process selection of natural gas recovery unit
TRANSCRIPT
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Process Selection of Natural Gas Recovery Unit
Mrs. Lakshmi Venkatesh Mr. Umesh Yeole Asst General Manager Senior Engineer Process
Petrofac Engineering India Limited, 7th Floor, Ventura, Central Avenue Hiranandani Business Park, Powai,
Mumbai 400076
ABSTRACT:
In this paper, a review of process selection and configuration of the NGL recovery unit for different projects is presented. The recovery of light hydrocarbon liquids from natural gas streams can range from simple hydrocarbon (HC) dew point control to avoid retrograde condensation in the export pipeline to deep ethane/propane/LPG extraction. Typical LPG recovery processes includes lean oil absorption (in older plants), gas expansion refrigeration, using JT valves or turbo expanders, and mechanical refrigeration followed by distillation. The process selection depends on the gas composition, inlet pressure to the gas plant, the products to be recovered and the extent of recovery that is desired. Optimum extent of recovery depends on economics, i.e. increased product value compared to additional CAPEX and OPEX.
This paper discusses typical configurations of the gas plant using the Turboexpander (TE) based processes. Differential pressure across the turbo-expander provides the driving force for NGL recovery. The effect of upstream gas pressure, mechanical refrigeration on product recovery and choice of the process based on the desired product recovery is discussed. The impact of impurities like carbon dioxide on the recoveries achievable due to the phenomenon of carbon dioxide freeze out at low operating temperature is also highlighted.
Three cases are presented to illustrate the principles using results from simulations:
Case 1: Considers a conventional turbo expander for the conditioning of residue gas so as to meet the export specifications.
Case 2: Considers over head recycle (OHR) process to maximize LPG recovery with no consideration for ethane recovery.
Case 3: Considers a gas plant with feed gas at a low pressure. Gas Sub-cooled Process (GSP) together with feed gas compression and mechanical refrigeration is used to achieve the required ethane and LPG recovery.
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INTRODUCTION
The recovery of hydrocarbon liquids from natural gas streams can range from simple HC dew point control to deep ethane extraction. The process selection, complexity, and cost of the processing facility depends on the gas composition, inlet pressure to the gas plant, recovered product specifications and the extent of recovery that is desired.
The term NGL (natural gas liquids) is a general term which applies to liquids recovered from natural gas and as such refers to ethane and heavier products. The term LPG (liquefied petroleum gas) describes hydrocarbon mixtures in which the main components are propane, iso and normal butane, propene and butenes.
Gas processing plants are designed to satisfy one of the following requirements listed in the order of the severity of processing
• Simple HC dew point control of the sales gas
• LPG recovery with Ethane rejection
• Ethane recovery
The processing requirement is market and geographically dependent.
This paper discusses typical configurations of the gas plant using the Turboexpander (TE) based process. Differential pressure across the turbo-expander provides the driving force for separation of heavier hydrocarbons.
The selection of a process for a specific application is presented by means of three case studies. The effect of upstream gas pressure, mechanical refrigeration and of impurities like carbon dioxide on the recoveries achievable is also highlighted.
TURBO EXPANDER BASED PROCESS
In a turbo expander based process, the feed gas pressure is used to produce required refrigeration by gas expansion across a turbo expander. The turbo expander recovers useful work from gas expansion. Typically the expander is linked to a centrifugal compressor to recompress the residue gas from the process. Because the expansion is near isentropic, the turbo expander lowers the gas temperature significantly more than expansion across a Joule-Thomson (JT) valve.
The turbo expander process has been applied to a wide range of process conditions and varies in complexity based on the severity of the hydrocarbon recovery desired.
Conventional Expander Process
A simplified flow scheme of Conventional Expander Process is given in Figure 1.
In this process, feed gas is first cooled by rejecting heat to process streams, and sent to the low temperature separator. Vapors from the separator are routed to a turbo expander where the low temperature required for the dropout of hydrocarbons is obtained as a result of expansion refrigeration. This is followed by distillation using a demethaniser or deethaniser
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column, as the case may be, for light ends rejection. The raw NGL stream consisting of ethane, LPG and heavier components is sent for downstream fractionation and product treatment.
