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Integrated process development for an optimum gas processing plant
chemical engineering research and design 1 2 4 ( 2 0 1 7 ) 114–123
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Chemical Engineering Research and Design journal homepage: www.elsevier.com/locate/cherd
Integrated process development for an optimum gas processing plant Abd El-Rahman Sayed a,b,∗ , Ibrahim Ashour b,c , Mamdouh Gadalla d a
Egyptian Natural Gas Company (GASCO), New Cairo, Egypt Department of Chemical Engineering, Faculty of Engineering, Minia University, Egypt c Environmental Engineering, Zewail City of Science and Technology, 6th October, Egypt d Department of Chemical Engineering, The British University in Egypt (BUE), Egypt b
a r t i c l e
i n f o
a b s t r a c t
Article history:
The aim of this work is to develop and optimize an integrated process for natural gas plant
Received 29 October 2016
in Egypt instead of flaring these gases and losing their revenues. The natural gas is sour wet
Received in revised form 6 March
feed gas containing mercury and some of volatile organic compounds with a capacity of
2017
around 21 million cubic feet per day. These impurities require sophisticated gas treatment
Accepted 23 May 2017
processes that can handle and control the pollutants to acceptable limits.
Available online 17 June 2017
The design of new gas plant will be performed through firstly, the design methodology and cascade configuration of gas plant units based on feed gas composition. Secondly, integrated
Keywords:
development and optimization of gas treatment process model is achieved using Aspen
Natural Gas
HYSYS simulation program. Thirdly, modeling of natural gas liquids extraction unit and
Gas sweetening
fractionation train is conducted based on the required marketable products specifications.
Gas dehydration
Finally, Aspen process economic analyzer program is used to calculate the expected capital
Sulfur Recovery Unit
expenditures of the plant. Optimizing the plant configuration accounts for best selection of
BTEX removal
treatment units and processing equipment, including mercury removal unit, sulfur recovery
NGLs recovery
unit, BTEX recovery unit, etc. The preliminary capital expenditures of the gas conditioning and processing plant will be around 48 MUSD. © 2017 Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved.
1.
Introduction
gas treatment units, natural gas liquids recovery unit and marketable products’ fractionation train.
taining a large quantity of methane along with heavier hydrocarbons
Currently these gases are flaring which in a way this poses a negative impact on the environment protection and thus leads to the loss of
such as ethane, propane, etc. In addition, it contains a considerable amount of non-hydrocarbons, for instant nitrogen, carbon dioxide and
the revenues of these gases and the products which can be recovered from these gases.
Natural gas is a gas obtained from a natural underground reservoir, con-
hydrogen sulfide (Younger, 2004). The sold natural gas specifications should meet the standard specifications of sales gas (ES: 1606, 2008). Therefore, the raw gas from the reservoir, which is at certain pressures
2.
and is generally saturated with water vapor, will need considerable treatment and processing. The current study presents an optimum and integrated process
2.1. Optimizing and integrating the design of gas processing plant
design for gas plant to treat a sour wet feed gas that contains mercury and some of volatile organic compounds (VOCs) with around capacity 21 MMSCFD. The developed process design contains all mandate
∗
Design methodology
For gas processing plant, the optimal design will be achieved with focusing around (Paradowski et al., 2005);
Corresponding author at: Egyptian Natural Gas Company (GASCO), New Cairo, Egypt. E-mail address: [email protected] (A.E.-R. Sayed). http://dx.doi.org/10.1016/j.cherd.2017.05.031 0263-8762/© 2017 Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved.
chemical engineering research and design 1 2 4 ( 2 0 1 7 ) 114–123
• Configuration of gas treatment processes cascade, which will be more economic and has high operability. • Selection of the proper acid gas removal process; type of solvent, H2 S vs. CO2 selectivity, requirement to remove other sulfur compounds. . .etc. • Selection of the proper gas dehydration process; liquid desiccant, solid desiccant or both. • Economic handling of separated impurities in the treatment units. • Selection of NGLs recovery technology; absorption or cryogenic process. • Environment pollutants control; all effluents streams contain pollutants within permissible limits as per Egyptian Environmental Law. In addition to the above list, the desired integration inside each unit in new plant means achieving the combination between needs and tasks of “opposite” kinds so that savings can be obtained (Gundersen, 2013). For example, the cooling & condensation are integrated with heating & evaporation as a kind of heat integration. Also, expansion integrated with compression is made as a kind of power integration. Moreover, the desired integration between different units in new plant can be achieved by using byproducts from unit as raw materials in other unit as a kind of chemical integration.
2.2.
Natural gas treatment units design methodology
Natural gas treatment units are deigned based on the impurities present in the process feed gas and permissible limits of these impurities in the sales gas. All acid gases like H2 S and CO2 , shall be removed to prevent corrosive damage to equipment and piping. Hydrogen sulfide is also toxic and total sulfur content is normally regulated. The conversion of H2 S to elemental sulfur is done in the event if it is economical, or required for environmental reasons. Also, all free liquids, both hydrocarbon and water, will be removed. Based on the composition of natural gas, the required gas treatment units are designed. Also, the impurities in natural gas control the sequences of gas treatment processes. In the event that the gas contains mercury, installing mercury removal unit upstream gas sweetening unit is undoubtedly considered the preferred location for sour gas and upstream dehydration package for sweet gas, which avoid emissions of mercury to the atmosphere and contamination of plant equipment. Such a location within the whole treatment plant shall minimize the environmental impact and hence is considered an optimization exercise. This location will ensure any NGLs produced are free from mercury. However, this location is more of a challenge for the mercury removal absorbent (Johnson Matthey, 2014). In the event of sour feed gas, sweetening unit is required to be installed downstream of mercury removal unit to remove H2 S from feed gas and avoid expensive material of construction for downstream units. H2 S removal technology and sulfur recovery process are selected based on flow rate of sour feed gas and the concentration of H2 S in it. The feed gas compositions and conditions, desired purity of the treated gas and the selectivity of H2 S over CO2 act as the evaluation criteria of acid gas removal technologies (NREL, 2009). Amine based solvent processes consider a simplified process flow scheme, representing a typical chemical solvent process. Acid gas to be treated is introduced into the bottom of tray or packed tower and is contacted counter-
115
currently with lean regenerated solvent. Sweet gas exits at the top of the tower (Okimoto and Gadeholt, 1998). Methyl Di Ethanol Amine (MDEA) is a tertiary amine that offers many advantages over other alkanol–amines. The difference in the rates of reaction with H2 S and CO2 gives MDEA a desirable feature over other amines, namely selectivity of H2 S over CO2 . This selectivity is advantageous when the acid gas stream is sent to a sulfur recovery unit (SRU), since CO2 is undesirable in SRU (BRE V.1401, 2013). For rich H2 S acid gas stream, the Claus process is the most widely used process technology for conversion of H2 S to elemental sulfur (Linde Process Plants, Inc.—Sulfur Process Technology, 2013). The conversion of H2 S to elemental sulfur in Claus process is achieved through two types of reactions; thermal and catalytic reactions. A typical Claus-plant can recover 95–98% of the sulfur of rich H2 S stream then incinerator may be required to burn residual H2 S to SO2 which is less toxic than H2 S (Okimoto and Gadeholt, 1998). After gas sweetening, the gas will be saturated with water. Therefore, it is required to dehydrate sweet feed gas. The selection of a proper dehydration method depends on the initial water content and required water dew point downstream of it (Kocken, 2016). Also, process character and economic factor are considered in choosing the proper dehydration method (Netusil and Ditl, 2011). Absorption by tri-ethylene glycol (TEG) is used widely for gas dehydration in the petroleum industry to dehydrate the natural gas, while adsorption dehydration can achieve very low water content and is best applied where very low dew point are required (<−50 ◦ C) (Netusil and Ditl, 2011). In the event of saturated gas with water vapor, it is economically using TEG. However, if water dew point is required for cryogenic temperatures of propane recovery, mole sieve dehydration package will be installed in downstream of TEG unit. In the event that the plant feed gas contains benzene, toluene, ethyl benzene and xylene (BTEX), the BTEX treatment will be in two locations; for acid gas stream released from amine stripper in sweetening unit and gases from TEG re-boiler. Removal of BTEX from acid gas is necessary to protect SRU catalysts from cracking and producing coke that blocks active sites (Petrofac Services Ltd., 2010). Deactivation of catalyst results in low sulfur recovery and subsequently frequent plant shutdowns to replace catalyst and very short catalyst life (Crevier et al., 2001). Axens CSM 31 is an advanced promoted alumina beads used as protective layer to avoid poisoning due to BTEX. Axens materials recommend CSM 31 for BTEX content <100 ppm. Also, Axens recommend BTEX removal facilities upstream to sulfur recovery unit for BTEX content >100 ppm (Roisin, 2015). Re-generable activated carbon beds used for BTEX removal depends on the fundamental principles of adsorption which takes place with acid gas in an upward flow (Morrow and Lunsford, 1997), where activated carbon beds with a low pressure drop of 0.2 bar are installed on the acid gas feed to adsorb the majority of BTEX (98%). Activated carbon has a bulk density of about 350 kg/m3 and surface area of 500 m2 /g (Burton et al., 2015). Regeneration is performed with downward flow of low pressure steam which eliminates the possibility of steam condensing and fluidization of the bed if the opposite flow regime were used. With the majority of the BTEX removed the acid gas flows to SRU. The steam strips the BTEX from the carbon. The resulting regeneration gas (steam and desorbed BTEX) is vaporized and sent
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chemical engineering research and design 1 2 4 ( 2 0 1 7 ) 114–123
Fig. 1 – Diagram of generic gas treatment sequences.
to the SRU incinerator in case of non-economic BTEX figures, otherwise, high BTEX figures are separated and acting added value as a product mixes with gasoline. After steaming, the BTEX Removal Vessel is purged with fuel gas to dry and cool the activated carbon bed ready for the next adsorption cycle. For outlet gases from TEG regenerating reboiler, BTEX treatment is required to avoid escape BTEX to atmosphere. In the event of BTEX will be higher than volatile organic compounds (VOCs) emission limits to atmosphere that causes health and environmental hazards (Egyptian Prime Minister’s Decree No. 964, 2015). Therefore, the vapors from the still column of the glycol reboiler are incinerated in the incinerator (Gordon Stewart, 2015). The gas treatment sequences can be illustrated as shown in Fig. 1.
