LOW-TEMPERATURE DISTILLATION METHOD FOR CAPTURING CARBON DIOXIDE

Description

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims the benefit of priority from Chinese Patent Application No. 202511777873.8, filed on Nov. 28, 2025. The content of the aforementioned disclosure, including any intervening amendments thereto, is incorporated herein by reference in its entirety.

TECHNICAL FIELD

This disclosure relates to industrial gas separation and recovery, and more particularly to a low-temperature distillation method for capturing carbon dioxide.

BACKGROUND

With the increasing demand for industrial energy conservation and emission reduction, the capture and recovery of carbon dioxide from industrial tail gas have become an important way to realize the recycling of carbon resource.

At present, the methods for separating carbon dioxide from the industrial tail gas mainly include pressure swing adsorption (PSA) method, chemical absorption method and low-temperature distillation method. However, these methods have certain limitations in terms of energy efficiency, recovery rate and operational stability.

Although the PSA method is simple to operate, the regeneration energy consumption of its adsorbent is high, and the recovery rate of carbon dioxide is low, making it difficult to stably obtain carbon dioxide with an industrial-grade purity.

The chemical absorption method consumes a large amount of chemical reagents. Moreover, the subsequent desorption process requires high energy and may also causes secondary pollution.

Although the traditional low-temperature distillation method can achieve separation of high-purity carbon dioxide, it struggles with high energy consumption for compression and cooling, and low mass transfer efficiency of column internals.

Methanol steam reforming (MSR) is applied in the fields of chemistry and energy. It is one of the three ways to produce hydrogen from methanol. The MSR method adopts two types of reaction systems. One is an isothermal reaction system based on heat exchange in a tubular reactor, and the other is an adiabatic reaction system that combines a catalyst bed with a heat exchanger. The MSR method is performed under a pressure of 1-5 MPa, and is capable of producing hydrogen with a purity greater than 99.99%. Compared with the conversion of natural gas and heavy oil into hydrogen, the investment and energy consumption of the MSR method have been reduced by 50%. The raw material of methanol is liquid at a normal temperature and normal pressure, facilitating transportation and storage. Therefore, the MSR method is suitable for a medium-scale hydrogen production of 100-1,000 m³/h.

The MSR reaction produces a gas mixture containing hydrogen (H₂), carbon dioxide (CO₂) and carbon monoxide (CO). To extract high-purity hydrogen, the PSA technology is often used for purification. This technology uses a zeolite molecular sieve or an activated carbon as an adsorbent, so as to selectively adsorb impurity gases other than hydrogen based on the difference in adsorption capacity of the adsorbents for gases under different pressures, thereby obtaining hydrogen with a purity of greater than 99.99% from an outlet. The gas desorbed from the adsorption column forms a tail gas containing a large amount of carbon dioxide and nitrogen, such that the tail gas is subjected to direct emission, not only causing waste of carbon resource, but also intensifying greenhouse effect.

However, how to efficiently capture carbon dioxide from the tail gas still faces the following challenges.

(1) Maintenance of System Operational Stability

Traditional distillation columns adopt bulk packing or fixed-diameter column sections, easily leading to uneven distribution of gas and liquid. When a gas velocity is too high, the liquid phase cannot form a uniform liquid film on the surface of the packing, thereby reducing the mass transfer efficiency. When a gas velocity is too low, the liquid phase will be retained and the separation time will be prolonged. When a flooding rate is lower than the reasonable range, the separation efficiency will be reduced. When a flooding rate is higher than the reasonable range, flooding may occur, causing separation failure or even equipment failure.

(2) Balance Between Energy Consumption and Separation Efficiency

If a reflux ratio is increased to improve the purity of carbon dioxide, the heat load of the reboiler will increase sharply, causing significant increase in energy consumption. It may also cause carbon dioxide at column bottom to be supercooled and solidified, thereby leading to equipment blockage and requiring shutdown for cleaning. On the contrary, if a cooling capacity is reduced in order to save energy, it will lead to incomplete separation and fail to meet the purity and recovery rate requirements of industrial-grade carbon dioxide product.

(3) Adaptability to Tail Gas Characteristics

The existing separation methods fail to fully consider the composition characteristic and fluctuation characteristic of the tail gas from the MSR section, making it difficult to achieve stable and efficient separation under a low energy consumption. Therefore, developing a method for capturing carbon dioxide based on the characteristics of the tail gas from the MSR section, which has both low energy consumption and high separation efficiency, and satisfies the purity requirements and impurity control requirements of industrial-grade carbon dioxide product under the premise of stable operation, is of great significance for industrial applications and environmental protection.

SUMMARY

In view of the deficiencies in the prior art, the present disclosure provides a low-temperature distillation method for capturing carbon dioxide in a tail gas from a methanol steam reforming (MSR) section.

A distillation method for capturing carbon dioxide, comprising:

• • (a) collecting a tail gas containing carbon dioxide and nitrogen from a methanol steam reforming (MSR) section, and pretreating the tail gas to obtain a feed gas; wherein the feed gas contains 84-98% by volume of carbon dioxide, 2-16% by volume of nitrogen, less than 1,000 ppm of hydrogen, less than 10 ppm of carbon monoxide and less than 1 ppm of methanol;
  • (b) subjecting the feed gas sequentially to a first-stage compression, a first-stage cooling, a second-stage compression, a second-stage cooling, a third-stage compression and a third-stage cooling to obtain a processed gas with a moisture content less than 5 ppm, wherein the third-stage compression is performed to reach a pressure of 1.4-1.8 MPa, and the third-stage cooling is performed to reach a temperature of −100° C. to −70° C.;
  • (c) feeding the processed gas into a distillation column packed with structured packing for distillation and separation, and collecting a liquid stream predominated by carbon dioxide from a bottom of the distillation column; wherein the distillation column are operated according to optimized process parameters, and the optimized process parameters are determined through simulation and optimization using an Aspen process simulation software;
  • wherein the simulation and optimization is performed through steps of:
  • establishing a distillation column model by using a RadFrac module and a Peng-Robinson model; setting, in the distillation column model, a carbon dioxide purity in the liquid stream to be equal to or larger than 99%; and setting, in the distillation column model, a theoretical-stage flooding percentage to be 60%-80%; and
  • determining the optimized process parameters utilizing a sensitivity analysis tool of the Aspen process simulation software with minimization of a total system energy consumption as an optimization objective; and
  • (d) pressurizing the liquid steam obtained from step (c) to 2.1-2.5 MPa to obtain an industrial-grade carbon dioxide product with a carbon dioxide purity of 99%, and transferring the industrial-grade carbon dioxide product to a vacuum-insulated buffer tank for storage.

In some embodiments, in step (a), the pretreating of the tail gas comprises removal of solid particles, a part of water vapor, a part of hydrogen, a par of carbon monoxide and all methanol.

In some embodiments, the tail gas collected from the MSR section contains 84-98% by volume of carbon dioxide, 2-16% by volume of nitrogen, 100-5,000 ppm of hydrogen, 10-1,000 ppm of carbon monoxide and 10-500 ppm of water.

In some embodiments, from the safety perspective, residual hydrogen in the pretreated feed gas entering the distillation column is as low as possible, generally not exceeding 1,000 ppm; during the distillation, carbon monoxide is prone to forming azeotropes with other impurities, thereby affecting the operation of the distillation column and product quality, so that the carbon monoxide content is reduced to below 10 ppm after pretreating; and water is prone to freezing at a low temperature, causing to block pipelines and devices, so that the water content is reduced to below 5 ppm after pretreating.

In some embodiments, the pretreatment is performed through conventional methods, comprising a two-stage filtration method, a silanized modified coal-based active carbon absorbent method, a palladium-based absorbent method and a copper-based absorbent method;

• • wherein the two-stage filtration method is performed by combining a metal sintered mesh filtration and a polytetrafluoroethylene (PTFE)-coated filter cartridge filtration; the metal sintered mesh is adopted to perform coarse filtration to intercept relatively large particles of 5 μm and above; and a PTFE-coated filter cartridge is adopted to perform fine filtration to enable efficient filtration of micro-sized particles of 1-3 μm, and also to filter out trace amounts of the water vapor;
  • the silanized modified coal-based active carbon absorbent method is adopted to remove methanol;
  • the palladium-based absorbent method is adopted to selectively absorb hydrogen/carbon monoxide; and
  • the copper-based absorbent method is adopted to remove carbon monoxide.

In some embodiments, a gas-liquid separator is additionally arranged following cooling by the first-stage cooler; residual water is deeply removed through a hydrophobic polymer membrane (such as a PTFE hollow fiber membrane) in the gas-liquid separator; and the PTFE hollow fiber membrane has an extremely low permeability to CO₂, ultimately reducing the water content to below 5 ppm; or water is removed through silica gel-activated alumina composite fixed bed adsorption.

In some embodiments, an impurity-removal purification module is arranged based on different conditions of the feed gas.

In some embodiments, in step (b), the first-stage compression is performed to reach a pressure of 0.4-0.9 MPa, and the first-stage cooling is performed to reach a temperature of −10° C. to 10° C.;

• • the second-stage compression is performed to reach a pressure of 1-1.4 MPa, and the second-stage cooling is performed to reach a temperature of −40° C. to −20° C.

