Glycol Loss Minimization for a Natural Gas Dehydration Plant under

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Glycol Loss Minimization for A Natural Gas Dehydration Plant under Upset Conditions Md Emdadul Haque, Qiang Xu, and Srinivas Palanki Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b04675 • Publication Date (Web): 08 Jan 2019 Downloaded from http://pubs.acs.org on January 16, 2019

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Industrial & Engineering Chemistry Research

Glycol Loss Minimization for A Natural Gas Dehydration Plant under Upset Conditions† Md Emdadul Haque, Qiang Xu*, and Srinivas Palanki* Dan F. Smith Department of Chemical Engineering Lamar University, Beaumont, Texas 77710, USA Abstract Low-temperature separation with mono ethylene glycol (MEG) injection process is a common dehydration technique for the natural gas processing. However, the MEG-based dehydration system frequently suffers the significant glycol loss during plant upset conditions, causing double penalties of economic loss and air emissions. Thus, it is very important to minimize the MEG loss in the dehydration process. In this paper, an efficient and effective methodology to reduce the MEG loss under upset conditions of a natural gas dehydration plant has been developed. First, a plant-wide steady-state simulation model is developed and validated at the normal operating condition. Next, the root cause analysis for MEG loss is performed by introducing various process upsets to the simulation model, which indicates that most MEG loss is due to the fluctuating temperature of the stripper column overhead. After that, a plant-wide dynamic simulation model is developed to help generate a new control strategy to regulate the stripper column operation and cope with other plant upset conditions, so as to minimize the MEG loss and improve natural gas product quality and plant operability. Simulation results indicate that plant MEG loss can be reduced by 37%.

Keywords: MEG Loss, natural gas dehydration; plant-wide dynamic simulation; control ___________________________________________________________________________ † For publication in Industrial & Engineering Chemistry Research. * All correspondence should be addressed to Prof. Qiang Xu or Prof. Srinivas Palanki; Phone: 409-880-7818; Fax: 409-880-2197; E-mail: [email protected]; [email protected])

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Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

1. Introduction

Natural gas is commonly used for residential, commercial, and industrial heating. Additionally, it is used for the generation of electricity. In the petrochemical industry, it is often used as a feedstock or raw material, e.g., in the production of ethylene. In the fertilizer industry, it is used for ammonia production. Hydrogen, sulfur and carbon black can be produced using natural gas 1, 2, 3. A schematic diagram of natural gas processing from wellhead to consumer end is shown in Figure 1. Natural gas may be classified according to the hydrocarbon content of the produced gas. Dry gas consists of methane as the major constituent with little or no C2 or heavier components (C2+), whereas wet gas contains C2+ constituents higher than 10 vol% 4.

Figure 1. Natural gas network and consumption.

It is necessary to clean the raw natural gas to meet pipeline transportation specifications and environmental clean-burning gas requirements 5, 6. Thus, gas processing is applied to purify the raw gas from materials identified as fuel impurities (e.g., sulfur), recover valuable components that can be used as petrochemical feedstocks (e.g., ethane), fuels (e.g., propane), or industrial gases (e.g., helium), as well as to help liquefy the natural gas to be transported or stored 7, 8. In a natural gas processing plant, unit operations for gas processing depend on the gas composition, the type of facility, and the product specifications. Among these operations, dehydration is a common technique to remove water from natural gas for the following reasons 1, 9:  water decreases the heating value of the produced gas, consequently increasing its volume;


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 the condensation of water in the pipelines causes slug flow;  water causes possible erosion and corrosion of the pipelines, especially when it is present with carbon dioxide and hydrogen sulfide (acid gases);  hydrates could be formed under specific conditions, which could plug valves, fittings or even pipelines. Dehydration is the process of removing water from a gas and/or liquid, so that no condensed water is present in the system. Inhibition is the process of adding chemicals to the condensed water, so hydrates cannot form. In some instances, inhibition may be the preferred process, particularly in sweet systems employing moderate refrigeration (-40.0˚C and above). Natural gas is commercially dehydrated in one of three ways: absorption (e.g., using glycol dehydration), adsorption (e.g., using molecular sieve, silica gel or activated alumina) and condensation (e.g., using refrigeration with glycol or methanol injection). Among these dehydration processes, glycol absorption is the most commonly used process to meet pipeline sales specifications and field requirements (gas lift, fuel etc.). Adsorption processes are used to obtain very low water contents (0.1 ppm or less) required in low temperature processing such as deep NGL (natural gas liquid) extraction and LNG (liquefied natural gas) plants. Condensation is commonly used as a dehydration process when moderate levels of refrigeration are employed or in pipeline transportation. An inhibitor such as mono ethylene glycol (MEG) or methanol is used to prevent hydrate formation, but the actual water extraction mechanism is condensation 10, 11, 12, 13.

This study focuses on the MEG dehydration process where liquids in the raw natural gas are separated by low temperature separation (e.g., refrigeration system) and MEG is injected as hydrate inhibitor. In such a process, a compression refrigeration technique is used for hydrocarbon



dew point control and NGL extraction. During dynamic operation of gas processing plant, many emergency situations may arise. One of the emergency situations is high pipeline pressure due to low demand at the downstream sites. This low demand leads to backpressure increases in the pipe line, which increases the plant pressure. As a result, the plant needs to cut flow to maintain its safety limit. This induced flow fluctuation will significantly affect the MEG regeneration system14. Meanwhile, this large flow fluctuation causes MEG carryover with gas and hydrocarbon from separator and MEG losses through vaporization from stripper column overhead. Another upset condition of the MEG dehydration process is due to refrigeration system failure. The refrigeration process may break down due to mechanical failure, refrigerant loss and lube oil carryover to the chiller (if the system uses screw compressor). During this breakdown, the plant cannot maintain temperature at the desired hydrocarbon dew point (e.g., fails to meet sales gas specification) and loses not only valuable C2+ components but also MEG due to this temperature fluctuation. This paper focuses on an industrial-scale natural gas treatment process to identify the root causes of glycol losses and proposes a mitigation plan for minimizing MEG loss. In particular, five upset conditions are considered in the following variables: (1) feed flowrate, (2) chiller temperature, (3) stripper column overhead temperature, (4) reboiler temperature, and (5) stripping gas flowrate. Both steady-state as well as dynamic models are developed using the Aspen simulation framework. Steady-state simulation results show that most of the glycol is lost via column overhead and some glycols are carried over with gas and hydrocarbons from low temperature separator (LTS). Dynamic simulations are utilized to develop a new control strategy for optimum column operation. The dynamic model is used to study plant upset conditions, minimize the glycol loss, improve product gas specifications and reduce plant operating costs. In


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particular, it is shown that the improved design has the potential to reduce glycol losses by 37%, lower emissions and improve plant operability.

