MXPA01005440A - Hydrocarbon gas processing - Google Patents

Abstract

A process for the recovery of ethane, ethylene, propane, propylene and heavier hydrocarbon components from a hydrocarbon gas stream is disclosed. In recent years, the preferred method of separating a hydrocarbon gas stream generally includes supplying at least portions of the gas stream to a fractionation tower having at least one reboiler, and often one or more side reboilers, to supply heat to the column by withdrawing and heating some of the tower liquids to produce stripping vapors that separate the more volatile components from the desired components. The reboiler and side reboilers (if any) are typically integrated into the feed stream cooling scheme to provide at least a portion of the refrigeration needed to condense the desired components for subsequent fractionation in the distillation column. In the process disclosed, the tower reboiling scheme is modified to use one or more tower liquid distillation streams from a point higher in the column than is used in the conventional reboiling scheme, providing colder stream(s) for the reboiler(s) that allow more effective cooling of the feed streams and thereby improve the efficiency with which the desired components are recovered. In addition, the tower liquid streams withdrawn from a higher point in the column contain larger quantities of the more volatile components, which when vaporized provide better stripping of undesirable components like carbon dioxide without reducing the recovery of the desired components. The heated distillation stream is returned to a lower point on the fractionation tower that is separated from the withdrawal point by at least one theoretical stage.

Description

PROCESSING OF HYDROCARBON GAS SPECIFICATION BACKGROUND OF THE INVENTION This invention relates to a process for the separation of a gas containing hydrocarbons. Ethane, ethane, propylene, propane and / or heavier hydrocarbons can be recovered from a variety of gases, such as natural gas, refinery gas and synthetic gas jets obtained from other hydrocarbon materials such as coal, petroleum crude oil, naphtha, bituminous shale, tar sands and lignite. Natural gas usually has a higher proportion of methane and ethane, that is, methane and ethane together comprise at least 50 mole percent of the gas. The gas also contains relatively less heavier hydrocarbon amounts such as propane, butanes, pentanes and the like, as well as hydrogen, nitrogen, carbon, carbon dioxide and other gases. The present invention generally deals with the recovery of ethylene, ethane, propylene, propane, and heavier hydrocarbons from such gas jets. A typical analysis of a gas jet that is processed in accordance with this invention may be, in approximate mole percent, 92.12% methane, 3.96% ethane and other C2 components, 1.05% propane and other C3 components, 0.15% iso-butane, 0.21% normal butane, 0.11% positive pentanes, with the rest constituted by nitrogen and carbon dioxide. Sulfur containing gases is also sometimes present. Historically cyclical fluctuations in the prices of natural gas and its constituents of natural gas liquids (NGL) have sometimes reduced the increasing value of ethane, ethylene, propane, propylene, and heavier components as liquid products. Competition for processing rights has forced facility operators to decrease the processing capacity and recovery efficiency of their existing gas processing facilities. The processes available to separate these materials include those based on gas cooling and cooling, oil absorption and refrigerated oil absorption. Additionally, cryogenic processes have become popular due to the availability of economic equipment that produces power while simultaneously expanding and extracting heat from the gas being processed. Depending on the pressure of the gas source, the gas wealth (ethane, ethylene, and heavier hydrocarbon content) and the desired end products, each of these processes or a combination thereof may be employed.

The cryogenic expansion process is now generally preferred for the recovery of natural gas liquids since it provides maximum simplicity with ease of commissioning, flexibility of operation, good efficiency, safety and good reliability. U.S. Patent Nos. 4,157,904; 4,171,964; 4,185,978; 4, -251, 229; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; Reissued North American Patent No. 33,408; and the co-pending request does not. 09 / 054,802 describe the relevant processes (although the description of the present invention in some cases is based on processing conditions different from those described in the cited North American Patents and patent applications). In a typical cryogenic expansion recovery process, a pressurized gas feed jet is cooled by heat exchange with other process jets and / or external cooling sources such as a propane refrigeration-compression system. As the gas cools, the liquids can be condensed and collected in one or more distillers as high pressure liquids containing some of the desired C? + Components. Depending on the richness of the gas and the amount of the liquids formed, high pressure liquids can expand to a lower pressure and fractionate. The vaporization that occurs during the expansion of the liquids results in additional cooling of the jet. Under certain conditions, pre-cooling high pressure liquids prior to expansion may be desirable in order to further decrease the temperature resulting from the expansion. The expanded jet, comprising a mixture of liquid and vapor, is fractionated in a distillation column (demethanizer). In the column, the cooled expansion jets are distilled to remove residual methane, nitrogen and other volatile gases as suspended vapor from the desired Cz components, C3 components and heavier hydrocarbon components as the liquid bottom product. If the feed gas is not completely condensed (typically it is not), at least a portion of the vapor remaining from the partial condensation can be passed through a working expansion machine or motor, or an expansion valve to a lower pressure at which the Additional liquids condense as a result of additional cooling of the jet. The pressure after the expansion is essentially the same as the pressure at which the distillation column is operated. The combined liquid-vapor phases resulting from the expansion are supplied as a feed to the column. In recent years, preferred processes for the separation of hydrocarbons involve feeding this expanded vapor-liquid jet to an average feed point in the column, with a top absorber section providing additional rectification of the vapor phase. The source of the reflux jet for the upper rectification section is typically a portion of the aforementioned steam that remains after partial condensation of the feed gas, although it is removed prior to the working expansion. An alternative source for the upper reflux jet can be provided by a recycled jet of waste gas supplied under pressure. Notwithstanding