Abstract:
Disclosed is an improved cryogenic demethanizer, for separating an inlet hydrocarbon gas having a mixture of hydrocarbon components into a residual lighter gas fraction and a heavier liquid fraction. The fractionation column in the demethanizer has a main body portion and an upper portion enlarged with respect to the main body portion. A packing which may be in the form of a plurality of contact trays, or random packing is located in the upper enlarged portion of the column. The invention may be used to retrofit existing cryogenic demethanizers, or used in new installations.

Description:
FIELD OF THE INVENTION 
     The present invention relates to an apparatus for and a method of separating or fractionating a hydrocarbon gas, comprising at least methane and ethane, into a residue gas comprising mainly methane and a heavier fraction comprising principally ethane and other heavy hydrocarbons, and to a method of retrofitting an existing cryogenic plant or apparatus so as to be capable of carrying out the method. The gas will usually be natural gas, received as either a gas or a liquid, and including ethane, propanes, butanes and higher or heavier hydrocarbons. More specifically, the invention relates to an improved fractionation column to separate ethane and heavier components, often referred to as “ethane plus” cryogenically from methane and heavier hydrocarbons. The column may be used in new facilities, or may be used as a retrofit with existing facilities. 
     BACKGROUND TO THE INVENTION 
     Cryogenic techniques and apparatus to separate methane and/or ethane (as well as other lighter gases such as carbon dioxide and nitrogen) from a gas containing a mixture of hydrocarbon gases have long been known. Such an apparatus is commonly called a “demethanizer”. While the term “demethanizer” will be used throughout, it is to be understood that the apparatus and methods described may also be used in other applications whereby the demethanizer would be operated as a deethanizer. Typically, the purpose of a gas processing facility is to receive a gas from a transmission line, efficiently cool and depressurize the gas, extract the more valuable heavier components (ethane and heavier hydrocarbons, referred to as “ethane plus”), reheat and recompress the gas, and feed it back into the transmission line. 
     Typically, in such an operation, an inlet gas is introduced into a process facility at a high pressure. The gas is then allowed to expand and cool in various stages, and the liquid and gas fractions are then separated into different streams. After several stages of expansion and cooling, the various hydrocarbon streams are introduced into a fractionation column, or demethanizer, at different heights in the column. The methane gas is then separated from the heavier components of the gas, with the methane component exiting the fractionation column through the top of the column as a residue gas, and the heavier components exiting the column at its lower portion (collected as a liquid). 
     In the fractionation column, there is typically located a packing, which may be in the form of a number of contact trays. However, any suitable packing or construction can be used that promotes contact between the vapour and liquid flows. Thus, conventional Raschig rings or other standard packing materials can be used in the column. This packing is designed to increase mass transfer contact between the falling liquids and rising gases within the column, which increases the efficiency of liquid-gas separation in the fractionation column. As well, there is usually located at the upper portion of the fractionation column an enlarged empty “disengagement section”. Typically, the stream entering the fractionation column at its upper portion (in the disengagement section) comprises between about 5% and about 50% liquid phase when it enters the fractionation column. The disengagement section allows the liquids entering the fractionation column space to separate or “deentrain” from the vapours with which it is mixed. The empty disengagement section is designed to alleviate potential problems of carryover of hydrocarbon liquids into the demethanizer overhead stream (which is supposed to be vapour). This results in a loss of hydrocarbon liquid product. Conventionally, the disengagement section is of a large diameter (about 10 to 18 feet) and is about ten to fifteen feet in height. 
     It has also been known that, to increase the efficiency of a standard demethanizer, to increase the amount of heavier hydrocarbons removed, additional packing or contact trays may be added in a separate column, which is connected with the fractionation column. In order to obtain a useful increase in efficiency, however, it has been thought that at least eight additional trays, and as many as twenty additional contact trays were required in this separate column. Therefore, the modification of existing facilities to obtain an increase in efficiency has previously been very expensive, since a whole new column (and all of the additional equipment associated with it) was required. 
     One known process for separating different components in a hydrocarbon gas stream is disclosed in U.S. Pat. No. 4,278,457 (to Campbell and assigned to The Ortloff Corporation). The claimed improvement in that patent is to divide the feed gas stream into two separate streams, one of which is cooled and then depressured through an expander, while the other of which is cooled to a greater degree and is then depressured through a simple expansion valve. The separate first and second streams are then supplied to the fractionation column at different feed points. 
     This U.S. patent discloses a number of different examples, which either operate at a flow rate of 6588 lb moles/hr or at a quarter of this flow rate, i.e. 1647 lb moles/hr. Interestingly, in three examples, where the input gas flow is separated into three separate inputs, and where a substantial portion, as much as 76%, of the input gas flow, passes either into the middle or bottom inlet of the column, the lower flow rate of 1647 lb moles/hr is given. Moreover, this patent does not discuss, in any way, details of the column design, in terms of diameter required at different heights, number of trays or number of trays between different inlets. 
     Even more particularly, in this Ortloff patent, there is no detailed direction as to the number of trays that might be required between a middle inlet and a top inlet to a fractionation column, where the major portion of the gas is supplied to the middle inlet, and a small portion is provided, substantially in the liquid phase, to the top inlet. 
     Supplying a substantial portion of the inlet gas to a middle inlet of the column gives a number of problems. In particular, this has a significant impact on the dimensions of the column. By far the largest portion of the feed gas is methane, which in the column travels upwards as a vapour. When a major portion of the supply gas is fed to one inlet, the column, above that inlet, will need to have a sufficient diameter to accommodate the upward flow of methane and heavier components in the vapour phase. To keep the methane velocities reasonable, and in known manner to prevent entrainment of the heavier component liquid droplets to be carried overhead out of the column with the methane, requires a large diameter for the column. Clearly, this diameter or cross-section will be related to the intended flow rate through the apparatus. 
     This requirement for a relatively large cross-section where there is a large flow of methane in the vapour phase, has often resulted in a second column being provided, as mentioned above. Thus, in view of the conventional teaching that a significant number of trays would need to be provided in any such enlarged upper section, this would often result in a top section for the column that was simply too large to be supported on top of the lower section. For this reason, such a section was often provided as a separate column. 
     Other hydrocarbon gas separation techniques and apparatus can be found in U.S. Pat. No. 3,702,541 (Randall et al.), U.S. Pat. No. 4,519,824 (Huebel) and U.S. Pat. No. 5,566,554 (Vijayaraghavan et al.). All of these patents disclose relatively complex techniques. What is also striking about all of these three proposals is that no portion of the inlet gas stream is taken off and fed separately to the fractionating column. Rather, it appears that the inventors, in all of these cases, have assumed that it is advantageous to achieve some separation of the inlet gas, before feeding this into the fractionation column. The assumption appears to be that if there is some initial separation, e.g. liquid/vapour separation, then this will improve the overall performance of the system. Thus, all of these proposals provide at least one separator in which vapour and gas phases are separated. 
     The Randall et al. patent is of interest, since the inlet gas flow is split into three separate streams, by way of liquid and vapour separators, to provide three separate inlets to the fractionation column, these being provided at various levels in the lower stripping section of the column. The column here is provided with an upper rectifying section of larger diameter. Gas taken off from the top of the column is separated into liquid and vapour fractions, and the liquid fraction is pumped back up to the top of the column, and fed into the top of the rectifying section. 
     Again, there is no discussion of the size problems where a substantial portion of inlet gas is fed in below the top of the fractionation column, resulting in a substantial flow upwards through the top part of the column. This problem becomes particularly acute, where a plant is designed for large flow rates. Thus, the present invention is intended to provide a plant or apparatus suitable for a flow rate as large as a thousand (1000) MMSCFD, approximately equivalent to 116500 lb moles/hr, i.e. a flow rate that is an order of magnitude or more greater than that in some of the prior art proposals discussed above. In conventional design practices, for such a large flow rate, the upper part of the fractionation column can require diameters approaching 20 feet, and it is impractical to support such large diameter sections on top of much smaller diameter lower sections. 