2.3. Design methodology of propane plus recovery unit and fractionation train The best technology to process gases depends on inlet gas characteristics and liquid recovery objective (Gaskin and Mehra, 1999). A major leap in NGLs recovery was the introduction of a turbo-expander technology, which is known as an Industry-Standard Single-stage (ISS) process module (Getu et al., 2013). The most known propane recovery processes schemes are developed by Ortloff (Ortloff, 2009) and IPSI (BECTEL, 2016) Companies. The gas sub-cooled (GSP), cold residue recycle (CRR) and Recycle vapor-split (RSV) were allowed by Orloff Company. The enhanced NGL recovery processes (IPSI.1 and IPSI.2) belong to IPSI Company. The GSP process scheme was developed as an improvement to the turbo-expander scheme. This process scheme uses a split-vapor feed as a reflux to the rectification section of the demethanizer column (Campbell and Wilkinson, 1981). In last two decades, all processes utilize advanced mechanical equipment’s which lead to improved efficiencies,
including using plate fin heat exchangers, higher efficiency expander re-compressor, minimize the required utilities and use of more flexible and accurate Process Simulation Programs (Gaskin and Mehra, 1999). The marketable products which can be produced from the feed gas are commercial propane, liquefied petroleum gases (LPG) and condensate. Based on the required marketable products, the lightest component that should be recovered is propane, therefore the design of NGLs recovery unit will be objective to propane recovery process. Subsequently, there is no need to use licensed propane recovery technology. Therefore, open art technology process scheme will be used as a NGLs recovery unit. The expander plant heat integration process scheme is selected to achieve high propane recovery with no or very low ethane recovery (ethane rejection). It demonstrates how to maximize heat integration for a turbo-expander process and implies that the available cooling is not being fully utilized. Also, there are some process variations due to feed gas and product specifications. The following controls to maximize heat integration. In each case, use the effective approach temperature:
• Control the outlet temperature of the gas/gas exchanger (feed side) such that the approach is 5 ◦ F [3 ◦ C]. For the initial guess, use the temperature required to condense a small portion of the feed (e.g. 10–20%). • Control the outlet temperature of the side reboiler such that the approach temperature is 10 ◦ F [5 ◦ C]. Use a temperature guess that is 20–30 ◦ F [10–15 ◦ C] warmer than that of the gas/gas exchanger. • Control the feed split fraction such that the approach temperature of the reboiler is 10 ◦ F [5 ◦ C]. For a turbo-expander process, the optimum feed split to the reboiler is typically
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between 20 and 50%. Use the average of these values (35%) as the initial split guess. If an inlet gas stream is very rich in liquids content, the amount of latent heat that must be removed in the form of expander shaft work can be significant. This high shaft work requires a lower column pressure and consequently higher recompression power. In this case, it can be more economical to remove some of the condensation energy via mechanical refrigeration (BRE V.1402, 2013). The refrigeration in gas processing is the propane cycle that has a low-end temperature of −32 to −37 ◦ C. For a conventional turbo-expander plant with an LTS temperature of −45.5 to −51 ◦ C, adding a propane chiller would require breaking the residue or gas/gas exchanger into two pieces, a hot gas–gas exchanger and a cold gas–gas exchanger with the propane chiller in between. The modeling of propane plus (C3+) recovery will be by manipulating process parameter until the achievement of high propane and low ethane recoveries are possible. These operating conditions are changeable depending on feed gas composition and conditions. This is another optimization exercise as for operating and designing parameters to maximize propane recovery. Once the natural gas liquids are recovered, they must be separated in a fractionation train before they are sold as final products. The technique used for this separation is distillation. Three rules of thumb for sequential distillation are: • Remove the most plentiful component first to reduce the size of the downstream processing equipment; • Remove the lightest components first to reduce the necessary pressure in subsequent columns for condensation of the light reflux components; and • Make the hardest separation last to reduce the amount of material that has to be processed in the tallest column with the highest reflux/boil-up ratios. With deethanized liquid streams, the light components tend to be the largest constituents, so the first two rules can be followed simultaneously. The NGL fractionation train is depropanizer and debutanizer columns.
3.
Problem definition and design basis
An actual plant in Egypt has sour wet feed gas with composition and conditions as shown in Table 1, the sample which contains higher propane plus content is considered the rich case of plant feed gas, and vice versa. The following meteorological design criteria shall be accommodated in the design of this plant: • Minimum temperature for design purposes shall be 4 ◦ C. • Maximum ambient air temperature is 50 ◦ C. • Ground temperature is 17 ◦ C in winter and 30 ◦ C in summer.
4.
Configuration of gas plant units
4.1.
Gas treatment units cascade
The feed gas composition contains high content of mercury, H2 S, water and traces of BTEX. Due to the presence of mercury in the feed gas, mercury removal unit (MRU) must be
Table 1 – Plant feed gas composition and conditions. Lean case Mole fractions Nitrogen CO2 H2 S Methane Ethane Propane i-Butane n-Butane i-Pentane n-Pentane n-Hexane Methyl cyclo pentane Benzene n-Heptane Toluene n-Octane Ethyl benzene m-Xylene p-Xylene o-Xylene n-Nonane n-Decane H2 O Total Mercury content, Nano gm/SCM Temperature, ◦ F (◦ C) Pressure [psig] Flow rate [MMSCFD]
0.003676 0.025637 0.084009 0.627313 0.118845 0.075869 0.010629 0.023265 0.006738 0.005623 0.004772 0.000000 0.000167 0.000344 0.000006 0.000447 0.000003 0.000004 0.000004 0.000005 0.000079 0.000000 0.012567 1.00 1155 131 ◦ F (55 ◦ C) 650 21.1
Rich case
0.012299 0.030874 0.056840 0.616847 0.125539 0.086503 0.012057 0.026178 0.006840 0.005499 0.003520 0.000006 0.000677 0.001599 0.000004 0.000506 0.000159 0.000076 0.000690 0.000296 0.000355 0.000070 0.012568 1.00
installed as the first treatment unit to ensure any NGLs produced are free from mercury. Also, feed gas contains H2 S which mandates installing gas sweetening unit in the downstream mercury removal unit. The H2 S content in feed gas is high while CO2 content is low or approach of required treat gas specification, therefore MDEA will be the best solvent for deep removing of H2 S over-than CO2 . The removed H2 S from feed gas will be directed to sulfur recovery unit but due to the presence of BTEX in acid gas with higher than 100 ppm content, BTEX removal unit should be installed upstream of sulfur recovery unit to separate BTEX, which will be burned in the incinerator. The sweet gas from amine sweetening unit will be saturated with water. Therefore, it is fed to TEG unit to remove water vapor from gas, and gases from TEG reboiler will be flowed to incinerator due to presence of BTEX in it. Mole sieve bed will be installed downstream TEG unit to carry out the required water dew point for NGLs recovery process. The treated gas flows to propane recovery unit to recover marketable products, which will be separated in fractionation train. The selected flow diagram of gas treatment processes will be as depicted in Fig. 2.
4.2.
Simulation basis of gas plant modeling
Today, using of simulation programs acts an important role in process development studies (Ebenezer and Gudmundsson, 2005). The modeling of all gas treatment unit, propane recovery unit and fractionation train is performed by using Aspen HYSYS process simulation program version 8.8. The proper property package for each unit is selected. For Claus process with incinerator, simulation was developed by using Aspen HYSYS version 9. Spreadsheet calculation is used for
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Plant Feed Gas
Mercury Removal Unit
Gas Sweetening Unit
Acid Gas
BTEX Removal Unit
Acid Gas W/o BTEX
Separated BTEX and HC
Elemental SRU
Sulfur Tail Gas
Separated Water vapor & other TEG Unit hydrocarbons (include BTEX)
Incinerator
Mole Sieve Beds Treated Feed Gas Fig. 2 – The sequences of gas treatment processes.
the design of mole sieve beds. Also, BTEX removal unit was designed based on private communication with specialized company (as the design of unit depends on the adsorbent criteria which differ from supplier to another). A fluid package is a combination of a component list and a collection of task or industry-specific property-derivation methods called a property package. Choosing a suitable property package is one of the most important considerations for a successful process simulation. Two of the key factors to consider are; specific system under consideration (what components are involved) and operating conditions (Aspentech-Optimize Hydrocarbon Processes with Aspen HYSYS, 2016; LaRue et al., 2013). Peng–Robinson property package is ideal for VLE calculations as well as calculating liquid densities for hydrocarbon systems. Several enhancements to the original PR model were made to extend its range of applicability and to improve its predictions for some non-ideal systems. However, in situations where highly non-ideal systems are encountered, the use of activity models is recommended. The acid gas–chemical solvents property package is used by the HYSYS acid gas cleaning workflow to simulate removal of acid gases such as hydrogen sulfide and carbon dioxide from process streams. The Sulsim (Sulfur Recovery) property package utilizes the property package created by Sulfur Experts and uses the same Gibbs free energy, enthalpy, and viscosity correlations. This enables to easily integrate the sulfur recovery simulation with other gas processes. Glycol property package contains the TST (Twu–Sim–Tassone) equation of state to determine the phase behavior more accurately and consistently for the TEG–water mixture.
