Energy Consumption and Optimization Analysis of Gas Production System of Condensate Gas Reservoir-Type Gas Storage

Abstract

This study investigates the energy consumption and losses associated with the gas production process in a condensate gas reservoir-type gas storage system. The energy consumption linked to each unit and key equipment was determined by HYSYS simulation, followed by a sensitivity analysis and exergy analysis. The findings reveal that the condensate oil stabilization tower is the primary energy-consuming equipment, responsible for 70.61% of the total energy consumption (3.82 × 10⁵ kJ·h⁻¹/1%). The temperature of the condensate reboiler is identified as the most significant influencing factor. Furthermore, the equipment exhibiting the highest exergy loss is the J-T valve (1.2 × 10⁷ kJ·h⁻¹), which contributes to 25.23% of the total loss. Consequently, to mitigate energy consumption in the gas production system, it is crucial to control the temperature of the condensate oil reboiler. Enhancing efficiency will rely on recovering the pressure energy loss associated with the J-T valve. The field gas gathering system lacks sub-unit energy consumption measurement and flow measurement for key process fluids. This study can provide methodological and data references for optimizing the operation of this condensate oil–gas reservoir-type storage facility.

1. Introduction

In the context of significant changes in the global energy landscape, natural gas serves as a crucial transitional energy source for low-carbon transformation, fulfilling the dual roles of ensuring energy security and promoting green development. In 2023, China’s apparent natural gas consumption reached 3.9 × 10¹¹ m³, accounting for 8.5% [1] of the primary energy consumption structure, and it continues to grow at an annual rate of nearly 10%. Based on geological reservoir characteristics, underground gas storage facilities are primarily classified into depleted oil and gas reservoirs, salt caverns, aquifers, and abandoned mines [1]. As of 2024, there are 692 gas storage facilities operating globally, with a working gas capacity of 4.17 × 10¹¹ m³, representing approximately 10.4% [2] of annual natural gas consumption, thereby achieving preliminary regional peak shaving and strategic reserve functions. The construction of gas storage facilities in China commenced in the 1990s, marked by the completion of the first batch of storage sites, including Dazhangtuo, Jintan, and Liuzhuang, which represented a significant advancement in China’s gas storage technology. By the end of 2024, a total of 46 gas storage facilities had been established nationwide, with a working gas capacity of 2.30 × 10¹⁰ m³, accounting for 5.8% [2] of annual natural gas consumption. This indicates a considerable gap compared to the internationally recognized standard of 12% to 15% reserve capacity. The inadequate reserve capacity of gas storage facilities not only creates peak shaving pressures during periods of high gas demand but also presents substantial challenges in supplying security amid fluctuations in international gas prices.

The surface process system of gas storage serves as a critical hub connecting underground reservoirs with pipeline networks. During the gas production period, it effectively meets demand by depressurizing and extracting gas. Reference [3] utilized HYSYS to simulate natural gas dehydration, selecting triethylene glycol as the adsorbent. They discovered that the optimal triethylene glycol circulation rate was 665.3 kg·h⁻¹, at which point the energy consumption was minimized to 371.7 kJ·h⁻¹. Reference [4] employed APEA to conduct economic analyses of both the natural gas dehydration method and the stripping gas method, resulting in a 38% reduction in operating costs following optimization. Reference [5] experimentally investigated the impact of RPB on TEG regeneration in natural gas dehydration, determining that a gas–liquid ratio of 4:5 yielded optimal regeneration performance. Reference [6] optimized the triethylene glycol dehydration process using the HYSYS model, achieving an 18.13% reduction in unit energy consumption by integrating a BP neural network with a genetic algorithm. Reference [7] reduced the natural gas water dew point from 279.15 K to 250 K through process simulation. References [8,9] made advancements in optimizing operating costs and carbon emissions, respectively. Reference [10] examined the discrepancies in natural gas molecular dehydration calculations between the current domestic SY/T 0076-2008 [11] and GPSA standards. Through practical case comparisons, it was found that the two can be used interchangeably at their inflection points and in anti-channeling calculations, thereby enhancing the accuracy and reliability of process equipment. Reference [12] improved the prediction accuracy and gas injection capability of the model through secondary development, thereby providing scientific support for decision-making in underground gas storage projects. Reference [13] employed a multi-objective particle swarm optimization method, achieving an accuracy rate of 92.8% in the dynamic model. This approach provided valuable insights for large-scale gas injection, thereby facilitating stable and efficient natural gas storage and supply in condensate underground gas storage reservoirs. It provides valuable references for the evaluation and optimization of natural gas storage and transportation performance. Other studies have compared the advantages and disadvantages of thermal oil heating systems, analyzing their applicable scenarios References [14,15,16]. Reference [17] summarized and analyzed recent progresses in H₂ production, storage, transportation, and application techniques.However, existing research predominantly focuses on individual units such as dehydration and condensation, lacking a systematic energy efficiency analysis of the entire gas production process. This limitation makes it challenging to reveal the energy coupling relationships between units and the potential for overall optimization.

