CRYOGENIC SEPARATION The present invention relates to cryogenic separation of light gases, in particular for recovering ethene (ethylene) or propene (propylene) from a mixture containing two or more light gases. Cryogenic technology has been employed on a large scale for recovering gaseous hydrocarbon components, such as C^C;, alkanes and alkenes from diverse sources, including natural gas, petroleum refining, coal and other fossil fuels. Separation of high purity ethene from other gaseous components of cracked hydrocarbon effluent streams has become a major source of chemical feedstocks for the plastics industry. Polymer grade ethene, usually containing less than 1% of other materials, can be obtained from numerous industrial process streams. Thermal cracking and hydrocracking of hydrocarbons are employed widely in the refining of petroleum to yield a slate of valuable products, such as pyrolysis gasoline, lower olefins and LPG, along with byproduct methane and hydrogen. Conventional separation techniques near ambient temperature and pressure can recover many cracking effluent components by sequential liquefaction, distillation, sorption, etc. However, separating methane and hydrogen from the more valuable C ₂ + aliphatics, especially ethene, ethane, propene, and/or propane requires relatively expensive equipment and processing energy. Primary emphasis herein is placed on a typical large scale cryogenic plant for recovering ethene from cracking gas.

Typical cryogenic systems are described in U.S. Patent Nos 3,126,267; 3,702,541; 4,270,940; 4,460,396; 4,496,380; 4,368,061 and 4,900,347.

It is an object of the present invention to provide an improved cryogenic separation system for separating light gases which is energy efficient and saves capital investment in cryogenic equipment. Accordingly, the invention resides in one aspect in a cryogenic separation system for separating a mixture

containing at least three volatile components each having different normal boiling points; comprising: a) first and second distillation towers, each having an upper reflux stage, a middle distillation stage and a lower reboiler stage; the second distillation tower being operatively connected to receive a first overhead vapor stream from the first distillation tower; b) compression means operatively connected to receive and adiabatically compress a second overhead vapor stream rich in at least one low boiling component from the second distillation tower reflux stage; c) means for passing adiabatically compressed vapor from the compressor means to the second distillation tower reboiler stage for condensing the compressed vapor and heating a liquid reboiler stream; d) flashing means for decreasing pressure of the condensed vapor to provide a partially vaporized flashed mixture stream rich in low boiling component; e) reflux fluid handling means operatively connected for receiving the flashed mixture stream, recovering a liquid portion and a vapor portion thereof, and for passing the liquid portion to the second distillation tower reflux stage; f) means for withdrawing an intermediate liquid stream rich in low boiling and medium boiling components from a middle stage of the second distillation tower and passing said intermediate liquid stream to the first distillation tower reflux stage; g) means for recovering at least one high boiling component from the first distillation tower reboiler stage; h) means for recovering at least one middle boiling component from the second distillation tower reboiler stage; and i) means for recovering the low boiling component.

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in a further aspect, the invention resides in a process for separating a hydrocarbon mixture containing an alkene, a corresponding alkane having the same number of carbon atoms as the alkene and at least one heavier hydrocarbon component comprising the steps of: a) feeding said hydrocarbon mixture to a first distillation tower having an upper reflux stage; b) recovering a first overhead vapor stream rich in alkene and alkane from the first distillation tower and passing said first overhead vapor stream to a middle distillation stage of a second distillation tower; c) recovering a second overhead vapor stream rich in alkene from the second distillation tower; d) adiabatically compressing the alkene-rich second overhead vapor stream and passing said compressed vapor to a reboiler stage of the second distillation tower to cool and condense the compressed vapor and heat a liquid reboiler stream; e) flashing the cooled and condensed vapor from the reboiler stage of the second distillation tower to provide a partially vaporized flashed mixture stream rich in alkene; f) recovering and separating the flashed mixture stream to provide a liquid portion and vapor portion; g) passing the liquid portion to a reflux stage of the second distillation tower; h) withdrawing an intermediate liquid stream rich in alkene and alkane from a middle stage of the second distillation tower; i) passing said intermediate liquid stream to the first distillation tower reflux stage; j) recovering the heavier component from the first distillation tower; k) recovering alkane from the second distillation tower reboiler stage; and