In some embodiments, in step (b), the first-stage cooling, the second-stage cooling and the third-stage cooling are each performed through heat exchange with a cooling medium; the cooling medium is selected from the group consisting of ethylene glycol and liquid nitrogen; and

• • step (b) further comprises:
  • before the second-stage compression, supplying the feed gas processed by the first-stage cooling to a gas-liquid separator for dehydration to reduce a water content to be 5 ppm or less.

In some embodiments, the cooling medium used in the first-stage cooling is ethylene glycol;

• • the cooling medium used in the second-stage cooling is liquid nitrogen; and
  • the cooling medium used in the third-stage cooling is liquid nitrogen.

In some embodiments, in step (c), the simulation and optimization is performed through steps of:

• • (c1) in the Aspen process simulation software, selecting the RadFrac module as a unit operation module and the Peng-Robinson model as a thermodynamic model to establish a distillation column model;
  • (c2) inputting preset process parameters into the distillation column model, wherein the preset process parameters comprise composition of the feed gas, operating pressure, temperature and column internal and packing parameters; and
  • based on the preset process parameters in combination with a preset pressurization pressure of a compressor used in the first-stage compression, the second-stage compression and the third-stage compression and a cooling medium flow rate in the first-stage cooling, the second-stage cooling and the third-stage cooling, simulating different feed gas states;
  • (c3) setting process constraint conditions, wherein the process constraint conditions comprise the carbon dioxide purity in the liquid stream ≥99%, and the theoretical-stage flooding percentage within a range of 60-80%; and
  • under the premise of satisfying the process constraint conditions, with the preset pressurization pressure in each of the first-stage compression, the second-stage compression and the third-stage compression and the cooling medium flow rate in each of the first-stage cooling, the second-stage cooling and the third-stage cooling as variables, for different feed gas compositions, analyzing an influence of different variable combinations on the total system energy consumption through multi-variate cross combination analysis by using the sensitivity analysis tool of the Aspen process simulation software;
  • (c4) performing collaborative optimization on process parameters of the distillation column based on analysis results obtained in step (c3) with minimization of the total system energy consumption as the optimization objective to determine the optimized process parameters, wherein the process parameters comprise theoretical stage number, reflux ratio and feed position.

In some embodiments, in step (c), the optimized process parameters of the distillation column comprise at least one of conditions (a1)-(a7):

• • (a1) structured packing is employed for internals of the distillation column, with a height of each packing layer being 0.1-0.3 m and a column section diameter being 0.07-0.6 m;
  • (a2) a theoretical stage number of the distillation column is 4-12, with each stage having a height of 0.25-1.2 m and a diameter of 0.07-0.6 m;
  • (a3) a feed pressure of the distillation column is 1.4-1.8 MPa;
  • (a4) a feed temperature of the distillation column is from −100° C. to −70° C.;
  • (a5) a column top temperature of the distillation column is −40° C.-0° C.;
  • (a6) a column bottom temperature of the distillation column is from −180° C. to −150° C.; and
  • (a7) a reflux ratio of the distillation column is 0.8-1.5.

In some embodiments, during the distillation and separation process in step (c), a monitoring and early-warning treatment is continuously preformed through steps of:

• • collecting data of key operation parameters of key process points of the distillation column in real time during the distillation and separation; wherein the key process points comprise column top, column bottom, feed inlet, reboiler outlet and condenser outlet; and the key operation parameters comprise temperature, pressure and flow rate;
  • inputting the data into the Aspen process simulation software, followed by comparison with the optimized process parameters to output a deviation; and in a case that the deviation exceeds a preset threshold, triggering an early-warning signal; and
  • adjusting process parameters according to the early-warning signal, wherein the process parameters comprise cooling medium flow rate, compressor object pressure and reflux flow rate of the distillation column; and under the premise of ensuring a feed temperature of the distillation column within a range from −100° C. to −70° C. and a feed pressure of the distillation column within a range of 1.4-1.8 MPa, adjusting other parameters to an optimal range to minimize the total system energy consumption.

In some embodiments, in a case that the column top temperature is higher than a predictive value, a reflux ratio requiring to be reduced is calculated through the Aspen process simulation software, and is automatically fed back to a reflux pump of the distillation column to adjust the reflux flow rate; and in a case that the distillation column pressure is higher than 1.8 MPa, a cooling medium flow rate requiring to be added is calculated through the Aspen process simulation software, and a linkage cooler adjusting valve is opened wider until other parameters are adjusted to the optimal range.

In some embodiments, the vacuum-insulated buffer tank comprises:

• • a liner;
  • a vacuum jacket wrapping the liner; and
  • a housing;
  • wherein the vacuum jacket is filled with a thermal-insulation material; and the housing is configured to wrap the vacuum jacket.

In some embodiments, the liner is made of stainless steel; and the liner has a mirror-polished inner wall;

• • a vacuum degree in the vacuum jacket is not greater than 10 Pa; and the thermal-insulation material is perlite; and
  • the housing is made of a carbon steel material.

Compared to the prior art, the present disclosure has the following beneficial effects.

(1) Strong Pertinence

The technical solutions adopted herein are specifically designed for MSR tail gas, so as to achieve efficient capture of industrial-grade carbon dioxide. It solves the problem that the product purity is not up to standard and the impurity content is high due to the poor compatibility of general process with gas source.

(2) Low Energy Consumption

The present disclosure proposes a process solution optimized through Aspen simulation, which integrates three-stage compression and utilizes liquid nitrogen and ethylene glycol for three-stage cascaded cooling. Compared to conventional low-temperature distillation processes, this technical solution reduces energy consumption by 15-20%. Furthermore, during the storage stage, the present disclosure employs a low-temperature storage technology with a pressure of 2.3-2.5 MPa, thereby eliminating a high-pressure compression step and further reducing the overall energy consumption of the system.

(3) High Separation Efficiency

The distillation column adopts a structured packing, coordinated with an optimized reflux ratio and operating pressure, so as to obtain an industrial grade carbon dioxide with a purity of 99% meeting an industrial standard.

(4) Operational Stability

By controlling the liquid flooding rate in the distillation column at 60-80%, the distillation is stably performed. At the same time, the stable liquid state of carbon dioxide is maintained with the help of a vacuum-insulated buffer tank, so as to provide ideal condition for storage and subsequent processing.

BRIEF DESCRIPTION OF THE DRAWINGS

In order to make technical solutions in the present disclosure clearer, the accompanying drawings needed in the description of the embodiment of the present disclosure will be briefly described below. It should be understood that the accompanying drawings described below are merely some embodiments of the present disclosure, and are not intended to limit this present disclosure. Other drawings can also be obtained by those skilled in the art based on the accompanying drawings without paying creative effort.

FIG. 1 is a flow chart of an industrial-grade carbon dioxide capture method, where FEED is a feed gas from a methanol steam reforming (MSR) section; COMPR1, COMPR2 and COMPR3 are a first-stage compressor, a second-stage compressor, and a third-stage compressor, respectively; COOLER1, COOLER2 and COOLER3 are a first-stage cooler, a second-stage cooler and a third-stage cooler, respectively; TOWER is a low-temperature distillation column; PUMP is a compression pump; and PRODUCT is an industrial-grade carbon dioxide product; and FIG. 1 illustrates a full process in which the feed gas sequentially undergoes three-stage compression, three-stage cooling, low-temperature distillation, compression and storage, and all equipment are each connected via pipelines to achieve a continuous production;

FIG. 2 is a schematic diagram of theoretical stage parameters for distillation in Embodiment 1;

FIG. 3 is a schematic diagram of distillation hydraulics calculation parameters in the Embodiment 1;

FIG. 4 is a schematic diagram of theoretical stage parameters for distillation in Embodiment 2;

FIG. 5 is a schematic diagram of distillation hydraulics calculation parameters in the Embodiment 2;

FIG. 6 is a schematic diagram of theoretical stage parameters for distillation in Embodiment 3; and

FIG. 7 is a schematic diagram of distillation hydraulics calculation parameters in the Embodiment 3.

DETAILED DESCRIPTION OF EMBODIMENTS

In order to facilitate understanding the technical solutions of the present disclosure, the technical solutions of the embodiments in the present disclosure will be described clearly and completely. Obviously, described below are merely some embodiments of the present disclosure, instead of all embodiments. Based on the embodiments in the present disclosure, all other embodiments obtained by those of ordinary skill in the art without paying creative labor shall fall within the scope of the present disclosure.

In some embodiments, a tail gas containing carbon dioxide and nitrogen is collected by connecting a gas collection device to a tail gas outlet of a methanol steam reforming (MSR) section. The tail gas mainly includes 84-98% by volume of carbon dioxide, 2-16% by volume of nitrogen and trace impurities, such as 100-5,000 ppm (0.01-0.5%) of hydrogen, 10-1,000 ppm of carbon monoxide, 10-500 ppm of water and 0-1 ppm of methanol.