2. Methodological Framework

Figure 2. shows the general methodological framework for this study, which includes four steps. In the first step, a plant-wide steady state model is developed for base case using the process flow diagram (PFD) of the natural gas treatment process. In this step, feed specifications and operating conditions 15 are used to perform steady-state simulations. After validation, simulationbased root cause analysis for MEG loss is conducted in the second step by changing the following variables: (1) feed flowrate, (2) chiller temperature, (3) stripper column overhead temperature, (4) reboiler temperature, and (5) stripping gas flowrate. After this analysis, the major root causes for MEG losses are identified. Note that the above analysis is based on steady-state simulation. Possible improvement plans to reduce the MEG loss have to be evaluated in the dynamic environment also where plant upsets can be handled via appropriate control strategies. In the third step, the steady-state plant-wide simulation model is converted to a dynamic model using appropriate parameters16, 17. This step also involves the modeling of the base control strategy and control parameters. In the fourth step, dynamic simulations are conducted to identify a new control strategy for the stripper column after analyzing the base model performance, so that it can control the stripper column at the desired operating conditions. The new control strategy includes: (1) a new ratio control for the stripping gas with stripper column feed, (2) a temperature control at stripper column overhead by manipulating rich MEG stream flow from LTS, and (3) a composition controller for stripper column bottom flow (lean MEG) cascaded with reboiler temperature. The



dynamic model is tuned to give the best responses and the model is verified under different disturbances to the inlet gas flow rate and chiller temperature. The improved control strategy is tested with different disturbances to evaluate its operability, which demonstrates the effectiveness of this strategy for MEG loss minimization.

Figure 2. Methodological framework.

3. Plant-wide Steady-state Modeling and Validation

3.1 Process description Figure 3 shows the process flow diagram of this dehydration system. There are four natural gas producing wells and the gas feed goes to two high-pressure separators (HP Separator) V-101 and V-102 at 90 bar and 32.2˚C. Well-1, Well-2 and Well-3 are connected to V-101 and Well-4 is connected to V-102. In HP Separators V-101 and V-102, liquids are separated, water streams goes to a water storage tank at atmospheric pressure and hydrocarbons (condensate) go to a medium pressure separator (MP Separator) V-104 at 21 bar and 26.7 ˚C. Gas streams from V-101 and V102 are mixed together and split into two equal parts and go to the tube side of two gas/gas heat exchanger (E-101 & E-102) for precooling with LTS gas stream. After pre-cooling, the streams are combined again and fed into tube side of a chiller (E-103, propane refrigerant used as the utility) at -7.6˚C. At the chiller, the gas is cooled to -26.1 ˚C and goes to LTS (V-103) for further separation. MEG is injected with the total injection rate of 0.80 m3/hr to the tube side of E-101, E-102 and E103 for the hydrate inhibition. Three MEG injection stream INJ-1, INJ-2 and INJ-3 are connected at MIX-103, MIX-104 and MIX-109 respectively. After separation at LTS, the gas stream goes to


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the shell side of the same pre-cooling gas/gas exchanger (E-101 & E-102) for heat exchange and finally go to the pipeline. The hydrocarbon stream from V-103 is first heated by a heater (E-106, propane refrigerant used as the utility) at 26.7 ˚C and then goes to V-104 at 21 bar and the MEGwater mixture is sent to stripper column (C-101) for regeneration.

Figure 3. Flowsheet of the NG dehydration Process.

In the stripper column (C-101), the MEG-water mixture (T7 stream, 279.25 kg/hr) of -23.4 ˚C exchanges heat with the column overhead vapor (VAP-2 stream, 231.56 kg/hr) of 109.7 ˚C to maintain column overhead temperature (T2 stream) at 100.0˚C and T9 stream temperature becomes 86.6 ˚C. There is further heat exchange from the total MEG-water mixture with the bottom flow of stripper column. After heat exchange at E-107, T9 stream of 86.6 ˚C connected at MIX-109 and mixed stream temperature is 9.2 ˚C heated up to 46.1 ˚C by BTM-1 stream which is 125.9 ˚C. E-104 and E-105 is connected through a heat stream to utilize the heat of E-105. After that, the MEG-water mixture is sent to a flash separator (V-105) at 6.18 bar and 46.1˚C for flashing off hydrocarbons that are carried over. From the flash separator, the MEG-water mixture goes to the stripper column as feed at 1.14 bar and 46.2˚C. In this column, the MEG-water is concentrated to 82 wt% of MEG and re-injected to injection point of heat exchangers E-101, E-102 and E-103 via a MEG transfer pump (P-101). In the MP Separator (V-104), where all hydrocarbon streams are connected, hydrocarbon vapor is generated as the pressure is reduced from 90 bar to 21 bar. The pressure of the hydrocarbon liquid just left the separators (V-101, V-102, and V-103) is 90 bar, then its pressure is reduced by valves and dumped in to V-104 at 21 bar. This vapor is used as the fuel gas for plant utility and a portion of this fuel gas is utilized as stripping gas for the stripper



column. The stripping gas is directly connected at the fire tube reboiler (R-101) of the column at a rate of 34.15 kg/hr to enhance the MEG separation. The stripping gas (any inert or dehydrated gas) can be introduced directly into the reboiler or between the reboiler and surge tank. In theory, adding gas to a packed unit between the reboiler and surge tank is superior and will result in lower stripping gas rates7, 10, 18, 19.

3.2 Steady-state model development A plant-wide steady-state model is developed with Aspen Plus V8.8 20 based on the existing plant PFD and operating conditions. The thermodynamic property method used for this simulation is the Cubic Plus Association (CPA) equation of state 21. This equation of state (EOS) is similar to the Soave-Redlich-Kwong (SRK) EOS and can handle polar effects between molecules. CPA is now popular for accurately representing glycol dehydration system, methanol and mercury partitioning system 22. The composition and condition of the feed is taken from existing process, which is shown in Table 1. In Table 2, the pressure and temperature specifications of the separators and heat exchangers at normal steady-state conditions are tabulated. All heat exchangers in the process, including stripper column reboilers duty, are also tabulated in Table 3. Required duty of both Gas/Gas exchangers (E-101 & E-102) are 1.35×107 kJ/hr, chiller is 8.76×105 kJ/hr and stripper column reboiler is 6.67×105.