This process is typically used when the intent is to meet Sales gas HC dew point constraints and achieve a nominal recovery of the LPG components.
Conventional Expander Process with refluxed Deethaniser
This process is a modification of the Conventional Expander Process. A simplified flow scheme of the process is shown in Figure 2. In this case, overhead vapors from deethaniser are partially condensed and used as reflux. The reflux stream rectifies vapor stream leaving the column by absorbing heavier components. With this process higher recovery of LPG components is obtained.
Conventional Expander Process with refluxed Deethaniser
This process is a modification of the Conventional Expander Process. A simplified flow scheme of the process is shown in Figure 2. In this case, overhead vapors from deethaniser are partially condensed and used as reflux. The reflux stream rectifies vapor stream leaving the column by absorbing heavier components. With this process higher recovery of LPG components is obtained.
Overhead Recycle (OHR) Process
A simplified flow scheme of Overhead Recycle (OHR) is given in Figure 3.
In this process, an absorber column is provided upstream of deethanizer column. Vapor stream from Deethanizer is condensed and used as reflux in absorber to rectify vapor leaving the expander.
This process is preferred when high recovery of LPG components is desired but the ethane is required to be rejected along with the overhead gas.
Gas Sub-cooled Process (GSP)
A simplified flow scheme of Gas Sub-cooled Process (GSP) is given in Figure 4.
In this process a part of the vapor from the low temperature separator is condensed and sub-cooled using demethaniser overhead vapors. This stream is then flashed to the tower operating pressure and used as reflux to the Demethaniser. Higher recovery of ethane+ components is obtained due to the colder reflux provided at the column top that serves to improve the recovery of ethane in the column bottoms.
Numerous variations of the processes described above are used in modern NGL recovery units that are licensed processes.
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The modifications usually provide higher recovery, better recovery efficiency, and better tolerance for CO2 in feed gas, improved operational flexibility in terms of desired product specifications and reduced lifecycle cost of the project.
FEED GAS CONTAMINANTS & PRETREATMENT
The feed gas to may contain a number of contaminants that will need to be removed prior to the gas processing.
Hydrogen Sulphide: Sour gas containing Hydrogen Sulphide must have an Acid gas removal unit for the removal of the Hydrogen Sulphide.
Carbon dioxide: Some quantity of Carbon dioxide if present is acceptable however it will impact the hydrocarbon recovery in two ways
• Sufficient ethane and heavier components must remain in the gas so that the heating value specification of the gas is achieved
• The phenomenon of CO2 freezing temperature must be considered in the top section of the demethaniser column and the column must be operated at a sufficiently high temperature to avoid the freeze out.
If the recoveries are limited by high CO2 in the feed gas, an upstream acid gas removal unit will need to be considered.
Nitrogen: This has same impact as CO2 on the heating value of the gas. Presence of large amounts of inert gases such as Nitrogen will affect the ability to condense the reflux stream in the demethaniser, reducing ethane recovery and overall efficiency of the process.
Mercury: Brazed Aluminium exchangers are often used in NGL recovery units in the cold section to improve the exchanger approach temperature and result is a better recovery of the cold in the process. Conventional shell and tube exchangers will limit the approach temperatures to approximately 5 Deg C. Lower approach temperatures of 2-3 Deg C are achieved with Brazed Aluminum exchangers which also have a smaller foot print than shell and tube exchangers. Mercury removal unit will be required for systems using aluminum plate-fin exchangers.
Water: Water removal is mandatory upstream of the NGL unit in order to avoid water freeze out and/or hydrate formation in the cold sections of the plant. For plants with LPG and ethane recovery, molecular sieve based dehydration units are employed and water content reduced to below 1 ppm.
IMPORTANT DESIGN PARAMETERS
Feed Gas Richness:
Richness of gas is defined in terms of cubic meters of recoverable hydrocarbons per thousand cubic meter of feed gas. Rich gas with a greater quantity of liquefiable hydrocarbons produces a greater quantity of products and hence greater revenue for gas processing facility. This also results in larger refrigeration duties and larger heat transfer area for exchangers.