     Another common characteristic of all of these earlier proposals is that there is no detailed investigation or consideration of the behaviour at the top of the fractionation column. Conventional teaching is that a substantial disengagement section needs to be provided above the top trays, to ensure adequate and complete separation of vapour and liquid phases, so as to ensure that liquid droplets are not carried over by the vapour flow leaving the top of the fractionation column. As detailed below, on an actual plant implementation of this invention, the size of the disengagement space was reduced from 10½ to 2½ feet and no liquid carry over has been observed during several months of operation. 
     SUMMARY OF THE INVENTION 
     In accordance with one aspect of the present invention, there is provided an apparatus for cryogenically separating an inlet hydrocarbon gas stream comprising at least methane and ethane into a residue gas stream comprising a major portion of the methane and a heavier hydrocarbon fraction comprising principally ethane and other heavier hydrocarbons, the apparatus comprising: 
     (1) a main inlet for the hydrocarbon mixture; 
     (2) a main stream connected to the main inlet and including a separator for separating liquid and vapour phases, the separator having a liquid phase outlet and a vapour phase outlet; 
     (3) a liquid phase stream connected to the liquid phase outlet and a vapour phase stream connected to the vapour phase outlet; 
     (4) a branch stream connected to the main inlet; 
     (5) a fractionation column including first, second and third inlets, the first inlet being connected to the liquid phase stream, the second inlet being connected to the vapour phase stream and the third inlet being connected to the branch stream, whereby the flow through the third inlet has substantially the same composition as the inlet gas, the second inlet being provided above the first inlet and the third inlet being provided above the second inlet; 
     (6) a first outlet means located at the bottom of the fractionation column for the heavier hydrocarbon fraction and a second outlet means at the top of the fractionation column for the residue stream; 
     (7) an outlet conduit connected to the second outlet means; 
     (8) first means for cooling the branch stream, including a first heat exchanger means provided between the outlet conduit and the branch stream for heat exchange therebetween, for cooling the incoming hydrocarbon gas branch stream and heating the residue gas stream in the outlet conduit; and 
     (9) second means for expanding and cooling the main stream. 
     Preferably, the fractionation column includes a body portion of relatively small diameter and an upper portion of relatively large diameter, and a transition section between the body portion and the upper portion, and the first inlet is provided towards the upper end of the body portion, the third inlet is provided in the upper portion of relatively large diameter, and the second inlet is provided in the transition section. 
     The fractionation column preferably includes packing comprising a plurality of trays. An important aspect of the present invention is the discovery that these trays can provide a tray efficiency much greater than usual. As such only a small number of trays is required, and there could be from one to six trays in the upper portion. Also, a disengagement zone much smaller than that usually provided can be used, as it has also been discovered that this still gives adequate separation and disengagement. 
     Another aspect of the present invention is directed to retrofitting an existing plant and provides a method of retrofitting an existing cryogenic apparatus for separating a compressed inlet hydrocarbon gas stream comprising at least methane and ethane into a residue fraction comprising a major portion of the methane and a heavier hydrocarbon fraction comprising principally ethane and other heavier hydrocarbons, said existing apparatus comprising: 
     (a) a main inlet for the hydrocarbon gas; 
     (b) means for expanding and cooling the inlet gas into a mixture of liquid and vapour phases connected to the main inlet; 
     (c) means for separating said liquid and vapour phases comprising: 
     (i) a fractionation column having a body portion and an upper portion above the body portion and generally enlarged with respect to the body portion, said upper portion being substantially empty and being originally intended to provide a disengagement zone; 
     (ii) at least one inlet means for supplying the vapour and liquid phases to the column provided on at least the body portion and connected to the means for expanding and cooling the inlet gas; 
     (iii) a first outlet means located at the bottom of the bottom portion for withdrawing the heavier hydrocarbon fraction from the column; 
     (iv) a second outlet means located at the top of the upper portion for withdrawing the residue stream from the column; and 
     (v) packing means in said body portion for increasing the amount of contact between the liquid and vapour phases; the method comprising: 
     (1) providing additional packing in the disengagement zone in the upper enlarged portion of the column to provide additional contact between the liquid and vapour phases; 
     (2) providing a branch stream between the main inlet and the upper portion; and 
     (3) a first heat exchange means in the branch stream, for cooling the branch stream so that at least a portion of the branch stream is in the liquid phase, whereby hydrocarbons discharging into the upper portion from the branch stream are at least partially liquid. 
     Again the additional packing in the upper portion of the column can comprise a plurality of trays, e.g. one to six, spaced vertically in said upper portion of the column. There can be three trays provided in an upper portion which is approximately 12½ feet in diameter and approximately 10 feet high, and the trays are positioned in the upper portions so as to leave a disengagement zone of less than three feet. 
     The upper portion can include a frustro-conical portion, and in this case one of the trays is provided extending into the frustro-conical portion. 
     Yet another aspect of the present invention provides a method of separating a hydrocarbon feed gas stream comprising at least methane and ethane into a residue fraction comprising a major portion of the methane and a heavier hydrocarbon fraction comprising principally ethane and other heavier hydrocarbons, the method comprising: 
     (1) passing the hydrocarbon gas stream through a first heat exchange means to cool the gas stream; 
     (2) separating the gas stream into first, second and third streams, with the second stream comprising a major portion of the gas flow; 
     (3) expanding the gas, after it has been cooled, to lower the temperature of the first, second and third streams, wherein the third stream, after cooling and expansion, is substantially in the liquid phase; 
     (4) providing a fractionation column comprising a lower portion having a relatively small diameter, an upper portion having a relatively large diameter, and a transition section therebetween and a packing within the lower and upper portions, wherein the packing comprises, for the upper portion, a small number of trays; 
     (5) supplying the third stream to the top of the upper portion, the second stream to the top of one of the transition section and one of the upper and lower portions adjacent the transition section and the first stream to the lower portion below the second stream; 
     (6) collecting the residue gas stream from the top of the upper portion and the heavier hydrocarbon fraction from the bottom of the lower portion, wherein the residue gas stream passes through the first heat exchange means to reheat the residue gas stream and cool the incoming gas. 
     Preferably, the method comprises: splitting the inlet hydrocarbon gas stream into a main stream and a branch stream, the branch stream comprising the third stream and the main stream being subsequently split into the first and second streams. Furthermore, a portion of the hydrocarbon feed gas stream can be split off into a further branch stream, and the further branch stream is then passed through reboiling means and then recombining the further branch stream with the main stream, the reboiler means, reboiling lighter hydrocarbon fractions from the fractionation column. 
     One embodiment of both the method and the apparatus of the present invention includes a static mixer, to which the vapour phase outlet of the expander and the branch or third stream are connected, wherein the second and third inlets to the fractionation column are combined and are connected to the outlet of the static mixer, the combined second and third inlets being provided at the top of the fractionation column, whereby the static mixer causes contacting between the branch stream and the liquid phase stream prior to the fractionation column. 
     Advantageously, the static mixer is sized to provide mass transfer substantially equivalent to one theoretical stage of contacting in a fractionation column. 
     In some instances, a separator may be utilized to separate the two phase stream leaving the expander. In this type of process configuration, the vapours from the separator are not directed to the demethanizer column but join directly with the overhead vapours leaving the demethanizer column. The liquids from the separator are generally sent as the top feed to the demethanizer column. This type of process results in a smaller column diameter but has the requirement of an additional separator vessel. Retrofitting this type of process is somewhat different since it is generally less effective, with respect to ethane recovery, to install trays in the overhead disengagement section. The static mixer approach is then useful since the third stream mentioned above, which is essentially condensed inlet gas, can be mixed with the expander outlet and sent through a static mixing device upstream of the separator vessel. This results in substantial additional cooling and absorption of the expander outlet stream and produces significantly more liquid feed to the demethanizer column which results in higher recoveries of ethane and heavier components. 
    