This paper proposes an exergy analysis approach grounded in the second law of thermodynamics to evaluate the gas production process in a condensate gas reservoir-type underground gas storage (UGS) facility. By employing Aspen HYSYS for process simulation, we systematically assess energy consumption patterns, key influencing factors, and equipment-level exergy destruction. The results pinpoint critical inefficiencies and prioritize optimization targets, offering both theoretical insights and practical methodologies for enhancing the energy efficiency of gas recovery operations in condensate gas reservoir-based underground gas storage systems.

The organization of this study is as follows. Section 2 presents a detailed description of the gas production process simulation, energy analysis methodology, and results. Exergy analysis methodology and results for gas production process and main units are described in Section 3. Section 4 discusses the optimization recommendations. Section 5 concludes this paper and demonstrates further work for this study.

2. Energy Consumption Analysis of Gas Production Process

2.1. Process Description

A condensate gas reservoir storage facility group consists of four storage facilities [18]. The extracted raw gas is a complex mixture of natural gas, condensate, and water. After being depressurized and cooled by downhole throttling valves, it enters the surface gas production system. In accordance with the stipulations of GB/T 50251 [19], the processed gas must meet the technical requirements of a water dew point of ≤268.15 K and a hydrogen sulfide content of ≤20 mg/m³.

Due to the inherent physical properties of condensate gas reservoirs, which exhibit multiphase coexistence and notable condensate characteristics, the gas production process within this reservoir group is primarily composed of four essential units: dehydration, condensate treatment, light hydrocarbon separation, and glycol regeneration.

The main components of the dehydration unit are illustrated in Figure 1. Initially, the feed gas enters the production separator for a three-phase separation process. It then passes through an air-cooler where its temperature is reduced to approximately 303.15 K, before entering the pre-cooling separator for a secondary three-phase separation. The gas phase is subsequently mixed with lean glycol and directed into a coil-wound heat exchanger, where it is cooled to around 273.15 K. Following this, the gas flows through a Joule–Thomson (J-T) valve, further cooling it to approximately 265.1 K, before entering the low-temperature separator to achieve gas–liquid two-phase separation, resulting in the production of low-temperature dry gas. Finally, the dry gas is routed back to the coil-wound heat exchanger for reheating, passes through the condensate oil heat exchanger for additional heating, and ultimately flows into the export valve manifold for pipeline export.

The condensate treatment unit primarily comprises the primary and secondary condensate flash separators, the condensate stabilizer column, the reboiler, the condensate heat exchanger, and the condensate–natural gas heat exchanger. The processing targets are the condensate mixtures discharged from the production separator and the pre-cooling separator in the dehydration unit.

The core equipment of the low-pressure gas and light hydrocarbon separation unit includes the low-pressure gas cold box, low-pressure gas cryogenic separator, light hydrocarbon separator I, and light hydrocarbon separator II. These components are primarily responsible for the thorough separation of the light hydrocarbon–glycol mixture discharged from the cryogenic separator. The glycol regeneration unit comprises the glycol regeneration tower, reboiler, glycol flash drums A/B, lean-rich glycol heat exchanger, and glycol circulation pump. This unit dehydrates and regenerates the rich glycol solution that has absorbed moisture through the high-temperature desorption process in the regeneration tower, producing a lean glycol solution with a moisture content of ≤20%, which is subsequently returned to the dehydration treatment unit for recycling.

2.2. HYSYS Model Development

This study models the surface gas production system utilizing Aspen HYSYS.

2.2.1. Equation of State Selection

The Peng–Robinson (PR) equation is chosen for hydrocarbon simulation:

$$p={\displaystyle \frac{RT}{V-b}}-{\displaystyle \frac{a}{V^2+2bV-b^2}}$$

(1)

where p represents pressure (Pa), V denotes molar volume (m³·mol⁻¹), $T$ indicates temperature (K), and R is the gas constant (8.314 J·(mol·K)⁻¹). The parameters a and b are specific to each substance.

2.2.2. Simulation Model of Gas Production Process

The composition of feed gas is shown in

Table 1. Feed gas composition.