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Leaner gases require more severe processing conditions such as higher differential pressure in turb-expander, lower temperatures to achieve high recovery efficiencies.
Feed Gas Pressure:
Differential pressure across the turbo-expander provides the driving force for NGL recovery. Typically an inlet pressure of above 50 barg is desired for most expander processes.
Inlet pressure above 80 barg will be an advantage for Propane+ recovery as much of the propane will be condensed in the low temperature separator.
For low inlet gas pressure, use of mechanical refrigeration or feed gas compression may be required depending on gas composition and the desired recovery.
Carbon Dioxide Content:
For propane-plus recovery, Carbon Dioxide (CO2) freezing will not be a concern. For ethane recovery applications, CO2 in the feed gas will normally split between the recovered ethane product and the residue gas. Higher CO2 content will affect specifications for both products. CO2 can also freeze in the low temperature sections of the process.
The process should be operated such that freezing of CO2 in low temperature sections is avoided. Increasing heavier content in liquid in demethaniser overhead sections will reduce the possibility of CO2 freezing. For designs that make use of feed gas to generate reflux for demethaniser, increasing the reflux flow can provide better operating margin for CO2 freezing.
Feed gas or product treatment for CO2 removal will be required if CO2 content in products is not acceptable.
Desired Product Recovery:
Recovered liquid hydrocarbons can be used as feed for downstream petrochemical processes. Alternatively NGL can also be used as a fuel, as propane, butane or LPG. Market value of these recovered liquid products might be much higher than the alternative value obtained for the same hydrocarbons lost in residue gas for use as sales gas. The optimum NGL recovery level is unique for each plant and will depend on the alternative NGL value, based on demand and plant location, i.e. proximity to markets.
A summary of some of the important design considerations while selecting a process and designing an NGL Recovery unit is given in Table 1.
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Table 1 Design considerations in process selection for NGL recovery unit Parameter Specification Design impact for turbo-expander based
processes Feed Gas Richness
Rich Gas Larger refrigeration duties, larger heat exchange surfaces and higher project lifecycle cost for a given recovery but more product to sell.
Lean Gas Requires more severe processing conditions (lower temperatures) to achieve high recovery efficiencies
CO2 in feed gas
> 2.0 mol% • CO2 Freezing in low temperature sections. Recovery efficiency will be limited by the CO2 freeze out temperatures.
• Impact on liquid product specification • Feed gas/ product treatment for CO2 removal.
H2S content in feed gas
>500 ppm mol
• Impact on Gas/Liquid product specifications to be checked. Typically, LPG products are required to pass Copper Strip corrosion test.
• Feed gas/ product treatment for H2S removal. Inert content in feed gas
> 2.0 mol % • Reduced ethane recovery due to reduced reflux flow
• Inert gas will reduce sales gas heating value necessitating lower hydrocarbon recovery to meet sales gas heating value.
Feed gas pressure
< 30 barg • Feed gas compression required. • Mechanical refrigeration likely to be required
60-85 barg No concern if low sales gas pressure specification. Mercury in feed gas
> 10 ng/Nm3 • Mercury removal unit/ Mercury guard bed required.
• Use of Brazed aluminum exchangers to be
reviewed Water content in feed gas
Saturated feed gas
Gas dehydration unit required
Residue gas Heating value
35.4-42.8 MJ/m3
Decides the amount of ethane and heavier components in residue gas. May limit the extent of ethane recovery.
Residue Gas BL pressure specification
>30 barg Residue gas compression required to achieve high BL pressure.
