product, with temperature over minus 3 °C is cooled and partially condensed. Partially condensed ethane fraction is followed to reflux tank, where it is separated to fluid and gaseous phase. Fluid is returned as reflux to the column, gas, discharged from the top of reflux tank is directed by pipeline 31 as product ethane. Bottom column flow is cooled in recuperative heat exchanger for ethane fraction, at 35 °C temperature and discharged by pipeline 32 as broad fraction of light hydrocarbons (BFLH), being fraction mainly consisting of propane and more heavy hydrocarbons. 43

BFLH from ethane and BFLH extraction unit 200 by pipeline 32 is fed at 1.8-2.2 MPa pressure to BFLH adsorption cleaning unit 210 from sulphur compounds (mercaptans) and methanol for stage (m) realization. 210 unit represents third adsorber system with stationary layer of adsorbent with use of adsorbents on the basis of wide-pore zeolite NaX, consisting of four adsorbers, furnace, recuperative heat exchangers, at that, two adsorbers are in adsorption stage and work in parallel, one is at adsorbent regeneration, one is at cooling; switch of adsorbers from one mode to another is performed in accordance of cyclogram. BFLH is fed to two adsorbers working in adsorption mode in parallel. Going through adsorbent from top to bottom, BFLH purified from mercaptans and methanol is supplied to tank of pure BFLH and next, to further processing to gas fractionation unit 220 by pipeline 33. Adsorber is cleared of liquid phase before being switched to adsorbent regeneration mode. To increase desorption effectiveness, zeolite regeneration and cooling is done by cooling gas (methane fraction) supplied from commercial gas header after booster compressors at 2.2-2.5 MPa pressure by pipeline 19. To eliminate dilution of methane in BFLH adsorber emptying is done by displacement of liquid from apparatus to the tank due to middle pressure nitrogen supplied by pipeline 34 at 2 MPa pressure. After adsorber emptying, adsorber is blowed through by methane fraction flow at 2.2-2.5 MPa pressure to remove liquid film from the surface of adsorbent grains and reduce power consumption at regeneration stage and its enhancement.

Blowdown methane fraction from adsorber is supplied to fuel network or to booster compression systems by pipeline 35. Desorbing gas is heated in recuperative heat exchanger and then in furnace to max. temperature 320 °C and fed to adsorber, working as desorber, then to filter, recuperative heat exchanger, where heat is transferred to cooling gas before furnace. After cooling is finished nitrogen blowdown is began. Saturated by mercaptans regeneration gas is directed to fuel network or to booster compressor systems by pipeline 36. Purified BFLH is discharged from unit and fed to gas fractionation unit 220 by pipeline 33.

Purified BFLH is fed to gas fractionation unit 220 by pipeline 33, realizing stage (n). Unit consists of two rectification columns for hydrocarbons separation, boilers, condenser-coolers, reflux tanks and pumps. Purified BFHL firstly is gone to depropanization column, on the top of which we get propane fraction in vapor phase, hydrocarbon vapors are condensed in air cooler and discharged to reflux tank. Part of propane fraction is used as reflux in depropanization column, and net part with propane content 97 % vol. is supplied by pipeline 37 to commercial propane storage park. Bottom product from depropanization column after heating in recuperative heat exchanger is fed to column-debutanizer as a feed. Butane fraction vapors from top of column-debutanizer are condensed in air cooler and directed to reflux tank, where from part of butane fraction is used as reflux in column-debutanizer, and net part with butane content 98 % vol. is discharged by pipeline 38 to commercial butane storage park. Bottom product of column- debutanizer - pentane-hexane fraction is cooled in air cooler and discharged from unit by pipeline 39. As heat source for boilers heating in two rectification columns of hydrocarbons separation (column-debutanizer and depropanizer column) AMT-300 oil is used or other similar heat carrier with warming temperature max.260 °C.

For realization of diagram mentioned above, several parallel process lines are provided at the facility.

Moreover there are several auxiliary units (not shown at fig. 1) corresponding to operation of (o), (q) and (r) stages: commercial gas compounding treatment of methane fraction and compressed and dried ballast gas with addition of methanol at winter time, storage of liquid nitrogen and helium in Dewar vessels in products park and their loading to liquefied gases transportation system; liquid methane transportation and storage, which is produced due to maintenance of heat pump compressors capacity by decrease of part of excess gas in processed gas.