Component	Mole Fraction%
EG	0.00
C1	79.54
C2	3.36
C3	0.92
i-C4	0.16
n-C4	0.20
i-C5	0.07
n-C5	0.05
C6	0.07
C5+*	0.04
CO2	2.06
N2	0.36
H2O	11.35

. This selection, in conjunction with the gas production process and the actual operating parameters of the gas storage facility, enables the establishment of a comprehensive gas production process, as illustrated in Figure 2. The adiabatic efficiency of the circulating pump is 75%. The number of theoretical stages for both the condensate stabilizer column and the glycol regeneration column is 4. The operating temperature of the condensate reboiler is 398.15 K, and that of the glycol reboiler is 458.15 K. The convergence of an entire process is achieved when global mass and energy conservation are satisfied. For individual unit operations, taking a column with a reboiler as an example, convergence is attained when internal iterations yield correction factors for vapor flow, liquid flow, and composition approaching unity within a default tolerance range of 0.99 to 1.01 (Component C₅+* is a hypothetical component representing the collective group of all heavy hydrocarbons with five or more carbon atoms in the process fluid, characterized by a molecular weight of 100 and a boiling point of 573.15 K).

2.2.3. Model Validation

A set of typical operational data from the surface gas production system of the gas storage facility was selected and input into the established HYSYS model for calculation and solution. The input simulation parameters for the HYSYS calculation are summarized in

Table 2. Comparison of simulation results with actual values.

Operating Parameter	Stimulation Results	Actual Values	Error
Operating pressure of the production separator/MPa	9.89	9.50	4.11%
Post-cooler temperature/K	303.15	306.18	0.99%
Pre-cooling separator pressure/MPa	9.6	9.45	1.59%
Temperature before the tubular heat exchanger/K	303.25	305.97	0.89%
Temperature after the tubular heat exchanger/K	271.15	272.61	0.54%
Pressure before the tubular heat exchanger/MPa	9.6	9.44	1.69%
Pressure before the tubular heat exchanger/MPa	9.48	9.37	1.17%
Low-temperature separator pressure/MPa	7.69	7.81	1.54%
Low-temperature separator temperature/K	263.15	265.37	0.84%
Export valve assembly pressure/MPa	7.18	7.31	1.78%
Export valve assembly Temperature/K	299.18	298.55	0.22%

. The inlet temperature of the feed gas is 339.15 K, and the inlet pressure is 9.89 MPa.

Process convergence is achieved when the input and output flow rates and properties of all unit modules satisfy the global mass and energy balance across the entire process network. This simulation results obtained from the HYSYS software (V12) were compared with field and actual operational results to verify the model’s rationality and reliability.

The maximum error recorded was 4.11%, while the minimum was 0.22%, both of which fall within the acceptable engineering margin of 10%. Thus, the model is deemed reliable.

2.3. Energy Analysis of Production Units

Through a quantitative analysis of the factors influencing the energy consumption of the gas storage production system, it has been determined that the primary factors affecting the energy consumption of the production unit include the temperature and pressure of the feed gas, the circulation volume of ethylene glycol (denoted as EG in the figure), and the reboiler temperatures of both the condensate stabilizer and the ethylene glycol regenerator. Utilizing the established model, a detailed quantitative analysis was conducted to identify the optimal parameters that influence energy consumption within the gas storage production system. The energy consumption of the energy-consuming device equipment is presented in

Table 3. Energy consumption of main devices.

Energy-Consuming Unit	Energy-Consuming Device	Energy Consumption (kJ·h−1)	Proportion of Total Energy Consumption %	Proportion of Unit Energy Consumption in Total Energy Consumption %
Condensate treatment	T-101	1.18 × 107	70.61	70.61
Light hydrocarbon processing	V-104	1.61 × 106	9.63	9.63
Ethylene glycol regeneration	T-100	1.90 × 106	11.37	11.62
	P-100	4.13 × 104	0.25
Dehydration	AC-100	1.36 × 106	8.14	8.14
System	-	1.67 × 107	100	100

.

Both the air-cooler fan and the ethylene glycol circulation pump are electrically driven, with the fan operating at a power rating of 148 kW and the pump at 4.49 kW. According to the GB/T-50441 [20] standard, the conversion efficiency of electricity to heat is 9.21 MJ·(kWh)⁻¹. Consequently, the heat generated from the electricity consumed by the air-cooler fan and the ethylene glycol circulation pump is 1.36 × 10⁶ kJ·h⁻¹ and 4.13 × 10⁴ kJ·h⁻¹, respectively.

2.4. Impact of Operating Parameters on Total Energy Consumption

In analyzing the impact of various operating parameters on the total energy consumption of the gas production system, the feed gas inlet temperature was set at 339.15 K, the inlet pressure at 9.89 MPa, the glycol circulation rate at 1350 kg·h⁻¹, and the temperatures of the condensate and glycol at 398.15 K and 458.15 K, respectively. A quantitative analysis was performed to assess the influence of these different operating parameters on total energy consumption, with results illustrated in Figure 3.