In some embodiments, the tail gas from the MSR section is pretreated to obtain a feed gas. The pretreating of the tail gas includes removal of solid particles, a part of water vapor, a part of hydrogen, a part of carbon monoxide and all methanol.

In some embodiments, from the safety perspective, residual hydrogen in the pretreated feed gas entering the distillation column is as low as possible, generally not exceeding 1,000 ppm. During the distillation, carbon monoxide is prone forming azeotropes with other impurities, thereby affecting the operation of the distillation column and product quality, so that the carbon monoxide is reduced to below 10 ppm after pretreating. Water is prone to freezing at a low temperature, causing to block pipelines and devices, so that water content is reduced to below 5 ppm after pretreating.

In some embodiments, the pretreatment can be performed according to conventional methods in the art. For example, a two-stage filtration method is performed by combining a metal sintered mesh filtration and a polytetrafluoroethylene (PTFE)-coated filter cartridge filtration. Moreover, a gas-liquid separator is arranged, such that the water in the feed gas is condensed into a liquid state by cooling for separation. Residual water is deeply removed by a hydrophobic polymer film (such as a PTFE hollow fiber membrane) integrated into the gas-liquid separator. The PTFE hollow fiber membrane has an extremely low permeability to CO₂, ultimately reducing the water content in the feed gas to below 5 ppm. Furthermore, a silanized modified coal-based active carbon absorbent is adopted to remove methanol. A palladium-based absorbent is adopted to selectively absorb H₂/CO, or a copper-based absorbent is adopted to remove CO.

In some embodiments, a flow chart of an industrial-grade carbon dioxide capture method is shown in FIG. 1. Referring to FIG. 1, FEED is a feed gas from the MSR section. COMPR1, COMPR2 and COMPR3 are a first-stage compressor, a second-stage compressor and a third-stage compressor, respectively. COOLER1, COOLER2 and COOLER3 are a first-stage cooler, a second-stage cooler and a third-stage cooler, respectively. TOWER is a low-temperature distillation column. PUMP is a compression pump. PRODUCT is an industrial-grade carbon dioxide product. FIG. 1 illustrates a full process in which the feed gas sequentially undergoes three-stage compression, three-stage cooling, low-temperature distillation, compression and storage, and all equipment are each connected via pipelines to achieve a continuous production.

In some embodiments, the feed gas is compressed through a compressor. The working principle of the compressor is that gas volume is decreased through a mechanical way, thereby enabling gas molecules to be compressed into a smaller space and causing a pressure and a temperature of the gas to increase simultaneously.

The compressors most commonly used in industry include a reciprocating compressor and a centrifugal compressor. In this embodiment, the reciprocating compressor is adopted. An exhaust valve opens when a pressure of gas inside a compressor cylinder rises to a preset pressure of an outlet pipeline plus a spring force of the exhaust valve. The compressed gas then enters a cooler. The cooler is a shell and tube heat exchanger or a plate heat exchanger. In this embodiment, the shell and tube heat exchanger is adopted.

After compression, hot gas enters from an end of the cooler, and flows inside the tubes. A cooling medium (ethylene glycol or liquid nitrogen) enters an inlet of an outer shell of the cooler, and flows outside the tubes but inside the shell, moving in an opposite direction of a gas pipeline, so as to maximize a heat exchange efficiency. A temperature is stably controlled through a temperature detecting element and dynamic regulation of a cooling medium flow rate. A temperature sensor (platinum resistance sensor or thermocouple sensor) is arranged on an outlet pipeline of the cooler to monitor an actual temperature of cooled gas in real time.

When an actual temperature detected by the temperature sensor is higher than a preset target temperature, a control system automatically increases the cooling medium flow rate corresponding to the cooler. The specific adjustments are differently performed according to different coolers.

(1) If the actual temperature is a temperature of the COOLER1 cooled by ethylene glycol, the cooling medium flow rate is increased by enhancing a power of an ethylene glycol circulation pump.

(2) If the actual temperature is a temperature of COOLER2/3 cooled by liquid nitrogen, the liquid nitrogen regulating valve is opened wider to increase supply volume of liquid nitrogen, and reduce the actual temperature to the target range by enhancing the heat exchange.

On the contrary, when the actual temperature detected by the temperature sensor is lower than the preset target temperature, the cooling medium flow rate is reduced to avoid excessive cooling and energy waste.

The present disclosure adopted three-stage compression and three-stage cooling, so as to adapt to a distillation pressure and a distillation temperature with a lower energy consumption. The pressure of the feed gas is gradually increased to an operating pressure range of the distillation column through the three-stage compression and three-stage cooling, avoiding impact caused by a significant difference between the feed gas pressure and the distillation pressure, and maintaining gas-liquid balance in the distillation column.

The three-stage compression and three-stage cooling have the following beneficial effects.

The total energy consumption of the compressor is significantly decreased to avoid drawbacks caused by a high single-stage compression ratio.

Given that the energy consumption for gas compression increases nonlinearly with the compression ratio, the larger the compression ratio, the faster the rise in energy consumption and the higher the compression energy required per unit volume of gas. If the feed gas is directly compressed from its initial pressure to 1.4-1.8 MPa through one/two-stage compression, it will result in an excessively high single-stage compression ratio, and a sharp surge in the energy consumption of the compressor. Simultaneously, the three-stage cooling employs the cooling medium in a cascaded manner, effectively reducing usage of high-energy-consumption medium such as liquid nitrogen.

In some embodiments, the low-temperature distillation column adopts a Mellapak 250X structured packing. Mellapak is a brand of the structured packing from Sulzer company. 250X represents a specific surface area of the structured packing 250 m²/m³, where X represents an inclined angle of 45°, so as to enable an optimal balance between a low pressure drop and a high throughput.

In some embodiments, a theoretical stage flooding rate is controlled at 60-80%, which is an optimal range of balancing separation efficiency and operating safety during the low-temperature distillation. The theoretical stage flooding rate is a ratio of an actual gas-liquid load in the distillation column to a maximum allowable gas-liquid load of the column internals, which determines the flow state and mass transfer efficiency of the gas phase and the liquid phase in the distillation column.

When the theoretical stage flooding rate reaches 100%, a flow velocity of the gas phase is too high, which impedes the downward flow of the liquid phase, thereby preventing normal countercurrent movements of the gas phase and the liquid phase. The liquid phase accumulates, and is eventually “lifted” upward by the gas phase, resulting in flooding. At this point, the distillation separation completely fails, and equipment failures may even occur due to sudden increase in the column pressure. Conversely, if the theoretical stage flooding rate is too low, the actual gas-liquid load is far below a maximum capacity of the equipment, indicating underutilization of the equipment. The gas-liquid load is too low, leading to inadequate contact on a surface of the structured packing. The liquid phase fails to form a uniform film on the surface of the structured packing. The flow velocity of the gas phase hinders sufficient collision and mass transfer with the liquid phase, thereby reducing the separation efficiency of carbon dioxide and nitrogen.

A target product obtained from the technical solution of this disclosure is the industrial-grade carbon dioxide product. According to the provisions of GB/T 6052-2011 “industrial liquid carbon dioxide” standard, in terms of a carbon dioxide purity, a superior product has a purity of 99.8% or higher, a first-grade product has a purity of 99.5% or higher and a qualified product has a purity of 99% or higher; and in terms of water content (by mass fraction), a superior product has a water content of 0.05% or less, a first-grade product has a water content of 0.2% or less and a qualified product has a water content of 0.4% or less. The industrial-grade carbon product satisfying the above standards can be obtained through the technical solutions of the present disclosure.

In some embodiments, carbon dioxide stored in the vacuum-insulated buffer tank has a purity of 99.99-99.997% and a water content of ≤10 ppm.

In some embodiments, the technical solutions of the present disclosure provide the vacuum-insulated buffer tank to store liquid carbon dioxide from column bottom. The liquid carbon dioxide has a temperature of −50° C. to −20° C. after pressurizing by the pump, and is stored at a low temperature under a pressure of 2.1-2.5 MPa.

In terms of structure and material, the vacuum-insulated buffer tank includes a liner, a jacket and a housing. A main material of the liner is SUS304 stainless steel (GB/T 24511). An inner wall of the liner is mirror-polished to reduce a surface roughness and effectively minimize an impurity adsorption and residue. The jacket is filled with a thermal-insulation material such as perlite and evacuated to a vacuum state (vacuum degree≤10 Pa) to ensure a long-lasting insulation performance. The housing is made of a carbon steel material such as Q235B material to provide necessary mechanical strength and support for the overall structure. The vacuum-insulated buffer tank is further provided with a safety valve and a liquid level meter.

The vacuum-insulated buffer tank can be customized according to actual production requirements.

In some embodiments, the feed gas flow rate is 292.83×10³ mol/h. A daily production of liquid carbon dioxide is approximately 310 tons (a density of liquid carbon dioxide is calculated at 467 kg/m³). Therefore, a vertical vacuum-insulated tank is employed, with a volume range of 5-50 m³ and a design pressure not less than 2.5 MPa, to meet the storage requirements of continuous production. In practical applications, those skilled in the art can make targeted adjustments and optimizations to tank volume, pressure rating, interface configuration and thermal-insulation material of the vacuum-insulated buffer tank as needed.

