Process for the purification of gases

FIELD OF THE INVENTION
The present invention is a process for the purification of gases and relates to the rejection of light components from a hydrocarbon gas mixture. Particularly, the present invention relates to the rejection of light components from a hydrocarbon gas stream using solvent absorption at non-cryogenic conditions. More particularly, the present invention relates to the use of a combination of solvent absorption, purification and recovery steps to provide a product gas having a reduced content of light components and a stream enriched in light components.
BACKGROUND OF THE INVENTION
Processes for removing light components such as hydrogen or nitrogen, or other light components from a hydrocarbon gas mixture represent a problem in petroleum processing. Typically, the hydrocarbon gas mixture is a refining off-gas or a natural gas and the presence of the light components in the mixture limit further processing, or prevent sales of the gas mixture without the removal of the light components. Often these light components include hydrogen, nitrogen, carbon dioxide, carbon monoxide, helium, argon, and mixtures thereof. A wide variety of hydrocarbon gas mixtures are found in petroleum refineries. Some streams are integral parts of specific processes. Often these streams are recycled from a fractionation column or a separator to a reactor. One such recycle stream may be an impure hydrogen stream which must be purified before returning the stream to the reactor. Other of these streams may be by-product streams of a major hydrocarbon conversion process such as a fluid catalytic cracking unit or an ethylene plant. Processes available for the rejection of the light components from these gas mixtures can employ solvent absorption, cryogenic separation, adsorption over molecular sieve adsorbents, or membrane separation. The choice of a suitable process depends upon many factors, some of which are the product purity that is desired, the gas recovery levels, available pressure losses, pretreatment requirements, off-gas composition, impact of reaction products remaining in the light component stream, and the turndown capability of the selected process.
When natural gas is produced from a gas well or is recovered as associated gas from an oil well, the natural gas contains a number of light components which can reduce its heating value. Typically, a natural gas stream comprises nitrogen, methane, ethane, carbon dioxide, inerts, and C.sub.3.sup.+ hydrocarbons. In order to improve the heat content of the product natural gas, the concentration of nitrogen and other inerts must be reduced. This reduction of inerts is often required to meet the quality specifications of pipeline companies that transport natural gas from a well head or natural gas processing plant to the consumer. Typically the natural gas must meet the following specifications:
Heat Content--900 to 1000 BTU
Total Inerts (N.sub.2 +CO.sub.2)--7 mol % maximum
Nitrogen--4 mol % maximum
Actual pipeline specifications vary somewhat depending upon the producer's contract for price and quality. In general, a specification for a higher heating value requires a reduced amount of light components such as nitrogen and carbon dioxide. Typically, natural gas at the well head can contain between 3 and 60 mol % nitrogen, between 0.1 and 10 mol % ethane, between 0.1 and 20 mol % C.sub.3 +hydrocarbons and between 0.1 and 20 mol % CO.sub.2 with the balance being methane. Light components, of which nitrogen is typically the major component, must be removed from natural gas to improve the heat content of the gas and to meet pipeline specification.
Compositions of the raw gas and the amount of impurities that can be tolerated in the product generally determine the selection of the most suitable process for purification.
Cryogenic technology, which consists of several process variations, has traditionally been employed to separate nitrogen from natural' gas. The operating principle of the process entails partially or fully liquefying the high nitrogen content feedstream under pressure and at very low temperatures (e.g., as low as -185.degree. C. (-300.degree. F.)). Afterwards, the partially or fully condensed stream is fractionated in one or two columns, which operate in tandem at two different pressures, to separate the feedstream into a rejected nitrogen stream overhead and a high methane content product stream at the bottom. After heat exchanges with the incoming feedstream, the rejected nitrogen stream is used for gas field operation, vented to atmosphere or reused. The bottom nitrogen-depleted product stream is regasified via process heat exchanges, recompressed, and delivered to the battery limits as an upgraded sales gas.
U.S. Pat. No. 4,936,888 relates to the use of cryogenic technology in an integrated dual distillation system for rejecting nitrogen in concentrations as high as 75 mol % or more from gaseous methane in a distillation system employing a high pressure fractionator and a low pressure fractionator. The low pressure fractionator accepts a feed predominantly comprising liquid nitrogen under conditions to produce a high purity nitrogen overhead stream and a high purity methane bottom stream.
U.S. Pat. No. 4,711,651 relates to a process for the separation of a high pressure gas stream such as refining gas in which the starting gas mixture is cooled and separated into a first vapor portion and a first liquid portion which is expanded to an intermediate pressure. The first vapor portion is further cooled and separated into a second vapor portion which may be further processed for ultimate recovery of a methane-rich product gas. Refrigeration is recovered from the mixed intermediate pressure stream.
U.S. Pat. No. 5,051,120 relates to the cryogenic processing of a feed containing nitrogen and methane. The method is used in treating natural gas which is contaminated with nitrogen. An improved stripping column is provided in which the methane product is produced at a higher pressure than would otherwise be possible.
Other process options have used non-cryogenic routes which employ solvent extraction. U.S. Pat. No. 4,832,718 relates to a process for processing a natural gas, a thermal or catalytic cracking gas or refining off-gas to produce a methane-rich product, a nitrogen-rich product, a hydrogen-rich stream or an olefin-rich product therefrom by solvent extraction. The process further relates to extractive flashing and extractive stripping using selective physical solvents. The process employs a solvent to separate N.sub.2 from hydrocarbons by extraction (i.e., absorption) at moderate temperatures (ambient to -30.degree. F.) and under pressure. As an N.sub.2 -rich gas stream and a solvent stream come in intimate contact in a packed column, the solvent dissolves nearly all of the hydrocarbons and a small amount of N.sub.2 in the feed gas stream, leaving behind a rejected gaseous N.sub.2 stream that contains the major fraction of the N.sub.2 previously present in the feed stream and a small amount of unrecoverable hydrocarbons. The dissolved hydrocarbons are subsequently released from the rich solvent by successive pressure reductions. The gases that evolve from the pressure reductions are recompressed and recombined into an upgraded sales gas stream. The regenerated lean solvent stream is recycled to the absorber. After refrigeration recoveries, the rejected N.sub.2 stream is available at a relatively high pressure and can be used for field applications, vented to the atmosphere, or reused.
U.S. Pat. No. 4,623,371 relates to the recovery of hydrocarbons from natural gas containing acidic compounds and from 3-75 mol % nitrogen to provide up to three products: nitrogen-gas product, C.sub.1 -rich gas product, and a C.sub.2.sup.+ liquid product. The process extracts a natural gas stream with a physical solvent to produce a nitrogen stream and a methane-rich solvent stream. The methane-rich solvent stream is flashed to provide a stripped solvent stream and substantially all of the C.sub.1.sup.+ hydrocarbons as a stream of flashed-off-gases. The stripped solvent is recycled to the extraction step.
U.S. Pat. No. 5,047,074 relates to a process for purging nitrogen from natural gas by passing the gas through an absorption column at elevated pressure and contacting the gas with an absorbent consisting primarily of a poly alpha olefin. The non-absorbed gas consisting predominantly of nitrogen is passed out of the absorption column to waste or to recovery. The resulting rich absorbent is desorbed to obtain the desorbed light hydrocarbons and the lean absorbent.
U.S. Pat. No. 5,019,143 relates to the use of a particular group of solvents including paraffins, naphthenes, C.sub.6 -C.sub.10 aromatic compounds and dialkyl ethers of polyalkylene glycol to contact the gas feed mixture in a demethanizing absorber with a reboiler operating between 50 to 400 psig and a temperature of +10.degree. to -40.degree. F. for the separation of ethylene and lighter hydrocarbons from the feed with a distillation column to regenerate the rich solvent. U.S. Pat. Nos. 4,511,381 and 4,526,594 relate to the separation of natural gas from natural gas liquids with physical solvents. Natural gas liquids include hydrocarbons heavier than methane. The physical solvent used for the absorption step in U.S. Pat. No. 4,511,381 is described in terms of a solvent having a relative volatility of methane over ethane of at least 5.0 and a hydrocarbon loading capacity at least 0.25 standard cubic feet of ethane per gallon of solvent. The rich solvent stream is flashed and the gas fraction is compressed, cooled, and condensed to produce the natural gas liquids. U.S. Pat. No. 4,526,594 relates to the selective extraction of a stream of hydrocarbons that are heavier than methane from a natural gas stream with a physical solvent by selectively rejecting the consecutively lowest molecular weight portion of the extracted stream by the use of a second extraction step with a physical solvent to reject hydrocarbons heavier than methane, flashing the resulting rich solvent to separate the selected heavier hydrocarbons, and subsequently de-ethanizing the selected heavier hydrocarbons.
U.S. Pat No. 4,883,514 relates to a process for the removal of nitrogen from nitrogen-rich gases which contain more than 3 mol % nitrogen. The nitrogen-rich gas stream is contacted with a lean oil comprised of paraffins, aromatic or cyclohydrocarbons or mixtures thereof having molecular weight between 75 and 250 at temperature no lower than -40.degree.F. to produce a nitrogen stream as an overhead product and a bottoms methane-rich oil stream. The bottoms methane-rich oil stream is flashed to recover a methane-rich overhead gas product and a lean oil rich bottoms stream, and the lean oil stream is recycled to the absorption step. The patent teaches that the contacting or absorption may take place in a methane extraction column which includes a reboiler and is operable in an extractive stripping mode.