its source, this jet of vapor is usually cooled to substantial condensation by means of heat exchange with other process jets, for example, vapors leaving the cold fractionating tower. "Part or all of the high-pressure liquid resulting from the partial condensation of the feed gas can be combined with this vapor jet before cooling in. The resulting substantially condensed jet then expands through an appropriate expansion device, such as a valve. During the expansion, a portion of the liquid will normally evaporate, resulting in cooling of the total jet.The instantly expanded jet is then supplied as a feed higher than the demethanizer. , the vapor portion of the expanded jet and suspended vapor of the demethanizer are combined in a top distillation section in the fractionator tower as the residual methane product gas.Alternatively, the cooled and expanded jet can be supplied to a distiller to provide jet streams. steam and liquid so that after d e this steam is combined with the vapors leaving the tower and the liquid is supplied to the column as a top column feed. The purpose of this process is to perform a separation that produces a waste gas that leaves the process which contains substantially all the methane in the feed gas with essentially none of the Cz components and heavier hydrocarbon components and a waste fraction. that leaves the demethanizer which contains substantially all of the components of C and the heavier hydrocarbon components with essentially no methane or more volatile components while meeting the facility's specifications for the maximum allowable carbon dioxide content. The present invention provides a means to provide a new installation or modify an existing processing facility to achieve this separation in significantly lower capital cost by reducing the size or eliminating the need for a pipeline to process systems for the removal of carbon dioxide. Alternatively, the present invention when applied to a new facility or as a modification to an existing processing facility can be used to recover more C2 components and heavier hydrocarbon components in the liquid bottom product for a given carbon dioxide concentration. in the gas feed than other processing schemes. In accordance with the present invention, it has been found that recoveries of C2 in excess of 84 percent can be maintained while maintaining the carbon dioxide content of the liquid bottom product within the specifications and essentially provide for the complete rejection of methane to the waste gas jet. The present invention, although it can be applied at lower pressures and warmer temperatures, is particularly advantageous when processing the feed gases at pressures in the range of 600 to 1000 psia or higher under conditions that require vapor temperatures leaving the column -48.8 ° C (-120 ° F) or colder. The present invention uses a modified kettle scheme that can be applied to any type of NGL recovery system. In a typical boiler or side boiler application in a distillation column, the entire downward flow liquid stream in the column is extracted from the tower and passed through a heat exchanger, then returned to the column essentially at the same point in the column. In this modified kettle system, a portion of downflow liquid is withdrawn into the column from a higher point in the column, i.e., separated from the point of return by at least one theoretical stage. Although the speed of the liquid flow may be lower, it is usually much colder and may have advantages in improving recovery or reducing the size of the exchanger. It has been found that when the present invention is applied to the prior art process for the recovery of NGL, the recovery of the C2 components and the heavier components is improved from one to two percent. The improvement in recovery is much greater, however, when it is desirable to reduce the carbon dioxide content in the recovered NGL product. The recovery of ethane in a typical NGL recovery facility also results in the recovery of at least part of the carbon dioxide contained in the feed gas since carbon dioxide falls between the methane and ethane in relative volatility. Therefore, as the recovery of ethane increases, so does the recovery of carbon dioxide in the NGL product. By applying the modified kettle scheme of the present invention, it has been found that it is possible to significantly improve the recovery of ethane in the NGL product compared to the use of the conventional boiler or side boiling systems when the column is boiled to meet the content of carbon dioxide desired in the NGL product. For a better understanding of the present invention, reference is made to the following examples and drawings. With reference to the drawings: Figure 1 is a flowchart of a prior art cryogenic natural gas processing facility; Figure 2 is a flowchart of an alternative adaptation of the prior art cryogenic natural gas processing facility; Figure 3 is a flow diagram illustrating how the processing facilities of Figures 1 and 2 can be adapted to be a natural gas processing facility in accordance with the present invention; Figure 4 is a flow diagram illustrating an alternative adaptation of Figures 1 and 2 to be a natural gas processing facility in accordance with the present invention; Figure 5 is a flow chart illustrating how an alternative prior art process can be adapted to be a natural gas processing facility in accordance with the present invention; Figure 6 is a diagram illustrating the modified kettle scheme of the present invention for a processing facility wherein the scheme includes a thermosyphon system; Figure 7 is a diagram illustrating the modified kettle scheme of the present invention for a processing facility where the scheme includes a pump system; Figure 8 is a diagram illustrating the modified kettle scheme of the present invention for a processing facility wherein the scheme includes a pump system; and Figure 9 is a diagram illustrating the modified kettle scheme of the present invention for a processing facility wherein the scheme includes a split column system. In the following explanation of the preceding Figures, tables are provided which summarize the flow rates calculated for the representative process conditions. In the tables that appear here, the values for flow rates (in moles of pounds per hour) have been rounded to the nearest number for convenience. The general jet velocities shown in the tables include all components without hydrocarbons and therefore are generally larger than the sum of the jet flow rates for the hydrocarbon components. The indicated temperatures are approximate values rounded to the nearest degree. It should be noted that the process design calculations made for the purpose of comparing the processes represented in the Figures are based on the assumption that there is no heat leak from (or on) the adjoining ones in (or from) the process. The quality of the commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.