Table 1. Feed Specifications of the NG Dehydration Process

Table 2. Normal Steady-state Operating Conditions of NG Dehydration Process

Table 3. Unit Heat Duty Requirements under Steady-state Operating Condition 8 ACS Paragon Plus Environment

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The Radfrac model template in Aspen Plus is used to model the stripper packed column. The column operates at atmospheric pressure. The MEG-water mixture enters at the top of the column and mass transfer with vapor of reboiler occurs in counter-current. The stripping gas is added in the reboiler to enhance mass transfer in the column. An external heat exchanger and reflux drum is used for better control of column overhead temperature instead of built-in model. MEG is recovered as bottom product and the reboiler duty is 6.67×105 kJ/hr. The separated water, along with stripping gas, is flared through vent line. Lean glycol (bottom flow of the column) is then injected to the injection point via MEG transfer pump (P-101). An external condenser and reflux drum is used to simulate the reflux coil effect of existing plant and liquids from reflux drum are sent back to the column. This strategy gives an opportunity to apply a more realistic and effective control strategy in the dynamic model. Aspen’s packing sizing tools are used for sizing the stripper column and also performed calculations according to heuristics described in the literature 16, 17, details of equipment sizing are described in section 5.1. Table 4 shows the detailed results of this model.

Table 4. Design Data of the Stripper Column

The Rich MEG stream (MEG-water mixture stream contains high wt% of water) is utilized to control column overhead vapor temperature and lean MEG (contains less wt% of water, bottom flow of the column, showed as red dotted line) temperature. Lean MEG temperature is cool down from 125.9˚C to 93.3˚C at injection point, here MEG temperature is kept higher than process gas temperature (32.2˚C) to control foaming problem

10, 14.

In the simulation, glycol water stream is



directly used in the tube side of the external condenser of the column, for simplicity purpose and heat stream (red dotted line) is connected from E-105 to E-104. With the developed steady-state model, the MEG carried over from the LTS is 0.018 kg/hr and from stripper column overhead is 0.913 kg/hr. MEG loss from flash separator is found to be very small compared to LTS and stripper column and is neglected. Thus, the total MEG loss from the system at normal operating condition is 0.931 kg/hr. Operating conditions and simulation results of this model are closely matched with existing process; e.g., the MEG loss rate is within 2% error which validates this steady-state model. Based on this result, a dynamic simulation is developed to understand the plant dynamics, which can be utilized to minimize this loss and also to handle any upset conditions by employing improved control strategy.

4. Simulation-based Root Cause Analysis for MEG Loss

Five upset scenarios usually occurred in plants have been investigated based on the developed simulation model, which include (1) feed flow variations, (2) chiller temperature variations, (3) stripper overhead temperature variations, (4) reboiler temperature variations, and (5) stripping gas flow variations. For all tests, both reboiler (R-101) and condenser (E-107) duties have been kept constant as normal plant behaviors to observe the temperature effect as well as the glycol loss through the vapor line (T3 stream).

4.1 Feed flow variations Inlet gas flow is varied from 12450.72 kmol/hr to 8715.72 kmol/hr to investigate the amount of MEG carried over or MEG loss by vaporization. Figure 4. shows the MEG loss tendency


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for reduced flow. In particular, Figure 4(a) shows the MEG loss from LTS and Figure 4(b) shows the MEG loss from column overhead. For the varied gas flow rate, MEG carried over with gas increases from 0.0120 kg/hr to 0.0136 kg/hr while MEG carried over with hydrocarbon increases from 0.0063 kg/hr to 0.0072 kg/hr. At the same flow fluctuation, MEG losses from column overhead increases from 0.913 kg/hr to 1.008 kg/hr. Both plots show a tendency of increasing MEG losses with decreasing inlet gas flow rates. The major reason is that the feed flow reduction caused the volume of separated MEG-water mixture at LTS is less compared to the normal expectation. As a result, less MEG-water mixture will flow from the LTS to the stripper column. As reboiler duty is kept the same, more MEG will be boiled up to the column top, which leads to the increased MEG loss and improper MEG concentration14.

Figure 4. MEG losses from (a) LTS and (b) stripper column overhead under NG feed flow variation.

4.2 Chiller temperature variations Figure 5 shows the chiller temperature fluctuation results. In this step, inlet gas flow is constant, and the chiller temperature varied from -27.8˚C to -16.1 ˚C and the respective mass flow rate of MEG from LTS and column overhead is recorded. From Figure 5(a), it is observed that MEG carried over with gas increases from 0.021 kg/hr to 0.068 kg/hr and MEG carried over with hydrocarbon increases from 0.011 kg/hr to 0.025 kg/hr. For the same temperature fluctuation, Figure 5(b) shows that the MEG loss increases in the column overhead from 0.186 kg/hr to 0.198 kg/hr. It is observed that at -17.8 ˚C, MEG rate drops from 0.1984 kg/hr to 0.1978 kg/hr because the reboiler temperature drops due to increasing feed rate of the stripper column. From the figure it is clear that both plots have tendency of MEG losses with decreasing chiller temperature. A



specific chiller temperature (-26.0˚C) is necessary for liquid recovery, dew point control and smooth column operations. If refrigeration system breakdown, separation efficiency will decrease at LTS and more MEG will be carried over with gas and hydrocarbon due to its hygroscopic nature and complete miscibility with water10. This tendency clearly observed from comparative analysis of Figure 4(a) and Figure 5(a), where the total carried over MEG is 0.018 kg/hr for the flow variation case and 0.032 kg/hr for the chiller temperature variation case. This improper separation affects the MEG-water mixture flowrate, which leads the condenser and reboiler temperature upsets at the stripper column.

Figure 5. MEG losses from (a) LTS and (b) stripper column under the chiller temperature variation.

4.3 Stripper overhead temperature variations For good separation of MEG from water, it is necessary to maintain 100.0˚C at the column overhead, which is the boiling point of water at 1 atm. Above this temperature, MEG is carried over with water and as a result MEG loss increases. Below this temperature, water accumulates in the column and this affects the Lean MEG concentration. Figure 6. shows this scenario clearly. Figure 6(a) shows that the MEG loss increases with increasing column overhead temperature and at the same time Figure 6(b) shows that MEG concentration is decreasing. So, maintaining the operating point of 100.0˚C and 82% glycol concentration can lead to a smooth operation and minimize MEG losses. The required MEG concentration for plant operation is 80~82 wt%. At this range, MEG’s freezing point is about -48.0˚C. As the plant is operating at -26.0˚C, there is enough safety margin for the plant, otherwise there is a chance of pipe blockage by MEG if plant is operated below this concentration level 23. 12 ACS Paragon Plus Environment

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Figure 6. (a) MEG loss and (b) MEG mass fraction at the stripper column bottom under the stripper overhead temperature variation.

4.4 Reboiler temperature variations The operating point for the reboiler is 125.9˚C, which maintains 82 wt% MEG concentrations. If the reboiler temperature fluctuates due to inadequate control system, it affects MEG concentration and column overhead temperature and as a result it increases the MEG losses through column overhead, as shown in Figure 7. Upset conditions like gas flow rate variations, chiller temperature variations and stripping gas rate have an impact on reboiler temperature. A good control system can handle this upset condition.

Figure 7. MEG loss at the stripper column overhead under reboiler temperature variation.

