Ethylene, propylene and other valable petrochemicals are produced by the thermal cracking of a variety of hydrocarbon feedstocks ranging from ethane to heavy vacuum gas oils. In the thermal cracking of these feedstocks, a wide variety of products are produced ranging from hydrogen to pyrolysis fuel oil. The effluent from the cracking step, commonly called charge gas or cracked gas, is made up of this full range of materials which must then be separated (fractionated) into various product and by-product streams followed by reaction (hydrogenation) of at least some of the unsaturated by-products.

The typical charge gas stream, in addition to the desired products of ethylene and propylene, contains C2 acetylenes, C3 acetylenes and dienes and C4 and heavier acetylenes, dienes and olefins as well as a significant quantity of hydrogen. In the majority of prior processes, the C2 acetylenes and C3 acetylenes and dienes and the C5 and heavier dienes, acetylenes and olefins are catalytically hydrogenated in fixed bed reactors using a series of commercially available catalysts. In a growing number of applications, the C4 acetylenes, dienes, and olefins are also catalytically hydrogenated in fixed bed reactors. These separate hydrogenation steps take place in one of two process sequences. In the first sequence, the charge gas is compressed to between 2.76 and 4.14 MPa (400 and 600 psia). It is then progressively chilled condensing the C2 and heavier components. Hydrogen is cryogenically recovered and methane is fractionated out of the stream. The remaining C2 and heavier stream enters a series of fractionation towers. The first tower produces an overhead stream containing the C2 acetylenes, olefins, and paraffins. This stream is sent to a fixed bed, vapor phase reactor where the C2 acetylene

is selectively hydrogenated using the hydrogen cryogenically separated earlier from the charge gas stream.

The second tower in this sequence produces an overhead stream containing the C3 acetylenes, dienes, olefins and paraffins. This stream is sent to a fixed bed, vapor or liquid phase reactor where the C3 acetylenes and dienes are selectively hydrogenated using the hydrogen cryogenically separated earlier from the charge gas stream.

The third tower in this first sequence produces an overhead stream containing the C4 acetylenes, dienes, olefins, and paraffins. This stream is then sent either to battery limits as a final product or to a fixed bed, liquid phase reactor where the dienes, acetylenes, and in some instances the olefins are hydrogenated using the hydrogen cryogenically recovered previously from the charge gas.

The bottoms of the third tower contains the C5 and heavier dienes, acetylenes, olefins and paraffins. This stream is sent to a series of two fixed bed, liquid phase reactors. In the first, the acetylenes and dienes are catalytically hydrogenated. The olefins are catalytically hydrogenated in the second reactor. Both reactors utilize the hydrogen cryogenically recovered previously from the charge gas. In some applications, the third tower produces an overhead stream containing both the C4 and C5 acetylenes, dienes, olefins, and paraffins. These are hydrogenated as discussed previously for the C4's alone, in a single fixed bed, liquid phase reactor. The C6 and heavier dienes, acetylenes, olefins and paraffins exit in the bottoms of the third tower and are hydrogenated as discussed previously in two fixed bed, liquid phase reactors.

In the second process sequence, the cracked gas is compressed to between 2.07 and 3.45 MPa (300 and 500 psia) and sent to a fractionation tower. The overhead of the tower is the C3 and lighter portion of the charge gas. It is sent to a series of fixed bed, vapor phase reactors where the C2 acetylene and a portion of the C3 acetylenes and dienes are hydrogenated using a small portion (typically less than 10%) of the hydrogen contained in the C3 and lighter stream. The unhydrogenated

portion of the C3 acetylenes and dienes as well as the C4 and heavier acetylenes, dienes, and olefins are hydrogenated in a fashion similar to that described above for the first process sequence. This still leaves over 90% of the hydrogen to be recovered cryogenically.

Also in such a system, it is necessary to fractionate out the C4 and heavier materials from the charge gas prior to the hydrogenation step.