In the art of separation of light hydrocarbons from heavy hydrocarbons, there are a number of mass-transfer operations which can be employed to perform physical separation. The primary difference between the mass-transfer operations is the nature of the phases involved in the operation. In distillation, a vapor phase contacts a liquid phase, and mass is transferred both from the liquid to the vapor and from the vapor to the liquid. The liquid is at its bubble point and the vapor in equilibrium is at its dew point. Mass is transferred simultaneously from the liquid by vaporization and from the vapor by condensation. The net effect is the increase in concentration of the more volatile component in the vapor and an increase in concentration of the less volatile component in the liquid.
Gas absorption involves the transfer of a soluble component of a gas phase into a relatively nonvolatile liquid absorbent. The liquid absorbent is below its bubble point and the gas phase is well above its dew point. Unlike distillation, in gas absorption the liquid and the vapor phases usually do not contain all of the same components.
Stripping or desorption is the opposite of absorption. In desorption, the soluble gas is transferred from the liquid to the gas phase because the concentration in the liquid is greater than that in equilibrium with the gas, and the concentration driving force is opposite to that for absorption.
Generally, when a absorption section is combined with a stripping section to perform a hydrocarbon separation as disclosed in U.S. Pat. No. 3,816,976 in Col. 4, lines 40-68, a stripping medium which has been heated is passed countercurrently to an enriched solvent stream in the stripping section wherein the light component and a portion of the more soluble component is stripped from the enriched solvent. In other combinations of adsorption and stripping, as disclosed in U.S. Pat. No. 4,242,108, the enriched solvent having the absorber is heated and passed to a separate stripper at a higher pressure wherein an absorbed light component is stripped from the solvent. Thus, in both of these examples, heat is added to the stripping section either by heating the enriched solvent or by heating the stripping medium to effectively remove the absorbed light component from the enriched solvent at the pressure of the absorption pressure or a pressure greater than the absorption pressure.
It is an object of this invention to provide a process for the rejection of a light component from a hydrocarbon gas stream using noncryogenic absorption with a hydrocarbon solvent. It is a further object of this invention to provide a process for the rejection of a light component from hydrocarbons which can achieve hydrocarbon recoveries of greater than 90 mol % and provides lower compression requirements, lower capital investment, and lower solvent losses than previous technology.
SUMMARY OF THE INVENTION
By the present invention a process is provided for the removal of a light component from a hydrocarbon gas mixture. The process does not require a selective solvent. Any hydrocarbon or hydrocarbon mixture with an appropriate freeze point can be used as a solvent for the process. Preferably, the freeze point of the solvent will be less than 0.degree. C. (32.degree. F.), more preferably less than -23.degree. C. (-10.degree. F.), and most preferably less than -34.degree. C. (-30.degree. F.). The process operates at mild low temperatures ranging from about -40.degree. C. (-40.degree. F.) to about ambient conditions allowing the equipment to be fabricated from carbon steel without the need for special alloys. In processing natural gas for the removal of nitrogen, hydrocarbon recoveries of greater than 99 mol % are achieved. The purity of the gas product stream or sales gas product is achieved by controlling a reflux stream returned to a purification zone. Recovery of natural gas liquids, C.sub.3 +, when present in the feedstream require no additional processing steps and at least a portion of the natural gas liquids in the feed gas mixture may be used as at least a portion of the solvent. The process of this invention provides an economic and flexible means for rejecting a light component such as nitrogen and/or hydrogen from a hydrocarbon gas feedstream without the need for cryogenic conditions. The process of this invention employs a unique combination of gas absorption and desorption or stripping wherein the stripping medium is not heated and preferably the stripping medium is cooled prior to its introduction to the purification zone.
In accomplishing the foregoing objectives and in accordance with the invention, the invention is a process for the rejection of a light component from a gas mixture comprising the light component and hydrocarbons and includes a series of process steps. The gas mixture is passed to a recovery zone under absorption conditions including an absorption pressure and an absorption temperature. The gas mixture is countercurrently contacted therein with a lean solvent to provide a first rich solvent stream and a rejected light component stream. The first rich solvent stream is passed to a purification zone at desorption conditions including a desorption temperature and a desorption pressure. A light gas stream is cooled to provide a cooled light gas stream and the cooled gas stream is passed to the purification zone. The first rich solvent is countercurrently contacted therein with the cooled light gas stream and a second rich solvent stream is recovered. The second rich solvent stream is passed to a solvent recovery zone to provide a product hydrocarbon stream.
The light component is selected from the group consisting of hydrogen, nitrogen, carbon dioxide, helium, argon, carbon monoxide, and mixtures thereof. The gas mixture is a natural gas or a refining off-gas stream with nitrogen and/or hydrogen. In a further step, lean solvent is recovered from the gas product to minimize solvent losses. At least a portion of the gas product is returned to the purification zone to provide the light gas stream as a reflux stream in the purification zone. The recovery zone and the purification zone may be located in a single absorber column.
In another embodiment, the present invention is a process for the catalytic dehydrogenation of a feedstream comprising hydrocarbons having from 3 to 5 carbon atoms per molecule. The process comprises a series of steps. The feedstream is admixed with a hydrogen recycle stream to form a combined feedstream. The combined feedstream is passed to a catalytic dehydrogenation reaction zone wherein the combined feedstream is contacted with a dehydrogenation catalyst at dehydrogenation conditions to obtain a reaction zone effluent stream comprising hydrogen and hydrocarbons having less than 5 carbon atoms per molecule. At least a portion of the reaction zone effluent stream is passed to a recovery zone under absorption conditions including an absorption pressure and an absorption temperature. The reaction zone effluent is countercurrently contacted with a lean solvent within the recovery zone to provide a first rich solvent stream and a rejected light component stream comprising hydrogen. The first rich solvent stream is passed to a purification zone at desorption conditions including a desorption temperature and a desorption pressure. A light gas stream is cooled to provide a cooled light gas stream and the cooled light gas stream is passed to the purification zone. In the purification zone, the first rich solvent stream is countercurrently contacted with the cooled light gas stream and a second rich solvent stream is recovered. The second rich solvent stream is passed to a solvent recovery zone to provide an effluent hydrocarbon stream. At least a portion of the rejected light component stream is passed to the catalytic dehydrogenation reaction zone as the hydrogen recycle stream.
Other embodiments include the combination of this invention with a catalytic reaction zone to reject light components. Examples of catalytic reaction zones suitable for this application include such petroleum refinery operations and petrochemical processes as catalytic reforming, fluid catalytic cracking, dehydrogenation and combinations thereof.
Other aspects of this invention include passing each of the second rich solvent stream, the high pressure liquid stream and the medium pressure liquid stream to separate hydraulic power recovery turbines to recover hydraulic energy from these streams before passing them to the respective flashing operations.

BRIEF DESCRIPTION OF THE DRAWINGS
FIG. 1 is a schematic flowsheet of the process of the instant invention with a 2-stage compressor.
FIG. 2 illustrates the process with a 3-stage compressor.
FIG. 3 illustrates the process with a 3-stage compressor and lean solvent recovery.
FIG. 4 is a chart which shows the effect of increasing the flow of light gas to the purification zone on the purity of the product gas, and the overall energy used in the process of the instant invention.

DETAILED DESCRIPTION OF THE INVENTION
The process of the present invention is useful for purifying methane- or ethylene-containing streams. In general, typical feedstreams will contain water up to saturation levels, less than 25 mol % C.sub.2 -C.sub.5 hydrocarbons, less than 30 mol % carbon dioxide and less than 2 mol % C.sub.6 and higher hydrocarbons and a number of other impurities such as nitrogen or hydrogen. Natural gas is a common source of such impurity-containing methane, and in natural gas the hydrocarbon impurities are primarily the saturated type such as ethane and propane. Additionally, natural gas can contain other components such as nitrogen, hydrogen, helium and argon, although such other components are not appreciably absorbed in the process of the present invention. Reactor effluent streams can contain an appreciable amount of hydrogen. Typically, the feedstream will contain from about 40 to 97 mol % methane, from about 3 to 60 mol % N.sub.2, from about 0.1 to about 15 mol % hydrogen, from about 0.1 to 10 mol % ethane, from about 0.1 to 20 mol % C.sub.3 +hydrocarbons and from about 0.1 to 20 mol % CO.sub.2.
The term theoretical tray or stage refers to an ideal stage of equilibrium separation. In a theoretical stage the light phase leaving the stage is in vapor/liquid equilibrium with the heavy phase leaving the same tray. The actual number of trays employed in any separation is typically greater than the number of theoretical trays, because vapor/liquid equilibrium typically is not achieved on a single actual tray. The number of stages or trays presented here are expressed in numbers of theoretical trays.
The process of the present invention will hereinafter be described with reference to the drawings. It should be also understood that the representative temperatures and pressure set forth herein, with relation to the description of the drawing and the examples, are illustrative only and are not to be considered limiting. The particular temperatures and pressures used in a particular separation will be dependent upon the nature and composition of the feed stream, upon the particular heat exchange surface area available, and upon the initial tempertures and pressures of the feed stream.