    
     BRIEF DESCRIPTION OF THE DRAWINGS 
     These and other advantages of the present invention will be more fully and completely understood, when the following detailed description of the preferred embodiment is read in connection with the following drawings, in which: 
     FIG. 1 is a schematic drawing of a conventional apparatus for separating methane from a hydrocarbon gas; 
     FIG. 2 is a schematic drawing of an apparatus in accordance with one aspect of the present invention; 
     FIG. 3 is a drawing of a theoretical model of an apparatus, similar to the apparatus of FIG. 1; 
     FIG. 4 is a drawing of a theoretical model of an apparatus, similar to the apparatus of FIG. 2; 
     FIG. 5 a  is a drawing of a theoretical model of an apparatus in accordance with another aspect of the present invention; 
     FIG. 5 b  is a drawing of a variant of the theoretical model of the apparatus shown in FIG. 5 a;    
     FIG. 5 c  is a drawing of a theoretical model which includes a static mixer for a plant with a second separator which separates expander vapours and bypasses them around the demethanizer column; 
     FIG. 6 is an elevational view of the top part of the fractionation column of FIG. 4, showing trays; and 
     FIG. 7 is a schematic plan view showing the trays of FIG.  6 . 
    
    
     DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENT 
     Referring to FIG. 1, there is shown schematically a typical conventional cryogenic separation plant, referred to generally by reference numeral  10 . The plant  10 , has an expansion and cooling area  15 , a fractionation column demethanizer  20 , and a recompression area  25 . While a conventional demethanizer will be described here in general terms, it is to be understood that this is not to be considered limiting to the present invention, which may be used with any existing conventional demethanizer, or as a new installation. Additionally, the process conditions associated with the demethanizer will also be described only generally, since they are not limiting, and any person skilled in the art will understand all of the equipment and conditions, and how they may be modified if desired. 
     A compressed inlet gas which may comprise methane, ethane, propane and heavier hydrocarbons, as well as smaller amounts of carbon dioxide, nitrogen and other gases, enters the plant  10  into the expansion area  15 , through an inlet  30 , where it is divided into two streams  35  and  40 . The inlet gas may be at a temperature of about 65° F. and at a pressure of about 400 to 1200 p.s.i.a. 
     The stream  35  then enters a heat exchanger  45 , where the gas is cooled through heat exchange, for example to a temperature of about −85° F. The stream  40 , which is split off from stream  35  before entering heat exchanger  45 , is directed through a series of heat exchangers  50 ,  55  and  60 . The heat exchangers  50 ,  55  and  60  provide reboiling for the fractionation column  20 , which is required to maintain the methane content of the ethane plus liquid recovered typically below 2.0 mole %. This will be discussed below. Correspondingly, the heat exchangers  50 ,  55  and  60  cool the incoming gas to a temperature of approximately −75° F. and a pressure somewhat lower than the inlet pressure, solely due to pressure losses in the various pipes, heat exchangers, etc. 
     The streams  35  and  40  are then recombined downstream from the heat exchanger  45 , into a combined stream  62 . The stream  62  is then directed into a low temperature separator  65 , where the liquid and gas phases are separated. The pressure in the separator  65  is approximately 550-1150 p.s.i.a. at a temperature of −85° F. The vapour phase leaves the low temperature separator  65  as an overhead stream  70 , while the liquid phase leaves separator  65  through a bottom stream  75 . The low temperature separator  65  is conventional, and is well known to those skilled in the art. 
     The vapour stream  70  is then depressured and further cooled in an expander  80 , and subsequently directed as a stream  85  into the fractionation column  20  at point  90 . The bottom stream  75  is typically reduced in pressure by passing through a valve (not shown), which causes flash evaporation or expansion to occur. 
     The fractionation column  20  has a main body  95 , and an upper enlarged portion  100 , and the inlet point  90 , for the stream  85 , is provided towards the bottom of the enlarged portion  100 . The upper enlarged portion  100  may be described as generally conical or “belled” in shape. The design of the fractionation column  20  is conventional, including the upper enlarged portion  100 , and is known to those skilled in the art. In the conventional design, the upper belled portion  100  of the fractionation column  20  is empty, and is designed to have a larger diameter than the main body  95 . This area is known as a disengagement zone, and is provided to allow the hydrocarbon mixture entering the fractionation column  20  as the stream  85  (which is largely a vapour) an area where liquids entrained in the vapour phase may disengage from the vapour phase. As previously stated, this vapour phase may contain from about 5% to about 50% liquid content. The upper enlarged portion  100  is often between about 10 to 15 feet in height, and about 10 to 18 feet in diameter, depending on the volume of gas being processed. 
     The liquid portion exiting the low temperature separator  65  through the stream  75  is normally directed to a lower point  105  in the fractionation column  20 , but above the location where reboiling streams  107 ,  108  and  109  leave the fractionation column  20  and are connected with heat exchangers  60 ,  55  and  50 , respectively. This is typically near the upper portion of the main body  95  of the fractionation column  20 , although this may vary from system to system. This reboiling system is required to maintain the methane content of the ethane plus stream leaving the fractionation column below 2.0 mole %. As will be appreciated, portions of the liquid phase falling through the fractionation column  20  are redirected out of the column  20 , through the reboiling streams  107 ,  108  and  109 , and are then passed through the heat exchangers  60 ,  55  and  50 , respectively. This aids in the cooling of the inlet gas passing through these heat exchangers. The stream  40 , now a cooled stream  40   b,  is recombined with stream  35  and are directed to the low temperature separator  65 . 
     Within the main body  95  of the fractionation column  20 , there is normally located a “packing”, which may be in the form of a series of “trays” or contact plates (typical to those trays seen in FIGS. 6 and 7 as will be later discussed). The contact plates are designed to increase the contact between the liquid and the vapour phases in the fractionation column, which in turn increases the efficiency of the ethane and higher hydrocarbon recovery. In a typical installation, there would be between 15 and 30 trays at a spacing of 2 feet between adjacent trays. As noted above, any suitable packing, such as Raschig rings can be used instead of trays. 
     In the fractionation column  20 , the vapour phase (mainly the methane content of the inlet gas) leaves the fractionation column  20  as an outlet stream  115 , where it is directed through the heat exchanger  45 , to aid in cooling to the inlet gas stream  35 . The stream  115  is then recompressed, for example in a brake compressor  120 , and is then further compressed to pipeline pressure utilizing a large compressor  125  and the compression section is generally indicated as  25 . The stream is then passed through an aftercooler  130  to lower the residue gas temperature to a level suitable for reentry into a gas transmission line; as is known, for some applications, it may be possible to omit the aftercooler  130 . 
     The more valuable heavier, liquid phase of the inlet gas (for example ethane and heavier) exits through the lower end of the column  20  as a stream  135 , and is collected and stored appropriately. 
     All of the features thus described are conventional, and will be readily known and understood by persons skilled in the art. 
     Referring to FIG. 2, in the present invention, the plant  10  is again shown schematically and modified by (1) the addition of a reflux section  150 , which allows the reflux of the vapours exiting the upper belled portion  100  of the fractionation column, and (2) the addition of trays or contact plates in the upper belled portion  100  of the fractionation column. These additional features provide increased efficiency of recovery of the heavier hydrocarbon components (ie. ethane and heavier). 
     For simplicity and brevity parts which are common between the apparatus of FIG.  1  and the apparatus of FIG. 2 are given the same reference numerals and description of these common components is not repeated. 
     In the present invention, the inlet gas stream  30  is initially split into three streams  35 ,  40  and  160  after entering the apparatus. The additional stream  160  is directed generally to the reflux section  150 , and more specifically, through an additional heat exchanger  165 , in which the stream  160  is cooled to a temperature of about −135° F. The heat exchanger  165  is conventional, and may be similar to the heat exchanger  45 . Again, some form of throttle or expansion valve (not shown) would be provided, so as to reduce the pressure and cause flash expansion of the hydrocarbon stream. 