Otherwise, the heat of the hydrogenation reaction would be excessive and there would be a high rate of hydrogenation catalyst fouling. Since such a fractionation occurs in a high hydrogen and methane environment, the energy requirements are high.

In most prior processes, the C2 and C3 acetylenes and C3 dienes are hydrogenated after the hydrogen separation/recovery step. The hydrogenation of the C4 and heavier acetylenes, dienes, and olefins always occurs after the hydrogen separation step and will consume up to 80% of the total available hydrogen. This hydrogenation also occurs in fixed bed catalytic reactors using catalysts chosen for the selectivity and degree of hydrogen saturation dictated by the particular process.

While widely practiced, both process sequences described above have a number of disadvantages. First, the cracked gas must be chilled and condensed in the presence of hydrogen. Due to the high partial pressure of the hydrogen, the mechanical refrigeration requirements to accomplish the condensation of the C2 and heavier material are high thereby increasing the energy consumption and capital investment in the process. Also, the hydrogen must be cryogenically separated to supply the hydrogen for the various downstream reactors which is both energy and capital intensive. Further, the hydrogenation steps occur in a series of fixed bed reactors requiring between 3 and 6 separate reactor systems thereby increasing the capital investment and complexity of the plant.

Summary of the Invention An object of the present invention is to hydrogenate in the liquid phase in a boiling point reactor the C2 to C5 and possibly heavier

acetylenes and dienes in a feed stream without hydrogenating the C2 and C3 olefins in the feed stream. Additionally, the C4, C. and some or all-of the heavier olefins may be hydrogenated still without hydrogenating the C2 and C3. olefins.

More specifically, an object of the present invention is to provide a system and method for hydrogenating the cracked gas in an olefin plant prior to the separation of hydrogen and methane from the cracked gas in a manner so as to hydrogenate the by-products, C2 acetylenes, C3 acetylenes and dienes and C4 and heavier acetylenes and dienes and, if desired, the C4 and heavier olefins, without significant hydrogenation of the ethylene and propylene. This involves the use of a combined reaction- fractionation step known as catalytic distillation hydrogenation upstream of the chilling and condensation of the C2 and heavier material to simultaneously carry out the reactions and separations in a manner so as to prevent or minimize the hydrogenation of the desired main products and to consume the hydrogen without the need for costly hydrogen separation.

The hydrogenation of the C4 and heavier acetylenes, dienes and olefins increases the hydrogen removal to between 70% and 100% and most typically 90% to 95%. This high removal of hydrogen reduces the hydrogen partial pressure thereby lowering the mechanical refrigeration requirements to chill and condense the C2 and heavier material thereby saving energy and capital investment. The cryogenic separation of the hydrogen from the cracked gas is eliminated. Since all of the hydrogenation reactions occur upstream of the hydrogen-methane separation steps, the hydrogen required for the hydrogenation reactions is already present in the charge gas. The elimination of the cryogenic separation of the hydrogen results in energy saving, lower capital investments and less complexity in the process. In the alternative, the present invention can be employed for hydrogenating the acetylenes and dienes without significant hydrogenation of olefins.

In the two processing sequences currently practiced, fouling in the fractionation towers bottoms typically occurs due to the presence of

acetylenes and dienes. The bottoms operating temperatures of these towers are limited to minimize the fouling tendencies but often spare equipment must be provided to ensure continuity of plant operation.

Hydrogenating the dienes and acetylenes prior to the fractionation towers eliminates the fouling tendencies in the fractionation tower bottoms.

Brief Description of the Drawings Figure 1 is a flow sheet for a conventional prior art olefin plant.

Figure 2 is a flow sheet for a portion of an olefin plant according to the present invention.

Figure 3 is a flow sheet for the remaining portion of an olefin plant according to the present invention illustrating the downstream processing of the olefin containing vapors.

Figure 4 is a flow sheet similar to the flow sheet of Figure 2 but illustrating an alternate embodiment of the present invention.