1) recovering an alkene product stream. The present invention is useful for separating mainly C ₂ -C ₄ + gaseous mixtures containing large amounts of ethene, ethane and/or propene/propane. Significant amounts of hydrogen and methane usually accompany cracked hydrocarbon gas, along with minor amounts of C ₃ + hydrocarbons, nitrogen, carbon dioxide and acetylene. The acetylene component may be removed before cryogenic operations. Typical petroleum refinery offgas or paraffin cracking effluent is usually pretreated to remove any acid gases and dried over a water-absorbing molecular sieve to a dew point of about 145°K to prepare the cryogenic feedstock mixture. A typical feedstock gas comprises cracking gas containing 10 to 50 mole percent ethene, 5 to 20% ethane, 10 to 40% methane, 10 to 40% hydrogen, and up to 10% C ₃ hydrocarbons. This feedstock is demethanized and may be depropanized and/or de-ethanized to concentrate the desired components in a feedstream suitable for use in the improved process described herein. In a preferred embodiment, dry compressed cracked feedstock gas at ambient temperature or below and at process pressure of at least 2500 kPa (350 psic) , preferably about 3700 Kpa (520 psig) , is separated in a chilling train under cryogenic conditions into several liquid streams and gaseous methane/hydrogen streams. The more valuable ethene stream is recovered at high purity suitable for use in conventional polymerization.

The invention will now be more particularly described with reference to the accompanying drawings, in which: Fig. 1 is a schematic process flow diagram depicting arrangement of unit operations for a typical hydrocarbon processing plant utilizing cracking and cold fractionation for ethene production; and

Fig. 2 is a detailed process and equipment diagram showing in detail an improved multi-tower distillation section for de-propanizing a cryogenic fraction and

splitting a C ₃ stream into propene and other product streams.

Referring to Fig. 1, the processing plant shown includes conventional hydrocarbon cracking unit 10 which converts fresh hydrocarbon feed 12 and optionally recycled hydrocarbons 13 to provide a cracked hydrocarbon effluent stream. The cracking unit effluent is separated by conventional techniques in separation unit 15 to provide liquid products 15L, C ₃ -C ₄ petroleum gases 15P and a cracked light gas stream 15G, consisting mainly of methane, ethene and ethane, with varying amounts of hydrogen, acetylene and C ₃ + components. The cracked light gas is brought to process pressure by a compressor 16 and cooled below ambient temperature by heat exchange means 17, 18 to provide feedstock for the cryogenic separation, as herein described.

The cold pressurized gaseous feedstock stream is separated in a plurality of sequentially arranged dephlegmator-type rectification units 20, 24. Each of these rectification units is operatively connected to accumulate condensed liquid in a lower drum portion 20D, 24D by gravity flow from an upper rectifier heat exchange portion 20R, 24R comprising a plurality of vertically disposed indirect heat exchange passages through which gas from the lower drum portion passes in an upward direction for cooling with lower temperature refrigerant fluid or other chilling medium by indirect heat exchange within the heat exchange passages. Methane-rich gas flowing upwardly is partially condensed on vertical surfaces of the heat exchange passages to form a reflux liquid in direct contact with the upward flowing gas stream to provide a condensed stream of cooler liquid flowing downwardly and thereby enriching condensed liquid gradually with ethene and ethane components. The preferred system provides means for introducing dry feed gas into a primary rectification zone or chilling

train having a plurality of serially connected, sequentially colder rectification units for separation of feed gas into a primary methane-rich gas stream 20V recovered at low temperature and at least one primary liquid condensate stream 22 rich in C ₂ hydrocarbon components and containing a minor amount of methane.

The condensed liquid 22 is purified to remove methane by passing at least one primary liquid condensate stream from the primary rectification zone to a fractionation system having serially connected demethanizer zones 30, 34. A moderately low cryogenic temperature is employed in heat exchanger 31 to refrigerate overhead from the first demethanizer fractionation zone 30 to recover a major amount of methane from the primary liquid condensate stream in a first demethanizer overhead vapor stream 32 and to recover a first liquid demethanized bottoms stream 30L rich in ethane and ethene and substantially free of methane. Advantageously, the first demethanizer overhead vapor stream is cooled with moderately low temperature refrigerant, such as available from a propylene refrigerant loop, to provide liquid reflux 30R for recycle to a top portion of the first demethanizer zone 30.