In the existing technologies, transportation methods are typically a vehicle transportation method integrated with a low-temperature atmospheric pressure storage tank, a medium-pressure storage tank and a low-temperature tank truck, as well as a pipeline transportation method integrated with a pressure-resistant storage tank.

(1) Low-Temperature Atmospheric Pressure Storage Tank

Design pressure of the low-temperature atmospheric pressure storage tank is close to an atmospheric pressure. The low-temperature atmospheric pressure storage tank can continuously run by cooperating it with a refrigeration unit. In-tank temperature is controlled within a range from −60° C. to −50° C., thereby allowing carbon dioxide to be in a low-temperature liquid state. The low-temperature atmospheric pressure storage tank requires a long-term consumption of refrigeration energy, and has a relatively high operating cost.

(2) Medium-Pressure Storage Tank

Design pressure of the medium-pressure storage tank is usually 2-3 MPa. The medium-pressure storage tank does not require a continuous refrigeration, but it needs to maintain the liquid state of carbon dioxide by regularly supplementing pressure to avoid an energy loss caused by pressure fluctuations.

(3) Low-Temperature Tank Truck

The low-temperature tank truck adopts a vacuum-insulated structure to transport liquid carbon dioxide. In this case, a single tank volume of the low-temperature tank truck is relatively small, generally ranging from 20 m³ to 50 m³. The low-temperature tank truck has a limited transportation capacity.

(4) Pipeline Transportation Integrated with a Pressure-Resistant Storage Tank

The pressure-resistant storage tank is typically made of a carbon steel material, and is connected to a pipeline system to form a point-to-point transmission system. Moreover, the pressure-resistant storage tank is cooperated with the pressure regulation system for using.

(5) High-Pressure Steel Cylinder

Working pressure of the high-pressure steel cylinder is not less than 15 MPa. The high-pressure steel cylinder is mainly used for small batch, intermittent usage scenario such as a laboratory and a small-scale food processing.

The present disclosure adopts the vacuum-insulated buffer tank to store the industrial-grade liquid carbon dioxide. The reasons are described below.

(1) Comparison with the High-Pressure Steel Cylinder

If the high-pressure steel cylinder in the prior art is adopted for storage, the liquid carbon dioxide at 2.1-2.5 MPa in the technical solutions of the present disclosure needs to be repressurized to 15 MPa, with an energy consumption of approximately 12-15 kWh/ton during the pressurization process. However, the present disclosure directly stores the liquid carbon dioxide under pressure of 2.1-2.5 MPa, avoiding additional pressurization.

(2) Comparison with Low-Temperature Atmospheric Pressure Storage Tank

The low-temperature atmospheric pressure storage tank in the prior art requires continuous operation of the refrigeration unit to maintain the in-tank temperature at −60° C. to −50° C., with an energy consumption of about 0.3-0.5 kWh/ton/day. The vacuum-insulated buffer tank in the present disclosure relies on a cooling capacity of a front-end low-temperature distillation. Cooling loss is reduced through jacket vacuum and thermal-insulation materials. Cooling loss energy consumption is only 0.1-0.2 kWh/ton/day, which is 40-60% more energy-efficient than the low-temperature atmospheric pressure storage tank. Calculated based on the feed gas flow rate of 292.83×10³ mol/h (i.e., daily production of liquid carbon dioxide is approximately 310 ton/day), daily energy savings is about 62-124 kWh, and annual energy savings is about 2.26×10⁴-4.52×10⁴ kWh.

In the technical solutions of the present disclosure, the liquid carbon dioxide from the column bottom has a low temperature ranging from −180° C. to −150° C. and a pressure of 1.4-1.8 MPa, so that the liquid carbon dioxide only requires a small-degree compression to 2.1-2.5 MPa, thereby entering the vacuum-insulated buffer tank. Compared to the prior art, there is no need to be additionally cooled as like the low-temperature atmospheric pressure storage tank, or be repressurized through the high-pressure steel cylinder. The technical solutions of the present disclosure are optimization designs based on a synergism between the cooling capacity and the pressure during the full process, rather than an independent choice.

In the existing technologies, the liners of conventional vacuum tanks are mostly made of ordinary stainless steel. Due to the lack of special treatment on the inner walls, their surface roughness is relatively high. Additionally, for ordinary vacuum-insulated tanks, the vacuum degree of the jacket is usually 10-100 Pa, which limits the thermal-insulation performance.

In contrast, 304 stainless steel is adopted herein as a liner material. The inner wall of the liner is mirror-polished to reduce a surface roughness thereof, not only minimizing the residue of liquid carbon dioxide but also lowering the risk of product contamination. Meanwhile, this present disclosure controls the vacuum degree of the jacket at ≤10 Pa and fills the jacket with thermal-insulation materials such as perlite, effectively reducing radiative heat transfer and significantly enhancing the thermal-insulation performance.

The total energy consumption in the technical solutions of the present disclosure includes the following aspects.

As shown in FIG. 1, the feed gas undergoes three-stage compression and three-stage cooling to reach the temperature and pressure for the low-temperature distillation separation process. During the process, process energy consumption is generated, including energy consumption of the compressors for three-stage compression, and energy consumption of the coolers for three-stage cooling. Energy consumption of the low-temperature distillation column includes energy consumption of the reboiler at the column bottom and energy consumption of the condenser at the column top. Energy consumption of product post-treatment includes energy consumption of the pump for pressurizing liquid carbon dioxide at column bottom. The total energy consumption is a sum of the above energy consumptions.

Unless otherwise specified, the experimental methods used in the following embodiments are all conventional methods, and the materials and reagent are commercially obtained.

Embodiment 1

This embodiment provided a low-temperature distillation method for capturing carbon dioxide, which was performed through the following steps.

(a) Obtaining of the Feed Gas

The tail gas was collected by connecting the gas collection device to the tail gas outlet of the MSR section. The tail gas mainly included 98% by volume of CO₂, 2% by volume of N₂ and trace impurities, such as 100-5,000 ppm (0.01-0.5%) of H₂, 10-1,000 ppm of CO, 10-500 ppm of H₂O and so on.

The tail gas was pretreated, including removal of dust, water, a part of hydrogen, a part of carbon monoxide and residual methanol, so as to obtain the feed gas containing carbon dioxide and nitrogen. The feed gas satisfied the following requirements of H₂<1,000 ppm, CO<10 ppm and essentially free of methanol.

(b) Three-Stage Compression and Three-Stage Cooling

(b1) First-Stage Compression and First-Stage Cooling

The feed gas was initially compressed from 3 bar (0.3 MPa) to 6 bar (0.6 MPa) through the first compressor, then cooled to 10° C. through ethylene glycol. Subsequently, the water content was reduced to below 5 ppm through the gas-liquid separator.

(b2) Second-Stage Compression and Second-Stage Cooling

The feed gas was then compressed to 11 bar (1.1 MPa) through the second compressor, followed by cooling to −30° C. through liquid nitrogen.

(b3) Third-Stage Compression and Third-Stage Cooling

Next, the feed gas was compressed to 18 bar (1.8 MPa) through the third compressor, and further cooled to −85° C. through liquid nitrogen, thereby completing the compression and cooling of the feed gas to obtain a processed gas.

(c) Optimization Design of Low-Temperature Distillation Method Parameters

(c1) Establishment of a Process Simulation Model

In the Aspen process simulation software, a RadFrac module was selected as a unit operation module and a Peng-Robinson (PR) model was selected as a thermodynamic model to establish a distillation column model. The Peng-Robinson model was expressed as:

P = R ⁢ T v - b - a v ⁡ ( v + b ) + b ⁡ ( v - b ) ; a = 0. 4 ⁢ 5 ⁢ 7 ⁢ 2 ⁢ 4 ⁢ R 2 ⁢ T c 2 [ 1 + ( 0 . 3 ⁢ 7 ⁢ 4 ⁢ 6 + 1 . 5 ⁢ 4 ⁢ 2 ⁢ 3 ⁢ ω - 0 . 2 ⁢ 6 ⁢ 9 ⁢ 9 ⁢ ω 2 ) ⁢ ( 1 - T / T c ) ] 2 / P c ; and ⁢ b = 0.0778 R ⁢ T c / P c ;

• • where P is a system pressure, expressed in a unit of Pa; v is a molar volume, expressed in a unit of m³/mol; R is a gas constant (8.314 J/(mol·K)); a and b are each a characteristic parameter; Tc is a critical temperature, expressed in a unit of ° C.; Pc is a critical pressure, expressed in a unit of Pa; and ω is an acentric factor.

(c2) Setting and Inputting of Process Parameters

Preset process parameters were input into the distillation column model, where the preset process parameters included composition of the feed gas, operating pressure, temperature and column internal and packing parameters.

(c3) Setting of Process Constraint Conditions

The process constraint conditions included the carbon dioxide purity in the liquid stream ≥99%, and the theoretical stage flooding percentage within a range of 60-80%. Under the premise of satisfying the process constraint conditions, with the preset pressurization pressure in each of the first-stage compression, the second-stage compression and the third-stage compression and the cooling medium flow rate in each of the first-stage cooling, the second-stage cooling and the third-stage cooling as variables, for different feed gas compositions, an influence of different variable combinations on the total system energy consumption was analyzed through multi-variate cross combination analysis by using a sensitivity analysis tool of the Aspen process simulation software.