4.5 Stripping gas flow variations Stripping gas process is regarded as the simplest and the most common route to obtain enhanced MEG regeneration performance. Introduction of stripping gas generally reduces the partial pressure of the water and leads to an increase in the final purity on the regenerated MEG 18. It is a physical separation process that works according to the thermodynamic principle of Raoult’s law. It relies on the use of a vapor stream (e.g., dry product gas) to remove one or more component from the liquid stream (e.g. water from MEG). The Raoult’s law of water is given as:

𝑥𝑤𝑎𝑡𝑒𝑟 =

𝑦𝑤𝑎𝑡𝑒𝑟 𝑃 𝑃𝑆𝑎𝑡 𝑤𝑎𝑡𝑒𝑟

=

𝑊𝑎𝑡𝑒𝑟 𝑃𝑎𝑟𝑡𝑖𝑎𝑙 𝑃𝑟𝑒𝑠𝑠𝑢𝑟𝑒 𝑊𝑎𝑡𝑒𝑟 𝑆𝑎𝑡𝑢𝑟𝑎𝑡𝑖𝑜𝑛 𝑉𝑎𝑝𝑜𝑟 𝑃𝑟𝑒𝑠𝑠𝑢𝑟𝑒


(1)


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where 𝑥𝑤𝑎𝑡𝑒𝑟 is the mole fraction of water in liquid phase, 𝑦𝑤𝑎𝑡𝑒𝑟 is the mole fraction of water in vapor phase, P is the system pressure, and 𝑃𝑆𝑎𝑡 𝑤𝑎𝑡𝑒𝑟 is the saturated pressure of water. The fugacity and the activity coefficients are absent in Eq. (1) because it is assumed that both liquid and vapor phases are in ideal state. This assumption is valid for the purpose of demonstrating the working principle of stripping gas mechanism although the mixture of water, MEG and methane are non-ideal in reality. At a given pressure and temperature in reboiler, the Raoult’s law of water (Eq. (1)) indicates the direct proportionality between the mole fraction of water in liquid phase and the partial pressure of water (i.e., 𝑥𝑤𝑎𝑡𝑒𝑟 decreases with reduction in 𝑦𝑤𝑎𝑡𝑒𝑟 𝑃 or vice versa). Next, the total pressure of the system (P) is given by the sum of partial pressure of all components as shown: 3

𝑃 ≈ ∑𝑖 = 1𝑦𝑖𝑃 = (𝑦𝑀𝐸𝐺𝑃) + (𝑦𝑤𝑎𝑡𝑒𝑟𝑃) +(𝑦𝑀𝑒𝑡ℎ𝑎𝑛𝑒𝑃)

(2)

Equation (2) is valid for any gas mixture but only methane is used in this case for the sake of simplicity. The partial pressure of MEG (𝑦𝑀𝐸𝐺𝑃) can be neglected from Equation (2) due to the higher vapor pressure relative to MEG. When the mixture is heated, much more water evaporates than MEG 24. Thus, Eq. (2) can be reduced to the following: 3

𝑃 ≈ ∑𝑖 = 1𝑦𝑖𝑃 = (𝑦𝑤𝑎𝑡𝑒𝑟𝑃) +(𝑦𝑀𝑒𝑡ℎ𝑎𝑛𝑒𝑃)

(3)

At constant pressure P (e.g., reboiler condition), an increase in the partial pressure of methane leads to the decrease in the partial pressure of water (Eq. (3)). This translates to a lower mole fraction of water in the liquid phase (since mole fraction of water is directly proportional to the partial pressure of water based on Raoult’s law Eq. (1)). Thus, stripping gas injection reduces the mole fraction of water in MEG and leads to a higher purity of the regenerated MEG streams.


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The composition of the stripping gas is shown in Table 5. Figure 8(a) shows the impact of stripping column reboiler duty and bottom MEG mass fraction under changes of the stripping gas flowrate. In this case inlet gas flow, the condenser duty and reboiler duty is kept constant. The observation shows that the higher the stripping gas flowrate the higher the concentration will be. Figure 8(b) shows the impact of the reboiler duty and MEG mass fraction again under the variations of the inlet gas flowrate. From these two figures, it can be seen that the synergic control of both stripping gas flowrate and inlet gas flowrate is very important.

Table 5. Composition of the Stripping Gas Fed to the Striping Column

Figure 8. Impacts of stripping column reboiler duty and bottom MEG mass fraction under changes of (a) the stripping gas flowrate and (b) the inlet gas flowrate for the base case.

Based on root cause analysis, the major findings include: (1) at the steady-state normal operating condition, the total MEG loss is 22.35 kg/d; (2) MEG carried over with gas and hydrocarbon is larger for the chiller temperature variation case (0.032 kg/hr) compared to the flow variation case (0.018 kg/hr); (3) The stripping gas flowrate has significant impact on MEG regeneration system and thus needs to be well regulated; and (4) most of the MEG losses are coming from upset operations of the stripper column. Therefore, an improved control strategy for the stripper column is needed.

5. Plant-wide Dynamic Modeling



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5.1 Equipment sizing Process equipment sizing is performed according to heuristics described in the literature 16, 17.

Feed and operating conditions are obtained from the industrial plant and the thermodynamic

properties are obtained from steady-state simulations. Two and three-phase separators are sized based on terminal velocity estimated from Stoke law’s

25.

The stripper column diameter is

calculated from an estimated flooding velocity and flooding velocity estimates from correlation of Leva16 and column height is obtained using heuristics and correlations17. The column diameter equation is: 4𝐺

𝐷𝑇 =

(4)

(𝑓𝑈𝑓)𝜋𝜌𝐺

where, 𝐷𝑇 is packed tower diameter, 𝐺 is mass flow rate of vapor, 𝑓 is a fraction of vapor flooding velocity, 𝑈𝑓 is vapor flooding velocity and 𝜌𝐺 is vapor density. The Leva flooding correlation fits the following equation: 𝑌 = exp [ ― 3.7121 ― 1.0371(𝑙𝑛𝐹𝐿𝐺) ―0.1501(𝑙𝑛𝐹𝐿𝐺)2 ―0.007544(𝑙𝑛𝐹𝐿𝐺)3]

(5)

where, 𝐹𝐿𝐺 is flow ratio parameter and computed as: 𝐿 𝜌𝐺

𝐹𝐿𝐺 = 𝐺

(6)

𝜌𝐿

where, 𝐿 and 𝐺 are mass flow rate of liquid and vapor. 𝜌𝐺 and 𝜌𝐿 are density of vapor and liquid respectively. 𝑌=

𝑈2𝑓𝐹𝑃 𝑔

(

𝜌𝐺

)𝑓(𝜌 )𝑓(𝜇 )

𝜌𝐻2𝑂(𝐿)

𝐿

(7)

𝐿

where, 𝑌 is flooding velocity factor, 𝐹𝑃 is packing factor. The density and viscosity functions are given by following equations:

(

𝑓(𝜌𝐿) = ―0.8787 + 2.6776

)

𝜌𝐻2𝑂(𝐿) 𝜌𝐿

―0.6313(

𝜌𝐻2𝑂(𝐿) 2 𝜌𝐿

)


(8)

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𝑓(𝜇𝐿) = 0.96𝜇0.19 𝐿

(9)

The area for the heat exchangers is estimated using recommended values for heat transfer coefficients 26 and the heat duties are obtained from simulations. The theoretical power for the pump is extracted from process simulations and efficiencies are applied to consider a more realistic performance. The results of equipment sizing are shown in Table 6.