Some results of design calculations are shown in tables 1-3 for plant performance 5.338 mln t/year, implementing method of deep treatment of natural hydrocarbon gas with high nitrogen content from East Siberian fields. As follows from table 1 , where quality of feed product and nonorganic and organic end products are shown, developed method provides high purity of end products by using fairly simple plant process diagram, in general, and optimum power consumption of recuperative heat exchange; thus, purity of organic products is in range 95.5 % (ethane in ethane fraction) - 98.8 % (butane in butane fraction), purity of produced nitrogen and helium is close to 100 %. Purity of produced methane (table 1) is significantly higher than standards and norms of Russian Federation for commercial fuel gas (table 2) and virtually complies with prospective export requirements for export fuel for the countries of the Pacific region. In table 3 plant material balance is shown, which indicates minimization of feed losses during production process. Large facility for natural gas processing from Eastern Siberia fields can include 7-10 described plants, thus provides facility flexibility due to optimal operation of separate plants with account of market conditions, as well as deviations in quality and quantity of feed hydrocarbon gas, incoming from different fields.

Thus, the claimed invention allows solution of task to provide complex method to extract valuable impurities from natural helium-rich hydrocarbon gas with high nitrogen content with preliminary drying and purification, and then, further, in sequential order (by boiling temperatures), extraction of pentane-hexane hydrocarbon fraction, butane fraction, propane fraction, ethane, commercial fuel gas, liquid nitrogen and liquid helium of high purification, with use of effective technologies of low temperature, rectifying, adsorbing, gravity and filter separation and minimization of power consumption for process realization due to optimum use of recuperative heat during heat exchange between hot and cold flows and use of turbodetanders and heat pumps.

Table 1

Name of parameter Value

Volume fraction of hydrocarbons, %

Methane, minimum 85.0 Ethane, max. 6.0 Propane, max. 3.0

Mole fraction of nitrogen, %, maximum 2.0

Mole fraction of carbon dioxide, max % 2.0

Mole fraction of oxygen , %, maximum 0.5

Weight concentration of hydrogen sulphide, mg/m³, maximum 6.0

Weight concentration of sour sulphur, mg/m³, maximum 16.0

Weight concentration of total sulphur, mg/m³, maximum 30.0

Weight concentration of mercury, mg/m³. maximum 20.0

Weight concentration of mechanical impurities, mg/m³, maximum 1.0

Table 2

Table 3

Claims

1 The method of complex extraction of valuable impurities from natural helium-rich hydrocarbon gas with high nitrogen content which includes the following stages:

(a) the first level of purification of feed stream of natural hydrocarbon gas from mechanical impurities and liquid using the method of gravity separation which results in the first flow of purified hydrocarbon gas;

(b) the second level of cleaning of first stream of purified hydrocarbon gas from impurities of hydrogen sulphide, carbon dioxide and methanol using regenerated absorbent in high pressure absorption tower which results in the second flow of purified hydrocarbon gas and rich absorbent flow;

(c) recovery of rich absorbent flow in the system of sequence expansion units with knock-out drums under condition of gradually increasing temperature of absorbent and with reduced pressure in expansion units and further in low pressure regeneration column with preparation of recovered absorbent, restoring to stage (b) after recuperator system and acid and expansion gases; acid gas is separated from methanol and partially hydrogen sulphide and carbon dioxide in washing tower using process water which results in acid water and low pressure acid gas flow, while expansion gas is forwarded into fuel network;

(d) stripping of acid water at the stage (c) from methanol, hydrogen sulphide and carbon dioxide with preparation of stripped process water, restoring to the stage (c), and stripping gases subject to further cooling, partial condensation and separation in order to derive industrial methanol and low pressure stripping gases flow, at that industrial methanol is additionally treated to obtain high-purity methanol;

