Salt solution salt separation method driven by heat energy of distilled ammonia waste liquid

Abstract

The invention relates to the technical field of industrial waste heat recycling and brine resource refining, in particular to a salt solution salt separation method driven by heat energy of distilled ammonia waste liquid, which comprises the following steps: the distilled ammonia waste liquid is subjected to clarification and sedimentation pretreatment and then is fed into a shell pass of an immersed pressure delay membrane distillation device; a salt solution with the pressure of 20-24 bar flows in a tube pass of the immersed pressure delay membrane distillation device; transferring water molecules in the distilled ammonia clear liquid into the salt solution in the form of steam to obtain a dilute salt solution; pressurizing the dilute salt solution, and separating salt by using a nanofiltration membrane to obtain a monovalent salt penetrating fluid and a divalent salt concentrated solution; and pressurizing the monovalent salt penetrating fluid, and performing reverse osmosis salt separation to obtain strong brine and fresh water. According to the invention, a PRMD-NF-RO multi-process efficient coupling system is constructed, collaborative optimization of triple functions of waste heat recovery, mechanical energy supply and fresh water recovery is realized, and the comprehensive energy consumption and cost of sea brine refining are reduced.

Description

Technical Field

[0001] This invention belongs to the field of membrane separation technology, specifically relating to a salt separation method for a salt solution driven by the thermal energy of ammonia stripping waste liquid generated in the ammonia-soda process for soda ash production. Background Technology

[0002] In the soda ash industry, the ammonia-soda process is widely used due to its mature technology and stable product quality. However, this process generates approximately 10m³ of waste gas per ton of soda ash produced. 3 The ammonia stripping wastewater is characterized by three high properties: high temperature (~95℃), high salinity (TDS>100,000mg / L), and high corrosivity (pH>10). Traditional treatment involves direct discharge into slag dumps, resulting in a massive waste of freshwater resources and low-quality heat energy, as well as severe environmental pressures and land resource occupation issues. How to efficiently and environmentally recover and utilize the water and heat from the ammonia stripping wastewater has become a common technical challenge restricting the sustainable development of the soda ash industry.

[0003] Meanwhile, in the field of sea brine refining, technological approaches are constantly evolving. Early methods employed brine refining 1.0 technology, which directly used sea brine to produce soda ash. However, this method resulted in high salt consumption and numerous impurities, severely impacting the quality of soda ash. Subsequently, brine refining 2.0 technology was developed, using nanofiltration (NF) membranes to pre-treat the sea brine for salt separation, effectively removing calcium... 2+ Mg 2+ The presence of divalent ions significantly improved the yield and quality of soda ash and reduced the consumption of solid salts. To further improve the utilization rate of NaCl, reverse osmosis (RO) technology was introduced to deeply concentrate the NF permeate, forming an NF-RO coupled process. However, both NF and RO are essentially high-energy-consuming processes that "exchange energy for quality," with energy consumption as high as 1.40~2.50 kWh per ton of water, becoming a bottleneck for their large-scale promotion.

[0004] Against this backdrop, the core idea of ​​brine refining 3.0 technology was born: to seamlessly couple the membrane separation process with renewable energy or industrial waste energy to achieve energy self-sufficiency or partial self-sufficiency, thereby breaking through the constraints of high energy consumption. The massive amount of low-quality heat energy contained in ammonia stripping waste liquid can theoretically provide an ideal energy source for this concept. However, existing technologies face three major technical challenges: 1) Energy form conversion barrier: The mechanical energy (high pressure) required for thermal energy and membrane separation are different forms of energy. Traditional heat-work conversion technologies (such as steam turbines and organic Rankine cycles) cannot be directly applied to low-quality heat sources at <100℃ and require complex intermediate conversion links, resulting in low efficiency. 2) Operating condition adaptability barrier: The composition of ammonia stripping waste liquid is complex and highly corrosive. Direct use in heat exchange can easily lead to scaling and corrosion of equipment, resulting in a very short lifespan for conventional heat exchange equipment. 3) System integration barriers: Although existing membrane distillation technologies (such as DCMD and VMD) can utilize waste heat, they can only produce fresh water and cannot directly provide the high pressure required to drive the NF-RO process. The energy conversion and membrane separation processes are disconnected, resulting in low overall system efficiency. Summary of the Invention

[0005] The purpose of this invention is to provide a salt solution separation method driven by the thermal energy of ammonia stripping wastewater. By constructing a highly efficient coupling system of a pressure delayed membrane distillation (PRMD) device, a nanofiltration (NF) device, and a reverse osmosis (RO) device, the low-quality thermal energy contained in the high-temperature, high-salt, and highly corrosive ammonia stripping wastewater generated during the ammonia-soda process is directly and efficiently converted into the mechanical energy required to drive the subsequent membrane separation process, while simultaneously recovering its freshwater resources. This significantly reduces the overall energy consumption and cost of sea brine refining and solves the environmental protection and resource waste problems faced by the soda ash industry.