(c4) Collaborative optimization was performed on process parameters of the distillation column based on analysis results obtained in step (c3) with minimization of the total system energy consumption as the optimization objective to determine the optimized process parameters, where the process parameters comprise theoretical stage number, reflux ratio and feed position.

(c5) The optimized process parameters of the distillation column were output, including optimal theoretical stage number, optimal reflux ratio and optimal feed position.

The distillation column adopted a Mellapak 250X structured packing, with a height of each packing layer being 0.1-0.3 m and a column section diameter being 0.07-0.6 m.

The simulation results of Aspen software were shown in FIG. 2. The theoretical stage number of the distillation column was eight, counting from the column top and proceeding downward. The height of each packing layer was 0.25 m. The feed gas was introduced into a fourth theoretical stage. The Mellapak 250X structured packing was adopted in the distillation column. The column section diameters of theoretical stages 2-3 were each 0.09 m. The column section diameter of the fourth theoretical stage was 0.28 m. The column section diameters of theoretical stages 5-7 were each 0.4 m. The pressure of the condenser was 14 bar. The theoretical stage pressure drop was 0.1 bar. The column bottom temperature was −163° C. The column top temperature was −28° C. The reflux ratio was 0.87.

Under the following conditions of column section diameter of 0.45 m, Mellapak 250X packing, flooding rate of 60-80%, feed gas (−85° C., 18 bar) and flow rate of 292.83×10³ mol/h, a volumetric flow rate was approximately 250 m³/h, which was calculated by 292830 mol/h×8.314×188.15 K/(1.8×10⁶ Pa). A column section area was about 0.159 m², which was calculated by π×(0.45 m/2)². A superficial gas flow velocity was approximately 0.43 m/s, which was calculated by 250 m³/h=0.159 m²÷3600 s/h. Under a constraint condition of the flooding rate ranging from 60% to 80%, an inlet velocity of the feed gas entering the distillation column was roughly 0.2-0.5 m/s.

The operational results of the Aspen process simulation were described as follows. The flow rate entering the distillation column was 645 L/min. The column section diameter of the fourth packing layer was 0.28 m, and the column section area thereof was about 0.0615 m². Under an inlet state, the flow rate was converted to 0.01075 m³/s, and the superficial gas flow velocity was about 0.175 m/s, which was calculated by 0.01075÷0.0615. Due to the presence of liquid phase flow, the superficial velocities of the gas phase and the liquid phase may vary. Through variable-diameter design, the flooding rate was ultimately maintained within a safe range. The simulation results indicated a carbon dioxide recovery rate of 99.8%, a purity of 99.99% and a flooding rate of 67.8-80%. A schematic diagram of the distillation hydraulic calculation parameters was shown in FIG. 3.

(d) Low-Temperature Distillation Separation

The engineering design and construction of a small-scale low-temperature distillation column was completed based on the optimized process parameters obtained from the optimization design in step (c). The low-temperature distillation separation experiment was conducted according to the optimized process parameters.

The processed feed gas from step (b) (−85° C., 18 bar) was fed into the distillation column with the height of each packing layer being 4.85 m and the column section diameter being 0.45 m for the low-temperature distillation separation, so as to yield the column bottom liquid stream mainly containing carbon dioxide.

During the distillation and separation, the operation of key equipment was continuously monitored, and data of key operation parameters of key process points of the distillation column was collected in real time. The key operation parameters included the column bottom temperature and pressure, the column top temperature and pressure, the feed temperature and the feed pressure, as well as the outlet temperature, pressure and material flow rate for each stage of the three-stage compression and each stage of the three-stage cooling. The data was input into the Aspen process simulation software, followed by comparison with the optimized process parameters to output a deviation. In a case that the deviation exceeded the preset threshold, triggering an early-warning. The process parameters were adjusted according to the early-warning signal.

During the actual operation, fluctuations in the feed gas composition (e.g., a carbon dioxide concentration dropping from 98% to 84%) or deviations in the distillation column parameters (e.g., a top temperature rising from −28° C. to −23° C.) could cause changes in gas-liquid load within the column, thereby shifting the flooding rate away from the optimal range of 60-80%. The flooding rate was dynamically corrected by fine-tuning the target pressurization pressure of the compressors or the cooling medium flow rate of the coolers. While enabling the feed temperature of the low-temperature distillation column to range from −100° C. to −70° C. and the feed pressure to range from 1.4 MPa to 1.8 MPa, other parameters were guided back to their optimal ranges. Through optimization control, the total energy consumption of the system was minimized while enabling the feed gas to satisfy the distillation feed conditions.

(e) Collection and Filling of Product

The column bottom liquid stream from step (d) was pressurized by the pump to 2.3 MPa and transferred to the vacuum-insulated buffer tank for storage, so as to yield an industrial-grade carbon dioxide product. It was confirmed by experiments that the industrial-grade carbon dioxide product had a purity of 99.99-99.997%, a recovery rate of 99.8% and a water content of ≤5 ppm.

A detailed comparison was made between the measured data and the simulation results of the Aspen software for the optimization design of low-temperature distillation process parameters in step (c), as shown in Table 1.

TABLE 1
	Value of	Value of
	small-scale	Aspen	Relative
Parameter	test	simulation	deviation	Conclusion

Column top	97.8%	98.1%	<0.5%	high
N2 product				consistency
purity (mol %)
Column	99.9915%	99.9903%	<0.1%	high
bottom CO2				consistency
product purity
(mol %)
Recovery rate	98.9%	99.5%	<0.6%	high
of key				consistency
composition
Optimal reflux	0.8571	0.8759	~2.3%	high
ratio				consistency
Feed tray	−61	−61.6	<1.0%	high
temperature				consistency
(° C.)
Reboiler heat	397.51	384.016	~5%	within the
load (kW)				acceptable
				range of
				the project

Therefore, the Aspen simulation program was verified through small-scale experiments, and the prediction results were accurate and reliable, truly reflecting the process flow. Therefore, the current model could be directly applied to the subsequent process scale-up design, providing the theoretical plate number, the reflux ratio, the heat load and other key design parameters for industrial-grade equipment design, and serving as theoretical basis for operational optimization.

Embodiment 2

This embodiment adopted the same tail gas as the Embodiment 1, containing 98% by volume of CO₂, 2% by volume of N₂ and trace impurities. Step (a) of this embodiment was identical to that of the Embodiment 1. The differences between other steps of this embodiment and those of the Embodiment 1 lied in the following aspects.

(b) Three-Stage Compression and Three-Stage Cooling

(b1) First-Stage Compression+First-Stage Cooling

The feed gas was compressed from 3 bar to 6 bar through the first compressor, then cooled to 0° C. through ethylene glycol. The water content was reduced to below 5 ppm through the gas-liquid separator.

(b2) Second-Stage Compression+Second-Stage Cooling

The feed gas was compressed to 11 bar through the second compressor, then cooled to −40° C. through liquid nitrogen.

(b3) Third-Stage Compression+Third-Stage Cooling

The gas was compressed to 18 bar through the third compressor, then cooled to −95° C. through liquid nitrogen, thereby completing the compression and cooling of the feed gas to obtain a processed gas.

(c) Optimization Design of Low-Temperature Distillation Process Parameters

In step (c), the processed feed gas had a temperature of −95° C. and a pressure of 18 bar. The simulation results of Aspen software were shown in FIG. 4. The distillation column had nine theoretical stages. The feed gas was introduced into a fifth theoretical stage. The column section diameters of theoretical stages 2-4 were each 0.09 m. The column section diameters of theoretical stages 5-8 were each 0.43 m. A schematic diagram of the theoretical stage parameters for distillation was shown in FIG. 4.

To verify the separation efficiency of distillation at the specified conditions (−95° C., 18 bar), the Aspen software was utilized to perform a simulation calculation under the premise of the feed gas flow rate of 292.83×10³ mol/h. The simulation results showed a carbon dioxide recovery rate of 99.89%, a purity of 99.997% and a flooding rate of 67.9-78.3%. A schematic diagram of the distillation hydraulic calculation parameters was shown in FIG. 5.

Embodiment 3

The tail gas processed in this embodiment mainly contained 84% by volume of CO₂, 16% by volume of N₂ and trace impurities. The relative proportions of the two main components differed from those in the Embodiment 1, which was attributable to different processes.

Despite the variation in tail gas composition, step (a) in this embodiment was performed through steps as described in the Embodiment 1.

(b) Three-Stage Compression and Three-Stage Cooling

(b1) First-Stage Compression+First-Stage Cooling

The feed gas was compressed from 3 bar to 6 bar through the first compressor, then cooled to 0° C. through ethylene glycol. The water content was reduced to below 5 ppm through the gas-liquid separator.

(b2) Second-Stage Compression+Second-Stage Cooling

The feed gas was compressed to 11 bar through the second compressor, then cooled to −40° C. through liquid nitrogen.