4.3.
Capital expenditures estimation
The cost estimation of gas treatment units, NGLs recovery unit and fractionation train will use Aspen process economic analyzer (APEA) program (Aspentech-Aspen Process Economic Analyzer, 2016). APEA uses expert system links to effect the automatic transfer of Aspen HYSYS process simulator output results.
With APEA, Equipment’s sizes can be revised and enter the values for un-sized equipment or develop sizes which built-in proprietary process simulator program. APEA projects the cash flow and operating cost of competing technologies and process configurations during conceptual design, automatically populating engineering models with details needed for economic evaluation, and subsequently leading to better-informed capital decisions. APEA evaluates alternate plant capacities and locations. The plant location can be changed (choosing from twenty-two different countries), and APEA’s plant relocation technology will automatically revise the design and cost basis parameters, including parity exchange rate, workforce rates, productivities, and construction practices. APEA estimates the updated supply cost and installed cost of each equipment with taking into account the plant location in the world. The cost of unit is the total costs of all its equipment’s and the cost of plant is the total costs of all its units.
5.
Results and discussion
5.1.
Modeling results of natural gas treatment units
Based on selected gas treatment flow scheme, the plant feed gas enters mercury removal unit. Presence of liquids in the plant feed gas depends on whether the gas is at its dew point and also is above ambient temperature. Therefore, to avoid liquids droplets in feed gas to a MRU, a vertical cylinder knock-out drum will be installed with a well-designed demister pad in the top. Then, well-designed filter-separator (F-S) will be installed in the downstream of the K-O drum. The piping from the F-S to the MRU, MRU Vessel and heads will be insulated with 75 mm thickness heat-traced insulation material (Selective Adsorption Associates Inc., 2012). The volume of mercury removal vessel is 1.5 m3 based on using PURASPEC 1173 catalyst bed type (Johnson Matthey Technical specification sheet, 2015) and the outlet mercury content less than 10 ng/sm3 (see Fig. S1 and Table S1 in supplementary content section).
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Fig. 3 – Process flow scheme of BTEX removal unit. The gas from MRU flows to MDEA contactor which is a part of amine sweetening unit. Amine solvent removes H2 S from feed sour gas at 640 psig and 75 ◦ C. The rich amine from MDEA contactor will be regenerated in the sweetening unit. The simulation modeling of amine unit will be done by using Aspen HYSYS program, the property package is acid gas, and valid phases are vapor and liquid. The simulator program calculates amine flow rate and required utilities loads. According to reference guidelines of sweetening unit design parameters (GPSA Section 21, 1998), the key parameters in MDEA sweetening simulation model are manipulated until achieving the required sweet gas specification, then try to optimize losses of amine and water as possible. The amine solvent flow rate will be between 5.5 to 7 SCF/Gal, it depends on the acid gas content in feed gas of sweetening unit. Also, the reboiler duty of amine regenerator will be between 1175 to 1197 BTU/Gal, which is complying with reference range between 800 to 1200 BTU/Gal of amine (see Fig. S2 and Table S2 in supplementary content section). Due to presence of BTEX in the separated acid gas from amine unit, the acid gas flows to BTEX removal unit. The water content in acid gas stream is very high at 8.37% and so a separate de-humidification would be required up stream of the BTEX removal unit. Also, the C5 and C6 hydrocarbon content is much higher. The C5 & C6 will adsorb as well as the BTEX and so the plant will have to handle this additional load
(Babcock and Wilcox Megtec, 2017). The BTEX removal unit would include:
• De-humidification (cooling, condensation & reheating of the gas stream) • Adsorber vessels to contain the activated carbon which will adsorb the BTEX and hydrocarbons • Heat exchanger to condense the desorbed BTEX, hydrocarbons and steam • Decanter to separate the BTEX/hydrocarbons from the steam condensate • Ancillary blowers and heat exchangers to condition the carbon after desorption
The adsorbent activated carbon with capacity 9000 kg and bulk density 425 kg/m3 will remove 98% of BTEX, hydrocarbons (C5 and C6) and around 1% of H2 S present in acid gas stream, as shown in Fig. 3 and Table 2. BTEX removal beds are frequently regenerated with steam; 120 min adsorption, 45 min regeneration and 15 min cool down/drying. Waste products from BTEX removal unit are spent activated carbon (disposal by incineration possible), liquid hydrocarbon product (for disposal to incinerator) and sour water product (for disposal to incinerator).
Table 2 – The calculations results of BTEX removal unit.
1 2 3 4
BTEX content in acid gas Hydrocarbons content in acid gas 1% of H2 S content in acid gas Adsorbent activated carbon (Babcock and Wilcox Megtec, 2017)
Kg/h Kg/h Kg/h Kg
Lean & high ambient temp. (50 ◦ C)
Rich & low ambient temp. (4 ◦ C)
8.44 184.6 26.869 9000
25.62 124.34 19.41
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The acid gas exits BTEX removal unit to sulfur recovery unit. The process simulation of sulfur recovery unit will be by Aspen HYSYS program with sulfur recovery as a fluid package. The temperature in main furnace of SRU unit will be 1197 to 1331 ◦ C and the temperatures in Claus converters 1, 2 and condensers will be 232.2, 218.8 and 135 ◦ C respectively. The overall sulfur recovery higher than 96% and produced sulfur is 58 and 42 t/day for lean and rich cases respectively (see Fig. S3 and Table S3 in supplementary content section). The sweet gas from amine sweetening unit will flow to TEG dehydration unit to remove water vapor. The simulation of TEG unit will be performed by Aspen HYSYS simulation program and glycol package as a fluid package. The TEG circulation rate will be between 2.17 to 2.54 gal TEG/lb H2 O absorbed in dehydration unit. Also, the heat duty of TEG reboiler will be between 1158 to 1210 BTU/Gal TEG (see Fig. S4 and Table S4 in supplementary content section). The tail gas from sulfur recovery unit, the separated BTEX in BTEX removal unit and gases from TEG re-boiler will be incinerated in the incineration unit. The simulation of incineration unit will be by Aspen HYSYS simulation program and Sulfur Recovery as a fluid package. The calculated incinerator temperature will be 1090 ◦ C and exit CO concentration is less than 61 ppm moles, which is complying with environmental law limitations (<100 ppm) (see Fig. S5 and Table S5 in supplementary content section).
5.2.
Mole sieve dehydration unit design
To achieve the proper water dew point for NGLs recovery, the exit gas from TEG dehydration unit flows to mole sieve dehydration unit. The design calculation of mole sieve is figured out according to Section 20—dehydration of GPSA engineering Data-book (GPSA Section 20, 1998). The calculated maximum velocity in mole sieve beds will be between 30 to 31 ft/s and water dew point of outlet dry gas from mole sieve will be −100 ◦ C (see Fig. S6 and Table S6 in supplementary content section).
5.3. Simulation modeling results of natural gas liquids (NGLs) recovery unit and fractionation train After feed gas dehydration in mole sieve, the feed gas became treated and can be processed to recover valued products from it. The treated gas specifications will be as revealed in Table 3. The selected process scheme named with expander plant heat integration, which achieves high propane recovery with no or very low ethane recovery. It is open art process technology scheme and considers mix of different technologies, which acts as an innovation for heavy feed gas (the inlet gas to NGLs recovery has Molecular Weight >23). In the selected NGL recovery process scheme as shown in Fig. 4, the feed gas splits into two streams and be cooled by cold residue gas and side cold stream from tray 9 in deethanizer. The cold feed gas from inlet gas exchangers passes through propane chiller and cold gas/gas exchanger to low temperature separator (LTS) at temperature and pressure are −49 ◦ C and 612.5 psig respectively. The vapor stream from LTS flows to turbo-expander and feeds the deethanizer in tray-1, and liquid from LTS flow through control valve to deethanizer at tray 8. The operating conditions at deethanizer column are −71.7 ◦ C first tray temperature, 48 ◦ C last tray temperature and 145 psig pressure. The recovered NGLs stream flows to depropanizer column to produce commercial propane product, the oper-
Table 3 – Treated feed gas composition and conditions.
