Hydrocarbon gas processing.

[0001] This invention relates to a process and a gas separation apparatus containing hydrocarbons. The inventors invoke the benefits under Title 35, US Code, Section 119 (e) of the previous US Provisionai Application. No. 61 / 186,361 filed on June 11, 2009. The inventors also invoke the benefits under Title 35, US Code, Section 120 as a continuation in part of the US Patent Application. No. 13 / 048,315 filed on March 15, 2011, and as a partial continuation of the US Patent Application. No. 12 / 781,259 filed on May 17, 2010, and as a partial continuation of US Patent Application. No. 12 / 772,472 filed on May 3, 2010, and as a partial continuation of US Patent Application. No. 12 / 750,862 filed on March 31, 2010, and as a partial continuation of US Patent Application. No. 12 / 717,394 filed on March 4, 2010, and as a continuation in part of the US Patent Application. No. 12 / 689,616 filed on January 19, 2010, and as a continuation in part of the US Patent Application. No. 12 / 372,604 filed on February 17, 2009. The assignees, SME Products LP and Ortloff Engineers, Ltd. were part of a joint investigation agreement that was in effect before the invention of the present application was made.

[0002] Propylene, propane and / or heavier hydrocarbons made from various gases, common natural gas, refinery and synthetic gas, obtained from other hydrocarbon materials, such as carbon, crude oil, naphtha, bituminous shale , bituminous sands and lignite. Natural gas usually contains a higher proportion of methane and ethane, that is, methane and ethane together constitute at least 50 mole percent of the gas. The gas also contains relatively smaller amounts of heavier hydrocarbons, such as propane, butane, pentanes and the like, as well as hydrogen, nitrogen, carbon dioxide and other gases.

[0003] The present invention relates in general to the recovery of heavier propylene, propane and hydrocarbons from said gas streams. A typical analysis of a gas stream to be processed in accordance with this invention would yield a result, in an approximate molar percentage, of 88.4% methane, 6.2% ethane and other C components.₂, 2.6% propane and other C components₃, 0.3% of e

isobutane, 0.6% normal butane and 0.8% pentanes, where the rest would be composed of nitrogen and carbon dioxide. Occasionally, sulfur-containing gases are also present.

[0004] Historically cyclical fluctuations in the prices of constituents, both natural gas and condensed natural gas (NGL), have sometimes reduced the increasing value of propane, propylene and heavier components, such as products liquids This has resulted in a demand for processes that can provide more efficient recoveries of these products and processes that can provide efficient recoveries with less capital investment. The processes available to separate these materials include those based on the cooling and cooling of the gas, the absorption of oil and the

<img file="AR080751A1_D0001.tif" />

C absorption of refrigerated oil. In addition, cryogenic processes have gained popularity, due to the availability of economic equipment that generates energy while simultaneously expanding and extracting heat from the gas under processing. Depending on the pressure of the gas source, the richness (content of ethane, ethylene and heavier hydrocarbons) of the gas and the desired final products, it is possible to use one of these processes or a combination of them.

[0005] Currently, the cryogenic expansion process is generally preferred for the recovery from condensed natural gas because it provides maximum simplicity together with ease of starting, operational flexibility, good efficiency, safety and good reliability. In U.S. Pat. N °: 3292380, 4061481, 4140504, 4157904, 4171964, 4185978, 4251249, 4278457, 4519824, 4617039, 4687499, 4689063, 4690702, 4854955, 4869740, 4889545, 5275005, 5555748, 5566554 577, 587 557 577 , 5983664, 6182469, 6578379, 6712880, 6915662, 7191617, 7219513, in newly published US Patent No. 33408 and in co-pending applications No. 11/430412, 11/839693, 11/971491; 12/206230; 12 / 689,616; 12 / 717,394; 12 / 750,862; 12 / 772,472; 12 / 781,259; 12 / 868,993; 12 / 869,007; 12 / 869,139; 12 / 979,563; and 13 / 048,315 relevant processes are described (although in some cases, the description herein is based on processing conditions different from those described in the cited US Patents).

[0006] In a typical cryogenic expansion recovery process, a pressurized gas feed stream is cooled by heat exchange with other process streams and / or external sources of \ ·^ζ refrigeration, such as a propane cooling-compression system.

As the gas is cooled, it is possible to condense and collect the liquids into one or more high pressure liquid separators containing some of the desired C3 + components. Depending on the richness of the gas and the amount of liquids formed, high pressure liquids can be expanded at a lower pressure and fractionated. The vaporization that takes place during the expansion of the liquids results in additional cooling of the current. Under certain conditions, it may be desirable to pre-cool the liquids under high pressure before expansion, in order to further reduce the temperature resulting from the expansion. The expanded stream, which comprises a mixture of liquid and vapor, is fractionated in a distillation column (desetanizer). In the column, the stream (s) cooled by expansion is (are) distilled (s) to separate residual methane, components C₂, nitrogen, and other volatile gases as head steam of the desired C3 components and heavier hydrocarbon components as liquid bottom product.

[0007] If the feed gas is not completely condensed (typically it is not), the remaining steam of the partial condensation can be passed through a machine or expansion work engine, or an expansion valve, up to a pressure lower than the additional liquids will be condensed as a result of additional cooling of the current. Then, the expanded current enters an absorption section in the column and contacts cold liquids to absorb components C₃ and the heaviest components of the vapor portion of the expanded stream. Liquids from the absorption section

<img file="AR080751A1_D0002.tif" />

C are then directed to the desetanization section of the column.