Description of the Preferred Embodiments Referring first to Figure 1 which illustrates a conventional prior art olefin plant such as the first process sequence previously discussed, a charge gas 10 is first compressed at 12 up to a pressure of 2.76 to 4.14 MPa (400 to 600 psia). The majority of the compressed gas then undergoes cryogenic treatment at 14 to separate hydrogen followed by separation of methane at 16. A small portion of the C3 and heavier material condenses in the compressor train and often bypasses the cryogenic demethanization and deethanization steps going directly to the depropanizer 30 as stream 31. The gas stream 18 is then deethanized at 20 with the C2 gas stream being hydrogenated at 22 and fractionated at 24 to produce essentially ethylene 26 and ethane 28. The bottoms from the deethanizer 20 are depropanized at 30 with the separated C3 stream 32 being hydrogenated at 34 and fractionated at 36 to produce essentially propylene 38 and propane 40. Likewise, the bottoms from the depropanizer 30 are debutanized at 42 with the C4 stream being

hydrogenated at 44 and the C6+ stream being hydrogenated at 46. As can be seen, nearly the entire feed stream is subjected to cryogenic treatment and the separation of hydrogen before any hydrogenations or fractionations are carried out. The separated hydrogen is then used downstream in the hydrogenation units 22,24, 44 and 46. This scheme with its cryogenic treatment and hydrogen separation has the disadvantages previously discussed.

Figure 2 illustrates the present invention where the charge gas 50 is compressed at 52 but only up to a pressure of 0.69 to 1.72 MPa (100 to 250 psia) and preferably to 1.21 MPa (175 psia). The compressed charge gas stream is fed into the feed zone 54 of a catalytic distillation tower 56. This catalytic distillation tower is a device which simultaneously carries out a catalytic reaction and distillation and comprises a stripping section 58 below the feed zone 54 and a rectifying/reaction section 60 above the feed zone 54. The stripping section 58 contains any desired distillation internals such as conventional trays 62 illustrated in Figure 2.

Reboiler 63 returns heated bottoms to the column.

The rectifying/reaction section 60 of the column 56 has the dual function of reacting (hydrogenating) selected components of the feed and distilling the components. Therefore, this section contains beds of a conventional hydrogenating catalyst 64. The criteria for this rectifying/reaction section is that conditions be created wherein the unsaturated hydrocarbons, with the exception of ethylene and propylene, are hydrogenated and wherein the requisite distillation is accomplished to separate essentially all of the C4 and lighter material as overhead and essentially all of the C, and heavier materials as bottoms. A portion of the C5 materials, 10 to 90% and typically 70%, exits the column overhead and the remaining portion, typically 30%, exits the column as bottoms. In some cases, all of the C5 will exit the tower overhead depending upon process, feedstock and byproduct requirements of the individual plants.

In order to selectively hydrogenate the C2 acetylenes, the C3 acetylenes and dienes and the C4 and heavier acetylenes, dienes and olefins while

leaving the ethylene and propylene unhydrogenated, the rectifying/reaction section 60 of the column 56 is operated such that there is a substantial concentration gradient of C4 and C6 materials relative to C2 and C3 materials in the liquid phase where the majority of the hydrogenation reaction occurs. In the preferred embodiment, this is accomplished by the use of a high liquid downflow, for example, by using a high reflux ratio and large intercondensing duties. The column reflux which is produced by overhead condensers 86 and 88 and column intercoolers or intercondensers 80 also removes the high heat of reaction.

As shown in Figure 2, the catalyst is separated into a series of discrete beds 66,68 and 70. Although three beds are shown, this is only by way of example and could be any number of beds depending on the dynamics of any particular plant. These catalyst beds are retained between the screens or perforated plates 72. Located between the catalyst beds are liquid collecting trays 74 which include vapor flow ports or chimneys 76. The liquid descending from a catalyst bed collects on the respective tray and drains into the sumps 78. The liquid is withdrawn from the sumps 78 as side streams through the intercondenser 80 and is then reinjected back into the column over the next lower catalyst bed through the distribution headers 82. This permits a portion of the heat of reaction to be removed in the intercondensers. By arranging the intercondensers in this fashion, the cooling medium can be cooling water while the cooling medium in the overhead condensers may need to be partly by use of mechanical refrigeration. Hence, the use of the intercondensers can significantly reduce the portion of the heat of reaction which needs to be removed by mechanical refrigeration.