As shown in Figure 3, the total energy consumption of the natural gas production process exhibits a positive correlation with multiple factors: the ethylene glycol circulation rate, the reboiler temperature of the ethylene glycol regenerator, and the reboiler temperature of the condensate stabilizer. Conversely, it shows a negative correlation with the feed gas inlet temperature and inlet pressure.

Figure 3A illustrates that as the feed gas temperature increases, the total energy consumption decreases. This phenomenon can be attributed to the reduced temperature difference between the separated condensate, the glycol-rich liquid, and the reboiler. Consequently, the heat required for condensate stabilization and glycol regeneration decreases, leading to lower thermal energy demand in the reboiler to drive the vaporization process.

Figure 3B illustrates that as the feed gas pressure increases, the inlet temperature influences the heat exchange between light hydrocarbon separator I and the thermal oil system, resulting in reduced total energy consumption. This phenomenon can be attributed to the decreased solubility of heavy components in the gas phase, which enhances gas–liquid separation efficiency. Consequently, fewer heavy hydrocarbons enter light hydrocarbon separator I, reducing the thermal energy required to drive the vaporization process.

Figure 3C illustrates that as the glycol circulation volume increases, total energy consumption also rises. This phenomenon can be attributed to the increased load on the glycol pump resulting from the higher glycol circulation volume, which consequently leads to an increase in total energy consumption.

Figure 3D shows that as the temperature of the glycol regeneration tower reboiler rises, total energy consumption also increases, while the water dew point remains constant. This can be attributed to the heightened temperature of the glycol regeneration tower reboiler, resulting in a greater temperature differential as glycol circulates through the reboiler. Consequently, the glycol absorbs more heat, thereby elevating the thermal load of the glycol regeneration tower reboiler.

Figure 3E indicates that as the temperature of the reboiler in the condensate stabilization tower increases, total energy consumption also rises. This phenomenon is due to the fact that a higher reboiler temperature results in a greater temperature differential as the condensate flows through the reboiler, leading to enhanced heat absorption by the condensate. Consequently, the heat load on the reboiler in the condensate stabilization tower is elevated.

The sensitivity of each parameter to total energy consumption was evaluated using the 1% variable perturbation method (Figure 3F). The results indicated that the unit variable sensitivity of the reboiler temperature in the condensate stabilization tower was the highest at 3.82 × 10⁵ kJ·h⁻¹/1%, followed by the glycol reboiler temperature at 1.60 × 10⁴ kJ·h⁻¹/1% and the glycol circulation rate at 1.22 × 10⁴ kJ·h⁻¹/1%. In contrast, the sensitivities of the feed gas inlet temperature and inlet pressure were relatively low, at 9.77 × 10³ kJ·h⁻¹/1% and 8.26 × 10³ kJ·h⁻¹/1%, respectively.

This discrepancy can be attributed to the energy proportions of the process units. The stabilizer reboiler bears the primary thermal load for the separation of condensate components, accounting for 70.61% of the total energy consumption of the gas production system. Even minor temperature fluctuations can lead to significant variations in heat demand. Conversely, the parameters of the feed gas influence energy consumption indirectly through phase equilibrium, characterized by a longer action path and higher energy transfer efficiency, resulting in lower sensitivity. Research indicates that precise control of the stabilizer reboiler temperature can facilitate efficient energy consumption regulation, thereby providing a prioritized basis for process optimization.

3. Exergy Analysis of Gas Production Process

3.1. Overall Exergy Model Establishment

In the analysis of system energy usage, three commonly utilized analytical models are the black box, white box, and gray box models. The selection of an appropriate model should be informed by the characteristics of the object under analysis and the specific requirements of the investigation. This study, grounded in practical engineering, examines the exergy status of each unit; therefore, the black box model is chosen. The reference state conditions for the exergy environment in this study are an environmental pressure (P₀) of 101.325 kPa and an environmental temperature (T₀) of 298.15 K.

In the analysis of system energy usage, three commonly utilized analytical models are the black box, white box, and gray box models. The selection of an appropriate model should be informed by the characteristics of the object under analysis and the specific requirements of the investigation. This study, grounded in practical engineering, examines the exergy status of each unit; therefore, the black box model is selected.