[0008] A stream of distillation steam is removed from the upper region of the detanization section and cooled by heat exchange relationship with the head steam stream from the absorption section, condensing at least a portion of the steam stream of distillation. The condensed liquid is separated from the cooled stream of distillation steam to produce a cold liquid reflux stream that is directed towards the upper region of the absorption section, where cold liquids can come into contact with the steam portion of the stream expanded as described above. The vapor portion (if any) of the cooled stream of distillation steam and the head steam from the absorption section combine to form the residual methane gas product and components C₂.

[0009] The separation that occurs in this process (producing a waste gas leaving the process that contains substantially all of the methane and components C₂ that is in the feed gas with essentially none of the components C₃ nor of the heaviest hydrocarbon components, and a fraction of the bottom that comes out of the degasser that contains substantially all of the C components₃ and the heavier hydrocarbon components essentially without methane, or C components₂ nor more volatile components) consumes energy to cool the feed gas, to reheat the detanization section, to reflux the absorption section, and / or to recompress the gaseous waste.

[0010] A novel means is now used to carry out the various steps described above, more efficiently and

<img file="AR080751A1_D0003.tif" />

by using fewer pieces of equipment. This is achieved by combining what until now were individual pieces of equipment in a common device, which reduces the area required for the processing plant and the cost of installation capitai. Surprisingly, applicants have discovered that this more compact arrangement also allows to significantly reduce the energy consumption required to obtain a high level of recovery, thereby increasing the efficiency of the process and reducing the operational cost of the installation. In addition, with this more compact arrangement, most of the pipe necessary to interconnect individual pieces of equipment in the designs of traditional facilities is also eliminated, thereby reducing the cost of capital and also connecting the connections of the associated flange pipes. As flange pipes are a source of potential losses of hydrocarbons (which are volatile organic compounds, VOCs, which contribute to greenhouse gases and which can also be precursors for atmospheric ozone formation), when these flange pipes are eliminated reduces the potential of atmospheric emissions that could damage the environment.

[0011] In accordance with the present invention, it has been found that C3 recoveries greater than 99.6% can be obtained while providing essentially complete rejection of the C components₂ to the gaseous waste stream. In addition, the present invention makes possible a separation of essentially 100% of the components C₂ and lighter components of the C3 components and other heavier components with better energy requirements compared to the prior art at the same time €

It maintains the same level of recovery. The present invention, even though it is functional at lower pressures and at higher temperatures, is particularly advantageous when feed gases are processed in a range of values ​​between 400 and 1500 psia [between 2758 and 10342 kPa (a) j ο higher, low conditions where temperatures at the head of the NGL recovery column of -50 ° F [~ 46 ° C] or lower are needed.

[0012] For a better understanding of the present invention, reference will be made to the following examples and figures, the figures are described below.

[0013] FIG. 1 is a flow chart of a natural gas processing plant of the technical background, in accordance with US Patent No. 5,799,507.

[0014] FIG. 2 is a flow chart of a natural gas processing plant in accordance with the present invention.

[0015] Finally, FIG. 3 to 21 are flow charts where the alternative means of applying the present invention to a natural gas stream is illustrated.

[0016] In the explanation of the aforementioned figures that will be provided below, tables are presented summarizing the flow rates calculated for representative process conditions. In the tables presented, the values ​​of the flows (in moles per hour) were rounded to the nearest integer value for reasons of convenience. The total flow rates listed in the tables include all components that are not hydrocarbons, so they are generally higher than the sum of the flow rates of the hydrocarbon components. The indicated temperatures

<img file="AR080751A1_D0004.tif" />

ο are approximate values ​​rounded to the nearest degree. It should also be borne in mind that the process design calculations made in order to compare the processes represented in the figures are based on the assumption that there is no heat loss from the environment to the process (or from said process to said process environment). The quality of commercial insulating materials allows this to be a very reasonable and typically very common assumption among those versed in the technique.

(0017] For reasons of convenience, the process parameters are detailed both in the traditional British units as in the units of the international system of units (SI). The molar flows indicated in the tables can be interpreted as either moles moles per hour or as kilograms moles per hour. The detailed energy consumption as horsepower (HP) and / or as thousands of units of British thermal units per hour (MBTU / hour) corresponds to the molar flow rates expressed as pounds moles per hour. The energy consumption reported as kilowatts (kW) corresponds to the molar flows defined in kilograms moles per hour. DESCRIPTION OF THE TECHNICAL BACKGROUND [0018] FIG. one It is a flow chart of a process where the design of a processing plant to recover C3 + components from natural gas is illustrated, using the prior art in accordance with US Patent No. 5,799,507. In this process simulation, the inlet gas enters the plant at 110 ° F [43 ° C] and 885 psia [6,100 kPa (a)] as current 31. If the inlet gas contains a concentration of sulfur compounds that could prevent the product streams from meeting the specifications, said sulfur compounds would be removed by

<img file="AR080751A1_D0005.tif" />

φ appropriate pretreatment of the gas supply (not shown). In addition, the feed stream is usually dehydrated to prevent the formation of hydrate (ice) under cryogenic conditions. For this, typically a solid desiccant is used.

[0019] Feed stream 31 is cooled in heat exchanger 10 by heat exchange with cold waste gas (stream 44), separator liquids rapidly expanded (stream 35a), and distillation liquids at -105 ° F [-76 ° C] (stream 43). The cooled stream 31a enters separator 11 at -34 ° F [-36 ° C] and 875 psia [6,031 kPa (a)] where the steam (stream 34) is separated from the condensed liquid (stream 35). The liquid from the separator (stream 35) expands at a pressure slightly higher than the operating pressure (approximately 375 psia [2,583 kPa (a)]) of the fractionation tower 15 via expansion valve 12, cooling stream 35a to -65 ° F [-54 ° C]. The current 35a enters the heat exchanger 10 to supply cooling to the gas feed as previously described, heating the current 35b to 40.56 ° C [41 ° C] before being supplied to the fractionation tower 15 at one point of feeding in the lower half of the column.