Table 6. Equipment Sizing Data

5.2 Dynamic modeling of the base case Two dynamic models are developed for this study, one is for base case, which reflects the same control strategy as existing plant, and the second is for the improved case, which uses three additional controllers on the stripper column. These additional controller loops are developed based on the results of the base case model. Conventional PI controllers are used in all control loops 27, 28. Figure 9 shows the plantwide control structure of the base case model. For the base case, stripper column is equipped with pressure controller, flow controller, level controller and temperature controller. The controlled and manipulated variables for stripper column for the base case are listed in Table 7.

Figure 9. Plant-wide dynamic model of the base case.

Table 7. CVs and MVs of all Control Loops for the Base Case

6. New Control Strategy and Demonstration



6.1 Comparison between the base and improved cases For the base case, the control strategy is taken from existing process, where a flow controller is used for each well flow rate control (inlet flow rate), and a level and pressure controller for each separator. For the stripping column, the pressure, temperature, and level controller are setup for maintain the column top pressure, temperature profile, and sump level, respectively. The stripped stream from the column overhead and the MEG-water flow from LTS meet at the condenser for heat exchange. Figure 10 shows a portion of the base control strategy for the separator section (See, Figure 10(a)) and the stripper column section (See, Figure 10(b)).

Figure 10. Control strategy for the (a) separator and (b) stripper column sections of the base case.

For the improved case, the inlet natural gas flow rate and separator control strategies are the same as the base case. It has been observed from the steady-state simulation that the stripper column in the base case is unstable, which leads to the MEG loss and also operating upset conditions. Thus, an improved control strategy for the stripper column is needed to minimize the MEG losses at normal operating condition with upset conditions. For this purpose, three new controller loops are added as shown in Figure 11:  A new temperature controller (E-107_TC) is added for the column overhead temperature control by manipulating the glycol-water flow rate. In the base case, the temperature control was maintained by operating a manual bypass valve, which potentially leads to large temperature fluctuations on the column overhead temperature.  A new ratio controller (C-101_FC) is added to control the stripping gas flow rate to the stripping column. For this modification, the stripping gas flow rate will change 18 ACS Paragon Plus Environment

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correspondingly with the flow rate upset of the glycol mixture to the striping column. This stripping gas regulation will help maintain the desired MEG concentration and column overhead temperature better.  A new cascade controller (C-101_CC) added to control the MEG concentration by manipulating the reboiler duty, while in the base case there is no such a composition controller. This will give better control on the MEG concentration.

Figure 11. Improved control strategy for the stripper column section.

6.1.1 Feed flow rate variations The inlet gas flow rate experiences a series of step decreases starting from 4,233.32 kmol/hr to 500 kmol/hr; and then experiences another series of step increases from 500 kmol/hr to 4,233.32 kmol/hr to observe the effectiveness of all added controllers for the improved case compared with the base case. Figure 12 shows the simulation results of various streams for the stripper column for both cases. Figure 12(a) shows the results of the overhead temperature response of the stripper column. It is observed that the base control strategy cannot maintain the overhead temperature to the specification of 100.0˚C under the designated feed flowrate variations; while the improved control strategy maintains this temperature fairly well. Figure 12(b) shows the results of the improved case; the MEG mass flow rate from the stripper column overhead decreases from 0.55 to 0.43 kg/hr when the inlet gas flowrate step decreases from for 4,233.32 to 500 kmol/hr, and then back to 0.59 kg/hr when the inlet gas flowrate is restored. However, for the base case, the MEG mass flow rate decreases from 0.92 to 0.63 kg/hr, and then back to 0.92 kg/hr under the same experience of the inlet gas flowrate upset. The shadowed area between the two curves in Figure



12(b) indicates the total savings of MEG loss from the base case control strategy upgraded to the improved control strategy, which is about 37%. Such a saving is significant.

Figure 12. Dynamic performance comparison for the base and improved cases under the feed flow variation: (a) stripper overhead temperature; (b) stripper overhead MEG loss; (c) MEG loss from LTS; (d) stripper bottom MEG mass fraction; and (e) stripper feed and bottom flowrates.

Figure 12(c) shows the MEG carryover flowrate with the outlet gas and hydrocarbon streams from the LTS. For the base case, MEG carryover flowrates for these two streams is 0.011 kg/hr and 0.0095 kg/hr, respectively. For the improved case, however, the MEG carryover flowrates are just 0.0057 kg/hr and 0.0049 kg/hr, respectively. Figure 12(d) gives the stripper reboiler temperature and MEG concentration trend for both cases under the same inlet gas flow rate variation. For the base case, the reboiler temperature is fairly constant, which causes the MEG concentration of the bottom stream having significant changes associated with the inlet gas flow rate changes. For the improved case, however, as the composition controller is cascaded to the reboiler duty controller, the MEG composition of the bottom stream could be well maintained and the reboiler duty could smartly vary with the inlet gas flow rate changes. Figure 12(e) shows the stripper column inlet and bottom flow flowrates for both cases. It shows better flow control for the improved case. In summary, Figure 12 demonstrates clearly that the improved control strategy can handle the flowrate variation condition appropriately and thus provide plant improved stability and reduce the MEG losses.

6.1.2 Chiller temperature variations


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The temperature of the chiller (See, E-103 in Figure 9) was set to experience a series of step changes from -26.0 ˚C to 0 ˚C to observe the controller effectiveness of all added controllers for the improved case compared with the base case. Figure 13(a) shows the response of chiller temperature for both cases. For the base case, the column overhead temperature fluctuated from 100.0˚C to 106.4˚C, which leads to an unstable operation and a cause for improper separation of MEG-water mixture. For the improved case, however, the newly added temperature controller overcome this situation with just very minor fluctuations.

Figure 13. Dynamic performance comparison for the base and improved cases under the chiller temperature variation: (a) stripper overhead temperature; (b) stripper overhead MEG loss; (c) MEG loss from LTS; (d) stripper bottom MEG mass fraction; and (e) stripper feed and bottom flowrates.