DESCRIPTION OF THE PREVIOUS TECHNIQUE Figure 1 is a process flow diagram showing the design of a processing facility for recovering C2 + components from natural gas using the prior art according to US Patent No. 4,157,904. Since this is a large facility designed for 1.0 trillion cubic feet of feed gas per day, the demetanizer, (fractionator tower) will be built in two sections, column 17 absorber and column 19 distiller. In this process simulation, the inlet gas enters the installation at 30 ° C (86 ° F) and 613 psia as jet 31. If the intake gas contains a concentration of sulfur compounds that can prevent product jets meet the specifications, the sulfur compounds are removed by appropriate pre-treatment of the feed gas (not shown). In addition, the feed jet is usually dehydrated to prevent the formation of hydrate (ice) under cryogenic conditions. The solid desiccant has typically been used for this purpose. The feed jet 31 is cooled in an exchanger 10 by heat exchange with the cold waste gas at -37.22 ° C (-99 ° F) (jet 37a), the liquids from the demetallizer boiler at -0.55 ° C (31 °). F) (stream 42), liquids from the bottom side boiler of the demetankator at -15 ° C (-5 ° F) (stream 41) and liquids from the top side demetanizer boiler at -37.22 ° C (-99 ° F) (jet 40). Note that in all cases the exchanger 10 is representative of any plurality of individual heat exchangers or a single multi-pass heat exchanger or any combination thereof. (The decision as to whether it uses more than one heat exchanger for the indicated cooling services will depend on a number of factors including, but not limited to, the admission gas flow rate the heat exchanger size, temperatures of jets etc.). The cooled jet 31a enters the separator 11 at -27.77 ° C (-82 ° F) and 603 psia where the steam (jet 32) is separated from the condensing liquid (jet 35). The steam (jet 32) of the separator 11 is divided into two jets, 33 and 34. The jet 33, which contains about 18 percent of the total steam is combined with the condensing liquid of the separator 11. The combined jet 36 passes through the exchanger 12 heat in relation to the heat exchange with the jet 37 of vapor suspended from the demethanizer resulting in cooling and substantial jet condensation. The jet 36a substantially condensed at 59.44 ° C (139 ° F) then expands instantaneously through an appropriate expansion device, such as an expansion valve 13, up to the operating pressure (about 333_psia of the absorber column 17). During the expansion, a portion of the jet evaporates, resulting in cooling of the total jet.In the process illustrated in Figure 1, the expanded jet 36b exiting from the expansion valve 13 reaches a temperature of 66.11 °. C (151 ° F) and is supplied to the distiller section 17a in the upper region of the absorber tower 17. The liquids separated therein become the feed greater than the theoretical stage 1 in the rectification section 17b. alternative for the distiller liquid (jet 35) according to US Patent No. 4,278,457 is indicated by a dotted line by means of which at least a portion of liquid is expanded to about 333 psia by the expansion valve 16, the cooling jet 35 to produce the jet 35a which is then supplied to the rectification section in the absorber tower 17 to a lower feed point or to the distillation tower 19 a top feeding point). The remaining 82 percent of the vapor from the separator 11 (stream 34) enters a work expansion machine 14 in which the machine energy is extracted from this portion of the high pressure feed. The machine 14 expands the steam substantially in isentropic form from a pressure of about 603 psia to a pressure of about 333 psia, with the working expansion cooling the expanded jet 34a to a temperature of about -51.66 ° C (-125). ° F). Typical commercially available extenders are capable of recovering in the order of 80-85% of the work theoretically available in an ideal isentropic expansion. The recovered work is often used to drive a centrifugal compressor (such as Article 15), which can be used for example to re-compress waste gas (jet 37c). The partially condensed and expanded jet 34a is supplied as a feed to the distillation column at a lower feed point (below the theoretical stage 7 in this case). The liquids (jet 38) from the bottom of the absorber column 17 at -52.77 ° C (-127 ° F) are supplied by the pump 18 for the distillation column 19 to an upper feed point (jet 38a). The present operation of the distillation column 19 (343 psia) is slightly greater than the operating pressure of the absorber column 17 so that the pressure difference between the two towers provides the driving force for the suspended vapors (jet 39) to - 51.66 ° C (-125 ° F) from the top of the distiller column 19 to flow the bottom feed point in the absorber column 17. The demethanizer in the absorbent tower 17 and the distillation tower 19 is a conventional distillation column containing a plurality of vertically separated rafts one or more compacted beds or a certain combination of racks and compaction. As is often the case in natural gas processing facilities, the absorber tower may consist of two sections. The upper section 17a is a separator wherein the partially vaporized upper feed is divided into its respective vapor and liquid portions and wherein the steam rising from the lower distillation or rectifying section 17b is combined with the steam portion (if the there would be) of the upper feed to form the cold waste gas distillation jet 37 that leaves the top of the tower. The lower rectifying section 17b and the distiller column 19 contain the rafts and / or compaction and provide the necessary contact between the liquids that fall downwards and the vapors that rise upwards. The distillation column 19 also includes kettles that heat and evaporate portions of the liquids flowing down the column to provide the distillers vapors flowing up the column. The liquid product (jet 43) leaves the bottom of the tower at 6.11 ° C (43 ° F), based on a typical specification of a methane to ethane ratio of 0.0237: 1 on a molar basis in the bottom product and is pumped to approximately 550 psia (jet 43a) in pump 20. (Discharge pressure of the pump is usually established by the ultimate destination of the liquid product.The liquid product generally flows to the storage and discharge pressure of the pump it is set to prevent any vaporization of the jet 43a as it is heated to ambient temperature).

The waste gas (stream 37) passes concurrently to the incoming-feed gas in: (a) heat exchanger 12 where it is heated to -37.22 ° C (-99 ° F) (stream 37a), (b) heat exchanger 10 where it is heated to 26.11 ° C (79 ° F) (jet 37b), and (c) heat exchanger 21 where 43.33 ° C (110 ° F) is heated (jet 37c). The residual gas is then recompressed in two stages. The first stage is the compressor 15 driven by an expansion machine 14, and the second stage is a compressor 22 driven by a supplementary power source. After the jet 37e is cooled to 46.11 ° C (115 ° F) (jet 37f) by the cooler 23 and up to 30 ° C (86 ° F) by the heat exchanger 21, the waste gas product (jet 37g) ) flows to the dispatch pipeline at 631 psia, enough to meet the requirements online (usually in the order of the admission pressure). A summary of the jet flow velocities and energy consumption for the process illustrated in Figure 1 is set forth in the following table: TABLE I (Figure 1) Summary of the Jet Stream- (Lb. Moles / Hr) Jet Methane Ethane Propane Butanes + Total Dioxides 31 121383 5218 1384 619 1054 131766 32 118997 4514 817 147 990 127561 35 2386 704 567 472 64 4205 33 22015 835 151 27 183 23599 34 96982 3679 666 120 807 103962 38 11021 4734 1353 616 462 18222 39 10516 304 12 1 90 11359 40 6568 6227 1444 625 891 15755 37 121278 788 43 4 682 124903 43 105 4430 1341 615 372 6863 Rescans * Ethane 84. 89% Propane 96. 90% Butans + 99. 