The steam (stream 34) coming from the separator 11 enters an expansion work machine 13 where the mechanical energy is extracted from this portion of the high pressure feed. The machine 13 expands the vapor substantially isentropically to the operating pressure of the fractionation tower 15, where the expansion work cools the expanded stream 34a to a temperature of approximately -100 ° F [-74 ° C]. Typical commercially available expanders have a capacity to

<img file="AR080751A1_D0006.tif" />

recovery in the order of 80-85% of the work available theoretically in an ideal isentropic expansion. Frequently recovered work is used to drive a centrifugal compressor (such as article 14) that can be used to recompress hot waste gas (stream 44a), for example. The partially condensed expanded current 34a is then supplied as feed in the fractionation tower 15 at a feed point in the upper half of the column.

[0021] The detanizer in tower 15 is a conventional distillation column containing a plurality of vertically separated trays, one or more filled beds or any combination of trays and fill. The dewatering tower consists of two sections: an upper absorption (rectification) section 15a containing the trays and / or the filling to provide the necessary contact between the vapor portion of the expanded current 34a that rises and the cold liquid that descends to condense and absorb C3 components and heavier components; and a lower depletion section 15b containing the trays and / or the filling to provide the necessary contact between the liquids that fall and the vapors that rise. The dewatering section 15b also includes at least one boiler (such as boiler 16) that heats and vaporizes a portion of the liquids that flow down the column to provide the depletion or drag vapors that ascend the column to deplete the liquid product, stream 37, methane, C components₂ and lighter components. The stream 34a enters the desetanizer 15 in a feeding position in the middle of the column located in the lower region of the absorption section 15a of the desetanizer 15. The liquid portion of the i

<img file="AR080751A1_D0007.tif" />

z expanded stream 34a intermingles with the liquids that descend from the absorption section 15a and the combined liquid continues to descend within the depletion section 15b of the detanizer 15. The portion of expanded stream steam 34a ascends through the absorption section 15a and makes contact with the cold liquid that falls by condensing and absorbing the C3 components and the heavier components.

[0022] A portion of the distillation vapor (stream 38) is removed from the upper region of the depletion section 15b. This stream is then cooled and partially condensed (stream 38a) in the heat exchanger 17 by heat exchange with the cold head current of the desetanizer 36 leaving the top of the desetanizer 15 at -109 ° F [-79 ° C ]. The cold-head current of the desetanizer is heated to approximately -33 ° F [-66 ° C] (current 36a) while cooling to current 38 from -30 ° F [-35 ° C] to approximately -103 ° F [- 75 ° C] (current 38a).

[0023] The operating pressure in the reflux separator 18 is kept slightly below the operating pressure of the degasser 15. This pressure difference provides the driving force that allows the distillation steam stream 38 to flow through the heat exchanger 17 and from there to the reflux separator 18 where the condensed liquid (stream 40) is separated from the uncondensed vapor (stream 39). The non-condensed steam stream 39 is combined with the heated dewatering head head 36a from the exchanger 17 to form the cold waste gas stream 44 at -37 ° C [-38 ° C].

[0024] The liquid stream 40 of the reflux separator 19 is pumped by the pump 15 at a pressure slightly higher than the pressure of

<img file="AR080751A1_D0008.tif" />

operation of the degasser 19. The resulting current 40a is then divided into two portions. The first portion (stream 41) is supplied as the upper cold feed of the column (reflux) to the upper region of the absorption section 15a of the detanizer 15. This cold liquid generates an absorption-cooling effect within the absorption (rectification) section 15a of the degasser 15, where the saturation of the vapors that ascend through the tower by vaporization of methane and liquid ethane contained in the stream 41 provides section cooling. Note that, as a result of it, both the steam that leaves the upper region (upper stream 36) and the liquids that leave the lower region (liquid distillation stream 43) of the absorption section 15a are colder than any of the supply currents (currents 41 and current 34a) of the absorption section 15a. This absorption-cooling effect allows the tower head (stream 36) to provide the necessary cooling in the heat exchanger 17 to partially condense the steam distillation stream (stream 38) without operating the depletion section 15b at a pressure significantly higher than the pressure of the absorption section 15a. This absorption-cooling effect also facilitates the condensation of the reflux current 41 and the absorption of the components C₃ and the heaviest components in the distillation vapor that rises through the absorption section 15a. The second portion (stream 42) of pumped stream 40a is supplied to the upper depletion section region 15b of the detanizer 15 where the cold liquid acts as reflux to absorb and condense the components C₃ and the heaviest components that are rising so that the steam distillation stream 38 contains minimal amounts of these components.

[0025] A stream of distilled liquid 43 from the degasser 15 is removed from the lower region of the absorption section 15a and sent to the heat exchanger 10 where it is heated while providing cooling of the incoming feed gas as described above. . Typically, the flow of said liquid from the degasser occurs through a thermosiphon circulation, but a pump could be used. The liquid stream is heated to -4 ° F [-20 ° C], partially vaporizing the stream 43a before returning it as a feed to the middle of the column to the detanizer 15, in the middle region of the depletion section 15b .