(e) compressing and drying of low pressure acid gases at the stage (c) in the blend of balance part of low pressure stripping gases at the stage (d) in the first system of adsorbers with stationary layer of adsorbent provided that adsorbers use adsorbent based on zeolite in order to obtain compressed and dried ballast gas that will be further injected into commercial fuel gas with blend of high-purity methanol or industrial methanol at the stage (d) on condition that heat value required for commercial fuel gas is ensured; whereas the rest part of low pressure stripping gases at the stage (d) is discharged on the flare, or low pressure acid gas at the stage (c) in the blend with full amount of low pressure stripping gases at the stage (d) is sent as raw material to the gas chemical enterprise

(f) the third level of drying and separation from mercury compounds of the second stream of purified hydrocarbon gas at the stage (b) in the second system of adsorbers with stationary layer of adsorbent provided that adsorbers apply drying of adsorbent gas on the basis of aluminum oxide and zeolite KA with spores diameter 3 A; adsorbent based on aluminum oxide, promoted by copper shall be applied in adsorbers intended for gas separation from mercury in order to obtain the third stream of dried and purified hydrocarbon gas that is partially used in recovery of adsorbents as a desorbing agent

(g) low temperature separation of the third stream of dried and purified hydrocarbon gas at the stage (f) in the first rectification column which results in the flow of deethanized gas at the bottom of the column and the first residue containing hydrocarbons C₂ at the top of the column;

(h) expansion and freezing of deethanized gas at the stage (g) with its partial condensation in the "cool box" in the system of recuperators using the flow of refrigerants of different temperatures;

(i) cryogenic denitrogenation of partially deethanized gas at the stage (h) in second rectification column and the column for nitrogen and helium escape to obtain nitrogen and helium blend at column top and the second residue, represented by methane fraction at bottom of second rectification column;

(j) removal of hydrogen from nitrogen and helium blend using catalytic gas-phase method for hydrogen oxidation by atmospheric oxygen to the state of water with consequent drying the stream of wet nitrogen and helium blend in the forth system of adsorbers with stationary layer of zeolite adsorbent, dehydration of nitrogen and helium blend to remove water, generated during hydrogen oxidation, cryogenic separation of helium in order to liquid nitrogen as a final product and liquid helium as a semi-finished product at this stage;

(k) cryogenic second-stage treatment of semi-finished product of liquid helium of 99.97 % consistency free of nitrogen, oxygen, argon and neon in the system of adsorbers operating in series with stationary layer of adsorbent using activated carbon as adsorbent to obtain liquid helium purified up to 99.9999 % as a final product at this stage;

(1) separation of the first residue at the stage (g) in the fourth rectification column which results in the generation of ethane at the top of the column and the forth residue at the bottom of the column, represented by broad fraction of light hydrocarbons (BFLH), mostly consisting of propane and heavier hydrocarbons;

(m) adsorption cleaning of BFLH from sulphur compounds, including mercaptans in the third system of adsorbers with stationary layer of adsorbent provided that adsorbers use adsorbent, based on zeolite to obtain purified BFLH and regeneration of zeolites of commercial methane fraction after booster compressor station free of mercaptans and discharge of regeneration gas, saturated by methanol and mercaptans to the booster compressor stations; (n) gas-fractionation of purified BFLH, coming from the stage (m), at least, in two rectification columns for hydrocarbons separation in order to obtain propane, butane, pentane- hexane fractions that will be later injected into tank parks for the storage of commercial products;

(o) preparation of commercial fuel gas by blending of methane fraction at the stage (i) and compressed dried ballast gas at the stage (e) with addition of industrial methanol at the stage (d), compression of gas with pressure from 1.9 MPa up to 5.6-7.5 MPa to be injected into gas- main pipeline, while regeneration gas supply at the stage (m) to the booster compressor stations is divided;

(p) formation of optimal set of streams of refrigerants of different temperatures and flow from heat pump for recuperators of "cold box", selected at the stages (g), (i), (j), (k), providing cost minimization to implement the method of high-refined natural gas in general;

(q) storage of liquid nitrogen and helium in Dewar vessels in the tank park and their injection into liquid gas transport systems;

(r) generation of liquid methane purified from nitrogen, its storage and transportation accordingly at the stage (i) either due to the increased capacity in compressors of heat pump at the stages (j) and (q), or owing to keeping the capacity unchanged while reducing the percent of excess nitrogen in processing gas.

2 The method in p. 1 differs in the fact that built-in coalescer filter is additionally installed in gravity separator at stage (a).