Total, Physical and Chemical Exergy of Streams

The expressions for total exergy, physical exergy, and chemical exergy are presented in Equations (2)–(4), respectively.

$$E_{\mathrm{x},\mathrm{tot}}=E_{\mathrm{x},\mathrm{ph}}+E_{\mathrm{x},\mathrm{ch}}$$

(2)

$$E_{\mathrm{x},\mathrm{ph}}=h-h_0-T_0(S-S_0)$$

(3)

where E_x,tot is the total exergy, E_x,ph and E_x,ch are the physical exergy and the chemical exergy, respectively, kJ·kg⁻¹; h is the enthalpy of the actual logistics situation, kJ·mol⁻¹; S is the entropy of the actual logistics situation, J·K⁻¹. h₀ is the enthalpy at the environmental reference state, kJ·mol⁻¹; S₀ is the entropy at the environmental reference state, J·K⁻¹. T₀ is the temperature under the environmental baseline, K.

$$E_{\mathrm{x},\mathrm{ch}}={\displaystyle \sum _i{x}_i{E}_i^{\theta }+R{T}_0}{\displaystyle \sum _i{x}_i\ln x_i}$$

(4)

In Equation (4), x_i is the mole fraction of component i in the mixture, %; E^θ_i is the standard chemical exergy of substance i, kJ·kg⁻¹; R is the universal gas constant.

In this context, for a single system, its exergy analysis is shown in Figure 4, E_in denotes the exergy introduced into the system by logistics for this gas storage facility; it specifically refers to the exergy brought in by the feed gas. E_pay represents the actual energy supplied, which encompasses the thermal energy provided by the gas storage for the electricity of the air-cooler and the heat transfer oil. E_loss indicates the loss of exergy, which in this gas storage facility includes the exergy carried away by vented gas, exergy losses from equipment, and exergy losses due to cooling. E_out refers to the system’s acquisition of exergy, which comprises the exergy of exported gas, exported condensate, and exported light hydrocarbons. The general calculation formulas for calculating exergy efficiency and exergy loss are Equations (5) and (6), respectively.

$${\eta }_{\mathrm{ex}}={\displaystyle \frac{E_{\mathrm{gain}}}{E_{\mathrm{pay}}}}$$

(5)

$$E_{\mathrm{loss}}=E_{\mathrm{pay}}-E_{\mathrm{gain}}$$

(6)

In Formula (5), E_gain denotes the effective utilization energy measured in kJ·h⁻¹. In Equation (6), E_loss represents the exergy loss of the equipment, kJ·h⁻¹.

For the gas production system, exergy outputs include the natural gas exergy of 3.21 × 10⁸ kJ·h⁻¹, exported condensate exergy of 2.62 × 10⁴ kJ·h⁻¹, and exported light hydrocarbon exergy of 6.23 × 10⁴ kJ·h⁻¹. The exergy inputs comprise raw gas exergy of 3.57 × 10⁸ kJ·h⁻¹, condensate reboiler heat exergy of 1.18 × 10⁷ kJ·h⁻¹, heat exergy from light hydrocarbon separator I of 1.60 × 10⁶ kJ·h⁻¹, ethylene glycol reboiler heat exergy of 1.90 × 10⁶ kJ·h⁻¹, circulation pump power exergy of 1.62 × 10⁴ kJ·h⁻¹, and air-cooler electricity of 148 kW. The system exhibits an efficiency of 86.69%, and the total exergy loss of the system is presented in

Table 4. Exergy loss of the gas production process.

Unit	Device	Exergy Loss of Device (kJ·h−1)	Proportion of Exergy Consumption in System (%)	Exergy Loss of Unit (kJ·h−1)	Proportion of Exergy Consumption for Unit in System (%)
Dehydration	V-100	1.85 × 106	3.89	2.51 × 107	52.82
	E-104	2.16 × 106	4.54
	V-101	8.03 × 105	1.69
	AC-100	6.93 × 106	14.57
	J-T Valve	1.20 × 107	25.23
	VLV-106	2.70 × 105	0.57
	V-102	1.11 × 106	2.33
Condensate treatment	VLV-107	1.76 × 106	3.70	1.88 × 107	39.44
	V-106	3.48 × 104	0.07
	VLV-101	1.21 × 103	-
	E-100	3.16 × 105	0.67
	VLV-108	9.97 × 104	0.21
	V-107	4.27 × 103	0.01
	VLV-102	1.02 × 105	0.21
	T-101	9.46 × 106	19.89
	E-102	2.40 × 106	5.05
	E-101	4.58 × 106	9.63
Light hydrocarbon processing	LNG-100	1.81 × 105	0.38	1.97 × 106	4.14
	V-103	2.46 × 104	0.05
	V-104	1.71 × 106	3.60
	V-105	3.31 × 102	-
	VLV-100	5.26 × 104	0.11
	VLV-105	1.23 × 103	-
Ethylene glycol regeneration	T-100	1.51 × 106	3.18	1.71 × 106	3.60
	E-105	5.20 × 103	0.01
	V-108A	45.21	-
	V-108B	43.33	-
	E-103	9.25 × 104	0.19
	E-106	1.03 × 105	0.22
Total				4.758 × 107	100

.

3.2. Exergy Analysis by Unit

For a process system composed of multiple process units or equipment, its analytical model is shown in Figure 5.