[0006] To achieve the above objectives, the present invention adopts the following technical solution: This invention provides a method for salt separation of a salt solution driven by the thermal energy of ammonia stripping waste liquid, comprising the following steps: (1) Add a polymeric surfactant to the ammonia stripping waste liquid, clarify and settle to obtain ammonia stripping clear liquid; the temperature of the ammonia stripping clear liquid is 92~95℃, the surface tension is >65mN / m, and the SS value is <500mg / L; (2) The ammonia distillate is fed into the shell side (hot side) of the submerged pressure delayed membrane distillation apparatus, and a salt solution with a pressure of 20~24 bar flows in the tube side (cold side) of the submerged pressure delayed membrane distillation apparatus. Driven by the temperature difference and concentration difference on both sides of the pressure delayed membrane, water molecules in the ammonia distillate migrate to the salt solution in the form of vapor to obtain a dilute salt solution and ammonia distillate concentrate. When the temperature of the ammonia distillate concentrate drops to 70~85℃, the ammonia distillate concentrate in the shell side is discharged. (3) Pressurize the dilute salt solution to 30~35 bar and send it into the nanofiltration membrane device to separate the salts and obtain monovalent salt permeate and divalent salt concentrate; (4) Pressurize the monovalent salt permeate to 60-70 bar and send it into the reverse osmosis unit to obtain concentrated brine and fresh water.

[0007] Preferably, the pressure delay membrane is a plasma-treated fluorinated PTFE curtain membrane with a liquid wetting pressure >25 bar, a temperature tolerance ≥90℃, strong alkali corrosion resistance (pH>10), a membrane surface energy ≤15 mN / m, a contact angle ≥135°, a roll-off angle ≤10°, an average pore size of 0.1~0.35 μm, and a membrane thickness of 90~180 μm. This design ensures that the membrane can operate stably for a long time without pressure wetting or chemical corrosion failure even under operating pressures as high as 20~24 bar in the tube side, providing a key technological guarantee for the direct conversion of thermal energy to mechanical energy.

[0008] The method for preparing the plasma-fluorinated PTFE curtain membrane includes the following steps: (a) Place the PTFE base film in a vacuum plasma reaction chamber and evacuate to a pressure ≤10×10 -5 bar; (b) Fluoride gas is introduced, and the pressure inside the reaction chamber is maintained at 20 × 10⁻⁶. -5 bar ~100×10 -5 bar; (c) Plasma treatment is performed on the PTFE base film to obtain plasma-fluorinated PTFE curtain film; the applied radio frequency power is 50~200W and the treatment time is 60~300 seconds.

[0009] Preferably, the PTFE base membrane is a porous hydrophobic membrane with an average pore size of 0.1~0.35μm, a thickness of 90~180μm, a porosity of 75~85%, and an initial contact angle of 120~125°.

[0010] Preferably, the fluorinated gas is selected from CF. 4 At least one of C2F6 or SF6.

[0011] Preferably, the concentrated brine and salt solution obtained in step (4) are sent to a pressure exchange device for pressure exchange, the salt solution is pressurized to 20~24 bar, and then sent to the tube side of an immersion pressure delayed membrane distillation device.

[0012] Preferably, in step (4), pressurizing the monovalent salt permeate to 60-70 bar means first sending the monovalent salt permeate and the divalent salt concentrate into a pressure exchange device to exchange pressure and increase the pressure of the monovalent salt permeate, and then using a booster pump to pressurize the monovalent salt permeate to 60-70 bar.

[0013] Preferably, the pressure delayed permeation membrane has aeration tubes uniformly distributed on its surface to provide microbubbles for flushing the membrane surface. The microbubbles have a size of 50~200μm and an aeration intensity of 0.5~1.5Nm. 3 / (m 2 ·h).

[0014] Compared with the prior art, the technical solution of the present invention has the following significant advantages: (1) Efficient direct utilization of low-quality heat sources: This invention realizes the direct conversion of low-quality heat energy below 95℃ into mechanical energy through pressure delayed membrane distillation (PRMD) technology, completely breaking the traditional multi-stage conversion mode of "heat energy → electrical energy → mechanical energy". During the operation of PRMD, a vapor pressure difference is formed on both sides of the hydrophobic membrane between the high-temperature ammonia waste liquid (92~95℃) and the high-pressure salt solution (20~24 bar). Water molecules migrate directly from the shell side to the tube side in the form of vapor, directly converting heat energy into pressure energy of the cold side fluid. This avoids the problems of low efficiency (<2%), complex equipment, and high investment in the utilization of low-temperature waste heat in the traditional organic Rankine cycle (ORC) system.