An ethene-rich stream is obtained by further separating at least a portion of the first demethanizer overhead vapor stream in an ultra-low temperature final demethanizer zone 34 to recover a liquid first ethene-rich hydrocarbon crude product stream 34L and a final demethanizer ultra-low temperature overhead vapor stream 34V. Any remaining ethene is recovered by passing the final demethanizer overhead vapor stream 34V through ultra low temperature heat exchanger 36 to a final rectification unit 38 to obtain a final ultra-low temperature liquid reflux stream 38R for recycle to a top portion of the final demethanizer fractionator. A methane-rich final rectification overhead vapor stream 38V is recovered substantially free of C ₂ + hydrocarbons. Utilizing the dual

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demethanizer technique, a major amount of total demethanization heat exchange duty is provided by moderately low temperature refrigerant in unit 31 and overall energy requirements for refrigeration utilized in separating C ₂ + hydrocarbons from methane and lighter components are decreased. The desired purity of ethene product is achieved by further fractionating the C ₂ + liquid bottoms stream 30L from the first demethanizer zone in a de-ethanizer fractionation tower 40 to remove C ₃ and heavier hydrocarbons in a C ₃ + stream 40L and provide a second crude ethene stream 40V, which is recovered as a vapor without substantial condensation or direct reflux according to the improved operating technique.

The present invention achieves improved operating economy and lower capital equipment requirements by passing overhead vapor stream 40V to a middle stage of distillation tower unit 50, commonly known as a C ₂ product splitter. Ethene-rich vapor is recovered from tower 50 via overhead 50V. Optionally, the polymer grade product is obtained by cofractionating the second crude ethene stream 40V and the first ethene-rich hydrocarbon crude product stream 34L to obtain a purified ethene product. The ethane bottoms stream 50L can be optionally recycled to cracking unit 10 , with recovery of thermal values by indirect heat exchange with moderately chilled feedstock in exchangers 17, 18 and/or 2OR. C ₃ + stream 40L may be sent to downstream fractionation facilities for recovery of other valuable components such as propene, butenes, etc.

Overhead vapor stream 50V is compressed adiabatically in compressor unit 60 to recover energy as a heat pump to reboiler 50B, after which the stream 50V is combined with an optional bypass stream from trim cooler 62 and depressurized by flashing means 64, partially condensing the ethene-rich stream. The partially condensed stream is fed to phase separator vessel 66 which recovers a liquid reflux stream 50R, which is fed to the reflux stage of the

tower 50, and an uncondensed vapor stream 69 which is combined with tower overhead stream 50V for recompression. Ethylene product is conveniently recovered as a liquid stream 68 from the compressor 60. A major advantage of this invention is realized by withdrawing a liquid C ₂ stream 4OR from tower 50 adjacent the inlet of stream 40V and passing liquid 4OR to an upper stage of tower 40 as reflux. The effective reflux ratio is maintained at less than 0.5, preferably 1:5 to 1:10 and most preferably at about 0.15 (wt. of liquid reflux/wt. of total overhead vapor) . This feature of the invention will be seen in the comparison of operating the present system with that of prior art distillation.

One of the major operating advantages for C ₂ cryogenic recovery systems is the enhanced separation of ethane and ethene that can be achieved in the same distillation column at lower pressure. The combination of the "umbilical" reflux arrangement between two adjacent towers permits greater savings in utility costs for this technique. An improved propene recovery fractionation system is shown in Fig. 2, wherein ordinal numbers correspond with their counterpart equipment in Fig. 1. The feedstock is exemplified by a propene-rich feedstream 130L, which feed has been de-ethanized to remove C ₂ - components and heavy cracking liquids to provide gaseous or liquid feedstock containing propene, propane and C ₄ + components, such as butenes and butanes. Multiple liquid or gas feedstreams may be employed, for instance, additional stream 130A. As depicted in Fig. 2, there are first and second distillation towers 140, 150, each having an upper reflux stage, middle distillation stages and a lower reboiler stage, with the second distillation tower 150 being operatively connected to receive a first overhead vapor stream 140V from the first distillation tower 140 at a middle stage. The system includes conventional means for controlling operating pressure in the second distillation tower at predetermined

pressure, as in a typical cryogenic fluid handling system by compressor, pump and valve control means.