In the model shown in Figure 5, integers 1 through n denote various units. Here, E_1in,i represents the exergy introduced by the material flow of unit 1, E_1pay,i signifies the supplied exergy of unit 1, and E_1out,i is the exergy carried out by the material flow of unit 1.

3.2.1. Dehydration Unit Exergy Analysis

The exergy loss of the dehydration unit is illustrated in Figure 6. In the hydrocarbon dew point control unit, the exergy of the feed gas entering the dehydration unit is 3.57 × 10⁸ kJ·h⁻¹, while the exergy of the lean glycol is 1.56 × 10⁴ kJ·h⁻¹, and the power consumption of the air-cooler fan is 148 kW. The exergy of the export gas is 3.21 × 10⁸ kJ·h⁻¹, with the exergy of the condensate prior to valve 107 being 4.45 × 10⁶ kJ·h⁻¹, and the exergy of the light hydrocarbons before valve 106 is 1.35 × 10⁶ kJ·h⁻¹.

The exergy analysis of the dehydration unit reveals that the J-T valve is a significant source of exergy loss. Analyzing this from the perspective of thermodynamic irreversibility mechanisms, the throttling process, which is a quintessential irreversible process, involves a sudden pressure drop due to abrupt changes in flow cross-section as high-pressure fluid passes through the valve. This leads to the phase transition of certain light hydrocarbon comp-nents (C₁~C₄) from the liquid phase to the gas phase. The phase transition is accompanied by a considerable increase in entropy generation, resulting in the exergy carried by the material flow being irreversibly transformed into unavailable environmental energy, thus becoming the primary contributor to the exergy loss within this unit. Regarding the air cooler, the thermal fluid’s temperature is reduced through air cooling, which results in a decrease in the exergy value of the thermal fluid as its temperature declines. In the spiral tube heat exchanger, heat is transferred from the thermal fluid to the cold fluid; the inherent temperature difference inevitably leads to an accumulation of entropy generation, which further contributes the dissipation of the system’s exergy. The above analysis indicates that the irreversible throttling and phase change processes of the J-T valve are the primary factors contributing to exergy loss, while the heat transfer irreversibility of the air-cooler and the coil-wound heat exchanger represents a secondary source of exergy loss.

3.2.2. Exergy Analysis of Condensate Treatment Unit

The exergy loss of the Condense treatment unit is illustrated in Figure 7. In the condensate treatment unit, the input exergy comprises condensate exergy of 2.69 × 10⁶ kJ/h and thermal oil exergy of 1.18 × 10⁷ kJ·h⁻¹. The output exergy includes exported condensate exergy of 2.62 × 10⁴ kJ·h⁻¹, high-temperature light hydrocarbons of 1.82 × 10⁶ kJ·h⁻¹, vented gas (52) of 1.28 × 10⁵ kJ·h⁻¹, and vented gas (56) of 5.78 × 10⁴ kJ·h⁻¹.

In the exergy flow analysis of the condensate treatment unit, the condensate stabilization tower functions as the primary equipment responsible for exergy loss. This irreversible loss arises from the coupling of multiple mechanisms, including heat transfer, mass transfer, and phase change processes. After the condensate absorbs heat from the reboiler, the C₁–C₄ light hydrocarbon components undergo a phase change due to the increase in temperature, transitioning from the liquid phase to the gas phase and being vented from the top of the tower. During this process, the finite temperature difference in heat transfer between the reboiler and the bottom material generates heat transfer entropy. Moreover, the migration of gas–liquid two-phase components leads to mass transfer entropy generation, while the change in latent heat during the phase change process results in phase change entropy generation. The synergistic accumulation of these three types of entropy generation results in significant irreversible dissipation of exergy, positioning the stabilization tower as the primary source of exergy loss within the unit.

The exergy loss of Valve 107 and the primary flash separator is relatively high. Valve 107 serves as a condensate pressure-reducing valve, reducing the pressure from approximately 10 MPa to about 1.6 MPa. This entire process is irreversible, resulting in an increase in entropy generation and a corresponding rise in exergy loss. The primary flash separator, due to its function of pressure reduction and separation, contributes to the generation of entropy from both heat transfer and mass transfer during the process. Consequently, the exergy loss of Valve 107 and the primary flash separator is significant. This analysis suggests that the multi-physical field coupling irreversibility of the condensate stabilization tower is a key factor contributing to the extremely low unit exergy efficiency, while the process irreversibility of both the throttle valve and the flash separator exacerbates the exergy loss.

