Patent Application
20210131728
The present invention relates to a process and an apparatus for separating a hydrocarbon used for separating and recovering ethane or propane from, for example, natural gas, petroleum-associated gas, or off-gas from a refinery or petrochemical plant.
Conventionally, separation of methane and hydrocarbons having 2 or more carbon atoms and separation of ethane and hydrocarbons having 3 or more carbon atoms have been carried out.
For example, as a process of recovering ethane or propane from natural gas, a process including cooling natural gas and distilling and separating a light component and an ethane (or propane) and a heavy hydrocarbon component in a demethanizer (in the case of propane recovery, a deethanizer) is widely used. In the process, a propane refrigeration system and a turboexpander are used to cool the natural gas to the temperature required for separation.
WO 2005/009930 A1 discloses a process of recovering ethane or propane from a feed gas such as natural gas using a distillation column. The process includes the following steps:
(a) a step of cooling and partly condensing the feed gas to separate into gas and liquid;
(b) a step of supplying the liquid obtained in step (a) to a distillation column;
(c) a step of expanding the gas obtained in step (a) by an expander, condensing a part of the expanded gas to separate into gas and liquid;
(d) a step of feeding the liquid obtained in step (c) to a distillation column;
(e) a step of dividing the gas obtained in step (c) into a first portion and a second portion;
(f) a step of feeding the first portion to the distillation column;
(g) a step of compressing and cooling the second portion to be condensed, and then reducing pressure and feeding to the distillation column as a reflux;
(h) a step of obtaining residual gas from the top of the distillation column and obtaining a heavy fraction from the bottom of the distillation column.
In the process described in WO 2005/009930 A1, the liquid obtained in step (c) is directly supplied to the distillation column. Therefore, there is room for improvement from the viewpoint of cold heat recovery, and a relatively large compressor power is required for the recovery of ethane or propane.
It is an object of the present invention to provide a process for separating hydrocarbons capable of recovering ethane or propane, including an improved cold heat recovery allowing reduction in compressor power. It is another object of the present invention to provide an apparatus for separating a hydrocarbon, suitable for carrying out this process.
An aspect of the present invention provides,
a process for separating hydrocarbons, wherein a feed gas containing at least methane and a hydrocarbon less volatile than methane is separated into a residual gas enriched with methane and lean in a hydrocarbon less volatile than methane and a heavy fraction lean in methane and enriched with a hydrocarbon less volatile than methane using a distillation column, the process comprising:
a) partially condensing the feed gas by cooling using the residual gas and another refrigerant as a refrigerant, followed by vapor-liquid separation;
b) depressurizing and supplying the liquid obtained from step (a) to the distillation column;
c) expanding a part or all of the gas obtained from step (a) by an expander to cause partial condensation, followed by vapor-liquid separation;
d) feeding the liquid obtained from step (c) to the distillation column after using it as the further refrigerant in step (a);
e) feeding a part or all of the gas obtained from step (c) to the distillation column; and
f) obtaining the residual gas from the top of the distillation column and the heavy fraction from the bottom of the distillation column.
Another aspect of the present invention provides,
a process for separating hydrocarbons, wherein a feed gas containing at least ethane and a hydrocarbon less volatile than ethane is separated into a residual gas enriched with ethane and lean in a hydrocarbon less volatile than ethane and a heavy fraction lean in ethane and enriched with a hydrocarbon less volatile than ethane using a distillation column, the process comprising:
a) partially condensing the feed gas by cooling using the residual gas and another refrigerant as a refrigerant, followed by vapor-liquid separation;
b) depressurizing and supplying the liquid obtained from step (a) to the distillation column;
c) expanding a part or all of the gas obtained from step (a) by an expander to cause partial condensation, followed by vapor-liquid separation;
d) feeding the liquid obtained from step (c) to the distillation column after using it as the further refrigerant in step (a);
e) feeding a part or all of the gas obtained from step (c) to the distillation column; and
f) obtaining the residual gas from the top of the distillation column and the heavy fraction from the bottom of the distillation column.
Another aspect of the present invention provides,
a separation apparatus for hydrocarbons, wherein a feed gas containing at least methane and a hydrocarbon less volatile than methane is separated into a residual gas enriched with methane and lean in a hydrocarbon less volatile than methane and a heavy fraction lean in methane and enriched with a hydrocarbon less volatile than methane, the separation apparatus comprising:
a distillation column discharging the residual gas from the top of the distillation column and the heavy fraction from the bottom of the distillation column;
a heat exchange means for partially condensing the feed gas by cooling, comprising a refrigerant flow path in which the residual gas flows as a refrigerant, and another refrigerant flow path in which another refrigerant flows;
a first vapor-liquid separator for vapor-liquid separation of the partially condensed feed gas obtained from the heat exchange means;
a line for supplying the liquid obtained from the first vapor-liquid separator to the distillation column via a pressure reducing valve;
an expander for expanding and partially condensing part or all of the gas obtained from the first vapor-liquid separator;
a second vapor-liquid separator connected to an outlet of the expander;
a line for supplying the liquid obtained from the second vapor-liquid separator to the distillation column via said another refrigerant flow path; and
a line for supplying part or all of the gas obtained from the second vapor-liquid separator.