[0015] (2) Comprehensive anti-corrosion design: The entire equipment has no metal parts, avoiding galvanic corrosion and ion precipitation. The chemical stability of the PTFE membrane (resistance to aqua regia, strong alkalis, and organic solvents) and the density of the fluorinated layer together form a "double insurance" protection system, enabling the system to operate stably without the need for chemical cleaning. Compared with ORC or traditional heat exchanger solutions that require frequent shutdowns for maintenance, this invention significantly reduces the frequency of maintenance and the cost of spare parts replacement, while ensuring the long-term consistency of freshwater recovery rate (>42%) and energy self-sufficiency rate (>15%).

[0016] (3) Significantly improved heat transfer efficiency: On the one hand, the high-temperature ammonia-evaporized waste liquid vaporizes on the surface of the PTFE hydrophobic membrane. The water vapor selectively permeates through the micropores and condenses on the cold side under the drive of the vapor pressure difference. This process achieves efficient transmembrane heat transfer in the form of latent heat, and its heat transfer intensity is much higher than that of single-phase flow. On the other hand, the microporous aeration pipe at the bottom of the membrane module continuously releases 50~200μm microbubbles. During the rising process, the boundary layer is disturbed, local turbulence is enhanced, and inorganic scale deposition is inhibited, which effectively reduces the thermal resistance of the shell-side liquid phase heat transfer and makes the membrane surface temperature closer to the temperature of the main waste liquid, thereby maintaining a high effective heat transfer temperature difference. The coupling of the two not only fully utilizes the inherent high heat transfer potential of pressure delayed membrane distillation, but also overcomes the problem of heat transfer performance decay caused by concentration polarization and scaling in high-salt and strong-alkali ammonia-evaporized waste liquid. Finally, it maintains stable and efficient heat and mass transfer performance during 8,400 hours of continuous operation.

[0017] (4) Significantly reduced flux decay rate: On the one hand, plasma fluorination treatment forms high-density –CF3 / –CF2 low-energy groups on the surface of the PTFE curtain membrane, increasing the water contact angle to ≥135° and the roll-off angle to ≤10°, which not only greatly enhances hydrophobic stability but also significantly reduces pollutants (such as Ca). 2+ Mg 2+The adhesion energy of colloids to the membrane surface inhibits the formation of irreversible fouling layers at the source. Secondly, the introduction of 50-200μm microbubbles for aeration in the submerged structure generates continuous microturbulence and shear disturbance on the membrane surface during bubble rise, effectively flushing the boundary layer, preventing inorganic salt crystallization and deposition, and disrupting the concentration polarization layer, thereby inhibiting the development of reversible fouling. More importantly, the high contact angle and high liquid wetting pressure (LEP>25bar) together construct a dual barrier against both "pressure wetting" and "non-pressure wetting," avoiding a sudden drop in flux due to localized wetting. Therefore, under the harsh conditions of high temperature, high alkali, and high salinity of ammonia wastewater, the system's flux decreases by only 3.1%~6.1% during 8,400 hours of operation, while Comparative Example 3, lacking the above synergistic protection, experienced a flux decrease of over 57% within 720 hours due to severe scaling and wetting failure.

[0018] (5) Breakthrough in thermal-mechanical energy conversion efficiency: In PRMD, water molecules permeate through the hydrophobic PTFE membrane in the form of vapor. Its flux depends only on the vapor pressure difference at the tube-side and shell-side interface (determined by the temperature difference) and is not affected by the cold-side hydraulic pressure. As long as the operating pressure is lower than the liquid wetting pressure (LEP) of the membrane, the membrane pores remain in the gas phase, and the mass transfer channels are unobstructed. Therefore, under given high-temperature waste liquid (92~95℃) conditions, the vapor pressure difference is basically fixed, and the water flux tends to be constant. At this time, the highest hydraulic pressure that the cold side can withstand is the maximum recoverable mechanical energy. This invention uses a plasma-fluorinated PTFE membrane with high LEP (>25 bar) and controls the cold-side operating pressure at 20~24 bar. Without sacrificing the water flux, the permeate-side freshwater is stored in the high-pressure chamber, and high-grade hydraulic energy is directly output to drive the subsequent NF / RO unit. In contrast, traditional ORC systems require multiple stages of conversion: "heat → working fluid evaporation → expansion and work → power generation → pump pressurization," each step accompanied by significant irreversible losses. In contrast, this invention achieves a direct path of heat energy → steam migration → high-pressure fresh water through PRMD, avoiding the Carnot limitations and electrical energy conversion losses of the heat engine cycle.

[0019] (6) Significantly improved system energy self-sufficiency: The PRMD unit utilizes the waste heat from ammonia stripping to drive water vapor transmembrane permeation and accumulates a high-pressure brine solution of 20-24 bar on the cold side. This high-pressure fluid can be directly used as the feed source for the NF and RO units without the need for pump pressurization, replacing the electrical energy consumed by traditional high-pressure pumps. Since the PRMD permeate flux is determined by the temperature difference and the output pressure is determined by the upper limit of LEP, the two are decoupled in mechanism, allowing the system to maximize energy recovery grade while maintaining high flux. Under the baseline operating conditions (Example 1), this self-powered mode achieves a system energy self-sufficiency rate of 15.8%; after further optimization of membrane performance and operating parameters (Example 3), the self-sufficiency rate is increased to 16.5%.