FIG. 1 is a schematic flowsheet of the process of the current invention. With reference to FIG. 1, a feedstream 41 at a feedstream pressure essentially equal to the absorption pressure is charged to a cold box 200 wherein the temperature of the feedstream is reduced to an absorption temperature ranging from about -40 to about -10.degree. C. (about -40 to about 10.degree. F.) and where a portion of any condensed liquid hydrocarbons may be removed in stream 42. Preferably, the condensed liquid hydrocarbons remain in the cooled feedstock. The cooled feedstock is then introduced to a recovery zone 201A via line 43. In the recovery zone, the cooled feedstock is countercurrently contacted with a lean solvent stream 60. The lean solvent enters the top of the recovery zone 201A.
The lean solvent may be any solvent material such as hydrocarbons having 9 to 14 carbon atoms per molecule selected from the group consisting of paraffins, isoparaffins, olefins, mono-olefins, naphthenes, olefinic naphthenes, alkylbenzenes, aromatics and mixtures thereof. In addition, hydrocarbons such as kerosines and diesels, can be used as the lean solvent. The above-mentioned solvent materials can be employed separately or as mixtures. Furthermore, the solvent or mixture can be admixed with natural gas liquids. The amount of natural gas liquids admixed with the lean solvent is preferably maintained at a concentration to depress the freeze point of the mixture of lean solvent to freeze point temperature of less than about -17.degree. C. (0.degree. F.). If desired, a portion of the lean solvent can be withdrawn from the process and flashed at conditions well-known to those skilled in the art to recover a portion of the natural gas liquids in the lean solvent, and return the lean solvent to the process of the Applicant's invention.
A nitrogen-rich gas is withdrawn from the top of the recovery zone in stream 44. The nitrogen-rich gas is then passed from line 44 through a cold box exchanger 202 and released at ambient conditions via line 45 at a temperature ranging from about 10.degree. to about 27.degree. C. (about 50.degree. to about 80.degree. F.). The nitrogen-rich gas may be vented to the atmosphere or used in other operations. A first rich solvent stream 46' is withdrawn from the bottom of the recovery zone 201A and passed to a purification zone 201B wherein the first rich solvent stream is countercurrently contacted with a reflux stream 57, and a second rich solvent stream 46 is withdrawn from the bottom of the purification zone. Preferably, the recovery zone will contain two or more equilibrium stages to achieve the desired separation of the light component from the feedstream. The purification zone will contain one or more equilibrium stages to achieve the desired displacement of light component from the second rich solvent. Any light component displaced from the first rich solvent stream is returned to the recovery zone in line 44'.
The second rich solvent stream is passed by line 46 to a first flash zone of a series of consecutive flash zones. Preferably, there will be between 1 and 4 separate flash zones and most preferably 2 to 3 separate flash zones. The number of separate flash zones selected will depend upon the feed pressure. The first flash zone will operate at the highest pressure. The temperature of the consecutive flashing zones ranges from about -46.degree. to 0.degree. C. (about -50.degree. to about 32.degree. F.). The second rich solvent stream is passed to a high pressure flash zone 204 wherein a temperature and a pressure effective to separate the second rich solvent stream into a high pressure off-gas 47 and a high pressure liquid 49 is maintained. Preferably, the pressure of the high pressure flash zone is maintained between about 1/5 and 1/2 of the pressure of the sales gas product 58, if the pressure of sales gas product is less than the feed gas pressure. If the pressure of the sales gas product is greater than the pressure of the feedstream, then the preferred pressure of the high pressure flash is 1/2 to 1/5 the pressure of the feedstream. The high pressure liquid is passed by line 49 to a medium pressure flash zone 206 wherein the high pressure liquid is flashed at conditions effective to produce a medium pressure off-gas in line 50 and a medium pressure liquid in line 51. Preferably, the pressure of the medium pressure flash zone is maintained between about 1/5 to about 1/2 of the pressure of the high pressure flash zone. The high pressure off-gas in line 47 is passed to a cold box exchanger 205 to recover refrigeration wherein the high pressure off-gas is heated from a temperature of between about -34.degree. and about -29.degree. C. (about -30.degree. and -20.degree. F.) to a temperature of between about 15.degree. and about 27.degree. C. (about 60.degree. and about 80.degree. F.) and passed by line 48 to a second stage 212B of a two-stage compressor. The medium pressure off-gas in line 50 at a temperature of about between about -37.degree. and about -32.degree. C. (about -35.degree. and about -25.degree. F.) is passed to a cold box exchanger 207 to recover refrigeration wherein the medium pressure off-gas is heated to a temperature of between about 15.degree. and about 27.degree. C. (about 60.degree. and about 80.degree. F.) and passed to a first stage 212A of the two-stage compressor 212 via line 53. A sales gas product at high pressure is withdrawn in line 58 from the two-stage compressor. The pressure of the sales gas product is typically dictated by the pressure of the pipeline to which the gas is returned. Pipeline pressures typically range from 4500 to about 13,800 kPa (650 to about 2000 psia). Typically, the high pressure product gas temperature will range between about 32.degree. and about 43.degree. C. (about 90.degree. and 110.degree. F.). Returning to the medium pressure flash zone 206, the medium pressure liquid in line 51 is passed to a low pressure flash zone 208 to provide a low pressure off-gas which is withdrawn in line 52, and a lean solvent stream which is withdrawn in line 55. The lean solvent in line 55 is passed to cold box exchanger 203 to reduce the temperature of the lean solvent to a range between ambient temperature and about -40.degree. C. (-40.degree. F.). The cooled lean solvent is passed to the top of the recovery zone via line 60. The low pressure flash is maintained at a temperature and pressure effective to separate the medium pressure liquid into the low pressure off-gas and the lean solvent. Preferably, the low pressure flash zone is maintained at a pressure of about 1/5 to about 1/2 the pressure of the medium pressure flash zone. The low pressure off-gas, at a temperature of between about -37.degree. and about -32.degree. C. (about -35.degree. and about -25.degree. F.), is passed in line 52 to a cold box exchanger 209 to recover refrigeration wherein the low pressure off-gas is heated to a temperature between about 15.degree. and about 27.degree. C. (about 60.degree. and about 80.degree. F.) and withdrawn in line 54. The low pressure off-gas is then passed to a three stage compressor 211 to compress the low pressure off-gas and provide a compressed light gas stream in line 56. The temperature of the compressed light gas stream in line 56 is typically between about 32.degree. and about 49.degree. C. (about 90.degree. and about 120.degree. F.). The compressed light gas stream is then passed to cold box 210 where the temperature of the light gas stream is reduced to a temperature of between about -32.degree. to about -26.degree. C. (about -25.degree. and -15.degree. F.) and the light gas stream is recycled to the purification zone 210B by line 57.
In accordance with the present invention, the absorption pressure is generally from about 345 to about 13,800 kPa (about 50 to about 2000 psia) and preferably from about 1030 to about 6900 kPa (about 150 to about 1000 psia), and most preferably from about 2070 to about 4200 kPa (about 300 to about 600 psia). Suitable operating temperatures are generally within the range of from about -40.degree. to about -0.degree. C. (about -40.degree. to about 32.degree. F.).
Typically, the feedstock temperature ranges between about 16 and about 32.degree. C. (about 60.degree. and about 100.degree. F.). The cooled feedstock temperature ranges between ambient and about -40.degree. C. (about -40.degree. F.) before it is charged to the recovery zone. The lean solvent temperature charged to the top of the recovery zone 201A is preferably at a temperature ranging from about ambient to about -40.degree. C. (about -40.degree. F.), wherein in the ambient temperature ranges from about 15.degree. to about 43.degree. C. (about 60.degree. to about 110.degree. F.).
The lean solvent preferred for the invention is a hydrocarbon having between 10 and 14 carbon atoms per molecule selected from the group consisting of natural gas liquids, paraffins, iso-paraffins, olefins, mono-olefins, naphthenes, olefinic naphthenes, alkylbenzenes, aromatics, kerosines, and diesels, and mixtures thereof. The criteria for using these solvents in the process of the instant invention are that the solvents are C.sub.9 -C.sub.14 hydrocarbons, preferably C.sub.10 -C.sub.14 hydrocarbons or mixtures thereof with boiling points of about 300.degree. F. or greater, and freezing points of about 0.degree. C. (32.degree. F.) or lower. Kerosines and diesels and petroleum distillates and paraffinic solvents with an initial boiling point above about 150.degree. C. (300.degree. F.) also may be used as solvents in the instant invention. If the freezing point of the solvent is higher than 0.degree. F., a portion of natural gas liquids comprising light hydrocarbons, ranging from C.sub.2 to C.sub.6 hydrocarbons may be admixed with the solvent to depress the freezing point of the solvent to less than about 0.degree. C. (32.degree. F.), the desired temperature range of operation. The more preferred lean solvent is a paraffin hydrocarbon having between 10 and 14 carbon atoms per molecule in pure form or in mixtures thereof with natural gas liquids. Preferably, the lean solvent in pure form or in mixtures has a freeze point of less than -17.degree. C. (0.degree. F.).