     After being cooled in the heat exchanger  165 , the gas is expanded by an expansion valve (not shown) and sent directly to the fractionation column  20  as a stream  170 , where it enters the fractionation column  20  at  172 , near the top of the upper belled portion  100 . 
     In the upper belled portion  100  of the fractionation column  20 , there is located additional packing, which may be in the form of a series of additional trays or contact plates  226 - 228  (FIG.  6 ). The additional trays  226 - 228  may be similar to the trays located in the main body portion  95  of the fractionation column  20 . These additional trays  226 - 228  are installed so as to leave a small disengagement zone between the upper most tray  228  and the upper limit of the fractionation column  20 . It is at this location where the stream  170  enters the upper belled portion  100 . The disengagement zone in the present invention may be as little as about two feet in height, in contrast to the prior art devices, which required a much larger disengagement zone, usually on the order of between about 10 to 15 feet in height. It has been surprisingly found that only a few additional trays  226 - 228  are required in the upper portion  100  of the column  20 , to achieve a substantial increase in the recovery of ethane and heavier components from the column. It is believed that this is due to an unexpectedly and unusual high tray efficiency, which should be above 60, and is expected to exceed 80. In practice, the use of as few as three additional trays  226 - 228  has been sufficient to significantly increase the recovery, as will be discussed below. 
     The residue gas leaving the fractionation column  20  through stream  115  is then redirected through the heat exchanger  165 , to cool the inlet gas passing through this heat exchanger. The residue gas is then split into two streams  180  and  185 . The majority of the residue gas, of the order of 70 to 90%, is sent via stream  180  to the heat exchanger  45 , also to aid in the cooling of the inlet gas stream entering this heat exchanger. The remainder of the residue gas exits the heat exchanger  165  as stream  185 . The stream  180 , after exiting the heat exchanger  45  is recombined with the stream  185 , to form stream  190 . 
     The combined stream  190  is then recompressed and cooled, as previously described, prior to exiting the process. 
     It was previously believed that it was unacceptable to include the additional contact plates in the upper belled portion  100 , since it was believed that there would be insufficient room for the incoming liquid and vapour phases to separate from each other. It was believed that this would result in substantial liquid carryover and loss of hydrocarbon liquid through the upper stream  115  leaving the fractionation column  20 . This would clearly be undesirable. Additionally, given that the upper enlarged portion  100  of the column  20  in existing facilities is of limited size (usually between 10 to 15 feet in height), it was believed that there was insufficient room to add an acceptable number of trays to obtain an increase in ethane plus recovery which was commercially acceptable. As previously mentioned, it was believed that a minimum of eight additional trays would be required to ensure sufficient mass transfer contact between the rising vapours and the falling liquids. This was due largely to the assumption that several theoretical stages of contact were required and that tray efficiencies would be quite low (e.g. 40-60%). 
     Therefore, if it was desired to increase the efficiency of the fractionation column  20 , and in particular when it was desired to add a reflux section, such as section  150 , to allow reflux of the residue gas, a separate contactor column would be constructed on the ground. In contrast, the present invention provides additional trays in the fractionation column. Therefore, such an installation would require a separate vessel between about 24 and 60 feet in height and a diameter of about 16 to 17 feet. Given the size of the separate vessel, it would not be feasible to place it on top of the fractionation column. 
     The cost of a second vessel is significant, and therefore, may often not be commercially feasible. Additionally, the extra equipment associated with the separate column, for example, a cryogenic pump to deliver the liquids from the second separator to the top of the demethanizer column, would significantly increase capital and operating costs. 
     In accordance with the present invention, it has surprisingly been found that the use of as few as three of these contact trays in the upper belled portion  100  of the fractionation column  20  can achieve almost the same level of efficiency as the use of eight to twenty trays, which requires a separate vessel, but at a much reduced cost. Surprisingly, a high tray efficiency of 82% has been obtained in a plant performance test, and this requires a much smaller number of trays to obtain the desired separation. Conventional design practices allow a tray efficiency of 45-60%. The fact that so few trays are required makes the size of the upper belled portion sufficiently small so that it can be supported by the main body  95  of the fractionation column  20 . 
     In an unoptimized actual plant test, it was found that the present invention, through modification of an existing facility, increased ethane recovery by about 8.4%, from about 77% to about 85.4%, and propane recovery was increased from about 98% to about 98.9%. These increased efficiencies are each significant, particularly given the relatively minimal cost and time required to modify an existing system to achieve the benefit. In an optimized trial basis it is expected that the ethane recovery will increase by 11%. In simulated comparisons where a grass roots or new facility is proposed, ethane recovery was increased by about 11% as compared to a conventional design as depicted in FIG.  1 . 
     The present invention enables the volume of the residue gas stream  115  to be reduced, because of the increased volume of liquid ethane and other substances removed from stream  85 . In an actual plant test, using the same residue gas compressor horsepower, this enabled inlet gas rates to be increased from 489 million standard cubic feet per day to 501 million standard cubic feet per day. 
     This aspect of the invention will be better understood by reference to FIG.  3 . This shows a theoretical schematic of the apparatus of FIG.  1 . This has been analyzed using conventional software, Hysim, a software program licensed by Hyprotech Ltd. from Calgary. The overall layout is similar to FIG. 1, with the exception of the items outlined below. Otherwise, like components are given the same reference numeral and their description is not repeated. 
     In FIG. 3, the stream  40  passes through a heat exchanger  140  and exchanges heat with ethane plus recovered from the column. It then flows through the heat exchanger  50 . 
     For the heat exchanger  50 , a stream  141  is taken from the bottom of the column  20 , passed through the heat exchanger  50  and then back to the fractionation column  20 . The liquid from the fractionation column is connected to a stream  146  for liquid, which is delivered to a pump  148 . The vessel  142  is shown for simulation purposes only. In practice the vessel  142  is part of the bottom of the fractionation column  20 . The pump  148  discharges the recovered ethane and other products through line  149 , which passes through the heat exchanger  140 . 
     The heat exchanger  45  of FIGS. 1 and 2 is now configured as heat exchangers  167   a,    167   b.  As shown, the stream  115  splits and then recombines prior to entering compressor  125  for flow through the heat exchanger  167   a,    167   b.  The combined stream  115  flows to compressors  120 ,  125 . Further, similar to the FIG. 2 embodiment, the line  40  and the line  35   b  combine, before entering the separation vessel  65 . 
     The separation vessel  65  has the outlet stream  75  for liquid, connected to the fractionation column  20  as before. The vapour line  70  is provided with a main branch stream  71  connected through the expander  80  and then stream  73  to the top of the fractionation column  20 . A bypass stream  72  is provided although as detailed below, it will often not be used. This enables the expander  80  to be bypassed, so that it can be serviced without shutting the whole processing plant down. 
     FIG. 4 shows a theoretical model of the plant or apparatus of FIG.  2 . Again, elements or components already identified and described are given the same reference numeral and description of them is not repeated, for simplicity and brevity. 
     As before, the principal difference between FIGS. 3 and 4, as for FIGS. 1 and 2, is the inclusion of a separate line  170  discharging into the top end of the column  20 . Here, the expansion valve  162  for the line  170  is shown. Additionally, the heat exchanger  165  is shown as two separate heat exchange elements  165   a  and  165   b.  As for FIG. 3, two separate heat exchange elements  167   a  and  167   b  are shown, approximately corresponding to the heat exchanger  45 . 
     For these two theoretical models, in FIGS. 3 and 4, theoretical performance results have been obtained, for temperature, pressure, molar flow rate and vapour fraction. These are set out in the following Tables 1 and 2. 
     