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Desired ethane recovery
Ethane rejection in sales gas
OHR process more suitable
Low Ethane recovery ~ 60-85%
Conventional turboexpander processes
>90% Increased complexity of process design, GSP process or its modifications most suitable
Desired LPG recovery
80-95% No concern if only LPG recovery is desired, any of the processes could be suitable
>95% Increased complexity of process design
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CASE STUDY 1 : SALES GAS DEW POINT CONTROL
This was a 200 MMSCFD gas plant with feed gas inlet pressure of 65 barg with feed composition as shown in Table 2. The specifications for the Sales gas were to achieve a cricondentherm of -10 Deg C and Heating Value between 35,400 and 45,000 kJ/m3 supplied at a pressure of 75 barg. Table 2 Gas Feed Composition
Component Mole % in feed gas
Nitrogen 2 Methane 74 Ethane 11 Propane 9 Isobutane 1 N-Butane 2 C5 1 Propane and butane as products were required with the following specifications. Quality-Propane w/w Quality-Butane w/w C2 and lighter <2% C3 and lighter <1% C3 >95% C4 >98% C4 <5% C5 and heavier <1%
Process Selection
Important consideration for process selection were • Main objective was to meet the Sales gas specifications • Ethane was to be rejected to Sales gas • No concern for recovery level of propane and butane • Feed gas was treated for CO2 removal and moisture removal at the upstream of gas
plant • Feed gas pressure of 65 barg was ideal for a turbo-expander based process • Sales gas compression was necessary to meet battery limit pressure of 75 barg
Two processing schemes were considered for the NGL recovery unit: • Scheme 1 : conventional expander with a de-ethanizer and • Scheme 2 : conventional expander with a refluxed de-ethaniser
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Process flow scheme were similar to the schematic shown in Figure 1 and Figure 2 respectively.
In scheme 1, reflux is not available for deethaniser column, so vapors from the turbo-expander are not rectified. This lowers the recovery of Propane and Butane.
In scheme 2, the overhead vapors from deethaniser are partially condensed and used as reflux. The reflux stream rectifies the vapor stream leaving the column by absorbing heavier components. With this process relatively higher recovery of Propane and Butane is obtained.
Results Summary
A comparison of achievable recovery is presented in Table 3. Scheme 2 provides better recovery efficiency with marginal increase in capital and operating cost.
Table 3 : Case 1 Result Summary
Process
Conventional Expander Process
Conventional Expander
with Refluxed
Deethaniser
Gas Feed Rate, MMSCFD 200 200
NGL unit inlet Pressure, barg 60 60
Sales gas cricondentherm, deg C -32.6 -40.16
Sales Gas heating Value, kJ/m3 42180 41580
Propane Quality, w/w 97.8 97.8
Propane Recovery, mol % 74.6 80.7
Butane Quality, w/w 98.4 99.2
Butane Recovery, mol% 93.2 97.7
Deethaniser Pressure, barg 13 13
Deethaniser reboiler Duty, kW 3974 3264
Sales Gas Compression power, kW 11730
11970
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Optimization of Recompression power
Operating pressure of deethaniser column is an important parameter for NGL recovery processes. Lower column operating pressure is desired from separation viewpoint as it increases differential pressure across turbo expander resulting in lower temperatures in column overhead section. Unfortunately, lower column pressure increases the recompression power requirement.
At higher column operating pressure, operating temperature is also higher. This lowers the temperature differential available in the heat exchangers and hence the heat integration is less efficient. As a result, column reboiler duty increases.
Effect of deethaniser operating pressure on Sales gas compressor power requirement and propane recovery is presented in Figure 5. Conclusions:
- Refluxed deethaniser process provides improved hydrocarbon recovery compared to conventional expander process.
- Process parameters were optimized to minimize energy consumption and to improve the recovery efficiency.
- Desired Sales gas specifications were met. Over 80 mol% propane recovery and 97 mol % butane recovery was achievable.
- Final process selection should be based on detailed cost benefit analysis considering CAPEX/OPEX for the facility and the value of the products.
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CASE STUDY 2: HIGH LPG RECOVERY, ETHANE REJECTION
This was a 500 MMSCFD gas plant with feed gas available at 74 barg at gas plant inlet. Three different feed cases were identified based on the various stages of the reservoir development. For the normal and lean case the plant was required to handle 500 MMSCFD of feed gas. In the rich case, the feed gas was about 220 MMSCFD.