[0020] (7) Significantly reduced overall energy consumption per ton of water: On the one hand, PRMD utilizes the waste heat (92~95℃) of ammonia stripping waste liquid to generate a 20~24 bar high-pressure, high-salt solution, which is directly used as NF / RO feed, saving some of the power consumption of the high-pressure pump; on the other hand, the system is equipped with two-stage high-efficiency pressure exchange devices on the NF and RO concentrate sides to recover more than 90% of the residual pressure for pre-pressurizing the new feed liquid. The combination of these two factors reduces the energy consumption per ton of water under the baseline condition (Example 1) to 1.19 kWh / m³. 3 The optimized operating condition (Example 3) reached 1.16 kWh / m³. 3 Compared to the traditional NF-RO process (1.40–2.50 kWh / m³), 3 It reduces energy consumption by 17-54%, achieving a fundamental shift in the membrane separation process from "high power consumption" to "waste heat driven + energy closed loop" without sacrificing recovery rate and water quality.

[0021] (8) High-efficiency freshwater recovery: This invention achieves both thermal energy to mechanical energy conversion and simultaneous high-efficiency recovery of freshwater resources from ammonia stripping wastewater. The high recovery rate stems from the intrinsic separation characteristics of the PRMD process and the synergistic effect of system integration optimization. Firstly, the PRMD uses a hydrophobic PTFE membrane as the mass transfer interface. Under high-temperature driving, water molecules in the wastewater selectively vaporize and permeate through the membrane pores, while high-concentration salts and Ca2+ are removed. 2+ Mg 2+ Organic pollutants are completely retained because they cannot enter the gas phase, achieving a near 100% desalination rate. Secondly, since the PRMD flux is determined by the temperature difference and is not suppressed by cold-side pressure (as long as the operating pressure is lower than LEP), the system can directly output permeate at a pressure of 20-24 bar while maintaining high flux, avoiding the water recovery efficiency loss caused by condensation reflux or low flux in traditional multi-effect evaporation or conventional membrane distillation. In addition, through microbubble dynamic flushing and the anti-fouling design of the fluorinated PTFE membrane, the risk of fouling and wetting on the membrane surface is effectively suppressed, ensuring flux stability during 8,400 hours of continuous operation (attenuation rate of only 3.1%-6.1%), thereby maintaining high and stable freshwater production. Therefore, the freshwater recovery rate reaches 42.7% under the baseline condition (Example 1), and is further improved to 43.8% under the optimized condition (Example 3).

[0022] (9) System integration benefits: This invention successfully constructed a high-efficiency coupling system of PRMD-NF-RO multi-process, realizing the synergistic optimization of the three functions of "waste heat recovery + mechanical energy supply + freshwater recovery", increasing the thermal energy-mechanical energy conversion rate to 4.20~4.40%, the system energy self-sufficiency rate to 15.5~16.5%, and the comprehensive energy consumption per ton of water to 1.16~1.20 kWh / m³. 3The freshwater recovery rate of ammonia stripping waste liquid was increased to 42.0-43.8% (i.e., the treatment load of ammonia stripping concentrate was reduced), and the annual flux decline rate was controlled at 3.1-6.1%. Attached Figure Description

[0023] Figure 1 This is a process flow diagram of a specific embodiment of the salt separation method for ammonia waste liquid driven by thermal energy according to the present invention.

[0024] In the diagram: 1. Sedimentation tank; 2. Submerged pressure delayed membrane distillation unit; 3. Booster pump No. 1; 4. Nanofiltration membrane unit; 5. PX pressure exchange unit No. 1; 6. Reverse osmosis unit; 7. PX pressure exchange unit No. 2; 8. Booster pump No. 2; 2-1 Shell side; 2-2 Tube side; 2-3 Aeration pipe. Detailed Implementation

[0025] The embodiments of the present invention will now be described with reference to the accompanying drawings.