Another aspect of the present invention provides,
a separation apparatus for hydrocarbons, wherein a feed gas containing at least ethane and a hydrocarbon less volatile than ethane is separated into a residual gas enriched with ethane and lean in a hydrocarbon less volatile than ethane and a heavy fraction lean in ethane and enriched with a hydrocarbon less volatile than ethane, the separation apparatus comprising:
a distillation column discharging the residual gas from the top of the distillation column and the heavy fraction from the bottom of the distillation column;
a heat exchange means for partially condensing the feed gas by cooling, comprising a refrigerant flow path in which the residual gas flows as a refrigerant, and another refrigerant flow path in which another refrigerant flows;
first vapor-liquid separator for vapor-liquid separation of the partially condensed feed gas obtained from the heat exchange means;
a line for supplying the liquid obtained from the first vapor-liquid separator to the distillation column via a pressure reducing valve;
an expander for expanding and partially condensing part or all of the gas obtained from the first vapor-liquid separator;
a second vapor-liquid separator connected to an outlet of the expander;
a line for supplying the liquid obtained from the second vapor-liquid separator to the distillation column via said another refrigerant flow path; and
a line for supplying part or all of the gas obtained from the second vapor-liquid separator.
According to one aspect of the present invention, there is provided a process of separating hydrocarbons capable of recovering ethane or propane, including an improved cold heat recovery allowing reduction in compressor power. According to another aspect of the present invention, there is provided an apparatus for separating a hydrocarbon, suitable for carrying out the process.
The following description and drawings are merely illustrative of preferred embodiments of the present invention, and the invention is not limited thereto. Note that, in a narrow sense, “reflux” means a liquid which is obtained by condensing the top gas of a distillation column to be returned to the distillation column again, but in addition to this, “reflux” broadly includes a liquid to be supplied to the top of a distillation column for the purpose of rectification. In this specification, “reflux” is used in the broad sense and also includes a liquid having a rectification effect for being supplied to a distillation column.
The present invention relates to an ethane recovery process and a propane recovery process. With respect to the embodiment 1 of the present invention, an example of the ethane recovery process will be described using the process flow diagram shown in
In this process, a feed gas containing at least methane and a hydrocarbon having a lower volatility than methane is separated into a residual gas enriched with methane and lean in the hydrocarbon less volatile than methane, and a heavy fraction lean in methane and enriched with a hydrocarbon less volatile than methane. For this purpose, a demethanizer 11 is used as a distillation column which discharges residual gas from the top of the column and discharges the heavy fraction from the bottom of the column. This process performs steps (a) to (f).
a) Step of Partially Condensing the Feed Gas by Cooling Using the Residual Gas and Another Refrigerant as a Refrigerant, Followed by Vapor-Liquid Separation.
In this step, a heat exchange means for partially condensing by cooling the feed gas is used, the heat exchange means including a refrigerant flow path in which the residual gas flows as a refrigerant, and another refrigerant flow path in which another refrigerant flows. In addition, a first vapor-liquid separator for vapor-liquid separation of the partially condensed feed gas obtained from the heat exchange means is used. The heat exchange means may include one or more heat exchangers. If the heat exchange means includes two or more heat exchangers, a refrigerant flow path through which the residual gas flows and another refrigerant flow path through which another refrigerant flows may be provided in the same heat exchanger, and they may be provided separately in different heat exchangers. Further, each of the plurality of heat exchangers may have a refrigerant flow path through which the residual gas flows. Each of the plurality of heat exchangers may have another refrigerant flow path. It is possible to use a plurality of refrigerants as another refrigerant, for example, one heat exchanger may have a plurality of “another refrigerant flow path” through which a plurality of refrigerant flows respectively.
The feed gas, e.g. natural gas, is cooled by heat exchange means and partially condensed. The partially condensed feed gas is separated in a first vapor-liquid separator 4, also called a low-temperature separator. To increase the recovery rate of ethane, the lower the temperature of the low-temperature separator 4, the more preferable. Further, the ratio at which the natural gas is condensed varies depending on the composition of the natural gas (the ratio of hydrocarbons having 2 or more carbon atoms), and is about 5 mol % or more and about 20 mol % or less. As the heat exchanger used for cooling the feed gas, it is possible to appropriately use a known heat exchanger such as a plate fin heat exchanger or shell and tube heat exchanger. The low-temperature separator 4 may be a vertical or horizontal vessel (a cylindrical vessel having end plates at both ends), and a mist eliminator may be provided inside the vessel in order to improve the separation efficiency of vapor and liquid.