Table 4 Feed Gas Composition Design Case Lean Normal Rich Flow rate, MMSCFD 500 500 220 Composition % mol Nitrogen 1 1 1. CO2 2.0 1.5 2.0 Methane 84.0 75.0 62.0 Ethane 7.0 16.0 18.0 Propane 2.0 4.0 10.0 i-Butane 1.0 0.5 3.0 n-Butane 1.0 1.5 2.0 i-Pentane 1.0 0.5 1.0 n-Pentane 0.0 0.0 1.0
LPG recovery of minimum 90% (mol) was required to be achieved with the following specifications of the LPG product.
Ethane 2.6 mol% max
Iso & Normal Pentane 0.4 mol% max
A residue gas compressor was used to recompress the gas upto 70 barg. This was followed by a gas reinjection compressor. Residue gas was used for reinjection back into the wells at 340 barg pressure.
Process Selection
Important consideration for process selection were • Main objective was to recover min 90% (mol) of LPG components present in feed gas • Ethane was to be rejected to residue gas • No concern for heating value of residue gas since it is to be used for reinjection • A molecular sieve dehydration unit at the upstream of gas plant was used to remove
moisture present in feed gas. The pressure drop in the initial section of the plant including dehydration and associated frictional loss was estimated to be about 5 bar
• Feed gas pressure of 69 barg was ideal for a turbo-expander based process
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Two processing schemes were considered for the NGL recovery unit: • Scheme 1 : conventional expander scheme with a refluxed de-ethaniser • Scheme 2 : Overhead recycle (OHR) process
Process flow schemes for NGL recovery unit were similar to the schematic shown in Figure 2 and Figure 3 respectively.
Results Summary
A comparison of achievable recovery is presented in Table 5. As seen from the results, desired LPG recovery of 90% can be obtained with OHR process. Therefore, OHR process was selected as a viable option for NGL recovery unit.
Table 5: Case 2 Results Summary
Design Case Normal Normal Lean
Rich (without Recycle)
Rich (with recycle)
Process
Conventional Expander
with Refluxed
Deethaniser
Overhead Recycle Process
Overhead Recycle Process
Overhead Recycle Process
Overhead Recycle Process
Gas Feed Rate, MMSCFD 500 500 500 220 500 NGL Unit Inlet Pressure, barg 68 68 68 68 68 LPG Quality, mol% 97.8 97.7 97.7 97.7 97.7 LPG Recovery, % 69 91.7 94.7 64.1 91.2 LPG product Rate Sm3/day 2333
3101 2191 2436 3422
Deethaniser pressure, barg 14 21.6 21.6 21.6 21.6
Absorber overhead temperature, deg C NA -58.3 -75.8 -29.3 -47.3
Deethaniser reboiler duty, kW 12620 9688 11243 5736 10339
Gas Compression power, kW 19923 13072 14285 9670 13424
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As seen from Table 5, for Rich Case (without recycle), only 64% LPG recovery was achieved. However, in this case sales gas flow rate was low and excess compression capacity was available. A fraction of the residue gas was recycled back to the inlet of NGL recovery unit and mixed with heavier feed gas. As a result molecular weight of the mixed gas was lowered. This resulted in lower temperature when the gas was expanded through the turbo expander. With this arrangement desired LPG recovery of 90% was achieved for Rich case (with recycle).
Absorber packing height has significant effect on the recovery of LPG components. As seen from Figure 6, as the packing height is increased, higher recovery of LPG is obtained.
Effect of Absorber operating pressure on compression power requirement and Propane recovery is presented in Figure 7. Column operating pressure of 20 barg was selected as the LPG recovery drops below 90% at higher pressures.
Effect of feed gas richness
As seen from Table 5 Rich gas with a greater quantity of liquefiable hydrocarbons produces a greater quantity of product and hence greater revenue for gas processing facility.
For rich gases, the operating temperature is higher, reducing the temperature differential available in heat exchangers. As a result the heat integration is less effective as is seen in Figure 8 where lower heat is recovered from the cold streams in the deethaniser. Figure 8 also demonstrates the effect of feed gas richness (i.e. higher LPG product rate) on the reboiler duty.