[0026] In the depletion section 15b of the detanizer 15, the feed streams are depleted of their methane content and components C₂. The liquid product stream 37 obtained as a result leaves the bottom of the tower at 201 ° F [94 ° C] based on a typical specification of a ratio of ethane to propane of 0.048: 1 on a molar basis of the product of the background. The cold waste gas (stream 44) goes into a countercurrent mode with respect to the incoming gas supply in the heat exchanger 10 where it is heated to 98 ° C [37 ° C] (stream 44a). Then, the waste gas is compressed again in two stages. The first stage comprises the compressor 14, driven by the expansion machine 13. The second stage comprises the compressor 20, driven by a supplementary energy source that compresses the waste gas (stream 44c) to the pressure of the sales line. After cooling to 120 ° F [49 ° C] in the discharge cooler 21, the waste gas stream 44d flows to the cafeteria

<img file="AR080751A1_D0009.tif" />

of gas for sales at 915psia [6,307 kPa (a)], a pressure that is sufficient to meet the requirements of the line (usually in the order of the intake pressure).

[0027] The following table presents a summary of the flow rates and energy consumption for the process illustrated in FIG.

Table I (FIG. 1)

Summary of flows: Lb. moles / hour [kg moles / hour]

<td>Current</td><td>Methane</td><td>Ethane</td><td>Propane</td><td>Butanes +</td><td>Total</td>

<td> 31</td><td> 19,419</td><td> 1,355</td><td> 565</td><td> 387</td><td> 21,961</td>

<td> 34</td><td> 18,742</td><td> 1,149</td><td> 360</td><td> 98</td><td> 20,573</td>

<td> 35</td><td> 677</td><td> 206</td><td> 205</td><td> 289</td><td> 1,388</td>

<td> 36</td><td> 18,400</td><td> 1,242</td><td> 3</td><td> 0</td><td> 19,869</td>

<td> 38</td><td> 2,759</td><td> 1,758</td><td> 15</td><td> 0</td><td> 4,602</td>

<td> 39</td><td> 1,019</td><td> 86</td><td> 0</td><td> 0</td><td> 1,116</td>

<td> 40</td><td> 1,740</td><td> 1,672</td><td> 15</td><td> 0</td><td> 3,486</td>

<td> 41</td><td> 1,044</td><td> 1,003</td><td> 9</td><td> 0</td><td> 2,092</td>

<td> 42</td><td> 696</td><td> 669</td><td> 6</td><td> 0</td><td> 1,394</td>

<td> 43</td><td> 1,388</td><td> 911</td><td> 365</td><td> 98</td><td> 2,796</td>

<td> 44</td><td> 19,419</td><td> 1,328</td><td> 3</td><td> 0</td><td> 20,985</td>

<td> 37</td><td> 0</td><td> 27</td><td> 562</td><td> 387</td><td> 976</td>

Recoveries

Propane

Butanes +

99,56%

100,00%

<img file="AR080751A1_D0010.tif" />

€

Compression of waste gas 9,868

Energy

Reflux pump

Totals

9,887

HP [16,223 kW] HP [31 kW]

HP [16,254 kW] * (Based on non-rounded flow values)

DESCRIPTION OF THE INVENTION [0028] In FIG. 2 a flow chart of a process according to the present is illustrated. The composition of the feed gas and the conditions considered in the process presented in FIG. 2 are the same as in FIG. 1 Accordingly, the process of FIG. 2 can be compared with the process of FIG. one [0029] In the simulation of the process of Figure 2, the intake gas enters the plant as stream 31 and enters a heat exchanger means in the cooling section of the feed 115a within the treatment assembly 115. This heat exchanger means may comprise a heat exchanger of the fin and tube type, a heat exchanger of the plate type, a heat exchanger of the welded aluminum type, or other type of thermal transfer device, including heat exchangers multipass and / or multiservices. The heat exchanger means is configured to provide heat exchange between the stream 31 flowing through a passage of the heat exchanger medium and the rapidly expanded separator liquids (stream 35a) and a waste gas stream from the section of condensation 115b inside the treatment set 115. The stream 31 is cooled while the rapidly expanded separator liquids and the gaseous waste stream are heated. A first portion (stream 32) of stream 31 is removed from the heat exchanger medium after stream 31 has been partially cooled to 25 ° F [-4 ° C], while the remaining second portion is still cooled (stream 33 ) such that it leaves the heat exchanger medium at -20 ° F [-29 ° C].

[0030] The separation section 115e has an internal head or other means for separating it from the detanization section 115d, so that two sections within the treatment assembly 115 can operate at different pressures. The first portion (stream 32) of stream 31 enters the lower region of separation section 115e at 875 psia [6,031 kPa (a)] where all the condensed liquid is separated from the steam before directing the steam into a mass and heat transfer medium within the separation section 115e. This mass and heat transfer means may also comprise a fin and tube type heat exchanger, a plate type heat exchanger, a welded aluminum type heat exchanger, or other type of thermal transfer device, including multipass and / or multiservice heat exchangers. The mass and heat transfer medium is configured to provide heat exchange between the vapor portion of the stream 32 flowing upward through a passage of the mass and heat transfer medium and the distilled liquid stream 43 from the absorption section 115c within the treatment assembly 115 flowing downward, such that the steam is cooled while the distilled liquid stream is heated. When the steam stream cools, a portion of it can condense and lower while the remaining steam continues to flow up through the medium of

<img file="AR080751A1_D0011.tif" />

mass and heat transfer. The mass and heat transfer medium provides continuous contact between the condensed liquid and the steam such that this also works by providing mass transfer between the vapor and liquid phases to provide partial steam rectification.