The overhead 84 from the column is cooled in the overhead condenser 86 with cooling water and in the condenser 88 with refrigeration and the resulting vapor and liquid separated at 90. The processing of the collected vapor in line 94 will be discussed hereinafter. The resulting liquid from separator 90 is pumped through line 96 back into the column as reflux. A number of trays are provided to fractionate out ethylene and

propylene from the liquid phase preventing these from entering the catalyst beds in high concentrations relative to the C4 and C5 material.

In the present invention, it is imperative to limit the loss of ethylene and propylene in the hydrogenation reaction because these are the principal products of an ethylene or olefin plant. However, under conventional conditions which would permit the hydrogenation of the C4 and heavier olefins, ethylene and propylene losses by hydrogenation would be unacceptably high. That is the primary reason why one of the currently practiced prior art process sequences described earlier only hydrogenates the C2 acetylenes and a portion of the C3 acetylenes and dienes upstream of the chilling and condensing step.

The hydrogenation in the column 56 occurs mostly in the liquid phase. The extent of the reaction is dependent upon the relative reactivity of the various components and the concentration of these components in the liquid phase at any particular point in the column. The C2 and C3 acetylenes and dienes are far more reactive than ethylene and propylene so that they react first and rapidly. However, the relative reactivities of ethylene, propylene and the C4 and heavier olefins, dienes and acetylenes are much closer. In order to react a significant quantity of the C4 and heavier olefins, dienes and acetylenes without any significant loss of ethylene and propylene, the concentration of the ethylene and propylene in the liquid phase must be minimized and the concentration and temperature profiles from top to bottom must be controlled. Since the hydrogenation occurs in a fractionation tower, this control can be accomplished by adjusting the overhead (external) reflux produced by the overhead condensers 86 and 88 and the side stream reflux from the intercondensers 80.

The feed 54 to the column at the previously mentioned pressure of 1.25 MPa (0.69 to 1.72 MPa) is in the temperature range of 25 to 120°C and preferably 70-90°C. At the feed point, the concentration of the hydrogen is the highest, the temperature (in the rectifying/reaction section) is the highest and the concentration of ethylene and propylene in the liquid

phase is the lowest. At this point, the concentration of C4 and C6 components in the liquid phase relative to the concentration of propylene is maintained in the range of 10 to 80 and preferably about 25 while the concentration of C4 and C5 in the liquid phase relative to ethylene is maintained in the range of 30 to 100 and preferably about 80. This low concentration of C2 and C3 in the rectifying/reaction section is achieved by a high liquid downflow ratio. This high liquid downflow ratio can be achieved by a high overhead reflux ratio and/or by the reflux created by the intercondensers 80. As will be explained later with respect to Figure 4, this high liquid downflow ratio can also be provided by the recycle and cooling of heavies from the bottom of the column. More specifically, the liquid downflow ratio provided by the overhead reflux 96, the intercoolers 80 and the recycle of heavies (160 in Figure 4) is equivalent to the liquid downflow that would be provided by an overhead reflux ratio in the range of about 0.2 to 10 without intercondensers and heavies recycle. This compares to a reflux ratio of less than 0.2 for a conventional column operated to achieve a similar overhead product specification. At the top of the rectifying/reaction section 60, where the temperature is 38 to 80°C and preferably 60°C and where the concentration of hydrogen is low because most of it has reacted, the ratio of C4 and C. components to C2 and C3 components is similarly high. The overhead reflux ratio and intercondenser temperatures are adjusted to maintain these operating parameters. With the hydrogenation of the C2 acetylenes, the C3 acetylenes and dienes and the C4 acetylenes, dienes and olefins, and a major portion of the C5 and C. acetylenes, dienes and olefins, 50 to 90% of the hydrogen contained in the cracked feed gas is reacted.