Lean feed gas requires lower temperatures to achieve high recovery efficiencies. However heat integration is more effective as column operating temperature is lower.
Conclusions:
‐ Desired LPG recovery was achieved with OHR process
‐ Absorber operating pressure was optimized to minimize energy cost
‐ CO2 in feed gas was not a concern as all the CO2 was rejected to sales gas and adequate margin over CO2 freezing temperature was available
‐ For Rich gas feed, operating changes such as residue gas recycle are required to achieve desired LPG recovery
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CASE STUDY 3: LOW PRESSURE FEED GAS, USE OF MECHANICAL REFRIGERATION & FEED GAS COMPRESSION
This was a 700 MMSCFD gas plant with 40 barg pressure at the inlet of the gas plant and feed composition as shown in Table 6:
Table 6 Feed Gas Composition Design Case Lean Normal Rich Flowrate 700 MMSCFD Liquid Content # 0.5 0.87 1.3 Composition % mol Nitrogen 1 1.0 1. H2S 0.0 0.0 0.0 CO2 2.0 2.0 2.0 Methane 84.0 74.0 62.0 Ethane 7.0 15.0 18.0 Propane 2.0 3.0 10.0 i-Butane 1.0 2.0 3.0 n-Butane 1.0 1.0 2.0 i-Pentane 1.0 1.0 1.0 n-Pentane 1.0 1.0 1.0
# Liquid content is defined as m3 of recoverable liquid hydrocarbon per 1000 m3 of feed gas
Ethane product was sent for further processing in downstream Ethane recovery units. Desired product specifications were as follows:
Residue Gas Specification CO2 content mole % < 2% Calorific value, KJ/m3 Gross (Min) 36000 Pentane & Heavier mole % < 0.1%
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Ethane Product Specification C1 mol % < 7% C3 and heavier mol% < 1% CO2 Content, mol % < 6% C2 content, mol % >85%
LPG Product Specification Vapour Pressure @ 37 Deg C (Max) 10 bar C5+ mol % (max) < 2%
Ethane recovery above 75% and LPG recovery over 96% was desired.
Process selection Important considerations in process selection of NGL recovery unit included:
1. Wide variation in hydrocarbon content in feed gas. The selected process should provide greater operational flexibility.
2. Feed gas contained on average 2 mol% CO2. Selected process should be tolerant to CO2.
3. High recovery of ethane and LPG products desired.
4. Boosting the feed gas pressure was necessary considering low pressure of feed gas and high required recovery of ethane.
5. Sales gas battery limit pressure was relatively low at 30 barg. Therefore, it was decided to eliminate residue gas compressor by maintaining higher operating pressure for demethaniser column.
6. Use of mechanical refrigeration was necessary to provide additional cooling duty in NGL separation unit, cooling medium for deethaniser overhead condenser.
Gas Sub-cooled process coupled with mechanical refrigeration was selected. Feed gas compression was used to boost the inlet gas pressure. This was followed by dehydration using molecular sieve bed and a mercury guard bed. Process flow scheme for NGL recovery unit was similar to the schematic shown in Figure 4. Feed gas was initially cooled by rejecting heat to process streams. This was followed by cooling the gas using propane refrigerant. A part of the vapor from the low temperature separator is condensed and sub-cooled using demethaniser overhead vapors. This stream is then flashed to the tower operating pressure and used as reflux to the Demethaniser.
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Results Summary:
Simulation results are summarized in Table 7. With this process over 75% (mole) Ethane and 97% (mole) LPG recovery was achieved.