Mole fractions Nitrogen CO2 Methane Ethane Propane i-Butane n-Butane i-Pentane n-Pentane n-Hexane n-Octane n-Heptane Benzene Total Mercury content, Nano gm/SCM Temperature, ◦ F (◦ C) Pressure [psig] Flow rate [MMSCFD]
Lean & high ambient temp. (50 ◦ C)
Rich & low ambient temp. (4 ◦ C)
0.0041 0.0169 0.7052 0.1334 0.0851 0.0119 0.0261 0.0069 0.0057 0.0043 0.0001 0.0002 0.0001 1.00 <10
0.0138 0.0219 0.6892 0.1373 0.0906 0.0118 0.0247 0.0051 0.0038 0.0014 0.0000 0.0002 0.0001 1.00
131.36 ◦ F (55.2 ◦ C) 620 18.73
108.4 ◦ F (42.45 ◦ C) 620 18.7
ating conditions at depropanizer column are 55 ◦ C first tray temperature, 85 ◦ C last tray temperature and 280 psig pressure. The liquid stream from the bottom of depropanizer column flows to debutanizer column to produce LPG product from the top and condensate from its bottom. The operating conditions at debutanizer are 55 ◦ C first tray temperature, 130 ◦ C last tray temperature and 145 psig pressure. The simulation modeling of NGLs recovery unit will be by Aspen HYSYS simulation program and Peng–Robinson as a fluid package. The duty of propane chiller in NGLs recovery unit is around 4 MMBTU/h, the power of residue gas compressor is 780 kW and reboiler duties of depropanizer and debutanizer columns are 2 and 35 MMBTU/h respectively. At deethanizer column, propane recovery is more than 94% and butane plus recovery is more than 99%, while ethane recovery is 1% (see Table S7 in supplementary content section). The products specifications should comply the certified specifications issued by Egyptian Organization for Standards and Quality (ES: 4753, 2011). The commercial propane product contains more than 95% propane component, the vapor pressure of LPG is higher than 10 kg/cm2 g and Reid vapor pressure of condensate is lower than 10 psi. The plant feed gas and products figures are revealed in Table 4.
5.4.
Cost estimation of plant supply
Based on the results of APEA program results, the preliminary capital expenditures of the gas treatment units are 30.92 MUSD and propane recovery unit & fractionation train are 17.81 MUSD. Therefore, the total preliminary capital expenditures of the overall gas processing plant are around 48.73 MUSD. The preliminary estimation of plant units cost is revealed in Table 5.
6.
Conclusions
This work has been focusing on developing an integrated process for gas processing plant that receives sour natural gas containing mercury. The research work has proposed an integrated process with some practical. As the plant feed gas
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Fig. 4 – Propane recovery unit and fractionation train process flow scheme.
Table 4 – Feed gas and products figures of NGLs recovery unit. Object
Variable
Lean & high ambient temp. (50 ◦ C)
Rich & low ambient temp. (4 ◦ C)
Unit
Inlet gas
Molar flow Pressure Temperature Molar flow Higher heating value Mass flow Master comp. mass fraction (Propane) Mass flow Pressure vapor pressure Std. ideal liquid volume flow Calculator (Reid VP at 37.8 ◦ C)
18.73 620.0 55.2 16.16 1101.6 24.7 0.952 102.1 160.8 308.5 2.4
18.70 620.0 42.4 16.19 1088.5 42.9 0.959 85.7 150.2 165.8 1.9
MMSCFD psig ◦ C MMSCFD Btu/SCF t/d
Sales gas Propane LPG Condensate
Table 5 – Preliminary capital expenditures of plant units. No. 1 2 3 4 5 6 7 8
9
Unit Mercury removal unit Amine sweetening unit BTEX removal unit Sulfur removal unit TEG dehydration unit Gas incineration unit Mole sieve unit Propane recovery unit including propane refrigeration package and turbo-expanders Fractionation train
Total capital cost of plant, USD
Capital expenditures, USD 3,973,180 6,993,767 3,266,375 9,647,978 2,576,324 1,149,882 3,310,354 12,236,359
5,577,450 48,731,669
contains mercury, the mercury removal unit is to be installed upstream the gas sweetening unit, then sour feed gas will be treated in an amine sweetening unit. The sweet gas that
t/d psig barrel/day psig
is saturated with water is dehydrated in TEG unit, and then a two-bed molecular sieve unit is necessary to achieve the proper water dew point for high propane recovery. For the current feed gas flow rate, the bed volume of mercury removal unit is 1.5 m3 , the amine flow rate in MDEA sweetening unit is around 192 USGPM, while the glycol flow rate in TEG dehydration unit is 4 USGPM. The separated acid gas in amine sweetening unit flows to BTEX removal unit with capacity 9 t activated carbon beds, then to SRU with two catalytic reactors. Providing that the separated BTEX from acid gas stream is not economic figure to recover, it will be mixed with tail gas from SRU and gases from TEG re-boiler, to burn in incinerator. This helps control releasing the flue gases to atmosphere complying with environmental law limits. Since the lightest component to be recovered is propane, the expander plant heat integration process is selected to achieve high propane recovery, which maximizes heat integration for a turbo-expander process. The recovered liquids
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chemical engineering research and design 1 2 4 ( 2 0 1 7 ) 114–123
will be separated in depropanizer and debutanizer columns into products before they are sold. As proposed by the research work, the recovered marketable products from treated gas in NGLs recovery unit will be around 25–43 t of propane, 86–102 t LPG and 166–308 BBls condensates. The residue gas after products separation will be around 16.2 MMSCFD. In addition to hydrocarbons products, elemental sulfur will be produced with around 42–58 t in sulfur recovery unit. The estimated capital expenditures of the gas processing plant is 48 MUSD.
Appendix A. Supplementary data Supplementary data associated with cle can be found, in the online http://dx.doi.org/10.1016/j.cherd.2017.05.031.
this artiversion, at
References Aspentech-Optimize Hydrocarbon Processes with Aspen HYSYS, [n.d.] [Online], available: https://www.aspentech.com/products/aspen-hysys.aspx (Accessed 14 May 2016). Aspentech-Aspen Process Economic Analyzer, available: http://www.aspentech.com/products/economic-evaluation/ aspen-process-economic-analyzer/. BECTEL—Gas Processing, available: http://www.bechtel.com/bhts/gas-processing/. BRE V.1401, BRE (Bryan Research & Engineering Inc.), 2013. Sour Gas Processing—Removing Undesirable Components from Gases and Liquids, vol. 1401, Texas, USA. BRE V.1402, BRE (Bryan Research & Engineering Inc.), 2013. Designing and Optimizing Hydrocarbon Recovery Processes, vol. 1402, Texas, USA. Babcock & Wilcox Megtec (2017), available at: http://www.babcock.com/megtec. Burton, Luke, Duncan, Chad, Diaz, Armando, Bagajewicz, Miguel, 2015. BTEX Prediction & Removal in Amine Units, Cited in http://focusky.com/scyk/mizo. Campbell, R.E., Wilkinson, J.D. (1981), Hydrocarbon gas processing. U.S. Patent 4278457. Crevier, P.P., Dowling, N.I., Clark, P.D., Huang, M., 2001. Quantifying the effect of individual aromatic contaminants on Claus catalyst. In: 51st Laurance Reid Gas Conditioning Conference, Norman, Oklahoma. Gordon Stewart, 2015. Disposal of H2 S, benzene, toluene, and xylene (BTEX), cited in: http://www.ggordonstewart.com/Disposal.html. ES: 1606/2008, 2008. Natural Sales Gases Egyptian Standards. Egyptian Organization for Standardizations & quality, Egypt. ES: 4753/2011, 2011. Liquefied Petroleum Gases Commercial Propane/butane Egyptian Standards. Egyptian Organization for Standardizations & quality, Egypt. Ebenezer, S., Gudmundsson, J.S., 2005. Removal of carbon dioxide from natural gas for LNG production, technical Project Work (Institute of Petroleum Technology Norwegian). University of Science and Technology, Trondheim, Norway. Egyptian Prime Minister’s Decree No. 964 of 2015 (2015)—the Amendment of Executive Regulation of the Egyptian Environment Law number 4 of 1994, available: http://www.eeaa.gov.eg. GPSA Section 20, 1998. GPSA Engineering Data Book, 11th edition (Electronic), FPS vol. II, Section-20 Dehydration, Tulsa, Oklahoma, USA. GPSA Section 21, 1998. GPSA Engineering Data Book, 11th edition (Electronic), FPS vol. II, Section-21 Hydrocarbon Treating, Tulsa, Oklahoma, USA. Gaskin, Thomas K., Mehra, Yuv R., 1999. Cryogenic or Absorption? When to Use Which for Processing Natural Gas. Advanced Extraction Technologies, Inc., USA.