3.2.3. Light Hydrocarbon Separation Unit Exergy Analysis

The exergy loss of the Light hydrocarbon processing unit is illustrated in Figure 8. In the light hydrocarbon processing unit, the input exergy comprises low-temperature light hydrocarbon exergy of 1.08 × 10⁶ kJ·h⁻¹, high-temperature light hydrocarbon exergy of 1.82 × 10⁶ kJ·h⁻¹, and thermal oil exergy of 1.61 × 10⁶ kJ·h⁻¹. The output exergy includes exported light hydrocarbon exergy of 6.10 × 10⁴ kJ·h⁻¹, rich ethylene glycol solution (before valve 105) exergy of 3.00 × 10³ kJ·h⁻¹, low-pressure gas export exergy of 2.20 × 10⁶ kJ·h⁻¹, and low-pressure external gas transmission (process stream 61) exergy of 2.20 × 10⁶ kJ·h⁻¹.

In the light hydrocarbon separation unit, the input exergy consists of low-temperature light hydrocarbon exergy, high-temperature light hydrocarbon exergy, and the exergy from light hydrocarbon separator I. The output exergy is represented by the light hydrocarbon export exergy. Exergy loss in this unit primarily occurs in light hydrocarbon separator I and the low-pressure cold box, with irreversibility arising from the accumulation of entropy generation due to multi-process coupling. Light hydrocarbon separator I is responsible for heating and separating the incoming light hydrocarbons; the produced natural gas is vented, wastewater is directed to the wastewater treatment system, and the light hydrocarbons are sent to light hydrocarbon separator II. Throughout this process, there are increases in heat transfer entropy, mass transfer entropy, and phase change entropy. The low-pressure cold box facilitates heat exchange between low-temperature and high-temperature light hydrocarbons, and the resulting increase in heat transfer entropy directly contributes to the decline of the system’s exergy value. Research indicates that the coupled irreversibility of heat transfer, mass transfer, and phase change in light hydrocarbon separator I is the primary contributor to exergy loss, while the irreversibility associated with the temperature difference in the low-pressure cold box further exacerbates exergy dissipation.

3.2.4. Exergy Analysis of the Glycol Regeneration Unit

The exergy loss of the Ethylene glyol regneration unit is illustrated in Figure 9. In the ethylene glycol regeneration unit, the input exergy comprises the exergy of a rich ethylene glycol solution, quantified at 1.78 × 10³ kJ·h⁻¹, the thermal oil heat exergy amounting to 1.90 × 10⁶ kJ·h⁻¹, and the exergy from the circulating pump, which is 1.62 × 10⁴ kJ·h⁻¹. Conversely, the output exergy is represented by the lean ethylene glycol solution, with an exergy value of 1.56 × 10⁵ kJ·h⁻¹.

The ethylene glycol regeneration unit is primarily based on the ethylene glycol regeneration column. This column heats the rich ethylene glycol solution using thermal oil to vaporize the water content into steam. The entire process results in an increase in heat transfer entropy, mass transfer entropy, and phase change entropy, leading to a significant exergy loss of ethylene glycol. The irreversible entropy generation caused by the coupling of multiple physical fields within the regeneration column is a critical factor contributing to low exergy efficiency. Notably, the substantial temperature difference between the high-temperature thermal oil and the low-temperature rich solution exacerbates the ineffective dissipation of exergy.

4. Optimization Recommendations

To investigate the impact of the state parameters of raw gas on the system’s exergy, the temperature, pressure, and flow rate of the raw gas were selected as independent variables, designating the system’s exergy loss as the dependent variable for quantitative analysis.

The exergy loss of the gas production system is positively correlated with the inlet temperature, inlet pressure, and inlet flow rate of the raw gas. Specifically, the exergy loss increases with higher inlet temperature, inlet pressure, and inlet flow rate. Figure 10a₄ illustrates the variation in the total exergy loss of the system resulting from a 1% change in each of the inlet temperature, inlet pressure, and inlet flow rate, respectively. It can be observed that the factors influencing the exergy loss of the gas production system are ranked as follows: inlet pressure > inlet flow rate > inlet temperature.

As demonstrated above, the primary equipment contributing to exergy loss in the gas production system includes the air cooler, J-T valve, spiral-wound heat exchanger, first-stage flash separator, condensate stabilization column, low-temperature cold box, light hydrocarbon separator I, and ethylene glycol regeneration column. The expression for the exergy efficiency of a heat exchanger is given in Equation (7); for devices with no net energy exchange, such as the J-T valve, the calculation formula is given in Equation (8); for the air cooler, the calculation formula is given in Equation (9). E_x,fan represents the power consumption of fans. The exergy loss and exergy efficiency of the main equipment are shown in

Table 5. Exergy loss of main devices.