FIG. 2 is a schematic flowsheet of the nitrogen rejection system of the current invention wherein a portion of the product gas is used to reflux the purification zone. A feedstock in line 120 is passed to a cold box exchanger 300 at a temperature of between about 15.degree. and about 32.degree. C. (about 60.degree. and about 90.degree. F.). In the cold box exchanger, the feedstock is cooled to a temperature ranging between about -40.degree. to about -29.degree. C. (about -40.degree. to about -20.degree. F.) to condense any liquids from the feedstock. A portion of the condensed liquids may be withdrawn in line 121, but preferably, at least a portion of the condensed liquid and the cooled feedstock are passed via line 122 to the bottom of a recovery zone 301A. In the recovery zone, the feedstock is countercurrently contacted with a lean solvent stream which is introduced at the top of the recovery zone at a temperature of between about -40.degree. to about -29.degree. C. (about -40.degree. and about -20.degree. F.) in line 133. A nitrogen-rich gas is withdrawn via line 123 and passed to cold box exchanger 302 wherein the temperature is raised from about -40.degree. to about -29.degree. C. (about -40.degree. to about -20.degree. F.) to a temperature of about 15.degree. to about 27.degree. C. (about 60.degree. to about 80.degree. F.) and released to the atmosphere via line 124. A first rich solvent stream is withdrawn from the recovery zone 301A and passed to a purification zone 301B via line 125'. In the purification zone the first rich solvent stream is countercurrently contacted with a light gas stream 138 which enters at the bottom of the purification zone to displace an additional amount of nitrogen-rich gas. The additional amount of nitrogen-rich gas stream comprising nitrogen and a portion of the light gas stream is withdrawn from the purification zone via line 123' and passed to the recovery zone. A second rich solvent stream is withdrawn from the bottom of the purification zone in line 125 and passed to a first flash zone in a series of consecutive flash zones. The second rich solvent stream at a temperature of between about -34.degree. and about -12.degree. C. (about -30.degree. to about 10.degree. F.) is passed to the first or high pressure flash zone 304 via line 125. In the high pressure flash zone 304, conditions are maintained at a temperature and pressure effective to produce a high pressure off-gas which is withdrawn in line 126 at a temperature typically ranging from about -40.degree. to about -15.degree. C. (about -40.degree. to about 5.degree. F.) and a pressure of about 1/5 to 1/2 the pressure of the product gas 140, to provide a high pressure liquid which is withdrawn via line 128. The high pressure liquid is passed to a medium pressure flash zone 306 which is maintained at a temperature and a pressure effective to produce a medium pressure off-gas which is withdrawn via line 129 and a medium pressure liquid which is withdrawn via line 130. The temperature of the medium pressure flash zone ranges from about -43.degree. to about -15.degree. C. (about -45.degree. to about 5.degree. F.) and a pressure of about 1/5 to 1/2 the pressure of the high pressure flash zone. The medium pressure liquid is passed to a low pressure flash zone 308 which is maintained at a temperature and a pressure effective to produce a low pressure off-gas which is withdrawn via line 131 and to provide a low pressure liquid, or lean solvent stream which is withdrawn via line 132. The temperature of the low pressure flash zone ranges from about -43.degree. to about 15.degree. C. (about -45.degree. to about 5.degree. F.) and a pressure of about 1/5 to 1/2 of the pressure of the medium pressure flash zone. The relative pressures in the flash zones are maintained at these ratios to maximize gas compression efficiencies between the consecutive flash zones. The lean solvent stream is passed via line 132 to cold box exchanger 303 wherein the lean solvent is cooled and passed to the recovery zone in line 133. The high pressure off-gas is passed from line 126 to cold box exchanger 305 wherein the high pressure off-gas is heated from a temperature ranging from between -25.degree. and -30.degree. F. to a temperature ranging between about 15.degree. to about 27.degree. C. (about 60.degree. to about 80.degree. F.) to recover refrigeration, and is withdrawn in line 127. Similarly, the medium pressure off-gas in line 129 is heated in cold box exchanger 307 to a temperature of between about 15.degree. to about 27.degree. C. (about 60.degree. to about 80.degree. F.) to recover refrigeration and withdrawn in line 134. Likewise, the low pressure off-gas in line 131 is heated in cold box exchanger 309 to a temperature of between about 15.degree. to about 27.degree. C. (about 60.degree. to about 80.degree. F.) to recover refrigeration and is withdrawn in line 135. The high pressure off-gas stream in line 127 is passed to a third stage 311C of a 3-stage compressor 311, the medium pressure off-gas stream in line 134 is passed to a second stage 311B of 3-stage compressor 311 and the low pressure off-gas in line 135 is passed to a first stage 311A of 3-stage compressor. The three stage compressor compresses these off-gas streams in stages 311A, 311B and 311C, respectively, to produce a compressed off-gas stream which is withdrawn in line 136. A portion of the compressed off-gas stream is withdrawn in line 140 as a product gas, and at least a portion of the compressed off-gas is passed via line 137 as a light gas stream to be returned to the purification zone. Preferably, the portion of the compressed off-gas which is countercurrently contacted with the first rich solvent stream in the purification zone ranges from 1 to 99 mol %, preferably 1 to 80 mol % and most preferably 1 to 60 mol % of the feedstock passed to the recovery zone depending on the amount of nitrogen in the feed and the degree of purification required. The portion of the compressed off-gas, or light gas stream returned to the purification zone or reflux can be expressed as a ratio of the molar flow of the light gas stream, to the feed molar flow. Preferably, the reflux-to-gas ratio in the purification zone should be maintained at a value above 0 and less than 0.8, preferably the reflux-to-gas ratio will be maintained at a level between 0.05 and about 0.5. This light gas stream is passed via line 137 to cold box exchanger 310 wherein the temperature is reduced from a temperature of about 32.degree. to about 49.degree. C. (about 90.degree. to 120.degree. F.) to a temperature of between about -32.degree. and about -26.degree. C. (about -25.degree. to about 15.degree. F.) and withdrawn via line 138. The light gas stream is then passed via line 138 to the bottom of the purification zone 301B as the reflux stream to countercurrently contact the first rich solvent stream.
It will be understood by one skilled in the art that the cold box exchangers illustrated in FIGS. 1 and 2 can be combined to take advantage of heat exchange benefits throughout the process.
FIG. 3 illustrates a schematic flowscheme of the present invention wherein some of the cold box exchangers have been combined to effect an efficient operation, and hydraulic power recovery turbines have been utilized to converse energy. The scheme uses a 3-stage compressor wherein at least a portion of the high pressure liquid from the high pressure flash zone is used in a solvent recovery step to remove lean solvent from a combined stream of the compressed low and medium pressure off-gases in the high pressure flash zone to reduce overall solvent losses from the process. The first and second stages of the compressor are used to provide the combined compressed low and medium pressure off-gas stream. The combined low and medium pressure off-gas stream is contacted with the high pressure liquid in the high pressure flash zone. In addition, a portion of the product gas is used to reflux the purification zone.
FIG. 3 is a schematic flowsheet of a nitrogen rejection scheme of the present invention having a recovery zone and the purification zone in a single absorber column, wherein the absorber column has an upper recovery zone and a lower purification zone, and using the high pressure flash to remove solvent from the low pressure and medium pressure off-gas. A natural gas feedstock at a temperature of between about 26.degree. and about 49.degree. C. (about 80.degree. and about 120.degree. F.) and a pressure of between about 690 to about 2410 kPa (about 100 to about 350 psia) is passed to the process via line 1. The feedstock is split into at least three streams to take advantage of heat and refrigeration recovery opportunities in the process. A first portion of the feedstock is passed via line 29 to exchanger 111 wherein the first portion is cross-exchanged with the medium pressure off-gas stream in line 17 to cool the feedstream and the cooled portion of the feedstream is returned via line 30 to line 3. A second portion of the feedstream in line 28 is exchanged with the high pressure off-gas stream 9 in exchanger 100 to produce a cooled feedstream 2 which is subsequently combined in line 3 with the remaining portions of the feedstream. A third portion of the feedstream is passed via line 31 to exchanger 101 wherein the feedstream is cooled to produce stream 31' which is subsequently combined with the other portions of the feedstream in line 3. Stream 3 is passed to a refrigeration unit 103 and a chilled feedstream is passed to a recovery zone 102A of an adsorber 102 via line 4. The chilled feedstream enters the absorber at a point between the recovery zone 102A and the purification zone 102B. The point at which the chilled feedstream enters the adsorber will depend upon the degree of separation to be obtained in each section. As a minimum, there will be at least two theoretical trays in the recovery zone and at least one theoretical tray in the purification zone. Within the recovery zone 102A, the chilled feedstream is countercurrently contacted with a lean solvent stream 27 introduced at the top of the recovery zone. A nitrogen-rich stream is withdrawn via line 5 and passed to exchanger 101 wherein the nitrogen-rich stream is heated to ambient conditions to recover refrigeration and released to the atmosphere via line 6. A first rich solvent stream in line 7' is withdrawn from the recovery zone 102A and passed to the purification zone 102B wherein any excess nitrogen is displaced from the rich solvent stream by the countercurrent contacting of the first rich solvent stream with a reflux stream 15 comprising light gas introduced at the bottom of the purification zone. The displaced nitrogen stream comprising nitrogen and light gases is returned to the recovery zone via line 5'. A second rich solvent stream is withdrawn from the bottom of the purification zone via line 7 and passed to a hydraulic power recovery turbine 104. The hydraulic power recovery turbine 104 recovers some of the hydraulic energy in the second rich solvent stream by reducing the pressure from the pressure of the purification zone which operates in a pressure range similar to the feedstream pressure to a pressure of the high pressure flash zone 105. The second rich solvent stream, now reduced in pressure, is passed via line 8 to high pressure flash zone 105. The high pressure flash is maintained at conditions of temperature and pressure effective to produce a high pressure liquid and a high pressure off-gas stream. The second rich solvent enters the high pressure flash zone and provides an off-gas stream which is removed via line 9 and a high pressure liquid which flows over an effective number of trays within the high pressure flash zone and is contacted with a compressed gas stream in line 21. The trays in the high pressure flash zone provide a contacting area with the high pressure flash zone to facilitate the removal of lean solvent from the compressed gas stream. Preferably 1 to 5 theoretical trays can be effective to reduce solvent losses from the high pressure flash zone. The compressed gas stream is a combination of a low pressure and a medium pressure off-gas stream which comprises solvent. The high pressure liquid stream is then passed via line 10 to a second hydraulic power recovery turbine 106 which recovers hydraulic energy from the high pressure liquid as the pressure of the high pressure liquid is reduced to the pressure of the medium pressure flash zone. The high pressure liquid is then passed from the second hydraulic power recovery turbine via line 16 to the medium pressure flash zone 107. The medium pressure flash zone is maintained at a temperature and a pressure effective to produce a medium pressure off-gas stream comprising solvent which is withdrawn out of line 17 and a medium pressure liquid which is withdrawn in line 18. The medium pressure liquid is passed to a third hydraulic power recovery turbine 108 which recovers hydraulic energy from the medium pressure liquid, and the medium pressure liquid, is subsequently passed via line 22 to a low pressure flash zone 109. The low pressure flash zone is maintained at a temperature and pressure effective to provide a low pressure off-gas comprising solvent which is removed via line 23 and a lean solvent which is removed via line 25. Lean solvent is passed via line 25 to lean solvent pump 110. The lean solvent is then pumped via line 26 to refrigeration unit 103 wherein the temperature of the lean solvent is reduced, and a cooled lean solvent is passed in line 27 to the top of the recovery zone 102A.