33% Power of C .V. Residual Compression 44,408 * (Based on unrounded rounds) As shown in Table 1, the prior art illustrated in Figure 1 achieves 84.89% recovery of ethane using the power in C.V. available residual compression (45,000 hp maximum) however, the concentration of carbon dioxide in the ethane product (methane, ethane, and carbon dioxide jet that results when the waste liquid product is subsequently fractionated to separate the C2 components and the lighter components of the C3 components and the heavier hydrocarbon components) is 7.59 mole percent which exceeds the facility owner's specification of 6.0 percent mole maximum. Thus, this installation design required the vision of a treatment system to remove carbon dioxide from hydrocarbons in order to produce a liquid tradable product. There are many options for removing carbon dioxide (treating the incoming feed gas, treating the total liquid product, treating the ethane product after fractionation, etc.), but all of these options would not only add to the capital cost of the product. installation (due to the cost of installing the treatment system) but also to the operating costs of the installation (due to the energy and chemical consumption in the treatment system). One way to keep the ethane product within the carbon dioxide specification is to operate the demethanizer in a way to distill the carbon dioxide from the liquid bottom product, by adding more boiling heat to the column that uses the servers lateral and / or the bottom kettle. Figure 2 depicts such an alternative operation scheme for the process depicted in Figure 1. The process of Figure 2 has been applied to the same feed gas composition and conditions as described above for Figure 1. However , in the process simulation of Figure 2, the operating conditions of the process have been adjusted to control the bottom temperature of the distillation column 19 so that the carbon dioxide content of the ethane product is within the specification. In the simulation of this process, as in the simulation for the process of Figure 1, the operating conditions were selected to maintain the level of recovery of ethane as high as possible without exceeding the power of C.V. of residual gas compression available. The feed jet 31 is cooled in exchanger 10 by heat exchange with the cold waste gas at -35.55 ° C (96 ° F) (jet 37a), the liquids from the demethanizer boiler at 10 ° C (50 ° F) (jet 42), liquids from the bottom side demetanizer boiler at 3.33 ° C (38 ° F) (jet 41) and liquids from the demetallizer top side boiler at -35.55 ° C (-32 ° F) (jet 40) . The cooled jet 31a enters the separator 11 at -57.77 ° C (-72 ° F) and 600 psia where the steam (jet 32) is separated from the condensed liquid (jet 35). The vapor (jet 32) of the separator 11 is divided into two jets, 33 and 34. The jet 33, which contains approximately 17 percent of the total steam is combined with the condensed liquid of the separator 11. The combined jet 36 passes through the exchanger 12 of heat in relation to the heat exchange with the jet 37 of suspended vapor from the demethanizer resulting in substantial cooling and condensation of the jet. The jet 36a substantially condensed at -91.11 ° C (-132 ° F) then expands instantaneously through the expansion valve 13. As the jet expands to the operating pressure of the absorber column 17 (326 psia), it cools to a temperature of approximately -102.22 ° C (-152 ° F) (jet 36b). The expanded jet 36b is supplied to the tower as the top feed. The remaining 83 percent of the steam from the separator 11 (jet 34) enters the work expansion machine 14 in which the mechanical energy is drawn from this portion of the high pressure feed. The machine 14 expands the steam substantially in isentropic form from a pressure of about 600 psia to the operating pressure of the absorbent tower 17 (326 psia) with the working expansion which cools the expanded jet 34a to a temperature of about -83.33 ° C (-118 ° F). The expanded and partially condensed jet 34a is supplied as a feed to the distillation column at a lower feed point. The liquids (jet 38) from the bottom of the absorber column 17 at -2.09 ° C (-120 ° F) are supplied by the pump 18 to the distillation column 19 at an upper feed point (jet 38a). The operating pressure of the distillation column (336 psia) is slightly higher than the operating pressure of the absorber column 17 so that the pressure difference between the two towers provides the driving force for the suspended vapors (jet 39). ) at -2.05 ° C (-118 ° F) from the top of the 19 distiller column to flow to the bottom feed point in the absorber column 17. The liquid product (jet 43) leaves the bottom of tower 19 at 0.97 ° C (56 ° F). This jet is connected to approximately 550 psia (jet 43a) in a pump 20. The waste gas (stream 37) passes concurrently to the incoming feed gas in: (a) the heat exchanger 12 where it is heated to -16.77 ° C (-961 ° F) (jet 37a), (b) heat exchanger 10 where it is heated to 1.22 ° C (70 ° F) (jet 37b), and (c) heat exchanger 21 where it is heated at 1.76 ° C (101 ° F) (jet 37c). The residual gas is then recompressed in two stages, the compressor 15 operated by the expansion machine 14 and the compressor 22 driven by a supplementary power source. After the jet 37e is cooled to 46.11 ° C (115 ° F) (jet 37f) by the cooler 23 and up to 30 ° C (86 ° F) by the heat exchanger 21, the waste gas product (jet 37g) flows to the dispatch pipeline at 631 psia. A summary of the jet flow velocities and energy consumption for the process illustrated in Figure 2 is set forth in the following table: TABLE II (Figure 2) Summary of the Jet Stream - (Lb. Moles / Hr) Jet Methane Ethane Propane Butane + Total Dioxide C. 31 121383 5218 1384 619 1054 131766 32 120263 4857 1037 233 1023 129517 1120 361 347, 386 31 2249 33 20745 838 179 40 176 22342 4 99518 4019 858 193 847 107175 8 6842 3841 1349 615 284 12953 9 6839 244 12 _ 1 56 7174 0 1886 6752 1588 645 1377 12248 7 121380 1621 47 5 826 125987 43 3 3597 1337 614 228 5779 Recoveries * Ethane 68. .94% Propane 96. .61% Butans + 99. .25% Power of C. V. Residual Compression 44.641 * (Based on unrounded rounds) The concentration of carbon dioxide in the ethane product for the process of Figure 2 is 5.95 mole percent, meeting the installation owner's specification of 6.0 percent mole maximum. However, note that the ratio of methane to ethane in the bottom product is 0.0008: 1 on a molar basis, against the permissible ratio of 0.0237: 1, indicating the degree of over-distillation required to control the carbon dioxide content of the liquid product at the required level. The comparison of the recovery levels displayed in Tables I and II shows that the operation of the process of Figure 2 in this way reduces the content of carbon dioxide in the ethane product which causes a substantial reduction in the recovery of liquids. The process in Figure 2 reduces the recovery of ethane from 84.89% to 68.94%, the recovery of propane from 96.90% to 96.61%, and the recovery of butans + from 99.33% to 99.25%. There are two factors in the work of the process of Figure 2 that result in less recovery of liquids from the bottom of the distiller tower 19 compared to the process of Figure 1. First, when the temperature at the bottom of the column 19 Distiller rises from 6 ° C (43 ° F) in the process of Figure 1 to 13.33 ° C (56 ° F) In the process of Figure 2, the temperatures at each point in the column increase in relation to their corresponding values in the process of Figure 1. This reduces the amount of cooling that the tower's liquid jets (jets 40, 41, and 42) can supply the feed gas in the heat exchanger 10. As a result, the cooled feed stream (jet 31a) entering separator 11 is more tempered (-57.77 ° C (-12 ° F) for the process of Figure 2 versus -63.33 ° C (-82 ° F) for the process of Figure 1), which in turn results in the lowest ethane retention in the absorber column 17 reflected by the ethane content of jet 38 (3841 Lb.