In the example shown in
The first feed gas cooler 1 has a refrigerant flow path through which the residual gas flows and has a refrigerant flow path through which the side stream F1 flows as the “another refrigerant flow path”. The second feed gas cooler 3 has a refrigerant flow path through which the residual gas flows as a refrigerant, and has a refrigerant flow path through which the liquid obtained from the second vapor-liquid separator (line 104) flows and a refrigerant flow path through which the side stream F3 flows as the “another refrigerant flow path”. The feed gas chiller 2 has a refrigerant flow path through which propane flows of the propane refrigeration system.
b) Step of Supplying the Liquid Obtained from Step (a) to the Distillation Column Under Reduced Pressure
In this step, a line 101 for supplying the condensate obtained from the low-temperature separator (first vapor-liquid separator) 4 to the demethanizer 11 is used. A pressure reducing valve 14 may be provided in this line. Typically, the pressure of the condensate is reduced by the pressure reducing valve 14 to a pressure obtained by adding a pressure loss at the time of feeding, to the operating pressure of the feed stage of the demethanizer (in the case of propane recovery, the deethanizer), and a part of the condensate is vaporized into a vapor-liquid two-phase flow. In addition, the temperature decreases with this vaporization (in the case of Example 1 corresponding to Embodiment 1, the temperature decreases to −84.6° C.)
c) Step of Expanding a Part or all of the Gas Obtained from Step (a) by an Expander to Partially Condense the Gas, Followed by Vapor-Liquid Separation
In this step, an expander which expands and partially condenses part or all of the gas obtained from the low-temperature separator (first vapor-liquid separator) 4, in particular a turboexpander 5, is used. A turbo expander outlet separator 7 connected to the turbo expander 5 outlet is also used as a second vapor-liquid separator.
In the present example, all of the low-temperature separator 4 outlet gas (line 110) is sent to the turbo expander 5, and typically the pressure at the outlet of the turbo expander 5 is reduced to a pressure obtained by adding a pressure loss at the time of feeding, to the operating pressure of the feed stage of the demethanizer (in the case of propane recovery, the deethanizer). At this time, due to the effect of the isentropic expansion, the outlet gas of the turboexpander 5 becomes extremely low temperature (in the case of Example 1, −85.2° C.) and partially condenses (in the case of Example 1, 27.9 mol % is liquefied). It is also possible to recover the energy which the gas loses during expansion as the power of the compressor 6.
The gas partially condensed at the outlet of the turboexpander 5 is separated in a turboexpander outlet separator 7 (second vapor-liquid separator).
The turboexpander outlet separator 7 may be a vertical or horizontal vessel (a cylindrical vessel having end plates at both ends) and may have a mist eliminator therein to increase the separation efficiency of vapor and liquid.
d) Step of Feeding the Liquid Obtained from Step (c) to the Distillation Column after Using it as the “Another Refrigerant” in Step (a);
In this step, a line for supplying the liquid obtained from the turboexpander outlet separator 7 to the demethanizer (distillation column) 11 via the above-mentioned “another refrigerant flow path” is used (lines 104 and 102). In the present example, “another refrigerant flow path” for flowing the liquid obtained from the turbo expander outlet separator 7 is one of the refrigerant flow paths provided in the second feed gas cooler 3 located most downstream based on the flow direction of the feed gas among the heat exchangers used for cooling of the step (a). In Example 1, the liquid in line 104 is warmed to −39.0° C. by being used as “another refrigerant”, resulting in a vapor-liquid two-phase flow.
e) Step of Supplying a Part or all of the Gas Obtained from Step (c) to the Distillation Column
In this step, a line for supplying a part or all of the gas obtained from the turboexpander outlet separator (second vapor-liquid separator) 7 to the demethanizer (distillation column) 11 is used.
In this example, all of the gas obtained from the turboexpander outlet separator (second vapor-liquid separator) 7 is supplied to the demethanizer 11 (line 103).
The demethanizer 11 equips, for example, trays or packings inside the column, and separates the high volatile component and the low volatile component by a distillation operation. The pressure of the demethanizer is preferably as high as possible as long as a predetermined ethane recovery rate can be achieved in order to reduce the power required for the compression of the residual gas downstream, and is preferably 1.5 MPa or more and 3.5 MPa or less from this viewpoint, and more preferably 2.5 MPa or more and 3.5 MPa or less.
In this example, three types of fluids are fed to the demethanizer 11. The top of the column is fed with condensate separated by a low-temperature separator 4 as a reflux via a pressure reducing valve 14 (line 101), below the feed location the outlet gas of the turbo expander outlet separator 7 is fed (line 103), and further below the feed location the liquid separated in the turbo expander outlet separator 7 is fed after heat exchange with the feed gas in a second feed gas cooler 3 (line 102). In
A reboiler 12 is installed at the bottom of the demethanizer to volatilize methane in the bottom liquid of the column, and heat is applied so that the concentration of methane in the bottom liquid of the column becomes equal to or lower than a predetermined value.
f) Step of obtaining a residual gas from the top of the distillation column and a heavy fraction from the bottom of the distillation column. From the top of the demethanizer, a residual gas containing methane as a main component, from which components such as ethane and propane have been removed, is separated and utilized for heat exchange with the feed gas. Thereafter, if necessary, the residual gas is compressed to a predetermined pressure by a compressor 6 driven by the turbo compressor and a compressor (residual gas compressor) 13 driven by a motor or the like. From the bottom of the demethanizer 11, ethane, propane and heavy components are separated as NGLs (natural gas-liquid). The obtained NGL is separated into respective components, for example, in an NGL separation step which is further provided downstream.
As the feed gas, a natural gas containing methane and hydrocarbons having lower volatility than methane is preferred. The raw material feed gas may be a petroleum-associated gas or an off-gas from a refinery or petrochemical plant.