The medium pressure gas is passed in line 17 to exchanger 111 wherein the medium pressure gas is heated to recover refrigeration and conveyed via line 19 to a second stage 115B of a 3-stage compressor 115. The low pressure off-gas in line 23 is passed to exchanger 112 to recover refrigeration which heats the low pressure off-gas. The heated low pressure off-gas is passed to the first stage 115A of the 3-stage compressor via line 24. A combined low pressure and medium pressure off-gas stream comprising lean solvent is passed from the second stage of the compressor 115B via line 20 to exchanger 100 wherein the combined gas is cooled. The cooled combined gas is then passed to the bottom of the high pressure flash zone 105 via line 21. The high pressure off-gas in line 9, essentially free of solvent, is passed to exchanger 100 wherein the high pressure off-gas is heated to a temperature ranging from about 15.degree. to about 27.degree. C. (about 60.degree. to about 80.degree. F.) and passed via line 11 to the third stage of the 3-stage compressor 115C. A compressed product gas is withdrawn from the third stage of the 3-stage compressor in line 12. At least a portion of the product gas is withdrawn via line 13 and passed to exchanger 112 wherein the portion is cooled from a temperature of about 32.degree. to about 49.degree. C. (90.degree. to about 120.degree. F.) to a temperature of about -40.degree. to about -23.degree. F. (about -40.degree. to about -10.degree. F.) in line 15 as the reflux stream. The reflux stream in line 15 is passed to the bottom of the purification zone 102B. The remaining portion of the product gas from line 12 is withdrawn as product gas in line 14.
It was discovered that a minimal amount of solvent was carried over as equilibrium losses in the high pressure flash zone, while increasing amounts of solvent were carried over in the medium and low pressure flash zones of the consecutive flash scheme. As a result, 1 to 3 theoretical trays, preferably at least 1 theoretical tray, was provided in the high pressure flash zone 105. The second rich solvent stream was introduced in the high pressure flash zone above the trays. The second rich solvent stream flashes and the resulting high pressure liquid is countercurrently contacted on the trays with a combined off-gas stream comprising the compressed low pressure and medium pressure off-gas streams which were introduced to the high pressure flash zone at a point below the trays. It is believed that this countercurrent contacting in the high pressure flash zone acts to remove solvent from the combined off-gas stream at the conditions of the high pressure flash to provide a high pressure off-gas essentially free of lean solvent. Preferably the high pressure off-gas essentially free of solvent, will contain less than 300 ppm, and more preferably less than 100 ppm, and most preferably less than 50 ppm solvent which permits the process scheme to reject nitrogen from the natural gas stream while maintaining the amount of solvent in the product gas and the overall solvent loss at a very low level, preferably less than 300 ppm and most preferably less than 100 ppm.
The hydraulic power recovery turbines 104, 106 and 108 are used to recover hydraulic energy from the second rich solvent stream, the high pressure liquid, and the medium pressure liquid before these streams are subsequently flashed in the consecutive flash zones. Energy recovered from these hydraulic power recovery turbines can be used to partially drive the lean solvent pump 110 in order to offset some of the hydraulic power requirements for the lean solvent pump.
Refrigeration unit 103 is preferably a propane refrigeration unit operating in a range of from about -46.degree. to about -12.degree. C. (about -50.degree. to about 10.degree. F.). In order to minimize overall processing costs, it is preferred to use equipment which could be constructed of carbon steel and employ a propane refrigeration system.
The process for the dehydrogenation of a hydrocarbon feed comprising C.sub.3 to C.sub.5 hydrocarbons consists of a reactor section, a product recovery section and a catalyst regeneration section. A detailed description of the dehydrogenation process including catalyst, processing conditions and product recovery for the production ethers from a dehydrogenation zone effluent is provided in U.S. Pat. No. 4,806,695, which is hereby incorporated by reference. Preferably the hydrocarbon feed is selected from the group consisting of propane, isobutane, isopentane and mixtures thereof. The hydrocarbon feed is admixed with a small amount of a hydrogen-rich gas and an unconverted recycle stream to provide a combined feed. The combined feed is heated and charged to the dehydrogenation reaction zone. In the dehydrogenation reaction zone, the admixture is contacted with a catalyst and at least a portion of the hydrocarbons are converted to the corresponding mono-olefins. This reaction is endothermic and heat must be provided to maintain conversion. At least a portion of the catalyst is removed and continuously reconditioned and returned to the reaction zone to maintain catalyst activity. The reactor effluent from the dehydrogenation reaction zone is cooled by cross exchange with the combined feed. The cooled reactor effluent may be passed to a compressor to raise the pressure of the reactor effluent. The cooled reactor effluent is passed to the product separation zone wherein the reactor effluent is separated into a hydrogen stream and a mono-olefin product stream.
The process of this invention is employed to reject the hydrogen and recover at least 95% and most preferably at least 99% of the hydrocarbon heavier than C.sub.3 hydrocarbons. The reactor effluent stream as a gas mixture comprising hydrogen and hydrocarbons is passed to an absorber column and a recovery zone therein. In the recovery zone, the gas mixture is countercurrently contacted with a lean solvent stream to provide a first rich solvent stream and to provide a rejected light component stream comprising hydrogen. The first rich solvent stream is passed to a purification zone in the absorber column at a point below the point where the gas mixture was introduced. A light gas stream is cooled to provide a cooled light gas stream and passed to the purification zone. The light gas stream is cooled to a temperature of between the absorption temperature and the desorption temperature, preferably the light gas stream will be cooled from about 3.degree. C. (5.degree. F.) to about 40.degree. C. (100.degree. F.) in order to be a vapor stream with a temperature in the range of about -40.degree. C. (-40.degree. F.) to about 40.degree. C. (100.degree. F.). Typically the light gas stream will comprise hydrocarbons having from 1 to 5 carbon atoms per molecule and depleting the light component or hydrogen. In the purification zone, the first rich solvent is countercurrently contacted with the cooled light gas stream to remove residual hydrogen from the first rich solvent and to provide a second rich solvent stream. The second rich solvent stream is passed to a solvent recovery zone wherein the lean solvent is separated from the hydrocarbon effluent stream. The hydrocarbon effluent now depleted of the light component is recovered. The light gas stream may be an external stream depleted in the light component to be rejected or at least a portion of the hydrocarbon effluent. At least a portion of the hydrocarbon effluent stream is cooled and recycled to the absorber column as the cooled light gas stream.
The hydrocarbon effluent stream may be recovered from the lean solvent by any means known to those skilled in the art such as successive multiple flashing steps, fractionation, selective solvent extraction, and combinations thereof. The lean solvent is returned to the absorber column. The hydrocarbon effluent stream comprising mono-olefins is admixed with a water soluble alcohol which preferably has less than 5 carbon atoms per molecule. The alcohol can be chosen from methanol, ethanol, primary and secondary propanol, the various butanols, and other alcohols. However, the preferred class of alcohols are C.sub.4 -minus aliphatic monocyclic alcohols with methanol and then ethanol being particularly preferred. The majority of the description of the invention is presented in terms of the reaction of isobutene with methanol since these are the preferred feed materials and this is the commercially predominant reaction. However, it is not intended to thereby lessen the scope of the inventive concept. This is especially true since there have been predictions that the expected large demand for ethers as anti-knock additives will lead to the use of large amounts of ethanol produced by fermentation in the etherification processes.