3 The method in p. 1 differs in the fact that water solution of amines is used as a regenerated absorbent at the stage (b).

4 The method in p. 3 is different in that water solution of amines contains 20-^40 % of the latter.

5 The method in p. 3 is different in that monoethanolamine, dietanolamine, triethanolamine, methyldethanolamine and their blends are used as amines at the stage (b).

6 Method in p. 3 differs in the fact that in water solution of triethanolamine 5-12 % piperazine is added at the stage (b).

7 The method in p. 1 differs in the fact that contact devices in the form of crossflow nozzle are applied in absorption, regeneration and washing towers at the stages (b) and (c).

8 The method in p. 1 is different in that at the stage (d) the stripping of acid water at the stage (c) to obtain stripped processed water, returning to the stage (c) and stripping gases is carried out by in the first additional rectification column while after purification of industrial methanol to achieve commercial high purity methanol is performed in the additional assembly of rectification columns or at least the second additional rectification column.

9 The method in p. 1 is different in that stripped process water, drained to the stage (c) is additionally fed with fresh water.

10 The method in p. 1 is different in that low pressure gas, generated at the stage (e) is compressed to pressure 2.0 mPa (absolute) and dried at the temperature 20-40 °C to the dew point of no more than - 20 °C.

11 The method in p. 1 is different in the fact that zeolite of KA type is used in adsorbers to remove moisture in drying process of low pressure acid gases at the stage (e).

12 The method in p. 1 differs in the fact that at the stage (e) two-layered adsorbent is used in adsorbers in drying process of low pressure acid gases to remove moisture along the gas flow, whereas granulated aluminum oxide is applied as an adsorbent in the first layer and zeolite of KA type - in the second layer.

13 The method in p. 1 differs in the fact that regeneration of adsorbents at the stage (e) is performed by the flow of low pressure nitrogen gas at the stage (j), heated to the temperature

300-350 °C, while freezing of adsorbents after regeneration is carried out by the flow of cool nitrogen gas at the stage (j).

14 The method in p. 1 differs in the fact that low pressure nitrogen gas after regeneration and freezing of adsorbent is discharged on the flare at the stage (e).

15 The method in p. 1 is different in that regeneration of adsorbents is performed by the flow of low and medium pressure gaseous agent, heated to the temperature 300-350 °C, while freezing of adsorbents after regeneration is made by the flow of cold gaseous agent of low and medium pressure.

16 The method in p. 15 differs in the fact that methane fraction which is drained from the denitrogenation stage (i) and further compressed, is used at stage (f) as a low and medium pressure gaseous agent for regeneration of spent adsorbents of gas drying and their freezing.

17 The method in p. 15 differs in the fact that at the stage (f) the part of the second flow of purified hydrocarbon gas at the stage (b) which is later freezed, condensated, separated from desorbed water and returned to the early stage (f), is used as a low or medium pressure gaseous agent for regeneration of spent adsorbents of gas drying, while the part of the third flow of dried and purified carbon dioxide at the stage (f) is applied as a low or medium pressure gaseous agent for the freezing of adsorbents of gas drying after regeneration.

18 The method in p. 15 is different in that at the stage (f) the consumption of low or medium pressure gaseous agent for freezing and regeneration of saturated adsorbents makes up 5-10 % of overall amount of the second flow of the purified hydrocarbon gas, coming to the stage (f).

19 The method in p. 1 differs in the fact that at the stage (f) the moisture content in the dried third flow of the dried and purified natural gas makes up no more than 0.1 ppmv (or 0.00001 % of volume), carbon dioxide content is no more than 2 ppmv, methanol - no more than 50 ppmv.

20 The method in p. 1 differs in the fact that at the stage (f) the obtained third flow of the dried and purified natural gas is subject to filtration in order to protect the equipment from possible entry of microscopic particles of adsorbent.

21 The method in p. 20 is different in that at the stage (f) the size of microscopic particles of adsorbent in the third flow of the dried and purified natural gas must be no more than 50 microns.

22 The method in p. 1 is different in that at the stage (h) the cryogenic equipment from the stages (i) and (j) is installed in the "cold box".