[0031] The second portion (stream 33) of stream 31 enters the separation section 115e within the treatment assembly 115 on the mass and heat transfer medium. All the condensed liquid is separated from the steam and mixed with all the liquid that condenses from the steam portion of the stream 32 flowing upward through the mass and heat transfer medium. The steam portion of the stream 33 is combined with the steam that exits the mass and heat transfer medium to form stream 34, which leaves the separation section 115e at -31 ° F [35 ° C]. The liquid portions (if any) of streams 32 and 33 and all the condensed liquid from the steam portion of stream 32 in the mass and heat transfer medium combine to form stream 35, which flows out of separation section 115e at -15 ° F [-26 ° C]. This expands at a pressure slightly higher than the operating pressure (approximately 383 psia [2,639 kPa (a)]) of the dewatering section 115d within the treatment assembly 115 by the expansion valve 12, cooling the current 35a to - 42 ° F [-41 ° C]. The stream 35a enters the heat exchanger medium in the cooling section of the feed 115a to supply cooling to the feed gas as described above, heating the stream 35b to 103 ° F [39 ° C] before supplying it to the section of detanization 115d within the treatment set 115 by a feeding point below the middle part of the column.

[0032] The steam (stream 34) from the separation section 115e enters an expansion work machine 13 where the mechanical energy is extracted from this portion of the high pressure feed. Machine 13 expands the vapor substantially isentropically at the operating pressure (approximately 380 psia [2,618 kPa (a)]) of the absorption section 115c, where the expansion work cools the expanded stream 34a to a temperature of approximately -98 ° F [-72 ° C]. Then, the partially condensed expanded current 34a is supplied as feed in the lower region of the absorption section 115c of the treatment assembly 115.

[0033] Absorption section 115c contains an absorption means consisting of a plurality of vertically spaced trays, one or more filled beds, or some combination of trays and padding. The trays and / or the filling in the absorption section 115c provide the necessary contact between the vapors that rise and the cold liquid that descends. The vapor portion of the expanded stream 34a is raised by rising through the absorption means of the absorption section 115c to be brought into contact with the cold liquid that falls to condense and absorb most of the components C3 and the heavier components. of these vapors. The liquid portion of the expanded stream 34a intermingles with the liquids that fall from the absorption medium in the absorption section 115c to form the distilled liquid stream 43, which is extracted from the lower region of the absorption section 115c to -102 ° F [-74 ° C]. The distillation liquid is heated to -9 ° F [-23 ° C] while cooling the steam portion of stream 32 in separation section 115e as described above, then supplying the heated distilled liquid stream 43a, to the detanization section 115d within the treatment assembly 115 by an upper feed point in the middle part of the column. Typically, the flow of said liquid from the absorption section 115c through the mass and heat transfer medium in the separation section 115e to the detanization section 115d occurs through a thermosiphon circulation, but a pump could be used.

[0034] The absorption section 115c has an internal head or other means for separating it from the detanization section 115d, such that the two sections within the treatment assembly 115 can operate with a slightly greater pressure on the detanization section 115d than that of the absorption section 115c. This pressure difference provides the driving force that allows a first stream of distillation steam (stream 38) to be extracted from the upper region of the detanization section 115d and directed to the heat exchanger means in the condensation section 115b within the treatment assembly 115 Similarly, this heat exchanger means may comprise a fin and tube type heat exchanger, a plate type heat exchanger, a welded aluminum type heat exchanger, or other type of thermal transfer device, including multipass and / or multiservice heat exchangers. The heat exchanger means is configured to provide heat exchange between the first distillation steam stream 38 flowing through a passage of the heat exchanger medium and a second distillation steam stream that rises from the heat section.

Absorción / L · absorption 115c within the treatment set 115. The second stream of distillation steam is heated while cooling to stream 38 and condenses it at least partially, which then leaves the heat exchanger medium and separates in their respective vapor and liquid phases. The vapor phase (if any) is combined with the second stream of heated distillation steam leaving the heat exchanger medium to form the gaseous waste stream that provides cooling in the cooling section of the feed 115a as described. previously. The liquid phase is divided into two parts, currents 41 and 42.

[0035] The first portion (stream 41) is supplied as cold feed from the top of the column (reflux) to the upper region of the absorption section 115c within the treatment assembly 115 by gravity flow. Said cold liquid causes an absorption cooling effect within the absorption (rectification) section 115a, where saturation of the vapors that rise up the tower by vaporization of methane and ethane liquids contained in stream 41 provides cooling to the section. This absorption cooling effect allows the second distillation steam stream to provide the necessary cooling in the heat exchanger medium of the condensation section 115b to partially condense the first distillation steam stream (stream 38) without operating at the detanization section 115d at a pressure significantly greater than that of the absorption section 115c. This absorption-cooling effect also facilitates the condensation of the reflux current 41 and the absorption of the components C₃ and the heaviest components in the distillation vapor that rises through the absorption section 115c. The second portion (stream 42) of the separated liquid phase in the condensation section 115b is supplied as cold feed from the top of the column (reflux) to the upper region of the detanization section 115d within the treatment assembly 115 by a gravity flow, such that the cold liquid acts as a reflux to absorb and condense the components C3 and the heavier components that flow from the bottom up such that the distillation steam stream 38 contains minimal amounts of said components.