Single stage compression is ordinarilty sufficient but, in the example shown in Fig. 2, multi-stage compression means 160A, 160B are operatively connected to receive a second overhead vapor stream 150V rich in at least one low boiling component (e.g. propene) from the second distillation tower upper reflux stage for adiabatic compression. Conduit means 161 is provided for passing adiabatically compressed vapor from the last stage compressor 160B to the second distillation tower reboiler stage 150B for condensing the compressed vapor and heating the liquid reboiler stream.

Flashing means is provided for decreasing pressure on the condensed vapor to provide a partially vaporized flashed mixture stream rich in low boiling component. This can be achieved in a single flashing unit; however, it is advantageous to achieve pressure reduction by way of a series of expansion turbines 164A, 164B operatively connected for fluid flow and mechanically linked to corresponding compressors to recover energy from the flashing expansion during the depressurizing steps. Intermediate separator unit 165 provides an intermediate vapor stream 165V for mixing with first stage compressed vapor stream 160C as feed to the second stage compressor 160B.

Reflux fluid handling means is provided by separator unit 166 operatively connected for receiving the flashed mixture stream 164V, recovering a liquid portion 150R and passing this liquid portion 150R to the second distillation tower 150 reflux stage. Pump means 14OP is operatively connected by conduits for withdrawing an intermediate liquid stream 140R rich in low boiling and medium boiling components (e.g. propene and propane) from a middle stage of the second distillation tower 150 and passing the

intermediate liquid stream to the first distillation tower 140 reflux stage. The desired reflux ratio (i.e. less than 0.5) may be controlled by convertional fluid handling means, pump 14Op, valve means, ratio controller, etc. Bottom conduit means 140L recovers at least one high boiling component (e.g. C ₄ +) from the first distillation tower reboiler stage, conduit means 150L recovers at least one middle boiling component (e.g. propane) from the second distillation tower reboiler stage; and conduit means 168 recovers the low boiling component (e.g. propene) from the compressor 160B.

In order to obtain full benefit from the "umbilical" configuration wherein reflux heat load for the primary distillation unit is provided by the rectification in the second distillation unit, it is desirable to provide conventional fluid control means for maintaining operating pressure in the first distillation unit not substantially greater than the second distillation operating pressure, usually less that 10-20% greater than the absolute second pressure. In the separation of propene from heavier hydrocarbons, lower pressure operation of the depropanizer tower permits a lower temperature operation in the reboiler stage thereof, thus avoiding undesirable reactions in this zone, especially polymerization of unsaturated C ₄ 's, such as butenes and diene.

Example A material balance with energy requirements is given for the production of polymer grade ethene according to the present invention and compared with conventional cryogenic distillation. In the following table, all units are based on steady state continuous stream conditions and the relative amounts of the components in each stream are based on 100 parts by weight of the feedstream. The utility requirements of de-ethanizer and C ₂ splitter tower operations are given.

STREAM COMPOSITION PER lOOkG OF STREAM RATE (Kg per 100 Kg)

Stream No.

(Fiα. 1) 30L 40V 4OR 40L 50L 68 5OR

Ethylene 66.61 71.61 59.49 0.00 0.08 99.89 99.89

Ethane 23.57 28.19 40.48 0.25 99.04 0.11 0.11

Propodiene 0.80 0.00 0.00 8.35 0.00 0.00 0.00

Propylene 7.36 0.20 0.03 74.24 0.87 0.00 0.00

Propane 1.66 0.00 0.00 17.16 0.01 0.00 0.00 lOO.OOKg lOO.OOKg lOO.OOKg lOO.OOKg lOO.OOKg lOO.OOKg lOO.OOKg

10 STREAM ENTHALPIES PER 100 Kσ OF STREAM RATE

(KJ per 100 Kg)