Getu, Mesfin, Mahadzir, Shuhaimi, Van Duc Long, Nguyen, Lee, Moonyong, 2013. Techno-economic analysis of potential natural gas liquid (NGL) recovery processes under variations of feed compositions. Chem. Eng. Res. Des. 9 (1), 1272–1283. Gundersen, Truls, 2013. what is Process Integration? In: International Process Integration Jubilee Conference, Gothenburg, Sweden. Johnson Matthey Technical Proposal for Mercury Adsorbent (2015), Cairo, Egypt. Johnson Matthey Brochure (2014), Purification Solutions for the Gas Processing Industry, available at: http://www.jmprotech.com/gas-processing-pursapecjohnson-matthey. Kocken Energy Systems International, 2016. Natural Gas Conditioning, Available http://www.kockenenergia.com/en/products-services/ dehydration/index.html. LaRue, Kelley, Grigson, Susan, Hudson, Hank, 2013. Sulfur plant configurations for weird acid gases. In: Laurance Reid Gas Conditioning Conference, Oklahoma, USA. Linde Process Plants, Inc, 2013. Sulfur Process Technology, Tulsa, Oklahoma, USA. Cited in http://www.linde-engineering.com. Morrow, David C., Lunsford, Kevin M., 1997. Removal and disposal of BTEX components from amine plant acid gas streams. In: Proceedings of the Seventy-Sixth GPA Annual Convention, Tulsa, OK: Gas Processors Association, pp. 171–173. NREL L, 2009. Survey and Down-Selection of Acid Gas Removal Systems for the Thermo-chemical Conversion of Biomass to Ethanol with a Detailed Analysis of an MDEA System, Task 1: Acid Gas Removal Technology Survey and Screening for Thermo-chemical Ethanol Synthesis. Nexant Inc., San Francisco, California, USA. Netusil, Michal, Ditl, Pavel, 2011. Comparison of three methods for natural gas dehydration, 2011. J. Nat. Gas Chem. 20 (2011), 471–476, Retrieved from: http://www.scribd.com/doc/162029680/Comparison-of-ThreeMethods-for-Natural-Gas-Dehydration. Okimoto, F.T., Gadeholt, G., 1998. Best Practices Guide Gas Sweetening Technology. SIEP RTS Library, Shell International Exploration and Production B.V., Research and Technical Services. Ortloff-NGL/LPG Recovery, available: http://www.ortloff.com/recovery/. Paradowski, Henri, Le-Gall, Andre, Laflotte, Benoit, 2005. Compare the Different Options for NGL Recovery from Natural Gas. Gas Processing Department—Technip, 92973 Paris La Défense, CEDEX, France. Petrofac Services Ltd. (2010), BTEX and Mercury Removal Option Selection Study, Sharjah—UAE. Roisin, Eric, 2015. Dealing with Aromatics in the Sulfur Recovery Unit. Middle East Sulfur Plant Operations Network, Abu Dhabi, UAE. Selective Adsorption Associates Inc., Langhorne, Pa. 19047 USA. available: http://www.mercuryadsorbents.com/FAQ.html. Younger, A.H., 2004. Natural Gas Processing Principles and Technology — Part I, vol. 1–5. University of Calgary, Canada, 366.
Glossary AC: Air cooler APEA: Aspen process economic analyzer BBls: Barrels BTEX: Benzene, toluene, ethyl benzene and xylene BTU/hr: British thermal unit per hour CO: Carbon monoxide CRR: Cold residue recycle GSP: Gas sub-cooled ISS: Industry-Standard Single-stage kW: Kilowatt LTS: Low temperature separator MDEA: Methyl Di-Ethanol Amine
chemical engineering research and design 1 2 4 ( 2 0 1 7 ) 114–123
MMSCFD: Million cubic feet per day MRU: Mercury removal unit MUSD: Million United States Dollars NGLs: Natural gas liquids ng/sm3 : Nano-gram per standard cubic meter PFS: Process flow scheme Ppm: Part per million Psig: Pound per square inch gauge RSV: Recycle vapor-split SRU: Sulfur recovery unit Std: Standard
TEG: Tri-ethylene glycol TST: Twu–Sim–Tassone USD: United States Dollars USGPM: United States gallon per minute VOCs: Volatile organic compounds W/o: Without
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Contents lists available at ScienceDirect
Chemical Engineering Research and Design journal homepage: www.elsevier.com/locate/cherd
Integrated process development for an optimum gas processing plant Abd El-Rahman Sayed a,b,∗ , Ibrahim Ashour b,c , Mamdouh Gadalla d a
Egyptian Natural Gas Company (GASCO), New Cairo, Egypt Department of Chemical Engineering, Faculty of Engineering, Minia University, Egypt c Environmental Engineering, Zewail City of Science and Technology, 6th October, Egypt d Department of Chemical Engineering, The British University in Egypt (BUE), Egypt b
a r t i c l e
i n f o
a b s t r a c t
Article history:
The aim of this work is to develop and optimize an integrated process for natural gas plant
Received 29 October 2016
in Egypt instead of flaring these gases and losing their revenues. The natural gas is sour wet
Received in revised form 6 March
feed gas containing mercury and some of volatile organic compounds with a capacity of
2017
around 21 million cubic feet per day. These impurities require sophisticated gas treatment
Accepted 23 May 2017
processes that can handle and control the pollutants to acceptable limits.
Available online 17 June 2017
The design of new gas plant will be performed through firstly, the design methodology and cascade configuration of gas plant units based on feed gas composition. Secondly, integrated
Keywords:
development and optimization of gas treatment process model is achieved using Aspen
Natural Gas
HYSYS simulation program. Thirdly, modeling of natural gas liquids extraction unit and
Gas sweetening
fractionation train is conducted based on the required marketable products specifications.
Gas dehydration
Finally, Aspen process economic analyzer program is used to calculate the expected capital
Sulfur Recovery Unit
expenditures of the plant. Optimizing the plant configuration accounts for best selection of
BTEX removal
treatment units and processing equipment, including mercury removal unit, sulfur recovery
NGLs recovery
unit, BTEX recovery unit, etc. The preliminary capital expenditures of the gas conditioning and processing plant will be around 48 MUSD. © 2017 Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved.
1.
Introduction
gas treatment units, natural gas liquids recovery unit and marketable products’ fractionation train.
taining a large quantity of methane along with heavier hydrocarbons
Currently these gases are flaring which in a way this poses a negative impact on the environment protection and thus leads to the loss of
such as ethane, propane, etc. In addition, it contains a considerable amount of non-hydrocarbons, for instant nitrogen, carbon dioxide and
the revenues of these gases and the products which can be recovered from these gases.
Natural gas is a gas obtained from a natural underground reservoir, con-
hydrogen sulfide (Younger, 2004). The sold natural gas specifications should meet the standard specifications of sales gas (ES: 1606, 2008). Therefore, the raw gas from the reservoir, which is at certain pressures
2.
and is generally saturated with water vapor, will need considerable treatment and processing. The current study presents an optimum and integrated process
2.1. Optimizing and integrating the design of gas processing plant
design for gas plant to treat a sour wet feed gas that contains mercury and some of volatile organic compounds (VOCs) with around capacity 21 MMSCFD. The developed process design contains all mandate
∗
Design methodology
For gas processing plant, the optimal design will be achieved with focusing around (Paradowski et al., 2005);
Corresponding author at: Egyptian Natural Gas Company (GASCO), New Cairo, Egypt. E-mail address: [email protected] (A.E.-R. Sayed). http://dx.doi.org/10.1016/j.cherd.2017.05.031 0263-8762/© 2017 Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved.
chemical engineering research and design 1 2 4 ( 2 0 1 7 ) 114–123
• Configuration of gas treatment processes cascade, which will be more economic and has high operability. • Selection of the proper acid gas removal process; type of solvent, H2 S vs. CO2 selectivity, requirement to remove other sulfur compounds. . .etc. • Selection of the proper gas dehydration process; liquid desiccant, solid desiccant or both. • Economic handling of separated impurities in the treatment units. • Selection of NGLs recovery technology; absorption or cryogenic process. • Environment pollutants control; all effluents streams contain pollutants within permissible limits as per Egyptian Environmental Law. In addition to the above list, the desired integration inside each unit in new plant means achieving the combination between needs and tasks of “opposite” kinds so that savings can be obtained (Gundersen, 2013). For example, the cooling & condensation are integrated with heating & evaporation as a kind of heat integration. Also, expansion integrated with compression is made as a kind of power integration. Moreover, the desired integration between different units in new plant can be achieved by using byproducts from unit as raw materials in other unit as a kind of chemical integration.
2.2.