[0036] The dewatering section 115d within the treatment assembly 115 contains a mass transfer medium consisting of a plurality of vertically spaced trays, one or more filled beds, or some combination of trays and fill. The trays and / or the filling of the dewatering section 115d provide the necessary contact between the rising vapors and the falling cold liquid. The detanization section 115d also includes a mass and heat transfer medium below the mass transfer medium. This mass and heat transfer means may also comprise a fin and tube type heat exchanger, a plate type heat exchanger, a welded aluminum type heat exchanger, or other type of thermal transfer device, including multipass and / or multiservice heat exchangers. The mass and heat transfer medium is configured to provide heat exchange between a heating medium that flows through a passage of the mass and heat transfer medium and a stream of distilled liquid that flows down from the transfer medium from

<img file="AR080751A1_D0012.tif" />

mass in the dewatering section 115d, such that the stream of distilled liquid is heated. As the distillation liquid stream is heated, a portion of it is vaporized to form the separation vapors, which rise as the remaining liquid continues to flow down through the heat and mass transfer medium. The heat and mass transfer medium provides continuous contact between the separation vapors and the distillation liquid stream, so that it also functions to provide the mass transfer between the vapor and liquid phases, so that separate the methane, the components of C₂and the lighter components of the liquid product stream 37. The liquid product that is obtained as a result (stream 37) leaves the lower region of the dewatering section 115d and leaves the treatment set 115 at 203 ° C [95 ° C ], [0037] The second stream of distillation steam that rises from the absorption section 115c is heated in the condensation section 115b while providing cooling to the stream 38 as described above. The second stream of heated distillation steam is combined with all steam that is separated from the first stream of cooled distillation steam 38 as described above. The waste gas stream that is obtained as a result is heated in the cooling section of the feed 115a while providing cooling to the stream 31 as described above, after which the waste gas stream 44 leaves the treatment set 115 a 104 ° F [40 ° C]. Subsequently, the waste gas stream is compressed again in two stages, the compressor 14, operated by the expansion machine 13, and the

VG '©

compressor 20, operated by a supplementary power source. After cooling to 120 ° F [49 ° C] in the discharge chiller 21, the waste gas stream 44c flows to the gas pipe for sales at 915 psia [6,307 kPa (a)], a pressure that is sufficient to meet the requirements of the line (usually in the order of admission pressure).

[0038] The following table presents a summary of the flow rates and energy consumption for the process illustrated in FIG. two.

Table II (FIG. 2)

Summary of flows: Lb. moles / hour [kg moles / hour]

<td>Current</td><td>Methane</td><td>Ethane</td><td>Propane</td><td>Butanes +</td><td>Total</td>

<td> 31</td><td> 19,419</td><td> 1,355</td><td> 565</td><td> 387</td><td> 21,961</td>

<td> 32</td><td> 4,855</td><td> 339</td><td> 141</td><td> 97</td><td> 5,490</td>

<td> 33</td><td> 14,564</td><td> 1,016</td><td> 424</td><td> 290</td><td> 16,471</td>

<td> 34</td><td> 18,870</td><td> 1,135</td><td> 348</td><td> 104</td><td> 20,683</td>

<td> 35</td><td> 549</td><td> 220</td><td> 217</td><td> 283</td><td> 1,278</td>

<td> 38</td><td> 2,398</td><td> 1,544</td><td> 13</td><td> 0</td><td> 4,015</td>

<td> 41</td><td> 1,018</td><td> 868</td><td> 8</td><td> 0</td><td> 1,924</td>

<td> 42</td><td> 737</td><td> 628</td><td> 5</td><td> 0</td><td> 1,394</td>

<td> 43</td><td> 1,112</td><td> 723</td><td> 353</td><td> 104</td><td> 2,320</td>

<td> 44</td><td> 19,419</td><td> 1,328</td><td> 3</td><td> 0</td><td> 20,984</td>

<td> 37</td><td> 0</td><td> 27</td><td> 562</td><td> 387</td><td> 977</td>

Recoveries *

Propane

99,63 %

Butanes + 100.00%

Energy

Compression of the residual gas 9,363 HP [15,393 kW] * (Based on non-rounded flow values) [0039] When comparing tables I and II, it is observed that the same recoveries are essentially maintained as with the technical background. However, by continuing with the comparison between Tables I and II, it is shown that product yields were obtained using significantly less energy than with the technical background. In terms of recovery efficiency (defined from the amount of propane recovered per unit of energy), this represents an improvement greater than 5% compared to the technical background process, which is represented in FIG. 1 [0040] The improvement in recovery efficiency provided herein by comparison with the technical background process, which is depicted in FIG. 1, is primarily due to three factors. First, the compact arrangement of the heat exchanger means in the cooling section of the feed 115a and the condensation section 115b in the treatment assembly 115 suppresses the pressure drop imposed by the interconnection cafeteria found in the plants of conventional treatment. The result is that in the present invention the gaseous residue flowing into the compressor 14 is at a higher pressure compared to the prior art, such that the gaseous residue entering the compressor 20 is at a significantly higher pressure, reducing

<img file="AR080751A1_D0013.tif" />

in that way the power required by the present invention to return the pressure of the pipe to the gaseous residue.

[0041] Second, the use of the mass and heat transfer medium in the detanization section 115d to simultaneously heat the distillation liquid leaving the mass transfer medium in the detanization section 115d while allowing that the vapors that are obtained as a result come into contact with the liquid and deplete its volatile components is more efficient than using a conventional distillation column with external boilers. Volatile components are separated from the liquid continuously, thereby reducing the concentration of volatile components in the separation vapors with greater speed and improving the separation efficiency for the present.

(0042) Third, the use of the mass and heat transfer medium in the separation section 115e to simultaneously cool the steam portion of the stream 32 while condensing the heavier hydrocarbon components of the steam provides a partial rectification of the current 34 before subsequently expanding and supplying it as feed to the absorption section 115c. As a result, a lower flow of reflux (stream 41) is necessary to rectify the expanded stream 34a to separate from it the components C3 and the heavier hydrocarbon components, as can be seen by comparing the flow rate of stream 41 in the Tables I and II.