30L 40V 4 OR 40L 50V 68 5 OR

^■ 7,098 +23,241 -17,138 +3,620 -11,735 +30,135 ^■ 20,378

Deethanizer Tower

Overhead Pressure, kPa 859.75

Overhead Temp, °K 222.7

Bottoms, Temp, °K 289.8 Reflux Ratio, Kg reflux/Kg overhead vapor 0.15

C ₇ Splitter Tower

Overhead Pressure, kPa 790.80

Overhead Temp, °K 214.4

Bottoms Temp, °K 235.6 Reflux Ratio, Kg reflux/Kg overhead vapor 0.70

Process Duties per 100 Kg of System Feed (KJ per 100 Kg)

Deethanizer Reboiler 34,766

Deethanizer Condenser (omitted) None C ₂ Splitter Reboiler 48,142

C ₂ Splitter Trim Cooler 35,504

C ₂ Splitter Heat Pump 25,443

It will be appreciated by one skilled in cryogenic engineering that the arrangement of unit operations allows reduction of reflux cooling requirements in the de¬ ethanizer zone as compared to conventional reflux type distillation units.

The low pressure, combined deethanizer/C ₂ splitter system requires 20% less process refrigeration than a conventional, high pressure separate dethanizer/C ₂ splitter system. In addition, the capital equipment cost for the combined deethanizer/C ₂ splitter system is less than a conventional system. The advantages of the combined low pressure deethanizer/C ₂ splitter can be classified into two areas: the advantages of low pressure deethanization, and the advantages of using the C ₂ splitter to reflux the deethanizer.

Operating the deethanizer at the lower overhead pressure (859.75 KPA vs 2983.33 KPA) facilitates the separation of ethane and propylene. The improved fractionation performance results from the inverse proportionality between the relative volatility of ethane to propylene and distillation pressure. The improved performance is manifested in a lower requirement for reflux in the above low pressure deethanizer tower. The performance reflux ratio for the low pressure deethanizer is maintained below 0.2 of an ethene recovery unit, preferably 0.15, while the required ratio for a conventional high pressure deethanizer is 0.38.

The reduced reflux requirement of the low pressure deethanizer results in two direct benefits: 1) a reduction in the process refrigeration required to condense the deethanizer overhead vapor. Since less reflux is required, less vapor needs to be condensed. This results in a direct utility savings in the operation of refrigeration system compressors; 2) a reduction in reflux pumping costs due to lower reflux volumes.

An additional benefit of the low pressure deethanizer is the ability to reboil the tower with condensing propylene refrigerant. The low pressure deethanizer requires a lower reboiler temperature than the high pressure deethanizer (289.8 "K vs. 344.4 °K) . The lower reboiler temperature of the low pressure deethanizer is approximately the condensing temperature (dew point temperature) of high pressure propylene refrigerant. Therefore, the low pressure deethanizer reboiler could be used to condense refrigerant, providing an energy credit to the refrigeration system.

Use of a liquid draw from the C ₂ splitter to provide reflux for the deethanizer results in a less expensive design than a conventional separate deethanizer/C ₂ splitter system. Both the combined and separate systems require the same distillation towers, towers reboilers, and C ₂ splitter

heat pump equipment. However, the conventional deethanizer/C ₂ splitter system requires a deethanizer overhead condenser and a deethanizer reflux drum, whereas these components are not required in the combined system of the invention. As a result, the total equipment cost for the combined system is lower than a conventional system. The liquid draw from the C ₂ splitter tower does not significantly affect the operation of the C ₂ splitter. The liquid rate in the C ₂ splitter tower is an order of magnitude higher than the liquid draw used for deethanizer reflux. The power requirement for the C ₂ splitter heat pump increases by less than 3% when the deethanizer reflux stream is withdrawn from the C ₂ splitter tower.

The increase in C ₂ splitter trim cooler duty is more than offset by the elimination of the deethanizer condenser. The two units requiring process refrigeration in the deethanizer/C ₂ splitter system are the deethanizer condenser and the C ₂ splitter trim cooler. The combined, low pressure deethanizer/C ₂ splitter system provides a 20% net reduction in overall refrigeration requirements over a conventional system.