[0026] Figure 1 A process flow diagram of a specific embodiment of the salt separation method for ammonia stripping waste liquid driven by thermal energy is shown below. Figure 1 As shown, the ammonia-containing waste liquid is sent to sedimentation tank 1, and a polyethyleneimine derivative surfactant is added for clarification and sedimentation to obtain a clear ammonia-containing liquid with a temperature of 92-95℃, a surface tension >65mN / m, and an SS value <500mg / L. This clear liquid then enters the shell side 2-1 of the submerged pressure delayed membrane distillation unit 2, while the salt solution is pressurized to 20-24 bar by the second PX pressure exchange device 7 and then enters the tube side 2-2 of the submerged pressure delayed membrane distillation unit 2. 50-200μm microbubbles (intensity 0.5-1.5Nm) are provided in the aeration pipe 2-3. 3 / (m 2 Under the scouring of the shell side, water vapor migrates to the tube side. When the temperature of the ammonia stripping concentrate drops to 70-85℃, it is discharged from the shell side and sent to the ash or slag yard slurry conditioning system. The dilute salt solution flowing out of the tube side is pressurized to 30-35 bar by the No. 1 booster pump 3 and then sent to the nanofiltration unit 4 for salt separation, obtaining monovalent salt permeate and divalent salt concentrate. After the monovalent salt permeate is energy recovered by the No. 1 PX pressure exchange unit 5, it is pressurized to 60-70 bar by the No. 2 booster pump 8 and enters the reverse osmosis unit 6 to produce fresh water and concentrated brine. The concentrated brine is returned to the No. 2 PX pressure exchange unit 7 to complete the energy cycle and then enters the soda ash salt system. After the divalent salt concentrate passes through the No. 1 PX pressure exchange unit 5, it is mixed with the ammonia stripping concentrate, white mud, etc., and then pressure filtered and landfilled.

[0027] This system utilizes the Biyuan RF-30 curtain-type MBR membrane module as the foundation for engineering modifications. This model was originally designed to replace imported PVDF membrane modules such as the Mitsubishi 60E0025SA, featuring standardized dimensions (2000×1250×30mm), compatible interfaces, and a mature integrated water collection structure. Based on this, the original PVDF hollow fiber membrane fibers were replaced with high-temperature resistant, strong alkali-resistant, and wetting-resistant PTFE hollow fiber membranes. Simultaneously, the original metal or ABS support frame and flow channel system were comprehensively upgraded to a high-polymer pressure-resistant and corrosion-resistant structure: the membrane frame uses glass fiber reinforced PPH (polypropylene homopolymer) composite material, and the water collection pipes and high-pressure connectors use PVDF socket fittings (nominal pressure PN25, temperature resistance up to 100℃), completely eliminating metal contact; the entire system has an effective membrane area of ​​1200m². 2 Processing capacity 800m 3 / d, shell-side operating temperature 70~95℃, pH 5~12; equipped with PVDF microporous aeration pipe, it can stably generate 50~200μm microbubbles. Through the synergistic design of the all-polymer pressure-resistant and corrosion-resistant structure and PTFE membrane, it achieves zero-distance heat exchange, low attenuation and long-term safe and stable operation under extreme conditions of ammonia stripping waste liquid.

[0028] The PTFE base membrane uses commercially available PTFE membranes as the substrate, such as GORE® Membrane Type GH (WLGore & Associates), Donaldson Synteq™ XPPTFE membrane (Donaldson), MF-Millipore™ PTFE membrane (Merck KGaA), or Pall PTFE Ultipor® VF membrane (Pall Corporation). Its basic parameters are an average pore size of 0.1~0.35μm, a thickness of 90~180μm, a porosity of 75-85%, and an initial contact angle of 120~125°. Before use, it needs to be plasma fluorinated. After treatment, the performance reaches a liquid wetting pressure >25bar, a contact angle ≥135°, a roll-off angle ≤10°, and a membrane surface energy ≤15mN / m. It is suitable for high-temperature and high-alkalinity ammonia-eating waste liquid environments at 92-95℃.

[0029] The surfactant used is a polyethyleneimine derivative, with a variety of commercially available options including ClearAmine™-PEI10 (manufactured by Merck, Germany), BASF's Lupasol® FG or Polymin® P, Dow Chemical's Polycat® SA-1, Ecolab's Nalco™ 7360, Kurita Water Industries' Clear-Trace™ PE-15, and SNF Floerger's FLoCat™ PEI-100. These are all high cationic charge density polymers with molecular weights ranging from 25,000 to 750,000 Daltons. The dosage is 5-15 mg / L. Under the high-temperature, alkaline conditions of the ammonia-containing wastewater, it forms a strong electrostatic bond with the anionic surfactant, increasing the surface tension of the wastewater from 50-55 mN / m to >65 mN / m and reducing the SS value to <500 mg / L. This effectively prevents the "non-pressure wetting" phenomenon of the membrane module, ensuring long-term stable operation of the system.

[0030] The sedimentation tank, booster pump, PX pressure exchange device, nanofiltration device, and reverse osmosis device can all be manufactured or purchased using existing technologies.

[0031] All reagents and materials used in the following examples and comparative examples were commercially available. Low-salinity wastewater and seawater brine were sourced locally. The thermal-mechanical efficiency, system energy self-sufficiency rate, energy consumption per ton of water, and freshwater recovery rate in the experimental methods are referenced in the following papers: (Temperature-enhanced Pressure Retarded Osmosis Powered by Solar Energy: Experimental Validation, Economic Consideration, and Potential Implication, DOI:10.1016 / j.cherd.2021.04.024; Theoretical Analysis of Pressure Retarded Membrane Distillation (PRMD) Process for Simultaneous Production of Water and Electricity, DOI:10.1021 / acs.iecr.7b03642). Example 1

[0032] (Baseline operating condition) Membrane fluorination treatment: The GORE® Membrane Type GHPTFE base membrane (supplied by WLGore & Associates, average pore size 0.25 μm, thickness 150 μm) was placed in a vacuum plasma reaction chamber and evacuated to a pressure of 8 × 10⁻⁶. -5 bar; CF input 4 Gas, maintaining the pressure inside the reaction chamber at 50 × 10 -5 bar; apply RF power of 120W, process for 180 seconds, and obtain plasma fluorinated PTFE curtain membrane with liquid wetting pressure of 27 bar, contact angle of 142°, roll-off angle of 8°, and membrane surface energy of 14.3 mN / m.