Natural gas treatment units design methodology
Natural gas treatment units are deigned based on the impurities present in the process feed gas and permissible limits of these impurities in the sales gas. All acid gases like H2 S and CO2 , shall be removed to prevent corrosive damage to equipment and piping. Hydrogen sulfide is also toxic and total sulfur content is normally regulated. The conversion of H2 S to elemental sulfur is done in the event if it is economical, or required for environmental reasons. Also, all free liquids, both hydrocarbon and water, will be removed. Based on the composition of natural gas, the required gas treatment units are designed. Also, the impurities in natural gas control the sequences of gas treatment processes. In the event that the gas contains mercury, installing mercury removal unit upstream gas sweetening unit is undoubtedly considered the preferred location for sour gas and upstream dehydration package for sweet gas, which avoid emissions of mercury to the atmosphere and contamination of plant equipment. Such a location within the whole treatment plant shall minimize the environmental impact and hence is considered an optimization exercise. This location will ensure any NGLs produced are free from mercury. However, this location is more of a challenge for the mercury removal absorbent (Johnson Matthey, 2014). In the event of sour feed gas, sweetening unit is required to be installed downstream of mercury removal unit to remove H2 S from feed gas and avoid expensive material of construction for downstream units. H2 S removal technology and sulfur recovery process are selected based on flow rate of sour feed gas and the concentration of H2 S in it. The feed gas compositions and conditions, desired purity of the treated gas and the selectivity of H2 S over CO2 act as the evaluation criteria of acid gas removal technologies (NREL, 2009). Amine based solvent processes consider a simplified process flow scheme, representing a typical chemical solvent process. Acid gas to be treated is introduced into the bottom of tray or packed tower and is contacted counter-
115
currently with lean regenerated solvent. Sweet gas exits at the top of the tower (Okimoto and Gadeholt, 1998). Methyl Di Ethanol Amine (MDEA) is a tertiary amine that offers many advantages over other alkanol–amines. The difference in the rates of reaction with H2 S and CO2 gives MDEA a desirable feature over other amines, namely selectivity of H2 S over CO2 . This selectivity is advantageous when the acid gas stream is sent to a sulfur recovery unit (SRU), since CO2 is undesirable in SRU (BRE V.1401, 2013). For rich H2 S acid gas stream, the Claus process is the most widely used process technology for conversion of H2 S to elemental sulfur (Linde Process Plants, Inc.—Sulfur Process Technology, 2013). The conversion of H2 S to elemental sulfur in Claus process is achieved through two types of reactions; thermal and catalytic reactions. A typical Claus-plant can recover 95–98% of the sulfur of rich H2 S stream then incinerator may be required to burn residual H2 S to SO2 which is less toxic than H2 S (Okimoto and Gadeholt, 1998). After gas sweetening, the gas will be saturated with water. Therefore, it is required to dehydrate sweet feed gas. The selection of a proper dehydration method depends on the initial water content and required water dew point downstream of it (Kocken, 2016). Also, process character and economic factor are considered in choosing the proper dehydration method (Netusil and Ditl, 2011). Absorption by tri-ethylene glycol (TEG) is used widely for gas dehydration in the petroleum industry to dehydrate the natural gas, while adsorption dehydration can achieve very low water content and is best applied where very low dew point are required (<−50 ◦ C) (Netusil and Ditl, 2011). In the event of saturated gas with water vapor, it is economically using TEG. However, if water dew point is required for cryogenic temperatures of propane recovery, mole sieve dehydration package will be installed in downstream of TEG unit. In the event that the plant feed gas contains benzene, toluene, ethyl benzene and xylene (BTEX), the BTEX treatment will be in two locations; for acid gas stream released from amine stripper in sweetening unit and gases from TEG re-boiler. Removal of BTEX from acid gas is necessary to protect SRU catalysts from cracking and producing coke that blocks active sites (Petrofac Services Ltd., 2010). Deactivation of catalyst results in low sulfur recovery and subsequently frequent plant shutdowns to replace catalyst and very short catalyst life (Crevier et al., 2001). Axens CSM 31 is an advanced promoted alumina beads used as protective layer to avoid poisoning due to BTEX. Axens materials recommend CSM 31 for BTEX content <100 ppm. Also, Axens recommend BTEX removal facilities upstream to sulfur recovery unit for BTEX content >100 ppm (Roisin, 2015). Re-generable activated carbon beds used for BTEX removal depends on the fundamental principles of adsorption which takes place with acid gas in an upward flow (Morrow and Lunsford, 1997), where activated carbon beds with a low pressure drop of 0.2 bar are installed on the acid gas feed to adsorb the majority of BTEX (98%). Activated carbon has a bulk density of about 350 kg/m3 and surface area of 500 m2 /g (Burton et al., 2015). Regeneration is performed with downward flow of low pressure steam which eliminates the possibility of steam condensing and fluidization of the bed if the opposite flow regime were used. With the majority of the BTEX removed the acid gas flows to SRU. The steam strips the BTEX from the carbon. The resulting regeneration gas (steam and desorbed BTEX) is vaporized and sent
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chemical engineering research and design 1 2 4 ( 2 0 1 7 ) 114–123
Fig. 1 – Diagram of generic gas treatment sequences.
to the SRU incinerator in case of non-economic BTEX figures, otherwise, high BTEX figures are separated and acting added value as a product mixes with gasoline. After steaming, the BTEX Removal Vessel is purged with fuel gas to dry and cool the activated carbon bed ready for the next adsorption cycle. For outlet gases from TEG regenerating reboiler, BTEX treatment is required to avoid escape BTEX to atmosphere. In the event of BTEX will be higher than volatile organic compounds (VOCs) emission limits to atmosphere that causes health and environmental hazards (Egyptian Prime Minister’s Decree No. 964, 2015). Therefore, the vapors from the still column of the glycol reboiler are incinerated in the incinerator (Gordon Stewart, 2015). The gas treatment sequences can be illustrated as shown in Fig. 1.
2.3. Design methodology of propane plus recovery unit and fractionation train The best technology to process gases depends on inlet gas characteristics and liquid recovery objective (Gaskin and Mehra, 1999). A major leap in NGLs recovery was the introduction of a turbo-expander technology, which is known as an Industry-Standard Single-stage (ISS) process module (Getu et al., 2013). The most known propane recovery processes schemes are developed by Ortloff (Ortloff, 2009) and IPSI (BECTEL, 2016) Companies. The gas sub-cooled (GSP), cold residue recycle (CRR) and Recycle vapor-split (RSV) were allowed by Orloff Company. The enhanced NGL recovery processes (IPSI.1 and IPSI.2) belong to IPSI Company. The GSP process scheme was developed as an improvement to the turbo-expander scheme. This process scheme uses a split-vapor feed as a reflux to the rectification section of the demethanizer column (Campbell and Wilkinson, 1981). In last two decades, all processes utilize advanced mechanical equipment’s which lead to improved efficiencies,
including using plate fin heat exchangers, higher efficiency expander re-compressor, minimize the required utilities and use of more flexible and accurate Process Simulation Programs (Gaskin and Mehra, 1999). The marketable products which can be produced from the feed gas are commercial propane, liquefied petroleum gases (LPG) and condensate. Based on the required marketable products, the lightest component that should be recovered is propane, therefore the design of NGLs recovery unit will be objective to propane recovery process. Subsequently, there is no need to use licensed propane recovery technology. Therefore, open art technology process scheme will be used as a NGLs recovery unit. The expander plant heat integration process scheme is selected to achieve high propane recovery with no or very low ethane recovery (ethane rejection). It demonstrates how to maximize heat integration for a turbo-expander process and implies that the available cooling is not being fully utilized. Also, there are some process variations due to feed gas and product specifications. The following controls to maximize heat integration. In each case, use the effective approach temperature:
• Control the outlet temperature of the gas/gas exchanger (feed side) such that the approach is 5 ◦ F [3 ◦ C]. For the initial guess, use the temperature required to condense a small portion of the feed (e.g. 10–20%). • Control the outlet temperature of the side reboiler such that the approach temperature is 10 ◦ F [5 ◦ C]. Use a temperature guess that is 20–30 ◦ F [10–15 ◦ C] warmer than that of the gas/gas exchanger. • Control the feed split fraction such that the approach temperature of the reboiler is 10 ◦ F [5 ◦ C]. For a turbo-expander process, the optimum feed split to the reboiler is typically
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between 20 and 50%. Use the average of these values (35%) as the initial split guess. If an inlet gas stream is very rich in liquids content, the amount of latent heat that must be removed in the form of expander shaft work can be significant. This high shaft work requires a lower column pressure and consequently higher recompression power. In this case, it can be more economical to remove some of the condensation energy via mechanical refrigeration (BRE V.1402, 2013). The refrigeration in gas processing is the propane cycle that has a low-end temperature of −32 to −37 ◦ C. For a conventional turbo-expander plant with an LTS temperature of −45.5 to −51 ◦ C, adding a propane chiller would require breaking the residue or gas/gas exchanger into two pieces, a hot gas–gas exchanger and a cold gas–gas exchanger with the propane chiller in between. The modeling of propane plus (C3+) recovery will be by manipulating process parameter until the achievement of high propane and low ethane recoveries are possible. These operating conditions are changeable depending on feed gas composition and conditions. This is another optimization exercise as for operating and designing parameters to maximize propane recovery. Once the natural gas liquids are recovered, they must be separated in a fractionation train before they are sold as final products. The technique used for this separation is distillation. Three rules of thumb for sequential distillation are: • Remove the most plentiful component first to reduce the size of the downstream processing equipment; • Remove the lightest components first to reduce the necessary pressure in subsequent columns for condensation of the light reflux components; and • Make the hardest separation last to reduce the amount of material that has to be processed in the tallest column with the highest reflux/boil-up ratios. With deethanized liquid streams, the light components tend to be the largest constituents, so the first two rules can be followed simultaneously. The NGL fractionation train is depropanizer and debutanizer columns.
3.
Problem definition and design basis
An actual plant in Egypt has sour wet feed gas with composition and conditions as shown in Table 1, the sample which contains higher propane plus content is considered the rich case of plant feed gas, and vice versa. The following meteorological design criteria shall be accommodated in the design of this plant: • Minimum temperature for design purposes shall be 4 ◦ C. • Maximum ambient air temperature is 50 ◦ C. • Ground temperature is 17 ◦ C in winter and 30 ◦ C in summer.
4.
Configuration of gas plant units
4.1.