Table 7 Results Summary Design Case Lean Normal Rich
Gas Feed Rate, MMSCFD 700 700 700 NGL unit inlet pressure, barg 77.7 78.7 78.7
CO2 in feed gas, mol% 2% 2% 2%
Ethane Quality, mol % 85.7 87.2 85.4
Ethane Recovery, % 77.8 77.7 77.7
LPG Quality, mol% 98.1 98 97.8
LPG Recovery, % 97.9 97.5 97.3 Feed Gas Compression Power, kW 25580 26260 26420
Refrigeration Duty, kW 3794 10190 17600 Demethaniser overhead temp, 0C ‐85.96 ‐76.4 ‐66.3
CO2 freeze out temperature, oC -117 -118 -118.6 Sales gas product Flowrate MMSCFD 614.0 551.1 468.9
Ethane Product Flowrate, MMSCFD 44.48 93.6 114.7
LPG Product Flowrate, Std M3/hr 131.4 197.1 479.4
Higher differential pressure across expander results in lower temperatures and therefore higher recovery. Demethaniser operating pressure was fixed so as to meet the residue gas battery limit pressure of 30 barg. As seen from Figure 9, boosting the feed gas pressure improves the recovery of ethane. Part of the inlet gas is used to generate reflux so maximum recovery achievable is limited by the vapor-liquid equilibrium in top section of the column. A reflux stream that is lighter, colder or higher reflux flowrate will provide improved stripping of demethaniser column vapors. This will improve ethane and LPG recovery.
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Effect of Feed Gas richness
As seen from Table 7 and Fig. 10 Rich gas with a greater quantity of liquefiable hydrocarbons produces a greater quantity of product. Higher refrigeration duty is required as the quantity of liquid to be condensed is more. Also expansion of rich gas in the turboexpander produces relatively higher temperature compared to the expansion of lean gas. As a result, operating temperatures with rich gas will be relatively higher. This results in lower temperature differential in process heat exchangers and therefore larger heat exchangers and higher equipment cost for a given recovery efficiency. Lean feed gas requires lower temperatures to achieve high recovery efficiencies. However heat integration is more effective as column operating temperature is lower. Less refrigeration duty is required as the quantity of liquid to be condensed is less.
Effect of CO2 in feed gas
It was observed that carbon dioxide in the feed gas splits between the recovered ethane product and the residue gas. Higher CO2 content had effect on specifications for both products.
CO2 can also freeze in the low temperature sections of the process such as turboexpander outlet or demethaniser overhead section.
As seen from Figure 11, higher CO2 in feed gas results in lower quality of Ethane product. Also, available margin over CO2 freezing temperature is also reduced.
In this case there is adequate margin between the operating temperature and the CO2 freezing temperature. In cases where adequate margin is not available, this can be obtained by increasing the CO2 recovery in demethaniser bottom stream or by increasing heavier content in liquid in demethaniser overhead sections.
Figure 12 demonstrates the impact of increasing the reflux flow by additional cold separator vapor or cold separator liquid to increase the margin of the operating temperature over CO2 freezing temperature.
Process operating conditions were specified so as to ensure the product specifications are met. For some of the cases this resulted in minor loss of efficiency due to higher operating temperature and pressure.
Conclusions:
- Use of mechanical refrigeration helps in achieving higher recoveries
- GSP has better tolerance for CO2 in feed gas
- Specifications and quality of both residue gas and ethane products are affected by the CO2 content of feed gas
- Although CO2 freezing was not a concern for the process under consideration, it was identified that reflux flow rate can be optimized to improve recovery and to maintain adequate margin over CO2 freezing temperature
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CONCLUSIONS
Typical configurations of the gas plant using turbo expander based process are described. Important parameters that impact the process selection and design of a NGL recovery unit are discussed. Three case studies are presented to illustrate the design principles and the considerations for process selection.
Operating pressure of light ends separation column is an important parameter for turbo expander based NGL recovery processes and is a tradeoff between the hydrocarbon recovery and the recompression power requirement.
Conventional expander process with refluxed deethaniser is suitable for LPG recovery of the order of 70-85%. OHR process provides more efficient recovery of propane and heavier hydrocarbons (>90%) than the conventional expander design.
For high ethane recovery in the range of 80-90% GSP is to be considered. Feed gas compression or mechanical refrigeration may be required to achieve the ethane recovery.
Maximum recovery achievable with split vapor processes like GSP is limited by the vapor-liquid equilibrium in the top section of the column. More than 90% recovery of ethane will necessitate opting for licensed designs that are modifications of the GSP process. This recovery can be improved by increasing reflux flow, reducing the reflux temperature or by making the reflux stream lighter.