       
         
               
               
               
               
               
             
               
               
               
               
               
             
           
               
                 TABLE 1 
               
               
                   
               
               
                   
                 Temperature 
                 Pressure 
                 Molar Flow 
                 Vapour 
               
               
                 Location 
                 ° F. 
                 p.s.i.a 
                 lb moles/hr 
                 Fraction 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                 INLET STREAM 
                 65.12 
                 766 
                 53650.57 
                 1 
               
               
                 30 
               
               
                 STREAM 40 
                 65.12 
                 766 
                 18666.8 
                 1 
               
               
                 exit from heat 
                 64.87 
                 762 
                 18666.8 
                 1 
               
               
                 exchanger 166 
               
               
                 exit from heat 
                 53.5 
                 758 
                 18666.8 
                 1 
               
               
                 exchanger 140 
               
               
                 exit from heat 
                 −73.04 
                 746 
                 18666.8 
                 .93 
               
               
                 exchanger 60 
               
               
                 STREAM 35a 
                 64.81 
                 761 
                 34983.77 
                 1 
               
               
                 stream 35b 
                 −88.18 
                 754 
                 34983.77 
                 .74 
               
               
                 stream 62 
                 −84.5 
                 746 
                 53650.57 
                 .82 
               
               
                 stream 70 
                 −84.5 
                 746 
                 44160.64 
                 1 
               
               
                 stream 75 
                 −84.5 
                 746 
                 9489.93 
                 0 
               
               
                 (upstream valve 
               
               
                 76) 
               
               
                 stream 75 
                 −132.68 
                 316.22 
                 9489.93 
                 .39 
               
               
                 (downstream 
               
               
                 valve 76) 
               
               
                 stream 71 
                 −85.47 
                 736 
                 44160.64 
                 1 
               
               
                 (upstream 
               
               
                 expander 80) 
               
               
                 stream 73 
                 −144.59 
                 290 
                 44160.64 
                 .85 
               
               
                 (upstream valve 
               
               
                 74) 
               
               
                 stream 73 
                 −140.57 
                 316.22 
                 44160.64 
                 .85 
               
               
                 (downstream 
               
               
                 valve 74) 
               
               
                 STREAM 115 
                 −145.05 
                 286 
                 49907.72 
                 1 
               
               
                 heat exchanger 
                 53.82 
                 278 
                 8889.79 
                 1 
               
               
                 167a exit 
               
               
                 (line 115) 
               
               
                 heat exchanger 
                 35.44 
                 276 
                 41017.93 
                 1 
               
               
                 167b exit 
               
               
                 (line 115) 
               
               
                 compressor 120 - 
                 38.7 
                 276 
                 49907.72 
                 1 
               
               
                 entry (stream 190) 
               
               
                 compressor 120 - 
                 70.46 
                 322.1 
                 49907.72 
                 1 
               
               
                 exit (stream 190) 
               
               
                 compressor 125 - 
                 69.47 
                 306 
                 49907.72 
                 1 
               
               
                 entry 
               
               
                 compressor 125 - 
                 242.16 
                 823 
                 49907.72 
                 1 
               
               
                 exit 
               
               
                 STREAM 141 
                 22.02 
                 290 
                 4876.1 
                 0 
               
               
                 (upstream heat 
               
               
                 exchanger 50) 
               
               
                 stream 141 
                 30.66 
                 290 
                 4876.1 
                 0.23 
               
               
                 (downstream heat 
               
               
                 exchanger 50) 
               
               
                 stream 144 
                 30.66 
                 290 
                 1133.71 
                 1 
               
               
                 stream 146 
                 30.66 
                 290 
                 3742.39 
                 0 
               
               
                 stream 149 - entry 
                 32.34 
                 384.17 
                 3742.39 
                 0 
               
               
                 to exchanger 140 
               
               
                 stream 149 - exit 
                 53.6 
                 379.17 
                 3742.39 
                 .01 
               
               
                 to exchanger 140 
               
               
                   
               
             
          
         
       
     
     
       
         
               
               
               
               
               
             
               
               
               
               
               
             
           
               
                 TABLE 2 
               
               
                   
               
               
                   
                 Temperature 
                 Pressure 
                 Molar Flow 
                 Vapour 
               
               
                 Location 
                 ° F. 
                 p.s.i.a 
                 lb moles/hr 
                 Fraction 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                 INLET STREAM 
                 65.06 
                 771 
                 55044 
                 1 
               
               
                 30 
               
               
                 STREAM 40 
                 65.06 
                 771 
                 16741.16 
                 1 
               
               
                 exit from heat 
                 64.93 
                 769 
                 16741.16 
               
               
                 exchanger 166 
               
               
                 exit from heat 
                 52.38 
                 765 
                 16741.16 
                 1 
               
               
                 exchanger 140 
               
               
                 exit from heat 
                 −80.06 
                 753 
                 16741.16 
                 .87 
               
               
                 exchanger 60 
               
               
                 STREAM 35a 
                 64.87 
                 768 
                 31763.96 
                 1 
               
               
                 stream 35b 
                 −77.08 
                 760.2 
                 31763.96 
                 .9 
               
               
                 stream 62 
                 −78.54 
                 753 
                 48505.13 
                 0.89 
               
               
                 stream 70 
                 −78.54 
                 753 
                 43066.88 
                 1 
               
               
                 stream 75 
                 −78.54 
                 753 
                 5438.24 
                 0 
               
               
                 (upstream valve 
               
               
                 76) 
               
               
                 stream 75 
                 −128.6 
                 301.23 
                 5438.24 
                 .39 
               
               
                 (downstream 
               
               
                 valve 76) 
               
               
                 stream 71 
                 −79.5 
                 743 
                 43066.88 
                 1 
               
               
                 (upstream 
               
               
                 expander 80) 
               