Gas treatment units cascade
The feed gas composition contains high content of mercury, H2 S, water and traces of BTEX. Due to the presence of mercury in the feed gas, mercury removal unit (MRU) must be
Table 1 – Plant feed gas composition and conditions. Lean case Mole fractions Nitrogen CO2 H2 S Methane Ethane Propane i-Butane n-Butane i-Pentane n-Pentane n-Hexane Methyl cyclo pentane Benzene n-Heptane Toluene n-Octane Ethyl benzene m-Xylene p-Xylene o-Xylene n-Nonane n-Decane H2 O Total Mercury content, Nano gm/SCM Temperature, ◦ F (◦ C) Pressure [psig] Flow rate [MMSCFD]
0.003676 0.025637 0.084009 0.627313 0.118845 0.075869 0.010629 0.023265 0.006738 0.005623 0.004772 0.000000 0.000167 0.000344 0.000006 0.000447 0.000003 0.000004 0.000004 0.000005 0.000079 0.000000 0.012567 1.00 1155 131 ◦ F (55 ◦ C) 650 21.1
Rich case
0.012299 0.030874 0.056840 0.616847 0.125539 0.086503 0.012057 0.026178 0.006840 0.005499 0.003520 0.000006 0.000677 0.001599 0.000004 0.000506 0.000159 0.000076 0.000690 0.000296 0.000355 0.000070 0.012568 1.00
installed as the first treatment unit to ensure any NGLs produced are free from mercury. Also, feed gas contains H2 S which mandates installing gas sweetening unit in the downstream mercury removal unit. The H2 S content in feed gas is high while CO2 content is low or approach of required treat gas specification, therefore MDEA will be the best solvent for deep removing of H2 S over-than CO2 . The removed H2 S from feed gas will be directed to sulfur recovery unit but due to the presence of BTEX in acid gas with higher than 100 ppm content, BTEX removal unit should be installed upstream of sulfur recovery unit to separate BTEX, which will be burned in the incinerator. The sweet gas from amine sweetening unit will be saturated with water. Therefore, it is fed to TEG unit to remove water vapor from gas, and gases from TEG reboiler will be flowed to incinerator due to presence of BTEX in it. Mole sieve bed will be installed downstream TEG unit to carry out the required water dew point for NGLs recovery process. The treated gas flows to propane recovery unit to recover marketable products, which will be separated in fractionation train. The selected flow diagram of gas treatment processes will be as depicted in Fig. 2.
4.2.
Simulation basis of gas plant modeling
Today, using of simulation programs acts an important role in process development studies (Ebenezer and Gudmundsson, 2005). The modeling of all gas treatment unit, propane recovery unit and fractionation train is performed by using Aspen HYSYS process simulation program version 8.8. The proper property package for each unit is selected. For Claus process with incinerator, simulation was developed by using Aspen HYSYS version 9. Spreadsheet calculation is used for
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Plant Feed Gas
Mercury Removal Unit
Gas Sweetening Unit
Acid Gas
BTEX Removal Unit
Acid Gas W/o BTEX
Separated BTEX and HC
Elemental SRU
Sulfur Tail Gas
Separated Water vapor & other TEG Unit hydrocarbons (include BTEX)
Incinerator
Mole Sieve Beds Treated Feed Gas Fig. 2 – The sequences of gas treatment processes.
the design of mole sieve beds. Also, BTEX removal unit was designed based on private communication with specialized company (as the design of unit depends on the adsorbent criteria which differ from supplier to another). A fluid package is a combination of a component list and a collection of task or industry-specific property-derivation methods called a property package. Choosing a suitable property package is one of the most important considerations for a successful process simulation. Two of the key factors to consider are; specific system under consideration (what components are involved) and operating conditions (Aspentech-Optimize Hydrocarbon Processes with Aspen HYSYS, 2016; LaRue et al., 2013). Peng–Robinson property package is ideal for VLE calculations as well as calculating liquid densities for hydrocarbon systems. Several enhancements to the original PR model were made to extend its range of applicability and to improve its predictions for some non-ideal systems. However, in situations where highly non-ideal systems are encountered, the use of activity models is recommended. The acid gas–chemical solvents property package is used by the HYSYS acid gas cleaning workflow to simulate removal of acid gases such as hydrogen sulfide and carbon dioxide from process streams. The Sulsim (Sulfur Recovery) property package utilizes the property package created by Sulfur Experts and uses the same Gibbs free energy, enthalpy, and viscosity correlations. This enables to easily integrate the sulfur recovery simulation with other gas processes. Glycol property package contains the TST (Twu–Sim–Tassone) equation of state to determine the phase behavior more accurately and consistently for the TEG–water mixture.
4.3.
Capital expenditures estimation
The cost estimation of gas treatment units, NGLs recovery unit and fractionation train will use Aspen process economic analyzer (APEA) program (Aspentech-Aspen Process Economic Analyzer, 2016). APEA uses expert system links to effect the automatic transfer of Aspen HYSYS process simulator output results.
With APEA, Equipment’s sizes can be revised and enter the values for un-sized equipment or develop sizes which built-in proprietary process simulator program. APEA projects the cash flow and operating cost of competing technologies and process configurations during conceptual design, automatically populating engineering models with details needed for economic evaluation, and subsequently leading to better-informed capital decisions. APEA evaluates alternate plant capacities and locations. The plant location can be changed (choosing from twenty-two different countries), and APEA’s plant relocation technology will automatically revise the design and cost basis parameters, including parity exchange rate, workforce rates, productivities, and construction practices. APEA estimates the updated supply cost and installed cost of each equipment with taking into account the plant location in the world. The cost of unit is the total costs of all its equipment’s and the cost of plant is the total costs of all its units.
5.
Results and discussion
5.1.
Modeling results of natural gas treatment units
Based on selected gas treatment flow scheme, the plant feed gas enters mercury removal unit. Presence of liquids in the plant feed gas depends on whether the gas is at its dew point and also is above ambient temperature. Therefore, to avoid liquids droplets in feed gas to a MRU, a vertical cylinder knock-out drum will be installed with a well-designed demister pad in the top. Then, well-designed filter-separator (F-S) will be installed in the downstream of the K-O drum. The piping from the F-S to the MRU, MRU Vessel and heads will be insulated with 75 mm thickness heat-traced insulation material (Selective Adsorption Associates Inc., 2012). The volume of mercury removal vessel is 1.5 m3 based on using PURASPEC 1173 catalyst bed type (Johnson Matthey Technical specification sheet, 2015) and the outlet mercury content less than 10 ng/sm3 (see Fig. S1 and Table S1 in supplementary content section).
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Fig. 3 – Process flow scheme of BTEX removal unit. The gas from MRU flows to MDEA contactor which is a part of amine sweetening unit. Amine solvent removes H2 S from feed sour gas at 640 psig and 75 ◦ C. The rich amine from MDEA contactor will be regenerated in the sweetening unit. The simulation modeling of amine unit will be done by using Aspen HYSYS program, the property package is acid gas, and valid phases are vapor and liquid. The simulator program calculates amine flow rate and required utilities loads. According to reference guidelines of sweetening unit design parameters (GPSA Section 21, 1998), the key parameters in MDEA sweetening simulation model are manipulated until achieving the required sweet gas specification, then try to optimize losses of amine and water as possible. The amine solvent flow rate will be between 5.5 to 7 SCF/Gal, it depends on the acid gas content in feed gas of sweetening unit. Also, the reboiler duty of amine regenerator will be between 1175 to 1197 BTU/Gal, which is complying with reference range between 800 to 1200 BTU/Gal of amine (see Fig. S2 and Table S2 in supplementary content section). Due to presence of BTEX in the separated acid gas from amine unit, the acid gas flows to BTEX removal unit. The water content in acid gas stream is very high at 8.37% and so a separate de-humidification would be required up stream of the BTEX removal unit. Also, the C5 and C6 hydrocarbon content is much higher. The C5 & C6 will adsorb as well as the BTEX and so the plant will have to handle this additional load
(Babcock and Wilcox Megtec, 2017). The BTEX removal unit would include:
• De-humidification (cooling, condensation & reheating of the gas stream) • Adsorber vessels to contain the activated carbon which will adsorb the BTEX and hydrocarbons • Heat exchanger to condense the desorbed BTEX, hydrocarbons and steam • Decanter to separate the BTEX/hydrocarbons from the steam condensate • Ancillary blowers and heat exchangers to condition the carbon after desorption
The adsorbent activated carbon with capacity 9000 kg and bulk density 425 kg/m3 will remove 98% of BTEX, hydrocarbons (C5 and C6) and around 1% of H2 S present in acid gas stream, as shown in Fig. 3 and Table 2. BTEX removal beds are frequently regenerated with steam; 120 min adsorption, 45 min regeneration and 15 min cool down/drying. Waste products from BTEX removal unit are spent activated carbon (disposal by incineration possible), liquid hydrocarbon product (for disposal to incinerator) and sour water product (for disposal to incinerator).
Table 2 – The calculations results of BTEX removal unit.
1 2 3 4
BTEX content in acid gas Hydrocarbons content in acid gas 1% of H2 S content in acid gas Adsorbent activated carbon (Babcock and Wilcox Megtec, 2017)
Kg/h Kg/h Kg/h Kg
Lean & high ambient temp. (50 ◦ C)
Rich & low ambient temp. (4 ◦ C)
8.44 184.6 26.869 9000
25.62 124.34 19.41
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The acid gas exits BTEX removal unit to sulfur recovery unit. The process simulation of sulfur recovery unit will be by Aspen HYSYS program with sulfur recovery as a fluid package. The temperature in main furnace of SRU unit will be 1197 to 1331 ◦ C and the temperatures in Claus converters 1, 2 and condensers will be 232.2, 218.8 and 135 ◦ C respectively. The overall sulfur recovery higher than 96% and produced sulfur is 58 and 42 t/day for lean and rich cases respectively (see Fig. S3 and Table S3 in supplementary content section). The sweet gas from amine sweetening unit will flow to TEG dehydration unit to remove water vapor. The simulation of TEG unit will be performed by Aspen HYSYS simulation program and glycol package as a fluid package. The TEG circulation rate will be between 2.17 to 2.54 gal TEG/lb H2 O absorbed in dehydration unit. Also, the heat duty of TEG reboiler will be between 1158 to 1210 BTU/Gal TEG (see Fig. S4 and Table S4 in supplementary content section). The tail gas from sulfur recovery unit, the separated BTEX in BTEX removal unit and gases from TEG re-boiler will be incinerated in the incineration unit. The simulation of incineration unit will be by Aspen HYSYS simulation program and Sulfur Recovery as a fluid package. The calculated incinerator temperature will be 1090 ◦ C and exit CO concentration is less than 61 ppm moles, which is complying with environmental law limitations (<100 ppm) (see Fig. S5 and Table S5 in supplementary content section).