The higher the concentration of the hydrocarbons having lower volatility than methane in the feed gas, the greater the difference between the methane concentration in the inlet gas of the turboexpander 5 and the methane concentration in the outlet gas of the turboexpander outlet gas separator 7, and accordingly the effect of improving the reflux tends to be produced. Therefore, when the concentration of hydrocarbons having lower volatility than methane in the feed gas is 5 mol % or more and 50 mol % or less, further, when the concentration is 10 mol % or more and 50 mol % or less, the effect of the present invention is particularly remarkable.
Further, since the lower the ethane concentration in the residual gas means a higher ethane recovery rate, the ethane concentration in the residual gas is preferably as low as possible, preferably 5 mol % or less, and more preferably 1 mol % or less.
The NGL is composed of hydrocarbons having lower volatility than the liquefied and recovered methane, and is sent to an NGL fractionation facility which is further provided downstream, for example, and is separated into products such as ethane, propane, and butane. In such a case, methane in NGL is preferably low to such an extent that the criteria of the ethane product can be satisfied, and is preferably 2 mol % or less, more preferably 1 mol % or less.
In the case of the propane recovery process, the same principle as in the above example is used, and a deethanizer is used instead of the demethanizer 11, and a residual gas containing methane and ethane as main components is separated from the top of the deethanizer, and propane and heavy components are separated as NGL from the bottom of the deethanizer.
With respect to the second embodiment of the present invention, an example of an ethane recovery process will be described using the process flow diagram shown in
In the embodiment 1, in the step (c), the entire amount of the gas (line 110) obtained from the step (a), i.e. from the low-temperature separator 4, is supplied to the turboexpander 5. In the embodiment 2, line 110 is divided and only a portion of the gas of line 110 (line 110a) is sent to turboexpander 5 for step (c). The distribution ratio of line 110 is determined in view of the required ethane recovery rate (line 110a:line 110b=70:30 (molar ratio) in the case of example 2 corresponding to the embodiment 2). The pressure at the outlet of the turboexpander 5 is reduced to a pressure obtained by adding a pressure loss at the time of feeding, to the operating pressure of the feed stage of the demethanizer (in the case of propane recovery, the deethanizer). At this time, due to the effect of the isentropic expansion, the outlet gas of the turboexpander 5 becomes extremely low temperature (in Example 2, −86.4° C.) and partially condenses (in Example 2, 24.7% is liquefied). It is also possible to recover the energy which the gas loses during expansion as the power of the compressor 6.
The remainder (line 110b) of the gas in line 110 is cooled and totally condensed by heat exchange with the residual gas obtained from the top of the demethanizer in the condenser 10 (in the case of Example 2, cooling to −90.8° C.), the pressure of the totally condensed liquid is reduced by the pressure reducing valves 15 and the totally condensed liquid is supplied to the demethanizer (distillation column) 11 (line 105). The pressure of the totally condensed liquid is reduced by the pressure reducing valve 15 to a pressure obtained by adding a pressure loss at the time of feeding, to the operating pressure of the feed stage of the demethanizer (distillation column) 11. Also, the totally condensed liquid is partially vaporized by decompression, resulting in a vapor-liquid two-phase flow, and the temperature decreases with vaporization (in the case of example 2, −94.2° C.).
For this purpose, the following apparatus are used:
line 110a, which feeds a portion of the gas obtained from the low-temperature separator (first vapor-liquid separator) 4 to the turboexpander 5;
condenser 10, which cools the remainder (line 110b) of the gas obtained from the low-temperature separator (first vapor-liquid separator) 4 by heat exchange with the residual gas to cause total condensation;
pressure reducing valve 15 for depressurizing the totally condensed liquid in the condenser 10; and
line 105, which connects the outlet of the pressure reducing valve 15 to the demethanizer (distillation column) 11.
As the condenser 10, a heat exchanger for exchanging heat between the gas in the line 110b and the residual gas can be used. The condenser 10 can be disposed upstream of the feed gas coolers 1 and 3 and the feed gas chiller 2 with reference to the flow direction of the residual gas.
In this example, four types of fluids are fed to the demethanizer 11. At the top of the column, the liquid from line 105 is fed as a reflux, below the feed location the outlet gas of the turbo expander outlet separator 7 is fed (line 103), below the feed location the liquid from the low-temperature separator 4 is fed after being decompressed with the pressure reducing valve 14 (line 101), and below the feed location the liquid from the turbo expander outlet separator 7 is fed after heat exchange with the feed gas (line 102).
With respect to the process flow, the embodiment 2 may be the same as the embodiment 1 except that the above points. However, the conditions such as the temperature and the pressure can be appropriately changed in accordance with the difference in the process flow.