[0043] The present invention has two additional advantages compared to the technical background, in addition to the increase in processing efficiency. First, the compact layout of the

<img file="AR080751A1_D0014.tif" />

Ό kr treatment assembly 115 of the present invention replaces six separate equipment units of the prior art (heat exchangers 10 and 17, separator 11, reflux separator 18, reflux pump 19, and fractionation tower 15 in Figure 1) with a single unit of equipment (the treatment set 115 in Figure 2). This reduces the requirements of the size of the plot of land, eliminates the pipeline of interconnection, and suppresses the power consumption by the reflux pump, reducing the cost of capitai and the operating cost of a treatment plant where the present one is used invention with respect to one of the prior art. Secondly, the elimination of the interconnection pipe means that a treatment plant where the present invention is used has much fewer connections with flanges compared to the prior art, reducing the number of potential sources of filtration in the plant. Hydrocarbons are volatile organic compounds (VOCs), some of which are classified as greenhouse gases, and some of which can be precursors for the formation of atmospheric ozone, so that this allows reducing the potential for atmospheric emissions that may occur. damage the environment.

Other embodiments [0044] Certain circumstances may favor the elimination of the cooling section of the feed 115a and the condensation section 115b of the processing assembly 115, and the use of one or more means for heat exchange external to the processing assembly to cool the feed and condense the reflux, as for example the heat exchangers 23 and 17 shown in Figures 14 to 21. An arrangement with said characteristics allows the processing set 115 to be more

<img file="AR080751A1_D0015.tif" />

small, which can reduce the total cost of the plant and / or shorten the production agenda. Note that in all cases the exchangers 23 and 17 are representative of either a plurality of individual heat exchangers or a single multi-pass heat exchanger, or any combination thereof. Each of said heat exchangers may comprise a heat exchanger of the type of tube and fins, a heat exchanger of the plate type, a heat exchanger of the welded aluminum type, or other type of heat transfer device, including multipass and / or multiservice heat exchangers. In some cases, it may be advantageous to combine the cooling of the feed and the reflux condensation in a single multi-service heat exchanger. With the heat exchanger 17 external to the processing assembly, the reflux separator 18 and the pump 19 will typically be necessary to separate the condensate stream 40 and supply at least a portion thereof to the absorption section 115c as reflux.

[0045] As described above for the embodiment of the present invention shown in Figure 2, the first distillation steam stream 38 is partially condensed and the condensate obtained as a result is used to absorb the valuable components C3 and heavier components of the vapors leaving the expansion work machine. However, the present invention is not limited to this embodiment. For example, it may be advantageous to treat only a portion of the steam from the outlet from the expansion work machine in this manner, or to use only a portion of the condensate as absorbent, in cases where other design considerations indicate that some portions

<img file="AR080751A1_D0016.tif" />

from the outlet of the expansion machine or condensate should be diverted to avoid the absorption section 115c of the treatment assembly 115. The conditions of the feed gas, the size of the plant, the available equipment, or other factors may indicate that the removal of the expansion work machine 13, or its replacement by an alternative expansion device (such as an expansion valve ), is feasible, or that the total (rather than partial) condensation of the first distillation steam stream 38 in the condensation section 115b within the treatment assembly 115 (Figures 2 to 13) or the heat exchanger 17 (Figures is possible or preferred) 14 to 21). It should also be noted that, depending on the composition of the feed gas stream, it may be advantageous to use external cooling to provide partial cooling to the first distillation steam stream 38 in the condensation section 115b (Figures 2 to 13 ) or heat exchanger 17 (Figures 14 to 21).

[0046] In certain circumstances, it may be advantageous to use an external separator vessel to separate the first and second cooled portions 32 and 33 or the cooled feed stream 31 a, instead of including the separation section 115e in the treatment assembly 115 Such as shown in Figures 8 and 18, a mass and heat transfer means can be used in the separator 11 to separate the first and second cooled portions 32 and 33 in the steam stream 34 and the liquid stream 35. Similarly, such a horn is shown in Figures 9a13y19a21y19a, the separator 11 can be used to separate the cooled feed stream 31a to give the steam stream 34 and the liquid stream 35.

[0047] The use and distribution of the liquid stream 35 from the separation section 115e or the separator 11 and the distilled liquid stream 43 from the absorption section 115c to exchange the heat of the process, the particular arrangement of the heat exchangers for cooling the feed gas (streams 31 and / or 32) and the first stream of distillation steam 38, and the selection of process streams for specific services for heat exchange should be evaluated for each particular application. For example, Figures 4, 6, 10, 12, 16 and 20 show the use of the distilled liquid stream 43 to supply a portion of the cooling of the first distillation steam stream 38 in the condensation section 115b (Figures 4 , 5, 10, and 11), the heat exchanger 10 (Figures 6 and 12), or the heat exchanger 17 (Figures 16 and 20). In such cases, a mass and heat transfer medium may not be necessary in the separation section 115e (Figures 4 to 6 and 16) or the separator 11 (Figures 10 to 12 and 20). In the embodiments shown in Figures 4 and 10, a pump 22 is used to supply the stream of distilled liquid 43 to the heat exchanger medium in the condensation section 115b. In the embodiments shown in Figures 5 and 11, the condensation section 115b is located below the absorption section 115c in the treatment assembly 115 such that the flow of the distilled liquid stream 43 occurs by A thermosiphon circulation. In the embodiments shown in Figures 6 and 12, a heat exchanger 10 external to the treatment assembly 115 is used and the cooling section of the feed 115a is located below the absorption section 115c in the assembly of treatment 115 such that the flow of the distilled liquid stream 43 occurs through a thermosiphon circulation.