The hydrocarbon effluent material may therefore be substantially pure normal butane, normal pentane, or a mixture of the corresponding isomeric and normal hydrocarbons. The preferred hydrocarbon effluent is a mixture of isobutane and normal butane such as is available from several sources in a petroleum refinery or as is available as field butanes. This variety of possible feed materials allows the production of a wide variety of ethers other than the preferred MTBE including methyl tertiary amyl ether, ethyl tertiary amyl ether, and ethyl tertiary butyl ether.
The etherification effluent is obtained by the catalyzed reaction of isobutylene with a stoichiometric excess of methanol in the liquid phase at a temperature of about 65.degree. to about 90.degree. C. A detailed description of processes, including catalyst, processing conditions, and product recovery for the production of MTBE from isobutylene and methanol are provided in U.S. Pat. Nos. 2,720,547 and 4,219,678. The preferred process is described in a paper presented at the American Institute of Chemical Engineers 85.sup.th National Meeting on Jun. 4-8, 1978 by F. Obenaus et al. The above references are hereby incorporated by reference. Other etherification processes of interest are the production of tertiary amyl ether (TAME) by reacting C.sub.5 iso-olefins with methanol and the production ethyl tertiary butyl ether (ETBE) by reacting C.sub.4 iso-olefins with ethanol. The isobutylene reactant is introduced into the etherification reactor as a mixture of trans butene, butadiene, isobutane and n-butane. The isobutylene constitutes about 45 mol. % of the C.sub.4 hydrocarbon mixture, and is the only C.sub.4 species which reacts with the methanol under the present conditions. The molar ratio of methanol to isobutylene is from about 1.05:1 to 1.5:1. The effluent from the etherification reactor comprises the other product MTBE, unreacted methanol, unreacted C.sub.4 's and small to trace amounts of dimethylether, TBA, and other reaction by-products. This etherification reactor effluent is passed to a distillation unit wherein the MTBE product is recovered from the bottom. The overhead effluent from the distillation unit is typically water washed or passed to a separate adsorber unit to recover the bulk of the unreacted methanol from the etherification reaction and return the alcohol to the etherification reaction zone. The ether is passed to a gasoline blending for the production of high octane unleaded gasoline or sold as a high octane gasoline additive.
Typically the reactor effluent from the dehydrogenation reaction zone is dried to remove water to a level which would prevent ice and/or hydrate formation at cryogenic conditions, compressed, and passed to a cryogenic separation zone. In the cryogenic separation zone, the hydrogen is rejected as a hydrogen-rich gas stream and a portion of the hydrogen-rich gas stream is returned to be admixed with the feed to the dehydrogenation reaction zone. In the present invention, the reactor effluent for the dehydrogenation zone is dried in a process step such as glycol dehydration to the extent necessary to prevent hydrate formation at non-cryogenic conditions. The dried reactor effluent is cooled to absorption conditions and is passed to the absorption column without the need for compression. At least a portion of the hydrogen rejected is returned to be admixed with the feed to the dehydrogenation zone. The hydrogen is rejected at the pressure of the reactor effluent which ranges from about 140 kPa (20 psia) to about 1400 kPa (200 psia), and preferably the pressure of the reactor effluent is between about 310 kPa (45 psia) and about 1400 kPa (200 psia). The hydrocarbon effluent stream recovered from the lean solvent may be withdrawn as a vapor stream or further compressed to the liquid phase for subsequent reaction or separation. The use of the present invention eliminates the cryogenic separation zone resulting in significant capital cost and operating cost savings.
The invention will be further clarified by a consideration of the following examples based on engineering design simulation models, which are intended to be purely exemplary of the use of the invention.
EXAMPLE I
Example 1 is based on the process for separating nitrogen from a mixture of natural gas and nitrogen in the scheme depicted in FIG. 3 wherein the purification zone was refluxed with a portion of the product gas. The feedstream was a natural gas available at a pressure of 2310 kPa (335 psia) with the following composition (mol %):
______________________________________ Nitrogen 13.71 Methane 85.99 Ethane 0.30 Water 13 ppm______________________________________
The above composition of the natural gas has a C.sub.2.sup.+ content which is low enough that natural gas liquids do not condense out from the feed in the nitrogen rejection process. If the concentration of C.sub.2.sup.+ material in the natural gas is greater than this amount, some prior art processes would require an additional liquid recovery step to remove the natural gas liquids. The instant invention does not require any additional recovery steps to remove condensed liquids. Any condensed liquids in the instant process become part of the solvent and aid in the separation.
The objective of the process was to reduce the nitrogen content of the product gas to less than 3 mol % while maintaining a high recovery of the nitrogen-depleted product gas. The results of an engineering design calculation based on a small natural gas upgrading facility with an inlet gas capacity of 0.54 million Nm.sup.3 per day (20 million SCF per day) are shown in Table 1 for a C.sub.9 paraffin solvent.
Using the C.sub.9 paraffin solvent, a solvent to feedstock gas molar ratio of 6.0 at a reflux-to-feedstock gas ratio of 0.07, the amount of nitrogen in the sales gas product was reduced to 1.7 mol %, while recovering 99.3 mol % of the natural gas in the feedstock. The separation required at least eight equilibrium stages in the recovery zone and at least two equilibrium stages in the purification zone. The energy requirements of the process expressed in BHP per mol of gas recovered was 1.35. The results of Example I indicated a total loss of solvent at a level of 498 ppm. No solvent recovery step was provided in Example I.
TABLE 1______________________________________NITROGEN REJECTION FROM NATURAL GASPRODUCT USED TO PURIFY RICH SOLVENTPROCESS VARIABLES C.sub.9______________________________________Solv/Gas Ratio (molar) 6.0Reflux/Gas Ratio (molar) 0.07Natural Gas Recovered (mol %) 99.3N.sub.2 In Nat. Gas Rec. (mol %) 1.7N.sub.2 In Gas Rejected (mol %) 95.3HP Flash Pressure (psia) 140.0MP Flash Pressure (psia) 42.0LP Flash Pressure (psia) 12.0Gas Compressor (stages) 3Absorber Stages (recovery/purification) 8/2Total Solvent Losses (ppm-wt) 498Solvent Temperature (.degree.C.) -34Gas Compression (BHP) 1945Propane Refrigeration (BHP) 188LS Pump (actual-recov) (BHP) 403TOTAL BHP REQUIRED BY PROCESS 2536BHP PER LB-MOL OF GAS RECOVERED 1.35______________________________________
EXAMPLE II
Example II is based on the feedstock of Example I and a C.sub.9 paraffin solvent and the flowscheme shown in FIG. 3 wherein the low pressure and medium pressure off-gases are compressed and returned to the high pressure flash zone in a solvent recovery step and at least a portion of the nitrogen-depleted gas product is used to reflux the purification zone. The results are shown in Table 2 in the column headed C.sub.9 paraffin. Although there was little change in the overall energy requirements from the results of Example I, in Example II the total solvent losses were reduced to 300 ppm from the 498 ppm of Example I which represented about a 40% reduction in the solvent losses compared to Example I.
TABLE 2______________________________________COMPARISON OF REFLUX SCHEMES INNITROGEN REJECTION PROCESSPROCESS VARIABLES C.sub.9 Paraffin______________________________________Solv/Gas Ratio (molar) 6.0Reflux/Gas Ratio (molar) 0.07Natural Gas Recovered (mol %) 99.3N.sub.2 In Nat. Gas Rec. (mol %) 1.7N.sub.2 In Gas Rejected (mol %) 95.3HP Flash Pressure (psia) 140.0MP Flash Pressure (psia) 42.0LP Flash Pressure (psia) 12.0Gas Compressor (stages) 3Absorber Stages (recovery/purification) 8/2Total Solvent Losses (ppm-wt) 300Gas Compression (BHP) 1888Propane Refrigeration (BHP) 233LS Pump (actual-recov) (BHP) 402TOTAL BHP REQUIRED BY PROCESS 2523BHP PER LB-MOL OF GAS RECOVERED 1.34______________________________________
EXAMPLE III
Example III is based on the feedstock of Example I and the flowscheme shown in FIG. 3. Table 3 shows the performance of the process using heavier, higher molecular weight solvent compared to the C.sub.9 paraffinic solvent of Example II. The performance for C.sub.10, a paraffinic solvent with 10 carbon atoms per molecule and C.sub.12, a mono-olefin with 12 carbon atoms per molecule are shown in Table 3. The results in the column headed C.sub.10 showed that the solvent losses could be reduced to 94 ppm, a 69% reduction compared to the C.sub.9 paraffin case with a 300 ppm solvent loss. The C.sub.12 solvent resulted in a 96% reduction in solvent losses with a solvent loss of 13 ppm.
Thus, the use of the higher molecular weight, lower volatility solvent significantly reduced the solvent losses while slightly reducing the solubility of the natural gas in the solvent in a scheme with a solvent recovery zone. The apparent reduction in gas solubility caused an increase in the solvent-to-gas ratio from 6 to 6.2 (a 3% increase) for C.sub.10 and from 6 to 7.1 (an 18% increase) for C.sub.12. Thus, the resulting loss in selectivity for the absorption of the light hydrocarbons from the use of the heavier solvent resulted in a slight increase of nitrogen in the hydrocarbon gas recovered. The resulting energy requirements remained essentially unchanged for C.sub.10. However, the energy requirement for the C.sub.12 solvent increased 7.5% over the C.sub.9 paraffin case.