With respect to the third embodiment of the present invention, an example of an ethane recovery process will be described using the process flow diagram shown in
In the embodiment 1, in step (e), the entire amount of gas (line 103) obtained from step (c), i.e. from the turboexpander outlet separator 7, is supplied to the demethanizer 11. In the embodiment 3, line 103 is divided and only a portion of the gas of line 103 (line 103a) is supplied to step (e), i.e., fed to demethanizer 11. The distribution ratio of line 103 is determined in view of the required ethane recovery rate (line 103a:line 103b=63:37 (molar ratio) in the case of example 3 corresponding to the embodiment 3). The remainder of the gas in line 103 (line 103b) is compressed (in the case of Example 3, 6.00 MPa), cooled by heat-exchanging with the residual gas obtained from the top of the demethanizer to cause total condensation (in the case of Example 3, −94.2° C.), and the totally condensed liquid is depressurized and supplied to the demethanizer 11 (line 105). The pressure of the totally condensed liquid is reduced by the pressure reducing valve 15 to a pressure obtained by adding a pressure loss at the time of feeding, to the operating pressure of the feed stage of the demethanizer (distillation column) 11. Also, the totally condensed liquid is partially vaporized by decompression, resulting in a vapor-liquid two-phase flow, and the temperature decreases with vaporization (in the case of example 3, −97.2° C.).
For this purpose, the following apparatus are used:
line 103a, which feeds a portion of the gas obtained from the turboexpander outlet separator 7 (second vapor-liquid separator) to the demethanizer (distillation column) 11;
compressor 8, which compresses the remainder (line 103b) of the gas obtained from the turbo expander outlet separator 7 (second vapor-liquid separator);
condenser (reflux condenser) 10, which cool the gas compressed by the compressor 8 by heat exchange with the residual gas to cause total condensation;
pressure reducing valve 15 for depressurizing the totally condensed liquid in the condenser 10; and
line 105, which connects the outlet of the pressure reducing valve 15 to the demethanizer (distillation column) 11.
As the condenser 10, a heat exchanger for exchanging heat between the gas in the line 103b and the residual gas can be used. The condenser 10 can be disposed upstream of the feed gas coolers 1 and 3 and the feed gas chiller 2 with reference to the flow direction of the residual gas.
In this embodiment, after the gas compressed by the compressor 8 is cooled by a heat exchanger (reflux cooler) 9 using propane refrigerant, the gas is cooled by heat exchange with residual gas in a reflux condenser 10 and totally condensed. The reflux cooler 9 can be provided as required and is not required if cooling by the reflux condenser 10 is sufficient.
In this example, four types of fluids are fed to the demethanizer 11. At the top of the column, the liquid from line 105 is fed as reflux, below the feed location part of the outlet gas of the turboexpander outlet separator 7 is fed (line 103a), below the feed location the liquid from the low-temperature separator 4 is fed (line 101) after being decompressed with the pressure reducing valve 14, and below the feed location the liquid from the turboexpander outlet separator 7 is fed after heat exchange with the feed gas (line 102).
With respect to the process flow, Embodiment 3 may be the same as Embodiment 1 except that the above points. However, the conditions such as the temperature and the pressure can be appropriately changed in accordance with the difference in the process flow.
Hereinafter, the present invention will be described in more detail based on Examples, but the present invention is not limited thereto.
A process simulation was carried out on an example of a case where ethane recovery was carried out using a hydrocarbon separation apparatus having the configuration shown in
The feed gas is heat-exchanged in the first feed gas cooler 1 with the residual gas of −39.0° C. and the side stream F1 of the demethanizer 11 of −33.5° C. to be cooled to −24.6° C. Thereafter, the feed gas is cooled to −37.0° C. by propane refrigerant in the feed gas chiller 2, and cooled to −62.9° C. in the second feed gas cooler 3 by heat exchange with the residual gas of −84.6° C., the side stream F3 of the demethanizer 11 of −76.1° C., and the condensate (line 104) of the turboexpander outlet separator 7 of −85.2° C. Here the first feed gas cooler 1 and the second feed gas cooler 3 are plate fin heat exchangers respectively, the feed gas chiller 2 is a shell and tube heat exchanger of kettle type.
Next, the feed gas is separated in the low-temperature separator 4. The low-temperature separator 4 is a vertical vessel having a mist eliminator therein (cylindrical container having a mirror plate at both ends).
The entire amount of gas at the outlet of the low-temperature separator 4 is sent to the turbo expander 5 and depressurized to 3.47 MPa. The outlet gas is cooled down to −85.2° C. by the effect of isentropic expansion and provides 529 kW of power to the compressor 6 driven by the expander. The gas at the outlet of the turboexpander 5 is separated in the turboexpander outlet separator 7. The turbo expander outlet separator 7 is a vertical vessel having a mist eliminator therein (cylindrical container having a mirror plate at both ends).
The condensate at −85.2° C. (line 104) separated by the turboexpander outlet separator 7 is fed to the demethanizer 11 (line 102) after the temperature is raised to −39.0° C. by providing cold heat to the feed gas in the second feed gas cooler 3.
The demethanizer 11 has 40 trays installed therein, and the gas at the outlet of the turboexpander outlet separator 7 is fed to the tray of the third stage from the top of the column (line 103). The liquid separated in the turboexpander outlet separator 7 passes through the second feed gas cooler 3 and is fed to the tenth stage from the top of the column (line 102). In addition, the liquid separated in the low-temperature separator 4 is decompressed to 3.29 MPa with the pressure reducing valve 14, and then fed to the first stage from the top of the column (line 101) as reflux.