(b3) Third-Stage Compression+Third-Stage Cooling

The gas was compressed to 18 bar through the third compressor, then cooled to −100° C. through liquid nitrogen, thereby completing the pressurization and cooling of the feed gas to obtain a processed gas.

(c) Optimization Design of Low-Temperature Distillation Process Parameters

In this case, the processed feed gas had a temperature of −100° C. and a pressure of 18 bar. The simulation results of Aspen software were shown in FIG. 6. The distillation column had eight theoretical stages, with a height of each packing layer being 1.2 m. The feed gas was introduced into a third theoretical stage. The column section diameter of a second theoretical stage was 0.14 m. The column section diameter of the third theoretical stage was 0.29 m. The column section diameters of theoretical stages 4-7 were each 0.4 m. A schematic diagram of the theoretical stage parameters for distillation was shown in FIG. 6.

To verify the separation efficiency of distillation at the specified conditions (−100° C., 18 bar), the Aspen software was utilized to perform the simulation calculation under the premise of the feed gas flow rate of 292.83×10³ mol/h. The simulation results showed a carbon dioxide recovery rate of 99.91%, a purity of 99.999% and a flooding rate of 60.5-79.9%. A schematic diagram of the distillation hydraulic calculation parameters was shown in FIG. 7.

Comparative Example 1

The tail gas processed in this comparative example mainly contained 97.74% by volume of CO₂, 2% by volume of N₂ and trace impurities, such as 2,000 ppm (0.2%) H₂, 500 ppm (0.05%) CO, 100 ppm (0.01%) H₂O and 1 ppm (0.0001%) methanol.

In step (a), the tail gas mainly containing CO₂ and N₂ was emitted from the MSR section, and directly used as the feed gas for the compression and cooling without impurity pretreatment. No gas-liquid separator was employed in step (b).

Through the Aspen Plus software simulation, it was showed that for this comparative example, the unprocessed tail gas was directly subjected to compression, cooling and low-temperature distillation to obtain a liquid carbon dioxide product with impurity concentrations of H₂ at 0.079 ppm, CO at 38.7 ppm, H₂O at 97 ppm and methanol at 0.9 ppm.

From the safety perspective, residual hydrogen in the processed feed gas should be kept as low as possible, generally not exceeding 1,000 ppm. During the low-temperature distillation, carbon monoxide may form azeotropes with other impurities, affecting the operation of the distillation column and product quality, so it generally needed to be reduced to below 10 ppm after pretreating. Water could freeze at a low temperature, blocking pipelines and equipment, so that it typically needs to be reduced to below 5 ppm after pretreating.

In this comparative example, the feed gas was subjected directly to compression, cooling and low-temperature distillation without pretreatment, causing a high risk of water freezing inside the column and potentially clogging the packing or pipelines and disrupting the distillation process. Other impurities also had some adverse effects on the distillation process.

Therefore, for this comparative example, the present disclosure provided a four-stage pretreatment method, as shown in Table 2.

TABLE 2
Pretreatment	Process method	Core reasons for choosing
First-stage	Polytetrafluoroethylene	1. The feed gas might have contained catalyst powder and
pretreatment	(PTFE)-coated filter	dust from the MRS, requiring the removal of solid
	cartridge filtration	impurities to prevent the pores of the follow-up adsorbent
		from being clogged.
		2. The PTFE-coated material exhibited good temperature
		resistance, allowing the replacement cycle to be extended to
		1-2 months, which met the requirements of industrial
		continuous production.
		3. This filtration unit provided a filtration accuracy of 1-5
		μm, enabling the complete interception of solid particles,
		and maintained a low operating pressure drop (≤0.1 bar),
		making it not affect the pressure coordination of subsequent
		units.
Second-stage	Silica gel-activated	1. Under the distillation condition of −85° C., water
pretreatment	alumina composite	would freeze, requiring the water content to be reduced
	fixed bed adsorption	to below 5 ppm. Specifically, silica gel was employed
		to preferentially adsorb the bulk of the moisture, while
		activated alumina was used for the deep removal of
		residual water (achieving levels below 5 ppm), with the
		two-step synergy fulfilling the dehydration
		requirement.
		2. The two adsorbents exhibited low adsorption
		capacity for CO2, resulting in minimal CO2 loss, which
		ensured that the target of 98% CO2 purity was not
		affected.
		3. The nitrogen used in the regeneration process of the
		adsorption bed could be recycled from the top stream
		nitrogen of the distillation column, thereby reducing
		energy consumption and costs.
Third-stage	Silanized modified	1. It had a high selectivity for methanol, failing to
pretreatment	coal-based activated	adsorb CO2 and N2, so as to deeply remove methanol.
	carbon adsorption	2. It had a high mechanical strength, which was suitable
		for a large handling capacity.
Fourth-stage	H2/CO selective	The feed gas containing 2,000 ppm H2 and 500 ppm CO
pretreatment	adsorption fixed bed (a	if directly fed into the distillation column, would
	palladium-based	increase the vent gas load at the column top and
	adsorbent or a copper-	potentially cause gas velocity fluctuations. The
	based adsorbent)	palladium-based adsorbent was used to preferentially
		adsorb H2, while the copper-based adsorbent
		chemically adsorbed CO enabling their respective
		reduction to below 500 ppm and 100 ppm.

Among the above pretreatment methods, all the treatment processes except for dehydration treatment were set in step (a) of the technical solutions of the present disclosure. As for the dehydration treatment, its setting position was relatively flexible. It could be set either in step (a) or after first-stage compression and cooling in step (b).

Comparative Example 2

This comparative example adopted the same tail gas as the Embodiment 1, and the same feed gas (98% CO₂+2% N₂) was obtained after the pretreatment in step (a).

The difference from the Embodiment 1 lied in that the parameters in step (b) of this comparative example were adjusted. The specific parameters were set below.

(b) Three-Stage Compression and Three-Stage Cooling

(b1) First-Stage Compression+First-Stage Cooling

The feed gas was compressed from 3 bar to 6 bar through the first compressor, then cooled to 0° C. through ethylene glycol. Subsequently, the water content was reduced to below 5 ppm through the gas-liquid separator.

(b2) Second-Stage Compression+Second-Stage Cooling

The feed gas was compressed to 11 bar through the second compressor, followed by cooling to −15° C. through liquid nitrogen.

(b3) Third-Stage Compression+Third-Stage Cooling

The feed gas was then compressed to 18 bar through the third compressor, and further cooled to −30° C. through liquid nitrogen, thereby completing the compression and cooling of the feed gas to obtain a processed feed gas.

The optimization parameters for the low-temperature distillation process adopted in step (c) were the same as in Embodiment 3.

To verify the separation efficiency of distillation at the specified feed conditions (−30° C., 18 bar), the Aspen software was utilized to perform a simulation calculation under the premise of the feed gas flow rate of 292.83×10³ mol/h. The simulation results showed a carbon dioxide recovery rate of 98.87% and a purity of 98.97%. At this temperature, the low-temperature distillation method failed to yield an industrial-grade carbon dioxide product.

Under the following conditions of the feed pressure 18 bar and the feed gas composition of 98% CO₂ and 2% N₂, the results of different low-temperature distillation temperatures and the unit capture cost per ton of carbon dioxide, in different embodiments and comparative examples, were compared, as shown in Table 3.

TABLE 3
Project	Feed temperature (° C.)	Feed pressure(bar)	Purity	Yield	Cost/t

	−20	18	98.02%	97.92%	53.1
Comparative	−30	18	98.97%	98.87%	80.3
example 2
	−35	18	99.09%	98.99%	84.0
	−40	18	99.33%	99.24%	87.2
	−45	18	99.56%	99.46%	89.7
	−50	18	99.73%	99.63%	91.9
	−55	18	99.85%	99.75%	93.9
	−60	18	99.92%	99.82%	95.7
	−65	18	99.96%	99.86%	97.4
	−70	18	99.98%	99.88%	99.1
	−75	18	99.98%	99.89%	100.8
	−80	18	99.99%	99.89%	102.4
Embodiment 1	−85	18	99.99%	99.89%	104.0
	−90	18	99.99%	99.89%	105.6
Embodiment 2	−95	18	99.99%	99.90%	107.2
	−100	18	99.99%	99.90%	108.7

Under the following conditions of the feed pressure 18 bar and the feed gas composition of 98% CO₂ and 2% N₂, the process energy consumption and capture cost at different distillation temperatures were simulated and compared through Aspen software. The results of the capture process energy consumption were shown in Table 4, where the energy consumption was expressed in a unit of kWh.