High CO2 in feed gas can affect specifications of residue gas as well as ethane product. The process should operate in a region that avoids CO2freezing in low temperature sections. Split vapor processes like the GSP have better tolerance for CO2 in feed.
Compressors constitute 30-40% of total equipment cost and major portion of operating cost of a gas processing facility. Selection and optimization of compressor function is important for overall economics of the processing facility.
Final process selection should be based on detailed cost benefit analysis considering CAPEX/OPEX for the proposed facility and the value of the products.
ACKNOWLEDGEMENT:
We take this opportunity to thank Mr. Nigel Paton and Mr. Barend Vljoen for their assistance in reviewing the paper and giving valuable comments.
REFERENCES:
1. GPSA Engineering Databook, 11th Edition
2. Pitman, R.N., Hudson, H.M., Wilkinson J.D., “Next generation processes for NGL/LPG recovery”, Proceedings of the 77th GPA Annual Convention.
3. Buck, L. L., U.S. Patent No. 4,617,039
4. Campbell, R.E., and Wilkinson, J. D., U.S. Patent 4,157,904
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FIGURES
Figure 1: Conventional Expander process
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Figure 2: Conventional Expander process with Refluxed deethaniser
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Figure 3: Overhead Recycle (OHR) Process
Figure 4: Gas Sub-cooled Process (GSP)
Figure 5
PR
%l
LPGRe
covery,m
ol%
5: Effect of
Figure 6:
60
65
70
75
80
85
8
Prop
ane Re
covery, %
mol
90
90.5
91
91.5
92
92.5
5.5
LPG Recovery, m
ol %
Deethanise
: Effect of A
10
De
Propane
Sales gas Power, KW
6.5
r Pressure o
Absorber Pa
12
eethaniser Col
Recovery, % m
compressor W
7.5 8Absorber P
on Compres
acked Bed H
14
lumn Pressure
mol
8.5 9.5Packed bed h
ssion Power
Height on L
16
e, barg
10.5height, m
r and Propa
PG Recover
8
9
18
11.5
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ane Recover
ry
8000
9000
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14000
15000
16000
Compression
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er, kW
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Fi
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igure 7: Effe
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18
LPG Recovery, m
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5000
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15000
20000
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30000
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50000
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Exchan
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19
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ure 8: Effect
20Absorbe
PG recovery mo
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23
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Power
10700
11200
11700
12200
12700
13200
13700
14200
14700
15200
24
Compression
Pow
er , kW
30
40
50
60
70
80
90
100
LPG Recovery, m
ol %
er
kW
23
Page 24
Figure 9: Feed Gas Compression Power Vs Ethane Recovery
Figure 10: Effect of Feed Gas Richness
60.0
65.0
70.0
75.0
80.0
85.0
90.0
23000 24000 25000 26000 27000 28000 29000 30000
C2 Recovery, m
ol %
Feed Gas Compressor Power kW
‐90
‐85
‐80
‐75
‐70
‐65
‐60
‐55
‐50
0
2000
4000
6000
8000
10000
12000
14000
16000
18000
20000
37 57 77 97 117
Dem
etha
niser O
verhead Temp, deg
C
Refrigeration Duty, kW
Ethane Product Rate, MMSCFD
Refrigeration Duty, kW
Page 25
Figure 11: Effect of CO2 content in feed gas
Figure 12: Effect of Reflux flow on CO2 freezing
82.0
83.0
84.0
85.0
86.0
87.0
88.0
89.0
90.0
91.0
10.0
20.0
30.0
40.0
50.0
60.0
70.0
80.0
0 0.5 1 1.5 2 2.5 3 3.5
Etha
ne Produ
ct Qua
lity, m
ol %
Margin over CO2 freezing
Tem
p, deg
C
CO2 in Feed, mol %
CO2 Freezing Margin
25.00
27.00
29.00
31.00
33.00
35.00
37.00
39.00
41.00
9000 10500 12000 13500
CO2 Freezing
Tem
perature M
argin
Reflux Flow, kgmol/h