Moles / Hr for the process of Figure 2 against 4734 Lb. Moles / Hr for the process of Figure 1). Second, the higher temperatures in the distillation column 19 causes the temperatures in the absorber column 17 to be higher, resulting in the less methane liquid entering the distillation column (6842 Lb. Moles / Hr in jet 38 for the process of Figure 2 against 11021 Lb. Moles / Hr for the process of Figure 1). When this liquid methane is subsequently evaporated by the side kettles and the main kettle is attached to the distillation column 19, the methane vapor helps to distill the carbon dioxide from the liquids flowing down the column. With less methane available in the process of Figure 2 to distill carbon dioxide, more of the ethane in the liquids must be vaporized to serve as the distillation gas. Since the relative volatiles for carbon dioxide and ethane are very similar, ethane vapor is a much less effective distillation agent than methane vapor, which reduces the distillation efficiency in the column.

DESCRIPTION OF THE INVENTION Example 1 Figure 3 illustrates a flow diagram of a process according to the present invention. The composition of feed gas and the conditions considered in the process presented in Figure 3 are the same as those in Figure 1. Therefore, the process of Figure 3 can be compared with that of the process in Figure 1 to illustrate the advantages of the present invention. In the simulation of the process of Figure 3, the intake gas enters at 30 ° C (86 ° F) and a pressure of 613 psia as jet 31. The feed stream 31 is cooled in exchanger 10 by heat exchange with the cold waste gas at -72.77 ° C (-99 ° F) (jet 37a), liquids from the demetallizer boiler at -1.11 ° C (30 ° F) (jet 42), liquids from the demetallizer side boiler at -20 ° C (-4 ° F) (jet 41) and a portion of the bottom liquids of the absorber column at -88.88 ° C (-128 ° F) (jet 45). The cooled jet 31a enters the separator 11 at -64.44 ° C (-84 ° F) and 603 psia where the vapor (jet 32) is separated from the condensed liquid (jet 35). The steam (jet 32) of the separator 11 is divided into first gaseous and second jets 33 and 34. The jet 33, which contains about 19 percent of the total steam, is combined with the condensed liquid (jet 35) to form the jet 36. The combined jet 36 passes through the heat exchanger 12 in heat exchange with respect to the cold waste gas (jet 37) where it is cooled to -94.44 ° C (-138 ° F). The resulting substantially condensing jet 36a then expands instantaneously through an appropriate expansion device such as an expansion va13 at the operating pressure (approximately 332 psia) of the absorber tower 17. During expansion, a portion of the jet is vaporized, resulting in cooling of the total jet. In the process illustrated in Figure 3 the expanded jet 36b exiting from the expansion va13 reaches a temperature of -101.66 ° C (-151 ° F) and is supplied to an absorbent column 17 as the upper column feed. The vapor portion (if any) of the jet 36b combines with the vapors rising from the upper fractionating stage of the column to form the distillation jet 37, which is drawn from an upper region of the tower. Returning to the second gaseous stream 34, the remaining 81% of the vapor from the separator 11 enters a working expansion machine 14 in whose mechanical energy is extracted from this portion of the high pressure feed. The machine 14 expands the steam substantially in an isentropic form from a pressure of about 603 psia to a pressure of about 332 psia, with the work expansion that cools the expanded -34a jet to a temperature of about -88.33 ° C (- 127 ° F). The expanded and partially condensed jet 34a is thereafter supplied as a feed for the absorber column 17 to a lower column feed point. Alternatively, as shown by the dotted line, the condensed liquid (stream 35) of the separator 11 can be instantly expanded through an appropriate expansion device, such as an expansion valve 16 to the operating pressure of the absorber tower 17, cooling the jet 35 to produce the jet 35a. The expanded jet 35a exiting from the expansion valve 16 can then be supplied to the absorber tower 17 at a lower column limitation point or to the distillation tower 19 at an upper column feeding point. The liquids (jet 38) from the bottom of the absorber column 17 enter the pump 18 at -88.88 ° C below (-128 ° F) and are pumped to a higher pressure (jet 38a) and divide into two portions. One portion (jet 44), which contains about 55% total liquid, is fed to the distillation column 19 at a higher feed point. The operating pressure of the distillation column 19 (342 psia) is slightly more alpha than the operating pressure of the absorber column 17 so that the pressure difference between the two towers provides the driving force for the suspended vapors (jet 39) at -86.11 ° C (-123 ° F) from the top of the distiller column 19 to flow to the bottom feed point in the absorber column 17. The other portion (jet 45), which contains the remaining 45% of the jet 38a of the pumped liquid, is directed to the heat exchanger 10 where part of the feed gas is supplied cooling as it is heated to -28.88 ° C (-20). ° F) and partially evaporated. The heated jet 45a is then supplied for the distillation column 19 to a medium column feed point, separated from the upper feed point where the jet 44 enters the column by at least one theoretical step. In this case, the partially evaporated stream flows to the same point in the column that was used for the return of the upper lateral boiler (theoretical stage 8 in the distillation tower 19) in the process of Figure 1, which is the equivalent of the seven theoretical stages lower than the liquid jet extraction point in the fractionation system (the upper feed point where the jet 44 enters the distillation column 19). The liquid product (jet 43) exits the bottom of tower 19 at 5.55 ° C (42 ° F). This jet is pumped to approximately 550 psia (jet 43a) in the pump 20. The waste gas (jet 37) passes concurrently to the incoming feed gas in: (a) heat exchanger 12 where it is heated to -72.77 ° C ( -99 ° F) (jet 37a), (b) heat exchanger 10 where it is heated to 26.66 ° C (80 ° F) (jet 37b), and (c) heat exchanger 21 where it is heated to 40.55 ° C ( 105 ° F) (jet 37c). The residual gas is then recompressed in two stages, the compressor 15 operated by the expansion machine 14 and the compressor 22 driven by a supplementary power source. After the jet 37e is cooled to 46.11 ° C (115 ° F) (jet 37f) by the cooler 23 and up to 30 ° C (86 ° F) by the heat exchanger 21, the waste gas product (jet 37g) ) flows to the dispatch pipeline at 631 psia. A summary of the jet flow rates and energy consumption for the process illustrated in Figure 3 is set forth in the following table: TABLE III (Figure 3) Summary of the Jet Stream - (Lb. Moles / Hr) Jet Methane Ethane Propane Butane + Dioxide Tot l de C. 31 121383 5218 1384 619 1054 131766 32 118694 4440 779 136 982 127126 2689 778 605 483 72 4640 33 22552 844 148 26 187 24154 34 96142 3596 631 110 795 102972 8 11906 4855 1357 616 557 19330 4 6548 2670 746 339 306 10632 5 5358 2185 611 277 251 8698 9 11800 362 13 1 156 12370 7 121277 725 40 4 653 124806 3 106 4493 1344 615 401 6960 Recoveries * Ethane 86.12% Propane 97.10% Butane + 99.41% Power of C.V. Residual compression 44,413 * (Based on unrounded rounds) A comparison of Tables I and III shows that, compared to the prior art, the present invention improves the recovery of ethane from 84.89% to 86.12%, propane recovery from 96.90% to 97.10%, and recovery of butans + from 99.33% up to 99.41%. The comparison of Tables I and III further shows that the improvement in yields was achieved using power requirements of C.V. equivalents (utility). By using the modified kettle approach, the liquid from the column flowing to the heat exchanger 10 (jet 45) is cooled more than the corresponding jet 40 of the process of Figure 1. This increases the available cooling for the intake gasgo. , since not only can you get considerably more performance from liquids with this scheme, but liquids are available at a cooler temperature level than would be possible with a conventional kettle scheme. The result is the increased C2 + component and heavier hydrocarbon component recoveries for the process in Figure 3 while essentially using the same amount of C.V. of residual gas compression as the process of Figure 1 of the prior art.