TABLE 4
Temperature	Compressor	Compressor	Compressor	Cooler	Cooler	Cooler
(° C.)	1	2	3	1	2	3	Pump	Reboiler	Condenser

−20	169	387	73	311	465	104	0	8	66
−30	169	387	73	311	465	1106	0	3	48
−35	169	387	73	311	465	1202	0	46	46
−40	169	387	73	311	465	1256	0	113	40
−45	169	387	73	311	465	1298	0	168	34
−50	169	387	73	311	465	1335	0	213	30
−55	169	387	73	311	465	1369	0	253	26
−60	169	387	73	311	465	1402	0	288	24
−65	169	387	73	311	465	1433	0	320	23
−70	169	387	73	311	465	1463	0	351	22
−75	169	387	73	311	465	1493	0	381	22
−80	169	387	73	311	465	1523	0	411	22
−85	169	387	73	311	465	1552	0	440	22
−90	169	387	73	311	465	1581	0	469	22
−95	169	387	73	311	465	1609	0	497	22
−100	169	387	73	311	465	1638	0	525	22

Based on different capture energy consumption levels, the unit capture cost per ton of carbon dioxide was calculated. The cost per ton was expressed in a unit of CNY, and subsequent costs were each expressed in a unit of ×10⁴ CNY. Electricity costs were calculated at 0.35 CNY/kWh in Xinjiang region. The cost of the distillation column was calculated based on construction cost of a single column. The costs for the compressors and the coolers were calculated based on three units each. The planned equipment operation lifespan was 20 years. Maintenance costs were calculated at 10×10⁴ CNY/year. Labor costs were calculated at 10×10⁴ CNY/person/year, with a design team of six personnel responsible for the system. Based on the above parameters, the unit capture cost per ton of carbon dioxide was calculated, with the results shown in Table 5.

TABLE 5
				Elec-	Elec-	Instal-
				tricity	tricity	lation
		Total		consumption	bill	of the
		energy		over 20	for 20	distillation				Mainte-
Feed		con-		years	years	column	Compressor	Cooler	Pump	nance	Labor
temperature		sumption	Cost/t	(×104	(×104	(×104	(×104	(×104	(×104	(×104	(×104
(° C.)	Purity	kWh/year	(CNY/ton)	kWh)	CNY)	CNY)	CNY)	CNY)	CNY)	CNY)	CNY)

−20	98.02%	1.25E+07	53.13	25073	8775	300	150	120	15	200	1200
−30	98.97%	2.03E+07	80.28	40588	14206	300	150	120	15	200	1200
−35	99.09%	2.14E+07	84.05	42741	14959	300	150	120	15	200	1200
−40	99.33%	2.23E+07	87.23	44560	15596	300	150	120	15	200	1200
−45	99.56%	2.30E+07	89.75	46,000	16100	300	150	120	15	200	1200
−50	99.73%	2.36E+07	91.91	47233	16531	300	150	120	15	200	1200
−55	99.85%	2.42E+07	93.86	48351	16923	300	150	120	15	200	1200
−60	99.92%	2.47E+07	95.69	49395	17288	300	150	120	15	200	1200
−65	99.96%	2.52E+07	97.42	50382	17634	300	150	120	15	200	1200
−70	99.98%	2.57E+07	99.11	51347	17971	300	150	120	15	200	1200
−75	99.98%	2.61E+07	100.77	52295	18303	300	150	120	15	200	1200
−80	99.99%	2.66E+07	102.40	53229	18630	300	150	120	15	200	1200
−85	99.99%	2.71E+07	104.01	54150	18953	300	150	120	15	200	1200
−90	99.99%	2.75E+07	105.61	55061	19271	300	150	120	15	200	1200
−95	99.99%	2.80E+07	107.18	55961	19587	300	150	120	15	200	1200
−100	99.99%	2.84E+07	108.74	56853	19899	300	150	120	15	200	1200

Comparative Example 3

This comparative example adopted the same tail gas as Embodiment 3, and the same feed gas (84% CO₂+16% N₂) was obtained after the pretreatment in step (a).

The difference from Embodiment 3 lay in that the parameters in step (b) of this comparative example were adjusted. The specific parameters were set below.

(b) Three-Stage Compressor and Three-Stage Cooling Processes

(b1) First-Stage Compressor+First-Stage Cooling

The feed gas was compressed from 3 bar to 6 bar through the first compressor, then cooled to 0° C. through ethylene glycol. Subsequently, the water content was reduced to below 5 ppm through the gas-liquid separator.

(b2) Second-Stage Compression+Second-Stage Cooling

The feed gas was compressed to 11 bar through the second compressor, followed by cooling to −15° C. through liquid nitrogen.

(b3) Third-Stage Compression+Third-Stage Cooling

The feed gas was then compressed to 18 bar through the third compressor, and further cooled to −55° C. through liquid nitrogen, thereby completing the compression and cooling of the feed gas to obtain a processed feed gas.

The optimization parameters for the low-temperature distillation process adopted in step (c) were the same as in Embodiment 3.

To verify the separation efficiency of distillation at the specified feed conditions (−55° C., 18 bar), the Aspen software was utilized to perform a simulation calculation under the premise of the feed gas flow rate of 292.83×10³ mol/h. The simulation results showed a carbon dioxide recovery rate of 96.74% and a purity of 96.82%. At this temperature, the low-temperature distillation method failed to yield an industrial-grade carbon dioxide product.

Under the following conditions of the feed pressure 18 bar and the feed gas composition of 84% CO₂ and 16% N₂, the results of different low-temperature distillation temperatures and the unit capture cost per ton of CO₂ were compared, as shown in Table 6.

TABLE 6
Project	Feed temperature (° C.)	Feed pressure(bar)	Purity	Yield	Cost/t

	−50	18	96.51%	96.43%	82.3
Comparative	−55	18	96.82%	96.74%	84.1
example 3
	−60	18	97.28%	97.20%	85.3
	−65	18	97.68%	97.60%	86.4
	−70	18	98.19%	98.11%	87.3
	−75	18	98.82%	98.74%	91.6
	−80	18	99.30%	99.21%	94.8
	−85	18	99.63%	99.55%	97.1
	−90	18	99.85%	99.77%	98.9
	−95	18	99.97%	99.89%	100.5
Embodiment 3	−100	18	99.99%	99.92%	102.0

Under the following conditions of the feed pressure of 18 bar and the feed gas composition of 84% CO₂ and 16% N₂, the process energy consumption and capture cost at different distillation temperatures were simulated and compared through Aspen software. The results of the capture process energy consumption were shown in Table 7, where the energy consumption was expressed in a unit of kWh.

TABLE 7
Temperature	Compressor	Compressor	Compressor	Cooler	Cooler	Cooler
(° C.)	1	2	3	1	2	3	Pump	Reboiler	Condenser

−50	170	391	75	306	459	1035	2	55	169
−55	170	391	75	306	459	1102	2	83	164
−60	170	391	75	306	459	1157	2	101	156
−65	170	391	75	306	459	1204	2	115	149
−70	170	391	75	306	459	1245	2	124	140
−75	170	391	75	306	459	1283	2	130	129
−80	170	391	75	306	459	1319	2	221	119
−85	170	391	75	306	459	1352	2	280	110
−90	170	391	75	306	459	1383	2	322	102
−95	170	391	75	306	459	1413	2	355	96
−100	170	391	75	306	459	1442	2	384	94

Based on different capture energy consumption levels, the unit capture cost per ton of carbon dioxide was calculated. The cost per ton was expressed in a unit of CNY, and subsequent costs were expressed in a unit of ×10⁴ CNY. Electricity costs were calculated at 0.35 CNY/kWh in Xinjiang region (with reference to the 2025 Xinjiang Grid latest electricity price grade classification and standards from https://www.toutiao.com/article/7522015259433746971/?upstream biz-doubao&source=m_redirect&wid=1761099025540). The cost of the distillation column was calculated based on construction cost of a single column. The costs for the compressors and the coolers were calculated based on three units each. The planned equipment operational lifespan was 20 years. Maintenance costs were calculated at 10×10⁴ CNY/year. Labor costs were calculated at 10×10⁴ CNY/person/year, with a design team of 6 personnel responsible for the system. Based on the above parameters, the unit capture cost per ton of carbon dioxide was calculated, with the results shown in Table 8.

TABLE 8
				Elec-	Elec-	Instal-
				tricity	tricity	lation
		Total		consumption	bill	of the
		energy		over 20	for 20	distillation				Mainte-
Feed		con-		years	years	column	Compressor	Cooler	Pump	nance	Labor
temperature		sumption	Cost/t	(×104	(×104	(×104	(×104	(×104	(×104	(×104	(×104
(° C.)	Purity	kWh/year	(CNY/ton)	kWh)	CNY)	CNY)	CNY)	CNY)	CNY)	CNY)	CNY)

−50	96.51%	2.07E+07	82.3	41366.94	14478.43	300	150	120	15	200	1200
−55	96.82%	2.12E+07	84.1	42386.63	14835.32	300	150	120	15	200	1200
−60	97.28%	2.15E+07	85.3	43078.8	15077.58	300	150	120	15	200	1200
−65	97.68%	2.18E+07	86.4	43675.35	15286.37	300	150	120	15	200	1200
−70	98.19%	2.21E+07	87.3	44193.51	15467.73	300	150	120	15	200	1200
−75	98.82%	2.33E+07	91.6	46653.86	16328.85	300	150	120	15	200	1200
−80	99.30%	2.42E+07	94.8	48491.92	16972.17	300	150	120	15	200	1200
−85	99.63%	2.49E+07	97.1	49817.34	17436.07	300	150	120	15	200	1200
−90	99.85%	2.54E+07	98.9	50858.55	17800.49	300	150	120	15	200	1200
−95	99.97%	2.59E+07	100.5	51748.36	18111.92	300	150	120	15	200	1200
−100	99.99%	2.63E+07	102.0	52632.13	18421.25	300	150	120	15	200	1200

In this case, the annual total energy consumption in Table 8 was calculated based on the sum of energy consumption from Table 7. The calculation formula was expressed as:

total energy consumption from Table 7 (kWh)×24 (h/day)×30 (days/month)×11 (months/year);

• • where the equipment operated for 11 months per year (30 days per month), with the remaining one month allocated for maintenance, and the theoretical annual electricity cost was derived by multiplying the total energy consumption (kWh) by the local electricity price (0.35 CNY/kWh).