5.2.
Mole sieve dehydration unit design
To achieve the proper water dew point for NGLs recovery, the exit gas from TEG dehydration unit flows to mole sieve dehydration unit. The design calculation of mole sieve is figured out according to Section 20—dehydration of GPSA engineering Data-book (GPSA Section 20, 1998). The calculated maximum velocity in mole sieve beds will be between 30 to 31 ft/s and water dew point of outlet dry gas from mole sieve will be −100 ◦ C (see Fig. S6 and Table S6 in supplementary content section).
5.3. Simulation modeling results of natural gas liquids (NGLs) recovery unit and fractionation train After feed gas dehydration in mole sieve, the feed gas became treated and can be processed to recover valued products from it. The treated gas specifications will be as revealed in Table 3. The selected process scheme named with expander plant heat integration, which achieves high propane recovery with no or very low ethane recovery. It is open art process technology scheme and considers mix of different technologies, which acts as an innovation for heavy feed gas (the inlet gas to NGLs recovery has Molecular Weight >23). In the selected NGL recovery process scheme as shown in Fig. 4, the feed gas splits into two streams and be cooled by cold residue gas and side cold stream from tray 9 in deethanizer. The cold feed gas from inlet gas exchangers passes through propane chiller and cold gas/gas exchanger to low temperature separator (LTS) at temperature and pressure are −49 ◦ C and 612.5 psig respectively. The vapor stream from LTS flows to turbo-expander and feeds the deethanizer in tray-1, and liquid from LTS flow through control valve to deethanizer at tray 8. The operating conditions at deethanizer column are −71.7 ◦ C first tray temperature, 48 ◦ C last tray temperature and 145 psig pressure. The recovered NGLs stream flows to depropanizer column to produce commercial propane product, the oper-
Table 3 – Treated feed gas composition and conditions.
Mole fractions Nitrogen CO2 Methane Ethane Propane i-Butane n-Butane i-Pentane n-Pentane n-Hexane n-Octane n-Heptane Benzene Total Mercury content, Nano gm/SCM Temperature, ◦ F (◦ C) Pressure [psig] Flow rate [MMSCFD]
Lean & high ambient temp. (50 ◦ C)
Rich & low ambient temp. (4 ◦ C)
0.0041 0.0169 0.7052 0.1334 0.0851 0.0119 0.0261 0.0069 0.0057 0.0043 0.0001 0.0002 0.0001 1.00 <10
0.0138 0.0219 0.6892 0.1373 0.0906 0.0118 0.0247 0.0051 0.0038 0.0014 0.0000 0.0002 0.0001 1.00
131.36 ◦ F (55.2 ◦ C) 620 18.73
108.4 ◦ F (42.45 ◦ C) 620 18.7
ating conditions at depropanizer column are 55 ◦ C first tray temperature, 85 ◦ C last tray temperature and 280 psig pressure. The liquid stream from the bottom of depropanizer column flows to debutanizer column to produce LPG product from the top and condensate from its bottom. The operating conditions at debutanizer are 55 ◦ C first tray temperature, 130 ◦ C last tray temperature and 145 psig pressure. The simulation modeling of NGLs recovery unit will be by Aspen HYSYS simulation program and Peng–Robinson as a fluid package. The duty of propane chiller in NGLs recovery unit is around 4 MMBTU/h, the power of residue gas compressor is 780 kW and reboiler duties of depropanizer and debutanizer columns are 2 and 35 MMBTU/h respectively. At deethanizer column, propane recovery is more than 94% and butane plus recovery is more than 99%, while ethane recovery is 1% (see Table S7 in supplementary content section). The products specifications should comply the certified specifications issued by Egyptian Organization for Standards and Quality (ES: 4753, 2011). The commercial propane product contains more than 95% propane component, the vapor pressure of LPG is higher than 10 kg/cm2 g and Reid vapor pressure of condensate is lower than 10 psi. The plant feed gas and products figures are revealed in Table 4.
5.4.
Cost estimation of plant supply
Based on the results of APEA program results, the preliminary capital expenditures of the gas treatment units are 30.92 MUSD and propane recovery unit & fractionation train are 17.81 MUSD. Therefore, the total preliminary capital expenditures of the overall gas processing plant are around 48.73 MUSD. The preliminary estimation of plant units cost is revealed in Table 5.
6.
Conclusions
This work has been focusing on developing an integrated process for gas processing plant that receives sour natural gas containing mercury. The research work has proposed an integrated process with some practical. As the plant feed gas
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Fig. 4 – Propane recovery unit and fractionation train process flow scheme.
Table 4 – Feed gas and products figures of NGLs recovery unit. Object
Variable
Lean & high ambient temp. (50 ◦ C)
Rich & low ambient temp. (4 ◦ C)
Unit
Inlet gas
Molar flow Pressure Temperature Molar flow Higher heating value Mass flow Master comp. mass fraction (Propane) Mass flow Pressure vapor pressure Std. ideal liquid volume flow Calculator (Reid VP at 37.8 ◦ C)
18.73 620.0 55.2 16.16 1101.6 24.7 0.952 102.1 160.8 308.5 2.4
18.70 620.0 42.4 16.19 1088.5 42.9 0.959 85.7 150.2 165.8 1.9
MMSCFD psig ◦ C MMSCFD Btu/SCF t/d
Sales gas Propane LPG Condensate
Table 5 – Preliminary capital expenditures of plant units. No. 1 2 3 4 5 6 7 8
9
Unit Mercury removal unit Amine sweetening unit BTEX removal unit Sulfur removal unit TEG dehydration unit Gas incineration unit Mole sieve unit Propane recovery unit including propane refrigeration package and turbo-expanders Fractionation train
Total capital cost of plant, USD
Capital expenditures, USD 3,973,180 6,993,767 3,266,375 9,647,978 2,576,324 1,149,882 3,310,354 12,236,359
5,577,450 48,731,669
contains mercury, the mercury removal unit is to be installed upstream the gas sweetening unit, then sour feed gas will be treated in an amine sweetening unit. The sweet gas that
t/d psig barrel/day psig
is saturated with water is dehydrated in TEG unit, and then a two-bed molecular sieve unit is necessary to achieve the proper water dew point for high propane recovery. For the current feed gas flow rate, the bed volume of mercury removal unit is 1.5 m3 , the amine flow rate in MDEA sweetening unit is around 192 USGPM, while the glycol flow rate in TEG dehydration unit is 4 USGPM. The separated acid gas in amine sweetening unit flows to BTEX removal unit with capacity 9 t activated carbon beds, then to SRU with two catalytic reactors. Providing that the separated BTEX from acid gas stream is not economic figure to recover, it will be mixed with tail gas from SRU and gases from TEG re-boiler, to burn in incinerator. This helps control releasing the flue gases to atmosphere complying with environmental law limits. Since the lightest component to be recovered is propane, the expander plant heat integration process is selected to achieve high propane recovery, which maximizes heat integration for a turbo-expander process. The recovered liquids
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will be separated in depropanizer and debutanizer columns into products before they are sold. As proposed by the research work, the recovered marketable products from treated gas in NGLs recovery unit will be around 25–43 t of propane, 86–102 t LPG and 166–308 BBls condensates. The residue gas after products separation will be around 16.2 MMSCFD. In addition to hydrocarbons products, elemental sulfur will be produced with around 42–58 t in sulfur recovery unit. The estimated capital expenditures of the gas processing plant is 48 MUSD.
Appendix A. Supplementary data Supplementary data associated with cle can be found, in the online http://dx.doi.org/10.1016/j.cherd.2017.05.031.
this artiversion, at
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Glossary AC: Air cooler APEA: Aspen process economic analyzer BBls: Barrels BTEX: Benzene, toluene, ethyl benzene and xylene BTU/hr: British thermal unit per hour CO: Carbon monoxide CRR: Cold residue recycle GSP: Gas sub-cooled ISS: Industry-Standard Single-stage kW: Kilowatt LTS: Low temperature separator MDEA: Methyl Di-Ethanol Amine
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MMSCFD: Million cubic feet per day MRU: Mercury removal unit MUSD: Million United States Dollars NGLs: Natural gas liquids ng/sm3 : Nano-gram per standard cubic meter PFS: Process flow scheme Ppm: Part per million Psig: Pound per square inch gauge RSV: Recycle vapor-split SRU: Sulfur recovery unit Std: Standard
TEG: Tri-ethylene glycol TST: Twu–Sim–Tassone USD: United States Dollars USGPM: United States gallon per minute VOCs: Volatile organic compounds W/o: Without
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