Demethanizer 11 is operated under the conditions of 3.27 MPa and −84.6° C. at the top of the column, and is operated under the conditions of 3.32 MPa and 39.8° C. at the bottom of the column. The temperature of the bottom of the column is determined by the equilibrium temperature at which the methane concentration in the NGL is less than or equal to 1 mol %, and 3.60 MW of heat is added from reboiler 12 in order to operate at that temperature. The residual gas separated from the top of the demethanizer 11 and the composition of NGL separated from the bottom of the column are as shown in Table 2. The flow rates are 12,553 kg-moles/hr (103 moles/hr) for residual gas and 1,147 kg-moles/hr (103 moles/hr) for NGL. Note that “NC4” represents normal butane, and “IC4” represents isobutane.
Of the ethane in the feed gas, 76.7% is recovered as NGL.
The residual gas leaving the top of the demethanizer 11 is heat-exchanged with the feed gas to reach 15.1° C. at the outlet of the first feed gas cooler 1. Thereafter, the residual gas is compressed to 3.25 MPa by the compressor 6 driven by the turbo expander, and is compressed to 3.77 MPa by the residual gas compressor 13. At this time, the required power of the residual gas compressor 13 is 1031 kW.
A process simulation was carried out on an example of a case where ethane recovery was carried out using a hydrocarbon separation apparatus having the configuration shown in
In Example 1, from a vapor-liquid separated condensate (line 104) by the turbo-expansion outlet separator 7, cold heat is recovered in the second feed gas cooler 3, and the condensate becomes a vapor-liquid two-phase flow (line 102). At this time, the methane fraction, which is a low boiling point component, is mainly vaporized, so that the concentration of methane in the vapor-liquid two phase stream of line 102 decreases. The higher the methane concentration in the reflux liquid of the demethanizer 11, the higher the reflux effect is, and hence in Example 1, the condensate of the low-temperature separator 4 (the methane concentration is higher than the vapor-liquid two-phase flow of line 102) is fed as a reflux liquid to the first stage of the demethanizer.
On the other hand, in the configuration shown in
In the demethanizer 11, the gas from the outlet of the turboexpander outlet separator 7 is fed to the tray of the fourth stage from the top of the column (line 103). The liquid separated in the low-temperature separator 4 is fed to the 14th stage from the top of the column after being decompressed to 2.82 MPa by the pressure reducing valve 14 (line 101).
With respect to the process flow, Comparative Example 1 is the same as Example 1 except for the above points.
In Table 3, the “Refrigeration Load” is the thermal load of the propane refrigeration system in the feed gas chiller 2. The lowering of the refrigeration load means the lowering of the propane refrigeration equipment capacity, and it is effective in lowering of the energy consumed in the propane refrigeration equipment and in lowering of the equipment cost of the propane refrigeration.
The “Reboiler Heat Load” is the heat load of the reboiler 12 at the bottom of the demethanizer column. The reduction means a reduction in the energy required for distillation, and there is an effect of a reduction in the cost of utilities supplied from the outside. Power of “Refrigeration Compressor” is a power consumed by the compressor in the propane refrigeration system. Power of “Residual Gas Compressor” is a power consumed by the residual gas compressor 13.
As apparent from Table 3, Example 1 can reduce the total compressor power and reboiler heat load even though the ethane recovery rate is about the same as that in the case where ethane recovery is performed in the configuration of Comparative Example 1.
A process simulation was carried out on an example of a case where ethane recovery was carried out using a hydrocarbon separation apparatus having the configuration shown in
The feed gas is heat-exchanged in the first feed gas cooler 1 with the residual gas of −39.0° C. and the side stream F1 of the demethanizer 11 of −39.3° C. to be cooled to −23.7° C. Thereafter, the feed gas is cooled to −37.0° C. by propane refrigerant in the feed gas chiller 2, and cooled to −60.4° C. in the second feed gas cooler 3 by heat exchange with the residual gas of −76.6° C., the side stream F3 of the demethanizer 11 of −77.7° C., and the condensate (line 104) of the turboexpander outlet separator 7 of −86.4° C. Here the first feed gas cooler 1 and the second feed gas cooler 3 are plate fin heat exchangers respectively, and the feed gas chiller 2 is a shell and tube heat exchanger of kettle type.
Next, the feed gas is separated in the low-temperature separator 4. The low-temperature separator 4 is a vertical vessel having a mist eliminator therein (cylindrical container having a mirror plate at both ends).
70 mole % of the outlet gas of the low-temperature separator 4 is sent to the turboexpander 5 (line 110a) and depressurized to 3.20 MPa. The outlet gas is cooled to −86.4° C. by the effect of isentropic expansion, and accordingly a part of the gas condenses into a vapor-liquid two-phase flow, and thereby 723 kW of power is provided to the compressor 6 driven by the expander. The gas (partially condensed) at the outlet of the turbo expander 5 is separated in the turbo expander outlet separator 7. The turbo expander outlet separator 7 is a vertical vessel having a mist eliminator therein (cylindrical container having a mirror plate at both ends).