               
                 stream 73 
                 −138.16 
                 301.23 
                 43066.88 
                 .87 
               
               
                 (upstream valve 
               
               
                 74) 
               
               
                 stream 73 
                 −138.16 
                 301.23 
                 43066.88 
                 .87 
               
               
                 (downstream 
               
               
                 valve 74) 
               
               
                 STREAM 160 
                 65.06 
                 771 
                 6538.87 
                 1 
               
               
                 stream 170 
                 −147.74 
                 757 
                 6538.87 
                 1 
               
               
                 (upstream valve 
               
               
                 162) 
               
               
                 stream 170 
                 −156.36 
                 300.7 
                 6538.87 
                 0.06 
               
               
                 (downstream 
               
               
                 valve 162) 
               
               
                 STREAM 115 
                 150.14 
                 294 
                 50877.09 
                 1 
               
               
                 stream 185a 
                 −111.76 
                 289 
                 4941.21 
                 1 
               
               
                 stream 185b 
                 60.79 
                 283.5 
                 4941.21 
                 1 
               
               
                 stream 180 
                 −111.76 
                 289 
                 45929.88 
                 1 
               
               
                 heat exchanger 
                 46.32 
                 281 
                 8995.2 
                 1 
               
               
                 167a exit 
               
               
                 heat exchanger 
                 40.72 
                 279 
                 36934.68 
                 1 
               
               
                 167b exit 
               
               
                 compressor 120 - 
                 43.61 
                 279 
                 50871.09 
                 1 
               
               
                 entry 
               
               
                 compressor 120 - 
                 73.77 
                 324.4 
                 50871.09 
                 1 
               
               
                 exit 
               
               
                 compressor 125 - 
                 73.08 
                 313 
                 50871.09 
                 1 
               
               
                 entry 
               
               
                 compressor 125 - 
                 243.32 
                 823 
                 50871.09 
                 1 
               
               
                 exit 
               
               
                 stream 141 
                 17.03 
                 298 
                 5313.04 
                 0 
               
               
                 (upstream heat 
               
               
                 exchanger 50) 
               
               
                 stream 141 
                 27.54 
                 298 
                 5313.04 
                 .21 
               
               
                 (downstream heat 
               
               
                 exchanger 50) 
               
               
                 stream 144 
                 27.54 
                 298 
                 1139.81 
                 1 
               
               
                 stream 146 
                 27.54 
                 298 
                 4173.23 
                 0 
               
               
                 stream 149 
                 31.56 
                 524.63 
                 4173.23 
                 0 
               
               
                 stream 149 
                 52 
                 519.63 
                 4173.23 
                 0 
               
               
                 (downstream heat 
               
               
                 exchanger 140) 
               
               
                   
               
             
          
         
       
     
     A review of the values given in these two tables will show many close similarities, which might be expected, in view of the very similar flow rates. The significant differences are in the conditions and flow rates for the streams entering the fractionation column. Thus, in Table 1, the stream  75  enters at a temperature of −132.68° F. and a flow rate of 9489.93 lb moles/hr and stream  73  at a temperature of −144.59° F. and a flow rate of 44,160.64 lb moles/hr. 
     When modified to provide the third stream, as indicated in Table 2, the flow rate for stream  73 , is decreased slightly to 43,066.88 at a temperature of −138.16° F. The bottom stream  75  is reduced to a greater extent to 5438.24 lb moles/hr and a slightly higher temperature of −128.6° F. However, there is now the additional top stream  170  which is introduced at a temperature of −156.36° F. and a flow rate of 6538.87. More significantly, while the vapour fractions for stream  73 ,  75 , differ little between the two examples, stream  170  is introduced almost entirely in the liquid phase, with a vapour fraction of just 0.06. The effect of this is to provide, at the top three trays, an upward flow of methane, originating principally from stream  73 , which meets a downward flow of liquid from stream  170 , which is introduced in the liquid phase at a significantly lower temperature. The effect of this is to create a downward flow from the top three trays, which would ensure that a significant portion of ethane and other heavier hydrocarbons are absorbed from stream  73  and carried down through the column, and are not carried upwards with the methane gas. 
     Reference will now be made to FIG. 5 a,  which shows schematically an apparatus or plant in accordance with a second aspect of the present invention. This has been designed as a complete new plant or facility, rather than as a modification to an existing facility. However, again for simplicity and brevity, parts, common with earlier Figures, are given the same reference numerals, and description of these common components is not repeated. 
     Here, the heat exchangers  50 ,  55 ,  60  are represented by the two sides of the heat exchangers, as heat exchange elements  50   a,    50   b  and  55   a,    55   b,    60   a,    60   b  for the respective reboiling streams  109 ,  108 ,  107 . To separately identify the inward end and outward end of the stream  40 , this is designated as  40   a  for the stream flowing to the heat exchangers, and stream  40   b  leaving the heat exchangers. 
     The heat exchanger configuration in this embodiment is somewhat different. Here, broadly corresponding to the heat exchanger  45  of FIG. 1, there are two heat exchange elements  200 ,  202  through which the stream  35  flows. The stream for the returned residue gas is again indicated as  115 , and after passage through a heat exchanger  206 , the stream is indicated as  185 . Due to the different configuration here, the designation  185  is used for the residue gas stream through to the compressor  120  and compressor  125 . The additional heat exchanger  206  is provided, for heat exchange with the stream  160  passing to the top of the fractionation column  20 . 
     The inlet stream  30  passes through a valve  208  to a stream  210 , connected to the heat exchange elements  200 ,  202 . Between the heat exchange elements  200 ,  202 , the additional stream  160  is branched off, and passes through the heat exchanger  206  to the top of the column  20 . 
     To enable the various streams to be identified at different points, primarily for identifying stream conditions as detailed below, various suffixes are used. Thus, the stream  160  is identified as  160   a  and  160   b  before and after the heat exchanger  206 . This stream  160  also includes a throttle expansion valve  162 , to cause flash expansion. The stream  185  is variously labelled as  185   a, b, c,  etc to identify the portions indicated on FIG.  4 . For the streams from the separator  65 , the liquid stream is identified as  75   b  after expansion through a valve  76 , and the vapour stream is identified as  70  and  73  before and after expansion in the expander  80 . 
     The plant of FIG. 5 a  is intended to handle 1.0 billion cubic feet of gas per day (BCFD), which enters through stream  30  at 613 p.s.i.a and 68° F. from a main gas transmission line. As before, the purpose of the apparatus is to receive gas from a transmission line, to efficiently cool and depressure the gas, to extract the valuable heavier components (“ethane plus”, i.e. ethane and heavier hydrocarbons), and then to recompress the gas back into the transmission line. The inlet gas primarily comprises methane although it contains other species including carbon dioxide (0.2-1.5%), nitrogen (0.5-1.5%), ethane (3-8%), propane (0.5-2%), and heavier hydrocarbons (0.5-5% ). The objective of such apparatus is to provide an extraction facility to remove ethane, propane, and heavier hydrocarbons in substantive quantities for subsequent resale. The FIG. 5 a  apparatus is intended to extract over 60% of the ethane and over 99% of the heavier propane plus components. 
     A portion of the inlet gas is split off and sent through the side heat exchangers  50 ,  55 ,  60  on the distillation column  220  to provide reboiling. The remainder of the inlet gas is sent through exchanger element  200  where it is chilled to −36° F. using residue gas in line  185 . Upon exiting exchanger  200 , the inlet gas is once again split and a portion of the gas is sent through heat exchanger  206 , where 84% of the stream is liquified. This stream then passes through the valve  162  and enters the top of the distillation or fractionation column  220 . The remainder or bulk of the inlet gas passes through a second inlet gas exchanger  202  and is chilled to −72° F. This gas is then mixed with stream  40   b  and the combined flow is sent to separator  65 . Vapours leaving separator  65  are expanded adiabatically in expander  80  from 595 p.s.i.a to 304 p.s.i.a causing the gas temperature to drop from −75° F. to −126° F. The expander  80  generates power from this expansion and the power is utilized in driving the brake compressor  120 . Liquids from the separator  65  are sent to the distillation column through the valve  76 . 
     The column, here indicated as  220 , comprises two sections or portions: a smaller bottom section or portion  222  primarily dedicated to providing reboiling, and a larger top section or portion  224  which accomplishes a majority of the ethane recovery out of the inlet gas. The top section  224  is quite large for a 1.0 BCFD feed gas rate and here is 18 foot in diameter, which currently is close to the limit of what can be constructed within a reasonable cost. The bottom section  222  of the column is much smaller and has a diameter of 9 feet. The bottom section  222  would include 10-15 theoretical trays, equivalent to 18-24 actual trays; the top section  224  includes three theoretical trays, equivalent to four actual trays. 
     The example process produces an incremental ethane recovery of 12.4%, i.e. 67.4% as compared to 55% in a conventional plant; and incremental propane and recovery of between 2-3%, i.e. 98.4% as compared to 96.1%. Higher ethane recoveries would be attainable by increasing the size of the heat exchangers. This would cool the inlet gas to a greater degree or extent than if less heat exchange is provided, effectively increasing ethane flow down the column  220  and reducing the amount of ethane carried over with residue gas. Increasing recompression of the residue gas could draw down the temperature and pressure in the column, also leading to increased recoveries. 
     It can also be noted that a plant following the Ortloff design in U.S. Pat. No. 4,278,457 cannot be accomplished in a single column, since conventional engineering design would require the top section to be at least a 30 ft height with a 17 ft diameter. This could not practically be supported on top of a 9 ft diameter bottom section. 
     This particular configuration would be able to produce a range of ethane recoveries from 50% up to about 95% by increasing the horsepower of compressor  125  to reduce the pressure of the columns  220 . Above 95% recovery the pressure in the distillation column would drop to much lower values and flooding may occur in the top section  224  of the column, thereby limiting the upper end of the recovery efficiency. 
     Residue gas leaving the top of the distillation column  220  is sent through exchanger  206 , cooling the gas flow, and is subsequently sent through heat exchange elements  202  and  200 , providing cooling for the inlet gas. This warms the residue or overhead gas up to 56° F. (at  185   b ) which is only slightly colder than the inlet gas temperature. The warm gas is then compressed at  120  to a pressure of 358 p.s.i.a using the power developed by the expander  80 . The compressed gas is further compressed at  125  up to a high enough pressure to put it back into the main gas transmission line. Gas leaving the recompressor  125  is cooled by water or air in heat exchanger  130  and is further cooled by a heat exchanger  212  with upstream residue gas. In general, the residue gas is required to leave the plant at the same temperature at which is entered. 
     The following Table 3 sets out process parameters that would be obtained in the plant of FIG. 5 a,  again based on a theoretical simulation of the plant. 
     