TABLE 3__________________________________________________________________________EFFECT OF INCREASING SOLVENT MOLECULARWEIGHT OF NITROGEN REJECTION PROCESS.sup.1PROCESS VARIABLES C.sub.9 C.sub.10 C.sub.12__________________________________________________________________________Solv/Gas Ratio (molar) 6.0 6.2 7.1Reflux/Gas Ratio (molar) 0.07 0.07 0.07Natural Gas Recovered (mol %) 99.3 99.3 99.4N.sub.2 In Nat. Gas Rec. (mol %) 1.7 1.7 1.9N.sub.2 In Gas Rejected (mol %) 95.3 95.3 95.4HP Flash Pressure (psia) 140.0 140.0 140.0MP Flash Pressure (psia) 42.0 42.0 42.0LP Flash Pressure (psia) 12.0 12.0 12.0Gas Compressor (stages) 3 3 3Absorber Stages (recovery/purification) 8/2 8/2 8/2Total Solvent Losses (ppm-wt) 300 94 13Solvent Temperature (.degree.C.) -34 -34 -34Gas Compression (BHP) 1888 1886 1913Propane Refrigeration (BHP) 233 233 233LS Pump (actual-recov) (BHP) 402 456 592TOTAL BHP REQUIRED BY PROCESS 2523 2575 2738BHP PER LB-MOL OF GAS RECOVERED 1.35 1.37 1.46__________________________________________________________________________ .sup.1 FIG. 3 with solvent recovery
EXAMPLE IV
This example is based on the feedstock of Example I and the flowscheme depicted in FIG. 3 to illustrate the advantage of returning a reflux stream to the purification zone on the performance of the process. The absorber for the Refluxed case has an additional two theoretical stages in the purification zone which are not present in the No-reflux case. The results of this comparison using the C.sub.9 paraffinic solvent are shown in Table 4 under the heading Refluxed and No-reflux. Both schemes produced a product gas with a natural gas recovery in excess of 99 mol % and had a solvent loss of 300 ppm. The no-reflux performance showed a slightly lower gas compression requirement (3%) at 1.30 BHP per mole of gas recovered. The nitrogen impurity in the product gas of the no-reflux case was 2.7%, compared to the nitrogen impurity of 1.7% in the refluxed case, representing an improvement of 40% in the reduction of nitrogen in the gas product. Therefore, without the purification zone, represented by the additional stages and the reflux operation, the process would have limited ability to control the purity of the gas product for changes in feedstock composition or changes in solvent composition.
TABLE 4______________________________________ADVANTAGE OF REFLUX TO PURIFICATIONZONE IN NITROGEN REJECTION PROCESSPROCESS VARIABLES RE-REFLUX FLUXED NO______________________________________Solv/Gas Ratio (molar) 6.0 6.0Reflux/Gas Ratio (molar) 0.07 0.0Natural Gas Recovered (mol %) 99.3 99.4N.sub.2 In Nat. Gas Rec. (mol %) 1.7 2.7N.sub.2 In Gas Rejected (mol %) 95.3 95.4HP Flash Pressure (psia) 140.0 140.0MP Flash Pressure (psia) 42.0 42.0LP Flash Pressure (psia) 12.0 12.0Gas Compressor (stages) 3 3Absorber Stages (recovery/ 8/2 8/0 purification)Total Solvent Losses (ppm-wt) 300 300Solvent Temperature (.degree.C.) -34 -34Gas Compression (BHP) 1888 1820Propane Refrigeration (BHP) 233 214LS Pump (actual-recov) (BHP) 402 405TOTAL BHP REQUIRED BY 2523 2439PROCESSBHP PER LB-MOL OF GAS 1.34 1.30RECOVERED______________________________________
EXAMPLE V
Table 5 illustrates examples of hydrocarbon solvents, such as paraffins, isoparaffins, naphthenes, olefinic naphthenes, mono-olefins, alkylbenzenes, and aromatics such as benzene, and toluene, and xylene which may be used in pure form or in mixtures in the process of this invention.
TABLE 5__________________________________________________________________________EXAMPLES OF PREFERRED SOLVENTS FOR USEAS A PURE SOLVENT OR IN MIXTURES THEREOFOR IN MIXTURES WITH OTHER HYDROCARBONSVISCOSITY SOLVENT CARBON NORMAL FREEZING S.G.100 F. NO. B.P. (F.) POINT (F.) 60/60 CS__________________________________________________________________________ParaffinsN-Nonane C.sub.9 303.5 -64.3 0.722 0.8093-Methyloctane C.sub.9 291.6 -161.7 0.725 --N-Decane C.sub.10 345.5 -21.4 0.734 1.0042-Methylnonane C.sub.10 332.7 -102.3 0.731 --N-Undecane C.sub.11 384.6 -14.1 0.744 1.236N-Dodecane C.sub.12 421.3 14.7 0.753 1.512NaphthenesN-Butylcyclopentane C.sub.9 313.9 162.4 0.789 0.908N-Pentylcyclopentane C.sub.10 356.9 -117.0 0.795 1.128N-Butylcyclohexane C.sub.10 357.8 -102.4 0.803 1.251Cis-Decahydronaphthalene C.sub.10 384.5 -45.5 0.901 --Tns-Decahydronaphthalene C.sub.10 369.2 -22.7 0.874 --N-Hexylcyclopentane C.sub.11 397.2 -99.0 0.801 1.415N-Pentylcyclohexane C.sub.11 398.6 -71.6 0.808 1.556N-Heptylcyclopentane C.sub.12 435.0 -63.0 0.805 1.748N-Hexylcyclohexane C.sub.12 436.5 -45.5 0.812 1.94N-Octylcyclopentane C.sub.13 470.3 -47.0 0.809 2.13N-Heptylcyclohexane C.sub. 13 472.8 -23.0 0.815 2.38N-Nonylcyclopentane C.sub.14 503.6 -20.0 0.812 2.57Mono-Olefins1-Decene C.sub.10 339.1 -87.3 0.745 0.8771-Undecene C.sub.11 378.8 -56.5 0.754 1.0851-Dodecene C.sub.12 416.0 -31.4 0.762 1.331-Tridecene C.sub.13 451.0 -9.5 0.769 1.61AlkylbenzenesN-Propylbenzene C.sub.9 318.6 -147.1 0.867 0.794N-Butylbenzene C.sub.10 362.0 -126.3 0.865 0.9471,3-Diethylbenzene C.sub.10 358.0 -119.0 0.868 --1,2-Diethylbenzene C.sub.10 362.2 24.2 0.884 --Tetrahydronaphthalene C.sub.10 405.7 -32.4 0.875 --N-Pentylbenzene C.sub.11 401.7 -103.0 0.863 1.1881-Methylnaphthalene C.sub.11 472.4 -22.9 1.024 2.209N-Hexylbenzene C.sub.12 439.0 -78.0 0.862 1.462N-Heptylbenzene C.sub.13 474.0 -54.0 0.861 1.778N-Octylbenzene C.sub.14 507.9 -33.0 0.860 2.14__________________________________________________________________________
EXAMPLE VI
Example VI is based on the feedstock of Example I and the flowscheme shown in FIG. 3. The lean solvent employed was a C.sub.10 paraffin solvent. The total solvent losses were maintained at 15 ppm. A series of cases were developed with a slightly different thermodynamic model than used in Examples I-IV to illustrate the improvement in product gas purity, as measured by the amount of nitrogen in the product gas, versus the ratio of reflux-to-feed mole ratio. The results of the cases are plotted in FIG. 4.
The process variables for the data shown in FIG. 4, are given in Table 6. Case D shows a case without the purification zone which is equivalent to a 0 reflux-to-feed molar ratio. A purification zone having two theoretical stages was used with the three cases labelled A, B, and C as the reflux-to-feed ratio was increased to 0.22. The number of stages in the recovery zone and the relative pressures of the consecutive flash zones were held constant. The amount of nitrogen in the product gas decreased from about 3.4 mol % to 1.7 mol % for a corresponding increase in the reflux to feed mol ratio from 0 in Case D to about 0.22 in Case A. The cost of this improvement in terms of the energy requirements of the scheme is reflected in the ratio of the total BHP per mol of gas recovered. As the reflux to gas ratio was increased from 0, or no reflux to 0.22, the ratio of BHP per mole of gas recovered increased from about 1.41 in Case D to about 1.67 in Case A. Thus, the process with the reflux, purification zone and solvent recovery elements permitted a 200% improvement in product purity for about an 18% increase in the total BHP per mol of gas recovered.