Figure 13(b) shows the MEG flowrate from stripper column overhead under the same temperature variation for the base case as well as the improved case. When there is a large fluctuation at the column overhead flow rate (i.e., the MEG rate is changed from 0.40 kg/hr to 1.12 kg/hr) the base case with the previous controllers takes a long time to reach the desired set-point; whereas the controller performance is much better for the improved case. Figure 13(c) shows the MEG carryover flowrate with the outlet gas from LTS. It is observed that MEG carryover flowrate fluctuated during the entire the variation test for the base case; however, in the improved case, the MEG carryover flowrate reaches the maximum 0.14 kg/hr approximately at 3.74 hr, which is much less than that of the base case under this variation, and even less than that under the normal operation condition of the base case. Thus, the improved controller strategy works well in the entire range of temperatures considered from 0˚C to -26.0˚C and the resultant MEG loss is much



less than the base case. Figure 13(d) shows the MEG concentration of bottom flow of the stripper column for both cases under the same temperature variation. For the base case, there is a significant drop in MEG concentration whereas the composition controller in the improved case is able to maintain the MEG concentration at the desired set-point. Figure 13(e) shows that in the base case, there are decreases in both column feed and bottom flowrates as an effect of the temperature variation, which causes a drop in reboiler temperature and thus a drop of the bottom flow MEG concentration (See, Figure 13(d)). Again, in the improved case, it demonstrates that the better control strategy keeps the MEG concentration at the desired level. It should be noted that out of five upset scenarios, the handling of the most two significant ones (feed flow and chiller temperature variation) with the developed control strategy with dynamic simulations have been shown. The other three upset scenarios including condenser temperature variation, reboiler temperature variation and stripping gas flow variation are also well controlled by the improved control strategy with the newly added condenser temperature controller (E-107_TC), composition controller (C-101_CC) and ratio controller (C-101_FC).

7. Conclusions

In this work, a new control strategy is developed for MEG loss minimization under upset conditions in a natural gas dehydration plant. First, the root causes of MEG loss are identified via simulation of a plant-wide steady-state model for natural gas dehydration process under upset conditions. Second, a plant-wide dynamic model is developed to study the dynamic behavior of the process. Through simulations, it is determined that most of the MEG is lost through the stripper column, which has an inadequate control strategy. An improved control strategy is developed,


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which includes the ratio control of stripping gas, composition control of bottom flow, and temperature control of the stripper column overhead. It is demonstrated that the improved control strategy works very well under various upset conditions and has the potential to reduce MEG losses by 37% under normal operating conditions. This shows the potential to reduce operating cost and MEG emissions of the plant, as well as improve the plant operability and product quality.

Acknowledgements

This work is supported by Graduate Student Scholarship and Anita Riddle Faculty Fellowship from Lamar University.



Abbreviations

CERE

Center of Energy Resources Engineering

CPA

Cubic Plus Association

CV

Controlled Variable

DEG

Di Ethylene Glycol

DTU

Technical University of Denmark

EOS

Equation of State

HP

High Pressure

LTS

Low Temperature Separator

MEG

Mono Ethylene Glycol

MP

Medium Pressure

MV

Manipulated Variable

SRK

Soave-Redlich-Kwong

TEG

Tri Ethylene Glycol

TREG

Tetra Ethylene Glycol

PFD

Process Flow Diagram

P&ID

Piping & Instrumentation Diagram

LNG

Liquefied Natural Gas

NGL

Natural Gas Liquid

ppm

Parts per Million


Page 24 of 52

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Nomenclature

Sets and indices i

index of components

Variables 𝐷𝑇

packed tower diameter (m)

𝑓

fraction of vapor flooding velocity

𝐹𝑃

packing factor (m2/m3)

𝐹𝐿𝐺

flow ratio

𝑔

gravitational acceleration (m/s2)

𝐺

mass flow rate of vapor (kg/hr)

𝐿

mass flow rate of liquid (kg/hr)

𝜇𝐿

viscosity of liquid (kg/m-s)

𝜌𝐺

vapor density (kg/m3)

𝜌𝐿

liquid density (kg/m3)

𝜌𝐻2𝑂(𝐿) density of water in liquid phase (kg/m3) 𝑈𝑓

vapor flooding velocity (m/s)

𝑥𝑤𝑎𝑡𝑒𝑟

mole fraction of water in liquid phase

𝑦𝑤𝑎𝑡𝑒𝑟

mole fraction of water in vapor phase

𝑃𝑠𝑎𝑡 𝑤𝑎𝑡𝑒𝑟

saturated pressure of water (bar)

𝑦𝑀𝐸𝐺𝑃

partial pressure of MEG (bar)

𝑦𝑤𝑎𝑡𝑒𝑟𝑃

partial pressure of water (bar)

𝑦𝑀𝑒𝑡ℎ𝑎𝑛𝑒𝑃

partial pressure of methane (bar)



𝑌

flooding velocity factor


Page 26 of 52

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References

(1) Mokhatab, S.; Poe, W.A.; Speight, J.G. Handbook of Natural Gas Transmission and Processing. Gulf Professional Publishing, United Kingdom, 2006. (2) Schoell, M. Isotope techniques for tracing migration of gases in sedimentary basins. J. Geol. Soc. Lond. 1983, 140, 415-422. (3) Tabak, J. Natural Gas and Hydrogen. Facts on File, USA, 2009. (4) Rojey, A.; Jaffret, C.; Cornot-Gandolphe, S.; Durand, B.; Julian, S.; Valais, M. Natural Gas: Production, Processing, Transport. Editions Technip, Paris, 1997. (5) EIA, Energy Information Administration. Natural Gas Processing: The Crucial Link between Natural Gas Production and its transportation to Market, 2006, 1-11. (6) Speight, J.G. Natural Gas: A Basic Handbook. Gulf Publishing Company, Houston, Texas, 2007. (7) Kidnay, A.J.; Parrish, W.R.; McCartney, D.G. Fundamentals of Natural Gas Processing, 2nd Ed., CRC Press, Taylor and Francis Group, LLC., 2011. (8) Mahmud, M.A.; Mazumder, M.; Xu, Q.; Khan, R.I. Sloshing Impact on Gas Pretreatment for LNG Plants Located in a Stranded Offshore Location. Ind. Eng. Chem. Res., 2018, 57, 5764-5775. (9) Speight, J.G. Gas Processing: Environmental Aspects and Methods. Butter-worthHeinemann Ltd., Oxford, England, 1993. (10) Campbell, J.M. Gas Conditioning and Processing, 8th ed., John Campbell and Company, Vol. 2., Norman, Oklahoma, 2004. (11) Gandhidasan, P.; Al-Farayedhi, A.; Al-Mubarak. 2001. A Dehydration of natural gas using solid desiccants. Energy. 2001, 26(9), 855-868. 27 ACS Paragon Plus Environment