Example 2 In those cases where the carbon dioxide content of the liquid product is a consequence, (due to the more demanding product specifications imposed by the customer as in the prior art process of the Figure 2 previously described, for example), the present invention offers very significant recovery and efficient advantages over the prior art process depicted in Figure 2. The operating conditions of the process of Figure 3 can be altered to reduce the dioxide content of carbon in the liquid product of the present invention as illustrated in Figure 4. The feed gas composition and the conditions considered in the process presented in Figure 4 are the same as those in Figures 1 and 2. Accordingly , the process of Figure 4 can be compared with that of the processes of Figures 1 and 2 to illustrate the advantages of the present invention. In the simulation of the process to Figure 4, the cooling of the intake gas and the separation scheme is essentially the same as that used in Figure 3. The main difference is that the controls of the installation have been adjusted to increase the proportion of the liquids from the bottom of the absorbent tower 17 (jet 45) which is heated in heat exchanger 10 and supplied to the distillation tower 19 at a medium feed point in the column. The installation controls have also been adjusted to raise the bottom of the 19-column distillator slightly (from 5.55 ° C (42 ° F) in the process of Figure 3 to 7.22 ° C (45 ° F) in the process of Figure 4) to maintain the ratio of methane to ethane in the bottom product in the molar ratio of 0.0237: 1 specified. The increased amount of the heated jet 45a entering the distillation tower 19 and the higher waste temperature both increase the distillation within the tower, which results in warmer temperatures for the process of Figure 4 in relation to the process of Figure 3 through the absorber column 17 and the distillation column 19, with the net effect of reducing the carbon dioxide content of the liquid product, jet 43, which leaves the distillation column 19. The more temperate column temperatures also result in a slight reduction in cooling that is available from the process jets that are applied to the column feed jets. In particular, this requires slightly reducing the proportion of the distiller feed gas (jet 32), which is directed to the heat exchanger 12 by the jet 33, thereby reducing the amount of jet 36b entering the upper feed point of the tower 17 absorber. A summary of the jet flow velocities and energy consumption for the process illustrated in Figure 4 is set forth in the following table: TABLE IV (Figure 4) Summary of the Jet Stream - (Lb. Moles / Hr) Jet Methane Ethane Propane Butane + Total Dioxide C. 31 121383 5218 13S4 619 1054 131766 32 118612 4421 770 133 980 127009 2771 797 614 486 74 4757 3 21943 818 143 25 181 23497 4 96669 3603 627 108 799 103512 8 11442 4976 1362 616 616 19052 4 5721 2488 681 308 308 9527 5 5721 2488 681 308 308 9527 9 11337 561 21 1 338 12297 7 121278 803 43 4 776 125011 3 105 4415 1341 6.15 278 6755 Recoveries Etano 84.61% Propane 96.96% Butane + 99.39% Power of C.V. Residual compression 44,573 * (Based on unrounded rounds) The concentration of carbon dioxide in the ethane product for the process of Figure 4 is 5.80 mole percent, well, below the specification required by the customer. The comparison of the recovery levels displayed in Tables I and IV shows that the present invention allows to achieve the required carbon dioxide content while maintaining almost the same liquid recovery efficiency as the process of Figure 1. Although the recovery of ethane decreases slightly from 84.89% to 84.61%, propane recovery and butane + recovery both increase slightly from 96.90% to 96.96% and from 99.33% to 99.39%, respectively. The comparison of Tables I and IV further shows that maintaining product yields was achieved using essentially the same requirements as the C.V. (utility) The comparison of the recovery levels displayed in Tables II and IV shows that the present invention allows to achieve much more liquid recovery efficiency than the process of Figure 2 when operating in a way to limit the carbon dioxide content of your liquid product. Compared to the process in Figure 2, the process in Figure 4 elevates ethane recovery from 68.94% to 84.61%, almost 15.7 percentage points higher. The recovery of propane and the recovery of butanes + also increases in some way from 96.61% to 96.96% to 96.96% and from 99.25% to 99.39%, respectively. The comparison of Tables II and IV further shows that the higher product yields were not simply the result of the increase in the requirements of the C.V. (utility) On the contrary, when the present invention is employed as in Example 2, it not only causes the recoveries of ethane, propane, and butanes + to increase over those of the prior art process, but also increases the recovery of liquids to 23 percent ( in terms of ethane recovered per power unit of CV consumed). As with the process of Figure 3, a significant benefit achieved by the embodiment of Figure 4 is that the modified kettle scheme provides colder column liquids for use in the cooling of incoming feed jets. This increases the available cooling for the intake gas, thus not only can a considerably higher performance of the liquid be obtained in this case, but at a cooler temperature level. At the same time, more methane is introduced into the lower part of the distillation column 19 than would otherwise be there when the boiling column finds the carbon dioxide content. (Note that jet 45 in the process of Figure 4 contains 5721 Lb. Moles / Hr of methane and is introduced in the theoretical stage 8 of column 19 distillator, while jet 40 in the process of Figure 2 contains only 1886 Lb. Moles / Hr methane and is introduced in the upper part of the 19 distillation column). The additional methane provided by the present invention in the process of Figure 4 helps to remove the carbon dioxide from the liquids flowing down into the distillation column. The amount of carbon dioxide in the NGL product can be adjusted by proper control of the amount of liquid removed to feed the modified kettle system instead of feeding the top of the distillation column.