TABLE 6__________________________________________________________________________NITROGEN REJECTION VS REFLUX-TO-FEED MOL RATIOPROCESS VARIABLES A B C D__________________________________________________________________________Solv/Gas Ratio (molar) 5.2 4.9 4.7 4.5Solvent Temperature (F.) -30 -30 -30 -30Reflux/Gas Ratio (molar) 0.22 0.14 0.07 0.0Natural Gas Recovered (mol %) 99.2 99.2 99.2 99.3N.sub.2 In Nat. Gas Rec. (mol %) 1.72 2.21 2.71 3.45N.sub.2 In Gas Rejected (mol %) 94.3 94.4 94.1 94.8HP Flash Pressure (psia) 140.0 140.0 140.0 140.0MP Flash Pressure (psia) 42.0 42.0 42.0 42.0LP Flash Pressure (psia) 12.0 12.0 12.0 12.0Gas Compressor (stages) 3 3 3 3Absorber Stages (recov/purif) 8/2 8/2 8/2 8/0Total Solvent Losses (ppm-wt) 15 15 15 15Gas Compression (BHP) 2140 2018 1892 1795Propane Refrigeration (BHP) 529 499 458 439LS Pump (actual-recov) (BHP) 463 444 421 409TOTAL BHP REQUIRED BY 3132 2961 2771 2643PROCESSBHP/LB-MOL OF GAS RECOVERED 1.67 1.58 1.48 1.41__________________________________________________________________________
EXAMPLE VII
Example VII is based on the process for the separation of a mixture of hydrogen and nitrogen from a refinery off-gas stream, such as an off-gas from a fluid catalytic cracking unit with the following composition:
______________________________________ Mol %______________________________________ H.sub.2 40.00 N.sub.2 9.53 CO 0.53 CO.sub.2 1.60 H.sub.2 S 3.41 CH.sub.4 22.68 C.sub.2 H.sub.4 9.09 C.sub.2 H.sub.6 9.29 C.sub.3 H.sub.6 2.83 C.sub.3 H.sub.8 0.48 N--C.sub.4 0.41 N--C.sub.5 0.15______________________________________
The process scheme used in the non-cryogenic absorption process is the scheme depicted in FIG. 3. The feed rate to the process was 0.364 million Nm.sup.3 per day (13.6 MMSCFD) at a pressure of about 1275 kPa (about 185 psia) and a temperature of about 38.degree. C. (100.degree. F.). The results of an engineering design simulation are shown in Tables 7 and 8. Table 7 shows the process material balance on a molar basis and Table 8 shows the process parameters. In Example VIII, the product gas is produced at a pressure essentially equal to the feed pressure. The hydrogen-rich gas is rejected at a temperature lower than the feedstream temperature at a recovery of about 95% mol and at a purity of 72.9%. The solvent losses were maintained at a low level of 30 ppm, equivalent to about 3320 liters (877 gallons) annually using a C.sub.10 paraffinic solvent admixed with natural gas liquids in a molar ratio of 85 to 15. The reflux gas returned to the purification zone was maintained at a level of 12 mol % of the feed rate to the process. The overall energy in BHP required by the process was 2,528, or about 1.68 BHP per pound mole of total gas recovered, hydrogen-free gas product and the hydrogen-rich gas.
TABLE 7______________________________________HYDROGEN/NITROGEN REJECTION FROMREFINERY OFF-GAS OVERALL MATERIAL BALANCEComponent, Feedstock Hydrogen-rich Gasmol % Gas Gas Product______________________________________H.sub.2 40.00 72.92 4.42N.sub.2 9.53 15.80 2.75CO 0.53 0.78 0.26CO.sub.2 1.60 0.07 3.25H.sub.2 S 3.41 0.69 6.36CH.sub.4 22.68 6.15 40.56C.sub.2 H.sub.4 9.09 0.87 17.98C.sub.2 H.sub.6 9.29 1.44 17.77C.sub.3 H.sub.6 2.83 0.92 4.93C.sub.3 H.sub.8 0.48 0.16 0.83N--C.sub.4 0.41 0.15 0.66N--C.sub.5 0.15 0.05 0.22Lb Mols/Hr 1500 779.0 721.0Pressure (kPa) 1270 1260 1270Pressure (psia) 170 168 170Temperature, .degree.C. 38 24 43______________________________________
TABLE 8______________________________________HYDROGEN/NITROGEN REJECTION FROMREFINERY OFF-GAS PROCESS VARIABLES______________________________________Solvent/Fuel Gas Ratio (molar) 5.8Solvent Temperature (.degree.C.) -34Solvent: C.sub.10 Paraffin/NGL (mixture mol %) 85/15Reflux/Fuel Gas Ratio (molar) 0.12Hydrogen Gas Recovered (mol %) 94.7Process Gas Compressor (stages) 3Absorber Stages (recov/purif) 8/2Equiv. Solvent Losses (ppm-wt/Gal. Year) 30/877.sup.(1)Gas Compression (BHP) 1,956Propane Refrigeration (BHP) 409.sup.(2)LS Pump (actual-recov) (BHP) 163TOTAL BHP REQUIRED 2,528BY PROCESS .sup.(1) One year = 8000 hr. .sup.(2) 1.1 MM Btu/hr
EXAMPLE VIII
Example VIII is based on a process for separating hydrogen from a gas mixture of hydrogen and hydrocarbons derived from the reactor effluent of a dehydrogenation reaction zone. The feed to the dehydrogenation reaction zone was a hydrocarbon with 3 carbon atoms per molecule and the reactor effluent comprised propylene and propane. The composition of this gas mixture is shown below. (All flows are shown in moles/hour.)
______________________________________ Gas H.sub.2 HydrocarbonComponent Mixture Rejected Product______________________________________H.sub.2 406.70 379.2 27.50N.sub.2 26.64 22.78 3.86CO.sub.2 13.32 0.00 13.32CH.sub.4 152.77 20.93 113.84C2H6 132.38 0.16 132.22C2.dbd. 132.38 0.07 132.31N--C3 1586.85 16.12 1570.73C3.dbd. 1586.85 13.54 1573.31C3.dbd. .dbd. 30.38 0.42 29.96N--C4 7.91 0.14 7.77I--C4 7.91 0.13 7.78C4.dbd. 7.91 0.14 7.77TOTALS 4092.00 453.63 3638.37Temp. (F.) 110 85 85Press. (psia) 75 74 80______________________________________
The rejection of hydrogen was carried out in an absorber column with a recovery zone above the point where the gas mixture entered the absorber column and a purification zone below the point where the gas mixture entered the absorber column. The gas mixture is cooled to a temperature of about -7.degree. C. (20.degree. F.) and charged to the bottom of the recovery zone and countercurrently contacted therein with a lean solvent having an inlet temperature of about -23.degree. C. (-10.degree. F.). A first rich solvent is removed from the recovery zone and passed to the purification zone wherein the first rich solvent is countercurrently contacted with a cooled light gas stream depleted in hydrogen. The cooled light gas stream is charged to the purification zone at a temperature of about 15.degree. C. (60.degree. F.) which was sufficient to maintain the light gas stream in the vapor phase as it entered the purification zone. A second rich solvent stream comprising the hydrocarbons is withdrawn from the purification zone and passed to a solvent recovery zone. In the solvent recovery zone, the second rich solvent is flashed in a high pressure flash zone to provide a high pressure vapor stream and a high pressure liquid stream. The high pressure liquid stream is passed to a low pressure flash zone to provide a low pressure vapor stream and a lean solvent stream. The lean solvent stream is refrigerated and passed to the recovery zone of the absorber column. The low pressure vapor is heated by cross exchange with the gas mixture and passed to a first stage of a 2-stage compressor to provide a first stage gas stream. At least a portion of the first stage a gas stream is passed to the high pressure flash zone to promote the recovery of the lean solvent. The high pressure vapor is passed to the 2nd stage of the 2-stage compressor to provide a compressed hydrocarbon stream. At least a portion of the compressed hydrocarbon stream is passed to a heat exchanger to cool the portion of the compressed hydrocarbon stream and the cooled portion as a vapor at desorption conditions is passed to the purification zone as the cooled light gas stream. The remaining portion of the compressed hydrocarbon stream is withdrawn as a hydrocarbon product. The following is a summary of the process variables in the process of Example VIII.
______________________________________PROCESS VARIABLES______________________________________Solv/Gas Ratio (molar) 3.55Solvent Temperature (.degree.F.) -10Reflux/Gas Ratio (molar) 0.16C3 Gas Recovered (mole %) 99.1H2 Rejected (mole %) 93.2Gas Compression (BHP) 7121Lean Solv. Pump (BHP) 325Refrigeration (BHP) 587Hydraulic Power Rec. (BHP) -123TOTAL BHP REQUIRED BY 7910PROCESS______________________________________
The above summary includes a hydraulic power recovery turbine on the second rich solvent stream before the high pressure flash to recovery energy. An additional 3140 BHP of compression was required to liquify the portion hydrocarbon stream withdrawn as the hydrocarbon product.
Other embodiments of the invention will be apparent to those skilled in the art from a consideration of this specification or practice of the invention disclosed herein. It is intended that the specification and examples be considered as exemplary only, with the true scope and spirit of the invention being indicated by the following claims.

Number	Name	Date
3816976	Stork et al.	Jun 1974
3977203	Hinton et al.	Aug 1976
4242108	Nicholas et al.	Dec 1980
4252548	Markbretter et al.	Feb 1981
4350511	Holmes et al.	Sep 1982
4511381	Mehra	Apr 1985
4526594	Mehra	Jun 1985
4623371	Mehra	Nov 1986
4711651	Sharma et al.	Dec 1987
4806695	Vora et al.	Feb 1989
4832718	Mehra	May 1989
4883514	Mehra	Nov 1989
4934146	Wilhelm et al.	Jun 1990
4936888	DeLong	Jun 1990
5019143	Mehrta	May 1991
5047074	MacGregor et al.	Sep 1991
5051120	Pahade et al.	Sep 1991

Process for the purification of gases

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