(In the embodiments shown in Figures 5, 6, 11, and 12 a reflux pump 19 is used to deliver the reflux to the positions on the point of the treatment assembly 115 where the condensed liquid phase of the current 38). In the embodiments shown in Figures 16 and 20, the thermosiphon circulation may be sufficient to allow the stream of distillation liquid 43 to flow through the heat exchanger 17, or the pump 22 may be necessary to circulate the current 43. Certain circumstances may favor the use of the distilled liquid stream 43 to cool the stream 32 in a heat exchanger external to the treatment assembly 115, such as the heat exchanger 10 shown in Figures 3, 9, 15 , and 19. In other circumstances they may favor not heating the distilled liquid stream 43 at all, and instead using the distilled liquid stream 43 as the reflux that is sent to the upper region of the detanization section 115d such as is shown in Figures 7, 13, 17 and 21. (For the embodiment shown in Figures 13 and 21, pump 22 may be necessary because the gravity flow of current 43 may not be possible).

[0048] Depending on the amount of heavier hydrocarbons in the feed gas and the pressure of the feed gas, it may be that the first and second cooled portions 32 and 33 entering the separation section 115e in Figure 2 and 14 or the separator 11 in Figure 8 and 18 (or the cooled feed stream 31a entering the section of separation 115e in Figures 3 to 7 and 15 to 17 or the separator 11 in Figures 9a13 and 19a21) do not contain all the liquid (because it is above

Ì3) e

his dew point, or because he is on his cricondenbara). In such cases, there is no liquid in stream 35 (such as shown by the cut lines). In such circumstances, the separation section 115e of the treatment assembly 115 (Figures 2 to 7 and 14 to 17) or the separator 11 (Figures 8a13 and 18a21) may not be necessary.

[0049] In accordance with the present invention, external cooling may be used to supplement the available cooling for the intake gas and / or the first distillation steam stream from the second distillation steam stream and the liquid stream distillate, in particular in the case of a rich intake gas. In those cases where additional cooling of the intake gas is desired, a mass and heat transfer means may be included in the separation section 115e (or a gas collection means in those cases where the first and second portions chilled 32 and 33 or the chilled feed stream 31a does not contain liquid) such horn is shown by the lines cut in Figures 3a7 and 15a17, This means that a mass and heat transfer medium can be included in the separator 11 as shown by the lines cut in Figures 9 to 13 and 19 to 21. This mass and heat transfer medium may comprise a fin and tube type heat exchanger, a plate type heat exchanger, a welded aluminum type heat exchanger, or other type of thermal transfer device, including multipass and / or multiservice heat exchangers. The mass and heat transfer medium is configured to provide heat exchange between a refrigerant stream (eg, propane) flowing through a passage of the mass and heat transfer medium and the steam portion of

<img file="AR080751A1_D0017.tif" />

the stream 31a flowing upward, such that the refrigerant further cools the vapor and condenses more liquid, which falls to form part of the liquid that was extracted in the stream 35. Such a horn is shown by the lines cut in Figures 2 , 8,14, and 18, the mass and heat transfer medium in the separation section 115e (Figures 2 and 14) or the separator 11 (Figures 8 and 18) may include means that allow to provide supplementary cooling with a refrigerant. Alternatively, conventional gas cooler (s) could be used to cool stream 32, stream 33, and / or stream 31 a with refrigerant before streams 32 and 33 enter the section of separation 115e (Figure 2 and 14) or separator 11 (Figures 8 and 18) or that current 31a enters separation section 115e (Figures 3 to 7 and 15 to 17) or separator 11 (Figures 9 through from 13 and 19 to 21). In cases where additional cooling of the first distillation steam stream is desired, the heat exchanger means in the condensation section 115b of the treatment assembly 115 (Figures 2 to 5, 7 to 11, and 13), the heat exchanger Heat 10 (Figures 6 and 12), or the heat exchanger 17 (Figures 14 to 21) may include means that allow to provide supplementary cooling with a refrigerant, such as is shown by the cut lines.

[0050] Depending on the type of thermal transfer devices selected for the heat exchanger medium in the cooling section of the feed 115a and the condensation section 115b (Figures 2 to 5, 7 to 11, and 13), It may be possible to combine said heat exchangers in a single multi-pass and / or multi-service thermal transfer device. In such cases, the transfer device

<img file="AR080751A1_D0018.tif" />

multipass and / or multiservice thermics will induce appropriate means to distribute, segregate, and collect at stream 31, stream 32, stream 33, first distillation steam stream 38, at all steam that separates from the cooled stream 38 , and the second stream of steam distillation to achieve the desired cooling and heating.

[0051] It will also be recognized that the relative amount of condensed liquid that is divided between streams 41 and 42 in Figures 2 to 6, 8 to 12, 14 to 16, and 18 to 20 will depend on several factors, including pressure of the gas, the composition of the feed gas, and the available power. The optimal division generally cannot be predicted without evaluating the particular circumstances of a specific application of the present invention. Certain circumstances may favor the feeding of all the condensed liquid to the upper region of the absorption section 115c in the stream 41 and nothing to the upper region of the dewatering section 115d in the stream 42, such as is shown by the cut lines for stream 42. In such cases, the stream of heated distilled liquid 43a can be supplied to the upper region of the dewatering section 115d to serve as reflux.

[0052] The present invention provides improved recovery of the C components₃ and the heaviest hydrocarbon components by amount of utility consumption necessary to operate the process. An improvement in the consumption of the utilities necessary to operate the process can be seen in the form of lower power requirements for compression or recompression, lower power requirements for external cooling, lower energy requirements for tower overheating, or a combination thereof.

[0053] As long as what is considered as preferred embodiments of the invention has been described, those skilled in the art will understand that it is possible to make further modifications and modifications thereto, for example to adapt the invention to different conditions, types of food or other requirements, without departing from the spirit of the present invention defined in the following claims.