Device	Exergy Loss (kJ·h−1)	Proportion of Exergy Consumption in the System Exergy Loss (%)	Exergy Efficiency (%)
AC-100	6.93 × 106	14.57	97.79
J-T Valve	1.20 × 107	25.23	96.57
E-104	2.16 × 106	4.54	78.61
T-101	9.46 × 106	19.89	19.89
LNG-100	1.81 × 105	0.38	66.93
V-104	1.71 × 106	3.60	10.78
T-100	1.51 × 106	3.18	20.45
Total	3.40 × 107	71.39	-

, the proportion of exergy destruction in major equipment relative to the total exergy destruction is shown in Figure 11. And the exergy flows of the main devices are shown in

Table 6. Exergy flow of main devices.

Device	Exergy In (kJ·h−1)	Exergy Out (kJ·h−1)	Exergy Pay (kJ·h−1)	Exergy Loss (kJ·h−1)
AC-100	3.494 × 108	3.430 × 108	5.30 × 105	6.93 × 106
J-T Valve	3.440 × 108	3.320 × 108	0	1.20 × 107
E-104	6.7173 × 108	6.6957 × 108	0	2.16 × 106
T-101	8.40 × 105	3.19 × 106	1.1813 × 107	9.46 × 106
LNG-100	4.595 × 106	4.414 × 106	0	1.81 × 105
V-104	8.40 × 105	7.30 × 105	1.60 × 106	1.71 × 106
T-100	2.00 × 104	4.10 × 105	1.90 × 106	1.51 × 106

.

$${\eta }_{\mathrm{ex}}={\displaystyle \frac{\Delta E_{\mathrm{x},\mathrm{cold}}}{\Delta E_{\mathrm{x},\mathrm{hot}}}}$$

(7)

$${\eta }_{\mathrm{ex}}={\displaystyle \frac{E_{\mathrm{x},\mathrm{out}}}{E_{\mathrm{x},\mathrm{in}}}}$$

(8)

$${\eta }_{\mathrm{ex}}={\displaystyle \frac{E_{\mathrm{x},\mathrm{out}}}{E_{\mathrm{x},\mathrm{in}}+E_{\mathrm{x},\mathrm{fan}}}}$$

(9)

From the perspective of energy conservation in accordance with the First Law of Thermodynamics, the condensate stabilization tower and ethylene glycol regeneration tower are the primary energy-consuming units in the gas production process, accounting for 70.61% and 11.62% of the total system energy consumption, respectively. Notably, the temperature of the condensate reboiler serves as a critical operational parameter that influences the system’s energy consumption. By optimizing the operating temperature of the reboiler, while ensuring compliance with the water dew point index of the export gas, it is feasible to simultaneously reduce the circulation volume of the glycol regeneration system. This optimization leads to a decrease in pump power loss during solution transportation and a reduction in the latent heat consumption of the regeneration tower.

Based on exergy analysis theory, enhancing the system’s exergy efficiency necessitates a dual optimization of key units and feedstock gas parameters. At the unit level, it is crucial to minimize exergy losses by addressing the primary mechanisms of exergy destruction in the condensate treatment unit, light hydrocarbon separation unit, and ethylene glycol regeneration unit. For the equipment exhibiting the highest exergy loss, specifically the J-T valve, it is recommended to replace it with an expander that utilizes the pressure differential to generate electricity. The aforementioned measures, through the synergistic effect of improving equipment energy efficiency and optimizing feedstock conditions, offer an effective solution for constructing a gas production process characterized by high exergy efficiency.

5. Conclusions

This paper establishes a HYSYS process simulation model for the surface gas production system of a gas storage facility. It systematically analyzes the main units and key equipment that influence energy consumption and exergy loss during gas production in the facility. The study identifies the primary energy-consuming and exergy-losing equipment within the gas production system and conducts a sensitivity analysis, leading to several key conclusions:

(1)

From an energy consumption perspective, the majority of energy consumption originates from the condensate stabilization tower, which accounts for 70.61% of the total energy consumption. An important factor affecting the total energy consumption of the system is the temperature of the condensate reboiler. Therefore, on-site attention should be focused on regulating the temperature of the reboiler in the condensate stabilization tower. Appropriately reducing the reboiler temperature according to different operating conditions helps to reduce energy consumption.

(2)

From the perspective of minimizing exergy loss, the primary source of exergy loss in the equipment is the J-T valve, which experiences a loss of 1.20 × 10⁷ kJ·h⁻¹. To mitigate pressure energy loss while ensuring the safe and stable operation of the system, the J-T valve may be replaced with an expander power generation device.

(3)

The inlet temperature of the air-cooler is set at 273.15 K in this paper. When extracting gas in cold regions, the required water dew point temperature for exported natural gas is lower; however, the ambient cold conditions contribute to a reduction in the energy consumption associated with the gas extraction cooling process. Consequently, analyzing the energy efficiency of gas extraction in cold regions is of paramount importance. Meanwhile, the optimization of energy consumption and exergy efficiency under multiple operating conditions is also a key focus of subsequent research.