       
         
               
               
               
               
               
             
               
               
               
               
               
             
           
               
                 TABLE 3 
               
               
                   
               
               
                   
                 Temperature 
                 Pressure 
                 Molar Flow 
                 Vapour 
               
               
                 Location 
                 ° F. 
                 p.s.i.a 
                 lb moles/hr 
                 Fraction 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                 INLET STREAM 
                 68 
                 598.28 
                 116546.84 
                 1 
               
               
                 30 
               
               
                 STREAM 40a 
                 67.87 
                 596.25 
                 24596.25 
                 1 
               
               
                 stream 40b 
                 −86.82 
                 590.74 
                 24596.25 
                 0.95 
               
               
                 MAIN STREAM 
                 67.87 
                 596.25 
                 91950.59 
                 1 
               
               
                 210 
               
               
                 stream 35 
                 −93.35 
                 590.74 
                 80970.13 
                 0.93 
               
               
                 STREAM 160a 
                 −36.0000 
                 593.49 
                 10980.47 
                 1 
               
               
                 stream 160b 
                 −137.19 
                 578.49 
                 10980.47 
                 0 
               
               
                 stream 160c 
                 −152.14 
                 329.83 
                 10980.47 
                 0.12 
               
               
                 stream 35 
                 −93.35 
                 590.74 
                 80970.13 
                 0.93 
               
               
                 stream 62 
                 −91.93 
                 590.74 
                 105566.38 
                 0.94 
               
               
                 stream 75b 
                 −122.27 
                 331.83 
                 6780.64 
                 0.23 
               
               
                 stream 70 
                 −92.28 
                 587.69 
                 98785.74 
                 1 
               
               
                 stream 73 
                 −132.76 
                 329.83 
                 98785.74 
                 0.94 
               
               
                 STREAM 115 
                 −140.57 
                 327.83 
                 111113.67 
                 1 
               
               
                 stream 185a 
                 −112.94 
                 322.83 
                 111113.67 
                 1 
               
               
                 stream 185b 
                 56.2 
                 320.02 
                 111113.67 
                 1 
               
               
                 stream 185c 
                 99.94 
                 313.49 
                 111113.67 
                 1 
               
               
                 stream 185d 
                 122.89 
                 358.67 
                 111113.67 
                 1 
               
               
                 stream 185e 
                 223.99 
                 626.71 
                 111113.67 
                 1 
               
               
                 stream 185f 
                 109.94 
                 621.63 
                 111113.67 
                 1 
               
               
                 stream 185g 
                 69.08 
                 613.51 
                 111113.67 
                 1 
               
               
                   
               
             
          
         
       
     
     Reference will now be made to FIG. 5 b,  which shows a further variant of the apparatus and method of the present invention. This again is for 1000 MMSCFD plant, and is similar in many respects to the configuration shown in FIG. 5 a.  For this reason, and again as before, description of common components is not repeated, and these common components are given the same reference numeral. 
     The difference in FIG. 5 b  is, in effect, that the two streams  73 ,  160   b  of FIG. 5 a  are now combined into a single stream ( 160   c ), before entering the top of the fractionation column  220 . Thus, in FIG. 5 b,  the stream  160   b  joins with stream  73  in a static mixer  214 . The combined stream downstream from the static mixer  214  is indicated at  160   c.  The static mixer  214  is a motionless device, i.e. without any moving parts, and of known construction. It causes swirling in the fluid flowed downstream, which increases the turbulence in the piping and consequently increases the mass transfer rate. If the static mixer has a sufficient length, it is possible to approach one theoretical stage of contacting, equivalent to one theoretical stage in the column  220 . It is expected that, for situations where it is not practical or possible to put trays in the top of the column  220  (e.g. in some retrofit applications where demethanizer columns are not belled or enlarged at the top at all) this configuration should provide a significant improvement at even less cost than retrofit with trays. 
     It can be noted that, as detailed above, the three tray modification gave an ethane recovery of 67.4%, whereas, as detailed in Table 4 below, static mixer  214  gives a recovery of 65.2%. Again, these are theoretical simulated results only and are predicated on the assumption that a full equilibrium stage can be achieved in the static mixer  214 . It is expected that actual test results would be close, although likely not quite so good. 
     A variant of the proposal of FIG. 5 b  is shown in FIG.  5   c.  Here, there is provided a separator tank or vessel  215 , similar to the separator  65 , to separate the gas and liquids from the static mixer  214 . Such a separator would have an outlet for liquid connected to the top of the column  220 , as shown for stream  160   c,  and an outlet for vapour or gas connected to the stream  115 , and indicated at  160   d.  Thus, FIG. 5 c  shows a typical application of a static mixer modification in which the lighter or vapour portion of the combined streams  73  and  160   b  is separated in the separator  215  and then flows around the demethanizer column  220  through the branch stream  160   d;  as this vapour flow can be a significnat part of the total flow, this can significantly reduce the flow rates at the top of the column  220 . Theoretical simulations of the improvement offered by the static mixer  214  indicated that an 8% incremental ethane recovery could be achieved by implementing static mixing in the manner shown. Again, this is under the assumption that one theoretical stage of mixing could be achieved in the static mixer. With proper design and contact time in the mixer it should be possible to closely approach the predicted performance and effect of the static mixer  214 . 
     