[0033] The ammonia-eradication waste liquid (temperature 92℃, TDS 110,000 mg / L, pH 10.3) was fed into a sedimentation tank, and ClearAmine™-PEI10 polyethyleneimine derivative surfactant (manufactured by Merck, Germany, dosage 8 mg / L) was added. After clarification and sedimentation, a clear ammonia-eradication liquid was obtained. This clear liquid was then fed into the shell side of a submerged pressure delayed membrane distillation unit. A brine solution was passed through (pretreated nearshore brine, TDS ≈ 35,000 mg / L, containing Na⁺ ≈ 10,500 mg / L, Ca…) 2+ ≈ 410 mg / L, Mg 2+ ≈ 1,290 mg / L, SO 42 (≈ 2,710 mg / L) is pressurized to 22 bar by a second pressure exchange device and then enters the tube side of the submerged pressure delayed membrane distillation unit; 120 μm microbubbles (intensity 1.0 Nm) are provided in the aeration tube. 3 / (m 2Under the flushing action of the tube side, water vapor migrates from the shell side to the tube side. When the temperature of the ammonia concentrate drops to 70-85°C, it is discharged from the shell side. Because 42.7% of the water has migrated to the tube side, the waste liquid volume is reduced by 42.7%, and the TDS is concentrated from 110,000 mg / L to approximately 191,000 mg / L (≈19.1wt%), significantly reducing the treatment load of the subsequent solid waste disposal system. The dilute salt solution flowing out of the tube side is pressurized to 32 bar by the first booster pump and then sent to the nanofiltration membrane device for salt separation, obtaining monovalent salt permeate and divalent salt concentrate, of which the divalent salt concentrate has a TDS ≥ 20,000 mg / L. The sodium chloride solution, with a concentration of mg / L and a volume approximately 15% of the original salt solution, is filtered and then mixed with white mud for landfill. The monovalent salt permeate, after energy recovery via the first pressure exchange unit, is pressurized to 65 bar by the second booster pump and enters the reverse osmosis unit, producing fresh water and concentrated brine. The concentrated brine is returned to the second pressure exchange unit to complete the energy cycle, increasing the TDS to approximately 140,000 mg / L (NaCl purity >98%), and then returned to the soda ash plant's salt processing system to replace some of the solid salt. The system operates continuously for 8,400 hours, with a thermal-mechanical energy conversion efficiency of 4.25% for the pressure delayed membrane distillation unit; the system's energy self-sufficiency rate is 15.8%; and the comprehensive energy consumption per ton of water is 1.19 kWh / m³. 3 The ammonia stripping wastewater freshwater recovery rate was 42.7%; the flux decline rate of the pressure delayed membrane distillation unit was 3.9%. Example 2

[0034] (Upper limit of temperature for ammonia stripping waste liquid) Membrane fluorination treatment: The Donaldson Synteq™ XPPTFE membrane (supplied by Donaldson, average pore size 0.22 μm, thickness 160 μm) was placed in a vacuum plasma reaction chamber and evacuated to a pressure of 7 × 10⁻⁶. -5 bar; C2F6 gas is introduced to maintain the pressure inside the reaction chamber at 60 × 10⁻⁶. -5 bar; applying RF power of 150W and processing time of 210 seconds, a plasma-fluorinated PTFE curtain membrane is obtained, with a liquid wetting pressure of 28 bar, a contact angle of 145°, a roll-off angle of 7°, a membrane surface energy of 13.8 mN / m, and enhanced heat resistance, allowing for long-term stable operation at 95℃.