The remaining 30 mole % of the outlet gas of the low-temperature separator 4 is sent to a condenser (reflux condenser) 10 (line 110b), heat-exchanged with the residual gas at the top of the demethanizer 11, and is cooled to −90.8° C. to be totally condensed. The pressure of the condensate is reduced to 3.00 MPa with the pressure reducing valve 15, and a part of the condensate is vaporized into a vapor-liquid two-phase flow, and the temperature is lowered to −94.2° C. as the condensate is vaporized. Thereafter, the two-phase flow is fed as a reflux liquid to the first stage from the top (line 105). Here, the reflux condenser 10 is a plate fin heat exchanger.
The demethanizer 11 has 40 trays installed therein, and the gas at the outlet of the turboexpander outlet separator 7 is fed to the tray of the fourth stage from the top of the column (line 103). Further, the condensate of −86.4° C. which is separated in the turbo expander outlet separator 7 (line 104) is temperature-elevated to −39.0° C. by the cold heat recovery in the second feed gas cooler 3 and thereby a part of the condensate is vaporized to become a vapor-liquid two-phase flow, and thereafter fed to the 20th stage from the top of the column (line 102). Further, the liquid separated in the low-temperature separator 4 is decompressed to 3.20 MPa with the pressure reducing valve 14 and thereby a part thereof is vaporized to become a vapor-liquid two-phase flow, and the temperature is lowered to −84.2° C. as the liquid is vaporized. Thereafter, the two-phase flow is fed to the 14th stage from the top of the column (line 101).
Demethanizer 11 is operated under the conditions of 3.00 MPa and −92.8° C. at the top of the column, and is operated under the conditions of 3.05 MPa and 31.5° C. at the bottom of the column. The temperature of the bottom of the column is determined by the equilibrium temperature at which the methane concentration in the NGL is less than or equal to 1 mol %, and 3.65 MW of heat is added from reboiler 12 in order to operate at that temperature. The residual gas separated from the top of the demethanizer 11 and the composition of NGL separated from the bottom of the column are as shown in Table 4. The flow rates are 12,444 kg-moles/hr (103 moles/hr) for residual gas and 1,256 kg-moles/hr (103 moles/hr) for NGL.
Of the ethane in the feed gas, 88.7% is recovered as NGL.
The residual gas leaving the top of the demethanizer 11 is heat-exchanged with the feed gas to reach 15.1° C. at the outlet of the first feed gas cooler 1. Thereafter, the residual gas is compressed to 3.17 MPa by the compressor 6 driven by the turbo expander, and is compressed to 3.77 MPa by the residual gas compressor 13. At this time, the required power of the residual gas compressor 13 is 1859 kW.
A process simulation was carried out on an example of a case where ethane recovery was carried out using a hydrocarbon separation apparatus having the configuration shown in
In the configuration shown in
In Comparative Example 2, cold heat recovery using the condensate of the turboexpander outlet separator 7 is not performed, and accordingly the temperature of the stream flowing into the low-temperature separator 4 is −52.0° C., which is 8.4° C. higher than in Example 2. Accordingly, the methane concentration in the gas (line 110) separated in the low-temperature separator 4 becomes lower as compared with Example 2, and eventually leads to a decrease in the reflux effect in the distillation column.
In the demethanizer 11, the liquid from line 105 is fed to the first stage from the top of the column as a reflux liquid. The gas at the outlet of the turbo expander outlet separator 7 is fed to the tray of the fourth stage from the top of the column (line 103). The liquid separated in the turboexpander outlet separator 7 is fed to the 14th stage from the top of the column (line 102). Furthermore, the liquid separated in the low-temperature separator 4 is fed to the 20th stage from the top of the column after being depressurized to 2.83 MPa with the pressure reducing valve 14 (line 101).
With respect to the process flow, Comparative Example 2 is the same as in Example 2 except for the above points.
As is apparent from Table 5, Example 2 can obtain a higher ethane recovery rate and can further reduce the total compressor power and reboiler heat load as compared with the case where ethane recovery is performed in the configuration of Comparative Example 2.
A process simulation was carried out on an example of a case where ethane recovery was carried out using a hydrocarbon separation apparatus having the configuration shown in
The feed gas is heat-exchanged in the first feed gas cooler 1 with the residual gas of −39.0° C. and the side stream F1 of the demethanizer 11 of −35.3° C. to be cooled to −22.6° C. Thereafter, the feed gas is cooled to −37.0° C. by propane refrigerant in the feed gas chiller 2, and cooled to −59.0° C. in the second feed gas cooler 3 by heat exchange with the residual gas of −68.0° C., the side stream F3 of the demethanizer 11 of −74.3° C., and the condensate (line 104) of the turboexpander outlet separator 7 of −86.8° C. Here the first feed gas cooler 1 and the second feed gas cooler 3 are plate fin heat exchangers respectively, and the feed gas chiller 2 is a shell and tube heat exchanger of kettle type.
Next, the feed gas is vapor-liquid separated in the low-temperature separator 4. The low-temperature separator 4 is a vertical vessel having a mist eliminator therein (cylindrical container having a mirror plate at both ends).