Other Modes Figure 5 is a flow diagram illustrating how the process and apparatus described and depicted in U.S. Patent No. 5,568,737 can be adapted to be a natural gas processing facility in accordance with the present invention. Figures 6, 7, 8, and 9 are diagrams showing part of the alternative methods for implementing the modified kettle scheme. Figure 6 shows a typical thermosyphon type application where the partial flow of liquid from the fractionating tower 50 to the boiler 57 can be controlled by the valve 58 in the liquid extraction line 61. The liquid portion not extracted from the column simply floods the stack 51 of the chimney on distributor 52 for compaction (or trays) 53 lower. The heated jet on line 61a of boiler 57 is returned to fractionating tower 50 to a lower panel which contains an appropriate feed distribution mechanism, such as a chimney pan 54 and distributor 55, for mixing the heated jet with the liquids from the downflow tower from the compaction 53 and supply the mixture to the compaction (or rafts) 56. Figures 7 and 8 show typical pumping adaptations where the downward flow of the general liquid is extracted on the extraction line 61 of liquids and is pumped to upper pressure by the pump 60. The liquid flow pumped into the line 61a is then divided by appropriate control valves 58 and 59 to reach the desired amount of liquid in the line 62 flowing to the boiler 57 The heated jet in line 62a of the boiler 57 is returned to the fractionating tower 50 to a lower point as previously described for the embodiment of the Figure 6. In the embodiment of Figure 7, the liquid that does not flow to the kettle (on line 63) is returned to the chimney pan 51 from which the liquid was initially drawn, where it can flood the chimney pan 51 on the distributor 52 for compaction (or trays) 53 lower. In the embodiment of Figure 8, the liquid that does not flow to the boiler (in line 63) is returned below the chimney pan 51 from which the liquid was initially extracted, directly to the distributor 52 which supplies the liquid to the compaction ( or bats) 53 lower. Figure 9 shows how the pumping system described for Figure 8 can be implemented in a split column approach, such as a top column 65 and bottom column 50, which is the same as that used in Figures 3 and 4. The person skilled in the art will recognize that the present invention gains some of its benefits by providing a cooler jet to the side boiler and / or kettles, allowing additional cooling of the column feed or feeds. This additional cooling reduces the utility requirements for a given level of product recovery, or improvement of product recovery levels for a given utility consumption, or a certain combination thereof. In addition, one skilled in the art will recognize that the present invention also benefits by introducing large amounts of lower methane into the demethanizer to aid in the distillation of carbon dioxide from downwardly flowing liquids. With more methane available to dehull liquids, correspondingly less ethane is needed for the dehilation, allowing greater retention of ethane in the liquid bottom product. Therefore, the present invention can be applied generally to any process dependent on cooling of any number of feed jets and supplying the resulting feed jets to the column for distillation. In accordance with this invention, the cooling of the demethanizer feed streams can be accomplished in many ways. In the processes of Figures 3 and 4, the feed jet 36 is cooled and substantially counteracted by the suspended steam stream 37 of the demethanizer, while the demethanizer liquids (jets 45, 41, and 42) are used only for gas jet cooling. In the process of Figure 5, the jet of high pressure waste feed 48 is also cooled and substantially condensed by the steam jet portions suspended from the distillation column (jets 46 and 37), while the liquid of the demethanizer (jets 40 and 42) are used only for t the gas jet cooling. However, the demethanizer liquids can be used to supply part or all of the substantial and cooling condensation of the jet 36 in Figures 3 to 5 and / or the jet 48 in Figure 5 in addition to or instead of the gas jet cooling. In addition, any jet at a cooler temperature than the cooling stream can be used. For example, a lateral extraction of steam from the demethanizer can be extracted and used for cooling. Other potential sources of cooling include, but are not limited to, instant high-pressure liquid separators and mechanical refrigeration systems. The selection of a cooling source will depend on a number of factors including, but not limited to, the composition of the intake gas and the conditions, installation size, size of the heat exchanger, temperature of the potential cooling source, etc. One skilled in the art will also recognize that any combination of the above cooling sources or methods can be used in combination to achieve the desired feed jet temperatures. According to this invention, the use of external cooling to supplement the available cooling for the intake gas of other process jets can be used, particularly in the case of a richer admission gas than that used in Examples 1 and 2. The use and distribution of demethanizer liquids for the heat exchange of the process, and the particular arrangement of the heat exchangers for the intake gas cooling should be evaluated for each particular application, as well as the choice of process jets for the services of heat exchange specific. The high pressure liquid in Figures 3 to 5 (jet 35) not all need to be combined with a portion of the separating vapor (jet 33) flowing to heat exchanger 12. Alternatively, this jet of liquid (or a portion thereof) can be expanded through a suitable expansion device, such as an expansion valve 16 and fed to a mid-point of feed in the lower distillation column (tower 17 absorber or tower 19 distiller in Figures 3 and 4), fractional tower 17 in Figure 5). The liquid jet can also be used for the cooling of inlet gas or other heat exchange service before or after the expansion step before flowing to the demethanizer. It will also be recognized that the relative amount of feed found in each branch of the column feed jets will depend on many factors, including the gas pressure, the feed gas composition, the amount of heat that can be extracted economically from the power and the amount of CV power available. More feed to the top of the column can increase the recovery while the decrease power recovered from the expansion machine increases with this the power requirements of C.V. of recompression. The lower implementation feed in the column reduces the power consumption of C.V. But it can also reduce product recovery. The positions of the middle column feed shown in Figures 3 and 4 are the preferred feed locations for the process operating conditions described. However, the relative locations of the feeds in the middle column may vary depending on the intake composition or other factors such as the desired recovery levels and the amount of liquid formed during the intake gas implement. further, two or more of the feed jets or portions thereof, may be combined depending on the relative temperatures in quantities of the individual jets, and the combined jet then fed to a supply position of the middle column. Figures 3 and 4 are the preferred embodiment for the compositions and pressure conditions shown. Although individual jet expansion is represented in particular expansion devices, alternative expansion means can be used anywhere. necessary. For example, conditions can guarantee the working expansion of the substantially condensed portion of the feed stream (36a in Figures 3 to 5) or the substantially condensed recycle jet (48b in Figure 5). Figures 3 and 4 represent a fractional tower built in two sections (17 and 19) due to the size of the installation. The decision to build the fractionator tower as a single container (such as 17 in Figure 5) or multiple containers will depend on a number of factors such as the size of the installation, the distance to manufacturing facilities, etc. While what is believed to be the preferred embodiments of the invention has been described, those skilled in the art will recognize that other additional modifications can be made thereto, for example to adapt the invention to various conditions, types of feeding, or other requirements without depart from the spirit of the present invention as defined by the following claims.