[0035] The ammonia-eating waste liquid (temperature 95℃, TDS 110,000 mg / L, pH 10.3) was sent to a sedimentation tank, and Lupasol® FG polyethyleneimine derivative surfactant (manufactured by BASF, dosage 10 mg / L) was added. After clarification and sedimentation, a clear ammonia-eating liquid was obtained. The remaining steps were the same as in Example 1. The system ran continuously for 8,400 hours. When the temperature of the concentrated ammonia liquid dropped to 70-85℃, it was discharged from the shell side with a volume reduction of 43.5% and TDS ≈ 195,000 mg / L, and entered the subsequent waste liquid treatment section. The tube-side outlet water temperature was 43℃. The thermal-mechanical energy conversion efficiency of the pressure delayed membrane distillation unit was 4.38%. The system energy self-sufficiency rate was 16.3%. The comprehensive energy consumption per ton of water was 1.17 kWh / m³. 3 The ammonia stripping wastewater freshwater recovery rate was 43.5%; the flux decline rate of the pressure delayed membrane distillation unit was 4.2%. Example 3

[0036] (Upper limit of operating pressure in the tube side of the pressure delayed membrane distillation unit) Membrane fluorination treatment: The MF-Millipore™ PTFE membrane (supplied by Merck KGaA, average pore size 0.20 μm, thickness 170 μm) was placed in a vacuum plasma reaction chamber and evacuated to a pressure of 6 × 10⁻⁶. -5 bar; CF input 4 A 1:1 volume ratio of SF6 mixed with gas was used to maintain the pressure inside the reaction chamber at 65 × 10⁻⁶. -5 bar; apply RF power of 180W, process for 240 seconds, and obtain plasma fluorinated PTFE curtain membrane with liquid wetting pressure of 28 bar, contact angle of 147°, roll-off angle of 6°, and membrane surface energy of 13.5 mN / m.

[0037] The ammonia-removed waste liquid (temperature 92℃, TDS 110,000 mg / L, pH 10.3) was sent to a sedimentation tank, and Polymin® P polyethyleneimine derivative surfactant (manufactured by BASF, dosage 9 mg / L) was added. After clarification and sedimentation, a clear ammonia-removed liquid was obtained. The remaining steps were the same as in Example 1, except that the tube-side operating pressure was set to 24 bar. During operation, the tube-side "pressure flow rate" increased by 60% compared to Example 1. The system operated continuously for 8,400 hours without membrane wetting. The thermal-mechanical energy conversion efficiency of the pressure delayed membrane distillation unit was 4.40%. The system energy self-sufficiency rate was 16.5%. The comprehensive energy consumption per ton of water was 1.16 kWh / m³. 3 The ammonia stripping wastewater freshwater recovery rate was 43.8%; the flux decline rate of the pressure delayed membrane distillation unit was 4.7%. Example 4

[0038] (Lower limit of aeration intensity) Membrane fluorination treatment: The PallPTFEUltipor® VF membrane (supplied by Pall Corporation, average pore size 0.26 μm, thickness 145 μm) was placed in a vacuum plasma reaction chamber and evacuated to a pressure of 9 × 10⁻⁶. -5 bar; CF input 4 Gas, maintaining the pressure inside the reaction chamber at 45 × 10 -5 bar; applying 100W of radio frequency power and processing time of 150 seconds, a plasma-fluorinated PTFE curtain membrane was obtained, with a liquid wetting pressure of 26.5 bar, a contact angle of 140°, a roll-off angle of 9°, and a membrane surface energy of 14.8 mN / m, exhibiting excellent antifouling properties.

[0039] The ammonia-containing waste liquid (temperature 92℃, TDS 110,000 mg / L, pH 10.3) was fed into a sedimentation tank, and Polycat® SA-1 polyethyleneimine derivative surfactant (manufactured by Dow Chemical, dosage 8 mg / L) was added. After clarification and sedimentation, a clear ammonia-containing liquid was obtained. The remaining steps were the same as in Example 1, except that the aeration intensity was reduced to 0.5 Nm. 3 / (m 2 •h) (the lower limit specified in claim 8), the microbubble size is maintained at 120 μm; the system operates continuously for 8,400 hours, and the thermal-mechanical energy conversion efficiency of the pressure delayed membrane distillation unit is 4.20%; the system energy self-sufficiency rate is 15.5%; the comprehensive energy consumption per ton of water is 1.20 kWh / m³. 3 The ammonia stripping wastewater freshwater recovery rate was 42.0%; the flux decline rate of the pressure delayed membrane distillation unit was 4.8%. Example 5

[0040] (Lower limit boundary of plasma fluorinated membrane liquid wetting pressure performance) Membrane fluorination treatment: The GORE® Membrane Type GHPTFE base membrane (supplied by WLGore & Associates, average pore size 0.30 μm, thickness 140 μm) was placed in a vacuum plasma reaction chamber and evacuated to a pressure of 10 × 10⁻⁶. -5 bar; CF input 4 Gas, maintaining the pressure inside the reaction chamber at 40 × 10 -5 bar; apply RF power of 60W, process for 90 seconds, and obtain plasma fluorinated PTFE curtain membrane with liquid wetting pressure of 25.2 bar, contact angle of 136°, roll-off angle of 10°, and membrane surface energy of 15.0 mN / m.