(12) Netusil, M.; Ditl, P. Comparison of three methods for natural gas dehydration. J. Nat. Gas. Chem. 2011, 20, 471-476. (13) Rouzbahani, A.N.; Bahmani, M.; Shariati, J.; Tohidian, T.; Rahimpour, M.R. Simulation, optimization, and sensitivity analysis of a natural gas dehydration unit. J. Nat. Gas. Sci. Eng. 2014, 2, 159-169. (14) Haque, M.E. Ethylene Glycol Regeneration Plan: A Systematic Approach to Troubleshoot the Common Problems. J. Che. Eng. IEB, Vol. ChE 27, No 1, 2012, 21-26. (15) Francis, A.W. Low Temperature Separation Systems, US Patent. 2,758,665. 1956. (16) Leva, M. Reconsider Packed-Tower Pressure-Drop Correlations, Chem. Eng. Prog., 1992, 88(1), 65-72. (17) Seader, J.; Henley, J. Separation Process Principles, 3rd ed. Wiley, 2011. (18) Piemonte, V.; Maschietti, M.; Gironi, F. A triethylene glycol-water system: a study of the TEG regeneration process in natural gas dehydration plants. Energy Sources, Part A Recover. Util. Environ. Eff 34, 456-464. http://dx.doi.org/10.1080/ 2012. (19) Mokhatab, S.; Poe, W.A.; Mak, J.Y. Handbook of Natural Gas Transmission and Processing. Elsevier Inc., http://dx.doi.org/10.1016/b978-0-12-801499-8.00007-9/ 2015. (20) AspenTech, Inc. Aspen physical property system. Aspen Plus V8.8, 2015. (21) Santos, L.C.D.; Abunaham, S.S.; Tavares, F.W. 2015. Cubic Plus Association Equation of State for Flow Assurance Projects. J. Ind. Eng. Che. Res. 2015, 6812-6824. (22) Dyment, J. Optimizing the dehydration process with advanced process simulation. Hydrocarbon Process. Nov. 2017, 63-64. (23) Gas Processors Suppliers Association, Engineering Data Book, 12th ed., Gas Processing Supply Association, Tulsa, Oklahoma, 2004.


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(24) Anyadiegwu, C.I.C.; Kerunwa, A.; Oviawele, P. Natural Gas Dehydration using Triethylene Glycol (TEG). Petroleum and Coal. 2014, 56(4): 407-417. (25) Stewart, M.; Arnold, K. Gas-Liquid and Liquid-Liquid Separators, Elsevier, Burlington (MA), 2008. (26) Seider, W.D.; Seader, J.; Lewin, D.R. Product and Process Design Principles, 3rd ed. Wiley, Danvers, MA, 2009. (27) Luyben, W.L. Principles and Case Studies of Simultaneous Design, John Wiley & Sons, Inc., New Jersey, 2011. (28) Luyben, W.L. Distillation Design and Control using Aspen Simulation, John Wiley & Sons, Inc., New Jersey, 2006.



Page 30 of 52

Table of Contents

Improved Control Strategy Validated by Dynamic Simulation

+ Process Upsets

Glycol Loss Minimization

Natural Gas Dehydration Plant

LIST OF TABLES

Table 1.

Feed Specifications of the NG Dehydration Process

Table 2.

Normal Steady-state Operating Conditions of NG Dehydration Process

Table 3.

Unit Heat Duty Requirements under Steady-state Operating Condition

Table 4.

Design Data of the Stripper Column

Table 5.

Composition of the Stripping Gas Fed to the Stripping Column

Table 6.

Equipment Sizing Data

Table 7.

CVs and MVs of all Control Loops for the Base Case


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Page 32 of 52

Table 1. Feed Specifications of the NG Dehydration Process Well-1

Well-2

Well-3

Well-4

Pressure (bar)

90

90

90

90

Temperature (˚C)

32.2

32.2

32.2

32.2

Flow Rate (kmol/hr)

1743.13

2241.17

4233.32

4233.32

Nitrogen (mol%)

0.18

0.16

0.22

0.12

CO2 (mol%)

0.09

0.05

0.08

0.07

Methane (mol%)

92.10

92.96

95.19

94.92

Ethane (mol%)

3.49

3.40

2.55

2.43

Propane (mol%)

1.53

1.41

0.39

0.56

i-Butane (mol%)

0.39

0.33

0.32

0.34

n-Butane (mol%)

0.45

0.30

0.06

0.14

i-Pentane (mol%)

0.22

0.17

0.10

0.17

n-Pentane (mol%)

0.13

0.10

0.03

0.13

n-Hexane (mol%)

0.16

0.12

0.11

0.22

n-Heptane (mol%)

0.31

0.25

0.22

0.26

n-Octane (mol%)

0.33

0.26

0.25

0.20

n-Nonane (mol%)

0.18

0.13

0.12

0.10

n-Decane (mol%)

0.10

0.07

0.06

0.05

n-Undecane (mol%)

0.05

0.04

0.02

0.03

n-Dodecane (mol%)

0.14

0.09

0.08

0.07

H2O (mol%)

0.17

0.18

0.20

0.20


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Table 2. Normal Steady-state Operating Conditions of NG Dehydration Process Equipment Name

Index

Operating Pressure

Operating Temperature

(bar)

(Inlet/Outlet, ˚C)

HP Separator

V-101

90

32.2/32.2

HP Separator

V-102

90

32.2/32.2

LTS

V-103

90

-26.1/-26.1

MP Separator

V-104

21.01

26.7/26.7

Flash Separator

V-105

6.18

46.1/46.1

Reflux Drum

V-106

1.15

100/100

Gas/Gas Exchanger

E-101

Hot Stream: 90

Hot Stream: 32.5/-7.6

Cold Stream: 90

Cold Stream: -26.1/10.3

Hot Stream: 90

Hot Stream: 32.5/-7.6

Cold Stream: 90


Gas/Gas Exchanger

E-102

Chiller

E-103

90

-7.4/-26.1

Glycol/Glycol Exchanger

E-104

6.87

9.2/45.9

Glycol/Glycol Exchanger

E-105

1.15

125.9/85.6

Heater

E-106

21.01

-44.3/26.7

Condenser

E-107

Hot Stream: 1.15

Hot Stream: 109.7/100

Cold Stream: 6.87


1.14

46.2/125.9

1.15/1.17

109.7/125.9

1.17/90

85.6/93.3

Stripper Column 

C-101

Top/Bottom

MEG Transfer Pump  Injection rate (m3/hr)

P-101

0.8



Table 3. Unit Heat Duty Requirements under Steady-state Operating Condition Equipment’s Duty