Claims (13)

1. CLAIMS í.
2. In a process for the separation of a gas jet, it contains methane, C2 components, C3 components and heavier hydrocarbon components in a fraction of volatile waste gas containing a greater portion of methane and a relatively smaller volatile fraction containing a greater portion of the components of C2, components of C3 and components of heavier hydrocarbons, in whose process (a) said jet of gas is. treats in one or more stages of heat exchange and at least one division step to produce at least a first feed jet that has been cooled under pressure to substantially condense all and at least one second feed jet that has been employee under pressure; (b) said first substantially condensed feed jet is expanded to a lower pressure whereby it is further cooled, and is then supplied to a fractionating tower at a higher feed point; (c) said second feed and cooling jet is expanded to the lower pressure, and the fractionating tower is then supplied to a feed point in the middle column; and (d) said first cooled expanded feed jet and the second expanded feed jet are fractionated into the lower pressure whereby the components of the relatively smaller volatile fraction are recovered; the improvement wherein (1) a liquid distillation jet is extracted from the fractionating and heated tower; (2) said heated distillation jet is returned to a lower point in the fractionating tower that separates from the extraction point by at least one theoretical stage; and (3) the quantities and temperatures of the feed jets for the fractionating tower are effective to maintain the suspended temperature of the fractionating tower at a temperature whereby the largest portions of the components in the relatively minor volatile fraction are recovered.
3. 3. The improvement according to claim 1 or 2, wherein the liquid distillation jet is pumped after being extracted from the fractionating tower.
4. The improvement according to claim 3, wherein (a) said pumped liquid distillation jet is divided into at least a first portion and a second portion; (b) said first portion is heated; and (c) said first heated portion is returned to a lower point in the fractionating tower and is separated from the extraction point by at least one theoretical step.
5. The improvement according to claim 1 or 2, wherein the liquid distillation jet is directed in heat exchange relationship with at least a portion of the gas jet or supply jets to supply cooling thereto and heats with this to the liquid distillation jet.
6. The improvement according to claim 3, wherein the pumped liquid distillation jet is directed in heat-exchange relationship with at least a portion of the gas jet or the supply jets, to supply cooling thereto and with this heat the liquid distillation jet pumped.
7. The improvement according to claim 4, wherein the first portion is directed in relation to the heat exchange with at least a portion of the gas jet or the supply jets, to supply cooling thereto and thereby to heat the first portion. -
8. 8. The improvement according to claims 1 6 2, wherein the quantity and temperature of the heated distillation jet and the heating supplied to the fractionating tower are effective to maintain the temperature of the bottom of the fractionating tower at a temperature to reduce the amount of carbon dioxide contained in the fraction less volatile The improvement according to claim 3, wherein the amount and temperature of heated distillation jet and the heating supplied to the fractionating tower are effective to maintain the bottom temperature of the fractionating tower at a temperature to reduce the amount of carbon dioxide contained in the relatively less volatile fraction. 10-. The improvement according to claim 4, wherein the amount and temperature of the first portion heated and the heating supplied to the fractionator tower are effective to maintain the bottom temperature of the fractionator tower at a temperature to reduce the amount of carbon dioxide contained in the relatively less volatile fraction . The improvement according to claim 5, wherein the amount and temperature of the heated jet and the heating supplied to the fractionator tower are effective to maintain the bottom temperature of the fractionating tower at a temperature to reduce the amount of dioxide carbon contained in the relatively less volatile fraction. The improvement according to claim 6, wherein the amount and temperature of the heated distillation jet and the heating supplied to the fractionating tower are effective to maintain the bottom temperature of the fractionating tower at a temperature to reduce the amount of carbon dioxide contained in the relatively less volatile fraction. The improvement according to claim 7, wherein the amount and temperature of the heated first portion and the heating supplied to the fractionating tower are effective to maintain the bottom temperature of the fractionating tower at a temperature to reduce the amount of carbon dioxide contained in the relatively less volatile fraction. SUMMARY A process for the recovery of ethane, ethylene, propane, propylene and heavier hydrocarbon components from a hydrocarbon gas jet is described. In recent years, the preferred method for separating a jet of hydrocarbon gas generally includes supplying at least portions of the gas jet to a fractionating tower having at least one boiler and often one or more side kettles, to supply heat to the column by extracting and heating part of the tower liquids to produce distillation vapors that separate the more volatile components from the desired components. The kettle and side kettles (if any) are typically integrated in the feed jet cooler scheme to provide at least a portion of the cooling necessary to condense the desired components for subsequent fractionation in the distillation column. In the process described, the tower boiler scheme is modified to use one or more liquid distillation jets from the tower from a point higher in the column than that used in the conventional boiler scheme, providing colder jets for kettles that allow more effective cooling of the feed jets and thereby improve the efficiency with which the desired components are recovered. In addition, the liquid jets of the tower extracted from a higher point in the column contain larger amounts of more volatile components, which when evaporated provide better distillation of undesirable components such as carbon dioxide without reducing the recovery of the desired components. The heated distillation jet is returned to a lower point in the fractionating tower that is separated from the extraction point by at least one theoretical stage.