The bottoms 98 from the column 56 contain a portion of the C5 material and essentially all of the Ce and heavier material. In the preferred embodiment, this bottoms product is sent to a second catalytic distillation hydrogenation column 100 for the production of hydrogenated pyrolysis gasoline. Alternately, the bottoms product can be burned in the plant fuel system or pumped and sent to a conventional fixed pyrolysis gasoline

hydrotreater as previously described under prior art. Also, in the preferred embodiment shown in Figure 2, the total net overhead 94 from the column 56, containing a portion of the C6 material and essentially all of the C4 and lighter material, is first compressed at 102 and sent to a hydrogen recovery membrane devices 104. Such membrane devices are commercially available for the separation of hydrogen. The intent of the membrane is to recover most of the hydrogen remaining in the overhead stream 94. The resulting hydrogen stream 106 is then fed to the pyrolysis gasoline hydrogenation column 100 along with the bottoms from the column 56. The compression step may or may not be required depending on the specific composition of the cracked gas, hydrogen membrane section, and operating condition of column 56. Alternately, a conventional fixed bed pyrolysis gasoline hydrotreater could be used without a membrane separator. In this case, the hydrogen now significantly reduced in stream 94 by the hydrogenation reactions occurring in column 56 would be cryogenically recovered as previously discussed.

Pyrolysis gasoline is a complex mixture of hydrocarbons ranging from C. compounds through materials with a boiling point of about 200°C.

The raw feed to the pyrolysis gasoline column 100 is highly unstable due to its high content of diolefins. Therefore, in the production of the pyrolysis gasoline, the feed is hydrogenated in the column 100. The column 100 is similar to the column 56 in that it has a typical bottom stripping section 108, a reboiler 110 and an upper rectifying/reaction section 112 containing the hydrogenation catalyst. It includes an overhead condenser 114 and separator 116 from which reflux 118 is returned to the column. The column may or may not include intercoolers or intercondensers similar to the intercondensers for column 56. In this column 100, the feed of the remaining C5 acetylenes, dienes and olefins and all of the C6 and heavier acetylenes, dienes and olefins is hydrogenated. This column operates between 0.21 and 0.86 MPa and preferably 0.34 MPa. The C. and lighter materials in the feed enter the catalyst bed where the acetylenes, dienes and olefins are hydrogenated. The Cg and heavier material exits from the

bottoms of column 100. The heat of reaction is removed by the reflux ux stream 118.

The reflux stream 118 also serves to control the selectivity of the hydrogenation reaction. There is a small amount of ethylene in stream 106 and, as has been pointed out, this ethylene is a valable product and its hydrogenation should be avoided. By the proper control of the column reflux 118, ethylene concentration in the liquid phase in the column can be minimized. This is a technique which is preferable to upgrading the membrane separation process to essentially exclude ethylene from passing through with the hydrogen. The passage of ethylene could be minimized by decreasing the pressure differential across the membrane and/or by increasing the membrane surface area. However, adding membrane surface area is a capital intensive cost and increasing the pressure differential is both energy and capital intensive. The ability to selectively hydrogenate in the column 100 permits a lower capital cost, less energy intensive process. The overhead vapor 120 from the column containing primarily C4 and lighter material is recycled to the feed for the process. The net overhead product condensed liquid is removed at 122 as pyrolysis gasoline.

Figure 3 illustrates the processing of the overhead stream 94 after it passes through the hydrogen separation step at 104 and emerges as stream 124. Alternately, this system can be used to process the stream 94 directly in the event that the membrane separation and pyrolysis gasoline portions of the process described above were not used. In that event, additiona I provisions would be made for cryogenic hydrogen separation.

The vapor stream 124 is chilled at 128 as required to liquify the C2 and heavier components. The methane overhead 130 is then separated in the demethanizer tower 132 from the C2 and heavier bottoms 134.