Simulation Results

(kJ/hr) E-101

1.35×107

E-102

1.35×107

E-103

-1.56×107

E-104

1.01×105

E-105

-1.01×105

E-106

8.76×105

E-107

8.21×104

Reboiler

6.67×105


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Table 4. Design Data of the Stripper Column Items

Specification

Section Starting Stage

1

Section Ending Stage

2

Column Diameter (m)

0.27072

Section Height (m)

1.143

Maximum Fractional Capacity

0.62

Maximum Capacity Factor (m/sec)

0.06215

Section Pressure Drop (bar)

0.005697

Average Pressure Drop/Height (mm-water/m)

45.94

Maximum Stage Liquid Holdup (m3)

0.0019

Maximum Liquid Superficial Velocity (m/sec)

0.005697

Surface Area (m2/m3)

205

Void Fraction

0.94

Packing Material and Size (mm)

Pall Ring-Metal, 25

Packing Factor (m-1)

157



Table 5. Composition of the Stripping Gas Fed to the Striping Column Component

Mol. %

CO2

0.10

Methane

88.6

Ethane

6.60

Propane

2.30

i-Butane

0.90

n-Butane

0.50

i-Pentane

0.30

n-Pentane

0.20

n-Hexane

0.20

n-Heptane

0.20


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Table 6. Equipment Sizing Data Index

Description

Dimension

V-101

HP Separator-Horizontal

96” ID × 27’-0” L

V-102

HP Separator-Horizontal

66” ID × 21’-0” L

V-103

LTS-Horizontal

96” ID × 20’-0” L

V-104

MP Separator: Vertical

42” ID × 7.75’-0” L

V-105

Flash Separator-Vertical

30” ID × 6’-0” H

V-106

Reflux Drum: Vertical

30” ID × 6’-0” H

C-101

Stripper Column

10.56” ID × 3.75’-0” H



Table 7. CVs and MVs of all Control Loops for the Base Case Control Loop Index

CV

MV

WELL-1_FC

Well-1 Flowrate

NG flowrate of Well-1

WELL-2_FC

Well-2 Flowrate


WELL-3_FC

Well-3 Flowrate


WELL-4_FC

Well-4 Flowrate


V-101_PC

HP Separator Pressure

Gas flowrate leaving from the separator

V-101_L1LC

HP Separator Hydrocarbon

Hydrocarbon flowrate of the

Liquid Level

separator

V-101_L2LC

HP Separator Water Level

Water flowrate of the separator

V-102_PC

HP Separator Pressure

Gas flowrate leaving from the separator

V-102_L1LC

HP Separator Hydrocarbon

Hydrocarbon flowrate of the

Liquid Level

separator

V-102_L2LC

HP Separator Water Level

Water flowrate of the separator

V-103_PC

LTS Pressure

Gas flowrate leaving from the LTS

V-103_L1LC

LTS Hydrocarbon Liquid Level

Hydrocarbon flowrate of the LTS

V-103_L2LC

LTS MEG-Water Level

MEG-Water flowrate of the LTS

V-104_PC

MP Separator Pressure

Vapor flowrate of the Separator

V-104_LC

MP Separator Liquid Level

Liquid flowrate of the Separator

V-105_PC

Flash Separator Pressure

Vapor flowrate of the Separator

V-105_LC

Flash Separator Liquid Level

Liquid flowrate of the Separator

V-106_PC

Reflux Drum Pressure

Vapor flowrate of the reflux drum

V-106_LC

Reflux Drum Liquid Level

Reflux flowrate

E-103_TC

Chiller Temperature

Chiller duty

C-101_S1PC

Stripper Column Pressure

Vapor flowrate of the column

C-101_S3TC

Reboiler Temperature

Reboiler heat duty

C-101_SumpLC

Reboiler Level

Bottom flowrate


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LIST OF FIGURES Figure 1.

Natural gas network and consumption.

Figure 2.

Methodology framework.

Figure 3.

Flowsheet of the NG dehydration process.

Figure 4.

MEG losses from (a) LTS and (b) stripper column overhead under NG Feed flow upset.

Figure 5.

MEG losses from (a) LTS and (b) stripper column under the chiller temperature variation.

Figure 6.

(a) MEG loss and (b) MEG mass fraction at the stripper column bottom under the stripper overhead temperature variation.

Figure 7.

MEG loss at the stripper column overhead under reboiler temperature variation.

Figure 8.

Impacts of stripping column reboiler duty and bottom MEG mass fraction under changes of (a) the stripping gas flowrate and (b) the inlet gas flowrate for the base case.

Figure 9.

Plant-wide dynamic model of the base case.

Figure 10.

Control strategy for the (a) separator and (b) stripper column sections of the base case.

Figure 11.

Improved control strategy for the stripper column section.

Figure 12.

Dynamic performance comparison for the base and improved cases under the feed flow variation: (a) stripper column overhead temperature; (b) stripper overhead MEG loss; (c) MEG loss from LTS; (d) stripper bottom MEG mass fraction; and (e) stripper feed and bottom flowrates.

Figure 13.

Dynamic performance comparison for the base and improved cases under the chiller temperature variation: (a) stripper column overhead temperature; (b) stripper overhead MEG losses; (c) MEG loss from LTS; (d) stripper bottom MEG mass fraction; and (e) stripper feed and bottom flowrates.



Figure 1. Natural gas network and consumption.


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Figure 2. Methodological framework.


Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47

Figure 3. Flowsheet of the NG dehydration process.


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Figure 4. MEG losses from (a) LTS and (b) stripper column overhead under NG feed flow variation.


(a)


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Figure 5. MEG losses from (a) LTS and (b) stripper column under the chiller temperature variation.


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Figure 6. (a) MEG loss and (b) MEG mass fraction at the stripper column bottom under the stripper overhead temperature variation.



Figure 7. MEG loss at the stripper column overhead under reboiler temperature variation.


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Figure 8. Impacts of stripping column reboiler duty and bottom MEG mass fraction under changes of (a) the stripping gas flowrate and (b) the inlet gas flowrate for the base case.


Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47

Figure 9. Plant-wide dynamic model of the base case.


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Figure 10. Control strategy for the (a) separator and (b) stripper column sections of the base case. 49 ACS Paragon Plus Environment


Figure 11. Improved control strategy for the stripper column section.


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Figure 12. Dynamic performance comparison for the base and improved cases under the feed flow variation: (a) stripper overhead temperature; (b) stripper overhead MEG loss; (c) MEG loss from LTS; (d) stripper bottom MEG mass fraction; and (e) stripper feed and bottom flowrates.



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Figure 13. Dynamic performance comparison for the base and improved cases under the chiller temperature variation: (a) stripper overhead temperature; (b) stripper overhead MEG loss; (c) MEG loss from LTS; (d) stripper bottom MEG mass fraction; and (e) stripper feed and bottom flowrates.


Glycol Loss Minimization for a Natural Gas Dehydration Plant under

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