23 The method in p. 1 differs in the fact that at the stage (k) the second-stage treatment of liquid helium semi-finished product to remove inclusions of nitrogen, oxygen, carbon dioxide, argon and neon is carried out at the temperature 70-100 K and pressure 4 MPa.

24 The method in p. 1 differs in the fact that at the stage (k) the regeneration of zeolite in the adsorber for drying of nitrogen and helium blend is performed by the flow of low pressure nitrogen gas at the stage (j), heated to the temperature 200-250 °C upon condition of discharge of desorption products on the flare, while the freezing of adsorbent after regeneration is carried out by the flow of low pressure cold nitrogen gas at the stage (j).

25 The method in p. 1 is different in that at the stage (k) the quality of treatment of liquid helium is estimated as per slippage of neon in the flow of liquid helium at the output of the last adsorber in the system of sequence adsorbers.

26 The method in p. 1 is different in that the cleaning of helium from hydrogen is performed not by catalytic gas-phase method for hydrogen oxidating by atmospheric oxygen to the state of water with consequent drying the stream of wet nitrogen and helium blend at the stage (j) in the forth system of adsorbers with stationary layer of zeolite adsorbent at the stage (k), but using adsoption liquid-phase method.

27 The method in p. 26 differs in the fact that zeolite of KA type is used in adsorbers when drying the wet nitrogen and helium blend.

28 The method in p. 26 differs in the fact that the regeneration of adsorbents when drying the flow of wet nitrogen and helium blend is made using the flow of low pressure nitrogen at the stage (j), heated to the temperature 300-350 °C, while the freezing of adsorbents after regeneration is carried out by the flow of low pressure cold nitrogen gas at the stage (j).

29 The method in p. 1 is different in that while drying the flow of wet nitrogen and helium blend, low pressure nitrogen is discharged on the flare after regeneration and freezing.

30 The method in p. 1 differs in the fact that zeolite of NaX type is used in adsorbers at the stage (m) in the course of adsorption cleaning of broad fraction of light hydrocarbons from sulphur compounds, including mercaptans.

31 The method in p. 30 is different in that at the stage (m) nitrogen is used for emptying of adsorber and blow-down of adsorbent layer prior to its regeneration.

32 The method in p. 30 differs in that at the stage (m) the regeneration of adsorbents is performed by methane fraction under the pressure nearly 0.08 MPa, heated to the temperature 300-350 °C, while the freezing of adsorbents after regeneration is carried out by the flow of cold methane fraction from the stage (i).

33 The method in p. 30 is different in that at the stage (m) the methane fraction is compressed and injected into fuel gas after regeneration and freezing.

34 The method in p. 1 differs in the fact that heat supply to cryogenic units at the stages (h), (i), (j), (k) is provided by recuperators, united into the "cold box", by hotter process flows or due to the warmth of gas flows condensation with higher temperature or by "heat pumps".

35 The method in p. 1 is different in that heat removal in devices of cryogenic stages (h), (i), (j), (k) is achieved in recuperators, united into the "cold box" by cooler process flows or due to the warmth of evaporation of liquid flows with lower temperature or by cooling cycles.

36 The method in p. 1 differs in the fact that at the stage (n) oil AMT-300 or other similar heating medium with heating temperature up to 260 °C is used as a heat source to warm up reboilers of two rectification columns for hydrocarbons separation (debutanizer and depropanizer).

37 The method in p. 1 is different in that at the stage (o) the obtained propane fraction contains no less than 97 % vol. of propane fraction, no less than 97 % of propane volume, butane fraction - no less than 98 % of butane volume, pentane-hexane fraction - no less 98 % of pentane volume and higher hydrocarbons.

38 The method in p. 1 differs in the fact that at the stage (j) methane and nitrogen blend free of ethane is used as a working substance; the ethane is generated in the additional side column, adjusted to the second rectification column, where the flow of methane and nitrogen blend is incidentally maintained to upgrade the quality of this fraction and to remove the accumulated impurities. 39 The method in p. 1 differs in the fact that enterprises for helium-rich gas conversion implement the conversion process in several, at least, two parallel production lines, while the number of production lines, including the stages of blow-down of ethane, BFLH and denitrogenetion is more than the number of production lines with the stages of helium generation.