In the prior art, as referenced in the “China carbon capture, utilization and storage (CCUS) cost assessment” by the Yan'an municipal ecology and environment bureau (source: https://mp.weixin.qq.com/s/-2XqZdRydyDXD8Dj7tiQuQ), it was indicated that in the low-temperature methanol washing (LTMW) method of the Yanchang petroleum CCUS integrated project, carbon dioxide originated from a pre-combustion process within coal gasification (i.e., a production process of syngas in coal gasification). It also indicated that the feed gas had a higher carbon dioxide purity. Compared to other carbon dioxide capture and transportation projects, the capture and operating costs of the Yanchang Petroleum CCUS integrated project decreased by approximately 26.4%, reaching only 26.5 USD/ton of carbon dioxide. The capture cost was 122.5 CNY/ton of carbo dioxide, equivalent to 17.52 USD/ton of carbon dioxide, and the transportation cost was 9.03 USD/ton of carbon dioxide.

The technical costs of each stage of the low-temperature distillation method provided herein were shown in Table 9.

TABLE 9
CCUS technical costs of each stage from 2025 to 2060

Year	2025	2030	2035	2040	2050	2060

Capture cost	pre-	100~180	90~130	70~80	50~70	30~50	20~40
(CNY/ton)	combustion
	capture
	post-	230~310	190~280	160~220	100~180	80~150	70~120
	combustion
	capture
	oxy-fuel	300~480	160~390	130~320	110~230	90~150	80~130
	combustion
	capture
Transportation cost	tank truck	0.9~1.4	0.8~1.3	0.7~1.2	0.6~1.1	0.5~1.1	0.5~1
(yuan/(ton ·	transportation
kilometer))	pipeline	0.8	0.7	0.6	0.5	0.45	0.4
	transportation

Storage cost (CNY/ton)	50~60	40~50	35~40	30~35	25~30	20~25
Note:
Costs included fixed costs and operating costs.
Data source:
[1] Feng, Wang et al., Technology and economy analysis of 660 MW coal-fired power unit with 1 Mt/a CO2 capture system [J]. Clean Coal Technology, 2016, 22(6): 101-105, 39.
[2] Jiajia, Liu et al., Modeling and optimization of the Whole Process of CO2 capture, transportation, oil displacement and storage [J]. Oil-gasfield Surface Engineering. 2018, 37(10): 1-5.
[3] Shuying, Zhu et al., Research on carbon dioxide emission reduction technology and cost in China's cement industry [J]. Environment Engineering. 2021, 39(10): 15-22.
[4] Chaofei, Nie et al., 2024. Strategy of clean and efficient use of China's modern coal chemical industry [J]. Resources & Industries, 26(3): 6-20.

In summary, the low-temperature distillation method of this disclosure exhibited significant advantages in terms of carbon capture costs. Compared with the LTMW method, this process achieved a reduction of approximately 15-20% in overall cost. Through process innovation, this disclosure enabled a lower energy and material consumption in the capture process, thereby achieving the considerable reduction in overall costs. The beneficial effects made this technology have strong market competitiveness and application prospects in the field of industrial-scale CCUS.

Comparative Example 4

This comparative example was substantially identical to the Embodiment 1. The difference from the Embodiment 1 lay in that a “two-stage pressurization and two-stage cooling” process in step (b) was adopted. The specific parameters were set below.

(b1) First-Stage Compression+First-Stage Cooling

The feed gas was compressed from 3 bar to 7 bar through the first compressor, then cooled to 0° C. through ethylene glycol. Subsequently, the water content was reduced to below 5 ppm through the gas-liquid separator.

(b2) Second-Stage Compression+Second-Stage Cooling

Next, the feed gas was compressed to 18 bar through the second compressor, followed by cooling to −85° C. through liquid nitrogen, so as to obtain a processed feed gas.

Low-Temperature Distillation Separation in Step (c)

The processed feed gas (−85° C., 18 bar) from step (b) was introduced into the distillation column for the low-temperature distillation separation.

However, the analysis results indicated that the process had the following drawbacks.

The two-stage pressurization in this comparative example caused higher compression ratios, which accelerated equipment wearing and aging and reduced system operational stability. In contrast, the technical solutions of the present disclosure adopted three-stage compression to distribute the total compression ratio (18/3=6) across three lower-stage ratios (2.0, 1.83, 1.64), so as to avoid the high temperature and the high mechanical load caused by the high single-stage compression ratio, thereby complying with compressor design specifications and significantly extending the service life of the equipment.

The analysis results were shown in Table 10.

TABLE 10
		Comparative
Comparative item	Embodiment 1	example 4	Core differences and basis
Compression	first stage: 3-6 bar	first stage: 3-7 bar	Compression ratio was a key parameter
stages and	(compression	(compression ratio ≈	affecting compressor service-life. The
compression ratio	ratio = 2.0)	2.33)	reciprocating compressor design code
	second stage: 6-11	second stage: 7-18	(GB/T 13279) recommends a single-stage
	bar (compression	bar (compression	compression ratio ≤3. Exceeding this
	ratio ≈ 1.83)	ratio = 2.57)	limit significantly increases equipment
	third stage: 11-18		load.
	bar (compression		In the Embodiment 1, the compression
	ratio ≈ 1.64)		ratio in all stages was each ≤2.0, within
			the safe range. Low compression ratios
			extended the replacement cycles for
			wearing parts and reduced maintenance
			frequency.
			In the Comparative example 4, the
			compression ratio of 2.57, while still safe,
			approached the recommended upper
			limit, carrying a higher risk of
			concentrated load. Furthermore, the
			friction coefficient between a piston ring
			and a cylinder wall increased, doubling a
			wearing rate per unit time.
Gas temperature	After the first-	After the first-stage	A flash point of typical compressor
after single-stage	stage adiabatic	adiabatic	lubricating oil was 160° C. According to
compression	compression, gas	compression, gas	the adiabatic compression temperature
	temperature ≈	temperature ≈ 140° C.	change formula, expressed as
	120° C.		T2 = T1 × (P2/P1){circumflex over ( )}[(k − 1)/k], the
			temperatures in the Embodiment 1 were
			farther from the flash point, avoiding
			oxidation and degradation of lubricating
			oil due to high temperatures.
			Although the temperature of 140° C. in the
			Comparative example 4 did not exceed
			the flash point, it accelerated the aging of
			sealing elements, causing a seal lifespan
			20% shorter than in the Embodiment 1.
Equipment design	designed service	designed service	Operational data from industrial
life and failure rate	life of over 10	life of over 7-8	carbon dioxide compression systems
	years, with an	years, with an	verified that for every reduction of 0.5
	annual failure	annual failure rate ≤5%	in single-stage compression ratio,
	rate ≤2% and a mean	and a mean time	equipment lifespan increased by 1.5-2
	time between	between failures	years. The low-load design in the
	failures (MTBF)	(MTBF) reaching	Embodiment 1 led to annual
	reaching over	about 6,000 hours	maintenance costs 40% lower than the
	12,000 hours		Comparative example 4, offering
			superior long-term operational
			economics. This design particularly
			suited the high demand for equipment
			stability in large-scale, continuous
			production.
Process	three-stage	The second stage	Low-temperature distillation imposed
compatibility and	compression +	compression ratio	strict requirements on the stability of
operational	gentle cooling	of 2.57 caused	the feed stream conditions.
stability	gradient,	larger fluctuations	In the Embodiment 1, the gentle
	continuous and	in the compressed	pressurization and cooling method
	stable variations	gas state, requiring	effectively minimized fluctuations in
	in gas pressure	more precise	feed parameters, thereby preventing
	and temperature	control of the	the distillation column flooding rate
		cooling rate to	from deviating beyond the optimal
		match the	range of 60-80%. In contrast, a high
		requirements of	compression ratio in the second stage
		subsequent	of Comparative example 4 was prone
		distillation.	to causing variations in the gas states,
			which increased the difficulty of
			controlling the distillation parameters.

The embodiments described are merely some preferable embodiments, and are not intended to limit this present disclosure. Although the preferable embodiments have been described above, the preferable embodiments should not be constructed as limitations of this present disclosure. For those of ordinary skill in the art, some modifications and equivalent replacements can be made to the embodiments described without departing from the scope of the present disclosure, and shall fall within the scope of the present disclosure.