The entire amount of gas at the outlet of the low-temperature separator 4 is sent to the turbo expander 5 and reduced to 3.07 MPa. The outlet gas is cooled down to −86.8° C. by the effect of isentropic expansion and provides 1259 kW of power to the compressor 6 driven by the expander. The gas at the outlet of the turboexpander 5 is separated in the turboexpander outlet separator 7. The turbo expander outlet separator 7 is a vertical vessel having a mist eliminator therein (cylindrical container having a mirror plate at both ends).
37 mol % of the outlet gas (line 103) of the turbo expander outlet separator 7 is pressurized by a compressor (low temperature compressor) 8 to 6.00 MPa, then cooled to −94.2° C. by a heat exchanger (reflux cooler) 9 by propane refrigeration and a condenser (reflux condenser) 10 for heat exchange with the residual gas on the top of the demethanizer 11 to be totally condensed. The pressure of the obtained condensate is reduced to 2.87 MPa with the pressure reducing valve 15, and a part of the condensate is vaporized into a vapor-liquid two-phase flow, and the temperature is lowered to −97.2° C. as the condensate is vaporized. Thereafter, the two-phase flow is fed as a reflux liquid to the first stage from the top (line 105). Here, the reflux cooler 9 is a shell and tube heat exchanger of kettle type, the reflux condenser 10 is a plate fin heat exchanger. When the outlet temperature of the reflux condenser 10 can be lowered to a temperature at which a predetermined ethane recovery rate can be achieved by only exchanging heat with the residual gas, the reflux cooler 9 may not be installed in order to reduce the load of propane refrigeration.
The demethanizer 11 has 40 trays installed therein, and a part of the gas at the outlet of the turboexpander outlet separator 7 is fed to the tray of the ninth stage from the top of the column (line 103a). Further, the condensate of −86.8° C. which is separated by the turbo expander outlet separator 7 (line 104) is temperature-elevated to −39.0° C. by the cold heat recovery in the second feed gas cooler 3 and thereby a part of the condensate is vaporized to become a vapor-liquid two-phase flow, and thereafter fed to the 18th stage from the top of the column (line 102). Further, the liquid separated in the low-temperature separator 4 is decompressed to 2.89 MPa with the pressure reducing valve 14, and thereby a part thereof is vaporized to become a vapor-liquid two-phase flow, and the temperature is lowered to −83.7° C. as the liquid is vaporized. Thereafter, the two-phase flow is fed to the 15th stage from the top of the column (line 101).
Demethanizer 11 is operated under the conditions of 2.87 MPa and −96.2° C. at the top of the column, and is operated under the condition of 2.92 MPa and 27.5° C. at the bottom of the column. The temperature of the bottom of the column is determined by the equilibrium temperature at which the methane concentration in the NGL is less than or equal to 1 mole %, and 3.35 MW of heat is added from reboiler 12 in order to operate at that temperature. The residual gas separated from the top of the demethanizer 11 and the composition of NGL separated from the bottom of the column are as shown in Table 6. The flow rates are 12,388 kg-moles/hr (103 moles/hr) for residual gas and 1,312 kg-moles/hr (103 moles/hr) for NGL.
Of the ethane in the feed gas, 95.5% is recovered as NGL.
The residual gas leaving the top of the demethanizer 11 is heat-exchanged with the feed gas to reach 15.1° C. at the outlet of the first feed gas cooler 1. Thereafter, the residual gas is compressed to 3.19 MPa in the compressor 6 driven the turbo expander, and is compressed to 3.77 MPa by the residual gas compressor 13. At this time, the required power of the residual gas compressor 13 is 1824 kW.
A process simulation was carried out on an example of a case where ethane recovery was carried out using a hydrocarbon separation apparatus having the configuration shown in
In the configuration shown in
In Comparative Example 3, cold heat recovery using the condensate of the turboexpander outlet separator 7 is not performed, and accordingly the temperature of the stream flowing into the low-temperature separator 4 is −44.1° C., which is 14.9° C. higher than in Example 3. Accordingly, the methane concentration in the gas (line 110) separated in the low-temperature separator 4 becomes lower as compared with Example 3, and eventually leads to a decrease in the reflux effect in the distillation column.
In the demethanizer 11, the liquid from line 105 is fed to the first stage from the top of the column as reflux. A part of the gas at the outlet of the turboexpander outlet separator 7 is fed to the tray of the ninth stage from the top of the column (line 103a). The liquid separated in the turboexpander outlet separator 7 is fed to the 12th stage from the top of the column (line 102). Furthermore, the liquid separated in the low-temperature separator 4 is decompressed to 2.82 MPa with the pressure reducing valve 14, and thereby a part thereof is vaporized to become a vapor-liquid two-phase flow, and the temperature is lowered to −64.0° C. as the liquid is vaporized. Thereafter, the two-phase flow is fed to the 15th stage from the top of the column (line 101).
With respect to the process flow, Comparative Example 3 is the same as Example 3 except for the above points.
As is apparent from Table 7, Example 3 can obtain a higher ethane recovery rate and can further reduce the total compressor power and reboiler heat load as compared with the case where ethane recovery is performed in the configuration of Comparative Example 3.
| Number | Date | Country | Kind |
|---|---|---|---|
| 2019-200772 | Nov 2019 | JP | national |