       
         
               
               
               
               
               
             
               
               
               
               
               
             
           
               
                 TABLE 4 
               
               
                   
               
               
                   
                 Temperature 
                 Pressure 
                 Molar Flow 
                 Vapour 
               
               
                 Location 
                 ° F. 
                 p.s.i.a 
                 lb moles/hr 
                 Fraction 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                 INLET STREAM 
                 68 
                 598.28 
                 116546.84 
                 1 
               
               
                 30 
               
               
                 STREAM 40a 
                 67.87 
                 596.25 
                 24596.25 
                 1 
               
               
                 stream 40b 
                 −90.97 
                 590.74 
                 24596.25 
                 0.94 
               
               
                 MAIN STREAM 
                 67.87 
                 596.25 
                 91950.59 
                 1 
               
               
                 210 
               
               
                 stream 35 
                 −92.57 
                 590.74 
                 80970.13 
                 0.93 
               
               
                 STREAM 160a 
                 −36 
                 593.49 
                 10980.47 
                 1 
               
               
                 stream 160b 
                 −134.41 
                 578.49 
                 10980.47 
                 0 
               
               
                 stream 160c 
                 −138.59 
                 330.06 
                 109643.09 
                 0.89 
               
               
                 stream 62 
                 −92.2 
                 590.74 
                 105566.38 
                 0.93 
               
               
                 stream 75b 
                 −122.56 
                 332.06 
                 6903.75 
                 0.23 
               
               
                 stream 70 
                 −92.56 
                 587.69 
                 98662.63 
                 1 
               
               
                 stream 73 
                 −132.93 
                 330.06 
                 98662.63 
                 1 
               
               
                 STREAM 115 
                 −138.74 
                 328.06 
                 111205.24 
                 1 
               
               
                 stream 185a 
                 −111.5 
                 322.86 
                 111205.24 
                 1 
               
               
                 stream 185b 
                 55.97 
                 320.25 
                 111205.24 
                 1 
               
               
                 stream 185c 
                 99.82 
                 313.72 
                 111205.24 
                 1 
               
               
                 stream 185d 
                 122.67 
                 358.7 
                 111205.24 
                 1 
               
               
                 stream 185e 
                 223.53 
                 626.71 
                 111205.24 
                 1 
               
               
                 stream 185f 
                 109.94 
                 621.63 
                 111205.24 
                 1 
               
               
                 stream 185g 
                 67.98 
                 613.51 
                 111205.24 
                 1 
               
               
                   
               
             
          
         
       
     
     Reference will now be made to FIGS. 6 and 7, which show details of the trays as shown in the top of the column  220 . As shown, the column  220  has a relatively narrow bottom or body section  222  and a top or upper section  224  of much greater diameter. The bottom section  222 , as indicated includes 10-15 theoretical trays, equivalent to 18-24 actual trays. 
     In the top section  224  of the column, there are three additional trays  13 ,  14  and  15 , here identified as  226 ,  227  and  228 . 
     The lowermost of the three trays, tray  226  has a central horizontal portion  230 , with upwardly extending lips  232  along opposite sides, which lips  232  provide a weir to maintain a desired fluid level on top of the central portion  230 . In known manner, the central portion would be provided with an array of bubble caps or valves. Side walls  234  include partially inclined sections  235 , which are connected to a lower floor  236 . Again, in known manner, the floor  236  can be provided with slots, permitting fluid to flow down through the slots to the tray below. 
     The tray  227  has a generally circular horizontal portion  240  and is provided with a slot  241  extending diametrically. Side walls  244  define the slot and include lips  242 , again providing a weir function. The side walls  244  include lowermost inclined sections  245 , to which a floor  246  is connected. Again, the floor  246  would be provided with slots for downflow of liquid, while the horizontal portion  240  would be provided with bubble caps or valves, to permit upward flow of vapour. 
     The uppermost tray  228  corresponds in many ways to the tray  226 , and includes a central horizontal portion  250 , lips  252  and side walls  254 . The side walls  254  again include inclined sections  255 , but here these incline outwardly and away from one another; for the tray  226 , the inclined sections  235  incline inwardly, to follow a frusto-conical transition section  225  between the top and bottom sections  222 ,  224 . 
     It will be appreciated that detailed design of the trays, bubble caps, louvres, mounting arrangements etc., can be largely conventional and follow known design practice. Such details do not form part of the present invention. 
     As noted above, conventional fractionation columns have a disengagement section, which typically may be 10 to 15 feet in height, with a diameter of 10 to 18 feet. Here, the top of the column has an upper cap  260 . The distance between the lower edge of the upper cap  260  and the horizontal portion  250  of the tray  228  is only 2 foot 4 inches. It has surprisingly been found that this gives adequate disengagement or separation of the vapour and liquid phases. 
     There are also other benefits associated with the present invention. Significantly, the carbon dioxide content of the residue gas stream  115  is reduced, which aids in alleviating the CO 2  freezing problems commonly encountered at the upper sections of the column. Instead, a higher proportion of the carbon dioxide gas in the inlet gas is recovered in the ethane plus stream leaving the bottom of the column. The reason for the increased recovery of carbon dioxide is that it has a boiling point quite close to the boiling point of ethane. Therefore, as there is a higher recovery of ethane in the present invention, there is also a higher recovery of carbon dioxide. In trials, the carbon dioxide content of the residue gas leaving the upper portion of the column was reduced from 0.54 mole % to 0.52 mole %, at the same inlet gas composition. 
     Additionally, because the present invention requires the addition of the reflux section  150 , there is an increased gas processing capability, since the inlet gas is initially further divided, and there is less gas flowing through the inlet heat exchangers. 
     It will be understood by persons skilled in the art that, although the present invention has been described in relation to a particular system, the invention may be implemented in any number of ways, particularly in the manner in which the various hydrocarbon streams are processed and delivered to the fractionation column. All such modifications are contemplated by the present invention. 
     Additionally, while the present invention has been described largely as a retrofit to existing facilities, it will be appreciated that the invention may be utilized in new facilities, and that advantages will be realized in such use. In particular, an increased recovery of ethane plus may be substantially achieved over conventional reflux designs, by the elimination of the high capital cost of a second separator which has conventionally been required. 
     It will be appreciated that various changes and modifications may be made within the spirit of the invention, and all such changes are included within the scope of the attached claims.