These bottoms 134 are then separated in the deethanizer tower 136 to produce a C2 overhead 138 and a C3 and heavier bottoms 140. The C2 overhead 138, which may first go through a drying step (not shown), is

then separated in tower 142 into ethane bottoms 144 and ethylene overhead 146. The bottoms 140 from the deethanizer 136 is then separated in tower 148 into a C4 and heavier bottoms 150 and a C3 overhead 152. This overhead 152, which may also then be dried, is fed to the tower 154 for the separation of propane 156 and propylene 158.

Figure 4 illustrates an alternate preferred embodiment of the present invention which incorporates recycles from the stripping section 58 of the column 56. In this embodiment, a recycle stream 160 from the stripping section 58 is recycle either to the column overhead 84 through line 161 and/or to the catalytic zone of the rectifying/reaction section 60 through line 163. Recycle via line 163 to the catalyst zone only is usually preferable. For example, this recycle may be a portion 162 of the bottoms 98 and/or a portion 164 from within the stripping section. This recycle 160 serves to recycle the heavies, C6 +, to the overhead or to the catalytic zone of the column. This increases the amount of dienes and acetylenes and perhaps some olefins which will be hydrogenated, thereby increasing the consumption of hydrogen. Also, it provides another control variable to increase the overhead temperature of the tower and/or of the catalyst bed.

Increasing the overhead temperature of the tower is desirable since it will decrease or eliminate the refrigeration requirements for generating the reflux. Increasing the temperature of the catalyst bed provides another variable to control the reaction rate of the catalytic reaction beds.

Although this embodiment achieves distillation internally in the column, it is not classic distillation since there is now some heavies in the overhead.

In that case, some further distillation would be provided downstream to make the final desired separations. The purpose of this embodiment is to improve the control of the reactions taking place in the tower 56 even though that also sacrifices some of the separation by distillation. When the catalytic distillation column is operated with a recycle of heavies to the overhead, the heavy stream is preferably cooled at 165. This cooling effect can be considerable, especially if a high recycle rate of these heavies is utilized. This reduces the reflux ratio of the catalytic distillation

column at equal liquid downflow rates. The reflux rate is further reduced if side cooling is utilized. The net effect of all of these cooling steps is to significantly decrease the reflux ratio. This can reduce refrigeration requirements as some of the cooling required to condense reflux may be provided at higher condensing temperatures. This can be another benefit of recycling heavies from the bottom section to the upper section of the column especially recycling to the vapor outlet 84 since that will raise the overhead temperature and reduce the refrigeration requirements.

When the catalytic distillation column is operated with bottoms recycle, the overhead reflux ratio is in the range of 0.05 to 0.4 and preferably, 0.1 to 0.2 when the bottoms recycle is directed to the top of the catalyst bed through line 163. When the bottoms recycle is directed to the overhead of the column through line 161, the reflux ratio is 0.2 to 10. But even with this lower overhead reflux ratio, a high liquid downflow ratio is maintained in the catalyst beds by the intercondensers and heavies recycle and cooling. The heavies recycle does not conform to what would be considered"classical"distillation since the recycle results in the loss of some of the net separation benefits of the distillation. However, this loss is outweighed by the benefit of minimizing ethylene and propylene concentrations in the liquid in the catalyst zone by the use of high liquid downflow rates and by the benefits of raising the temperature in the catalyst bed.

The ability of the present invention to remove 85 to almost 100%, typically 90%, of the hydrogen contained in the charge gas prior to chilling and condensation steps lowers the energy consumption and reduces capital costs. By using the hydrogen contained in the charge gas as the source of hydrogen for the various hydrogenation reactions, the need for the separate cryogenic separation of hydrogen is eliminated. By the proper control of the concentration profiles in the catalytic distillation hydrogenation column, the C4 and heavier olefins can be hydrogenated without any significant hydrogenation of either ethylene or propylene.

Therefore, the hydrogenation reactions are combined into one or two reactor systems.