[0041] The ammonia-removed waste liquid (temperature 92℃, TDS 110,000 mg / L, pH 10.3) was sent to a sedimentation tank, and Nalco™ 7360 polyethyleneimine derivative surfactant (produced by Ecolab, dosage 11 mg / L) was added. After clarification and sedimentation, a clear ammonia-removed liquid was obtained. The remaining steps were the same as in Example 1, except that the tube-side pressure was set to 24 bar. The system was basically stable for the first 6,000 hours of operation, but micro-wetting of local membrane elements occurred later. The system automatically reduced the pressure to 23.5 bar and then returned to stability. The overall flux decline rate of the pressure-delayed membrane distillation unit was 6.1% after 8,400 hours; the thermal-mechanical energy conversion efficiency of the pressure-delayed membrane distillation unit was 4.22%; the system energy self-sufficiency rate was 15.6%; and the comprehensive energy consumption per ton of water was 1.19 kWh / m³. 3 The freshwater recovery rate of ammonia stripping waste liquid was 42.5%. Example 6

[0042] (Preprocessing to strengthen boundaries) Membrane fluorination treatment: The Donaldson Synteq™ XPPTFE membrane (supplied by Donaldson, average pore size 0.24 μm, thickness 155 μm) was placed in a vacuum plasma reaction chamber and evacuated to a pressure of 7 × 10⁻⁶. -5 bar; C2F6 gas is introduced to maintain the pressure inside the reaction chamber at 55 × 10⁻⁶. -5 bar; apply RF power of 130W, process for 200 seconds, and obtain plasma fluorinated PTFE curtain membrane with liquid wetting pressure of 27.5 bar, contact angle of 143°, roll-off angle of 7.5°, and membrane surface energy of 14.0 mN / m.

[0043] The ammonia-removed waste liquid (temperature 92℃, TDS 110,000 mg / L, pH 10.3) was fed into a sedimentation tank, and Clear-Trace™ PE-15 polyethyleneimine derivative surfactant (manufactured by Kurita Water Industries, dosage 12 mg / L) was added. After clarification and sedimentation, a clear ammonia-removed liquid was obtained. The remaining steps were the same as in Example 1. The system operated continuously for 8,400 hours, and the thermal-mechanical energy conversion efficiency of the pressure delayed membrane distillation unit was 4.28%. The system energy self-sufficiency rate was 16.0%, and the comprehensive energy consumption per ton of water was 1.18 kWh / m³. 3 The ammonia stripping wastewater freshwater recovery rate was 42.9%; the flux decline rate of the pressure delayed membrane distillation unit was only 3.1%. Comparative Example 1

[0044] (Conventional Organic Rankine Cycle (ORC) Power Generation Driven Nanofiltration Membrane Unit - Reverse Osmosis System) Membrane fluorination: Not applicable (this system does not use membrane distillation technology). Ammonia-evaporized wastewater (92℃, TDS 110,000 mg / L, pH 10.3) under the same conditions was directly introduced into the ORC evaporator. Due to the high alkalinity, high salt content, and high calcium content of the wastewater, a dense CaCO3 / Mg(OH)2 scale layer formed on the evaporator tube wall within 72 hours, reducing the heat transfer coefficient by 45%. Simultaneously, the alkaline environment caused stress corrosion cracking in the 316L stainless steel evaporator, leading to leakage after 15 days of operation. Even with a pre-filter, micron-sized tar colloids still penetrated and adhered to the heat exchange surface on the working fluid side, causing a continuous decline in ORC efficiency. Ultimately, the actual thermoelectric efficiency was only 1.2% (theoretical 1.75%), and the system's energy self-sufficiency rate was only 3.8%. The energy consumption per ton of water reached as high as 1.52 kWh / m³. 3 Equipment maintenance costs increased by 2.3 times; the evaporator was replaced 3 times in 8,400 hours, and the ORC downtime totaled 920 hours. Comparative Example 2

[0045] (316L stainless steel shell and tube heat exchanger for preheating salt solution) Membrane fluorination: Not applicable (this system does not use membrane technology). Ammonia-removed waste liquid at 92℃ undergoes countercurrent heat exchange with a brine solution in a tube-and-shell heat exchanger. The brine solution is preheated from 22℃ to 45℃ before being fed into a fully electrically driven nanofiltration membrane unit-reverse osmosis system. The initial heat transfer coefficient is 850 W / (m²). 2 (·K), but after 320 hours of operation, the heat transfer coefficient plummeted to 358W / (m²). 2 • K), scaling significantly increases thermal resistance; after 500 hours, the heat exchange tubes undergo alkaline stress corrosion cracking due to the synergistic effect of high pH, ​​high Cl⁻, and high temperature, resulting in localized perforation and system shutdown; short-term energy consumption per ton of water is 1.42 kWh / m 3 However, it cannot operate for a long time. Comparative Example 3

[0046] (Conventional pressure delayed membrane distillation unit components) Membrane fluorination treatment: A commercial PTFE membrane (average pore size 0.35 μm, thickness 120 μm) was placed in a vacuum plasma reaction chamber and evacuated to a pressure of 15 × 10⁻⁶. -5 bar (beyond the scope of the claims); CF access 4 Gas, maintaining the pressure inside the reaction chamber at 30 × 10 -5 bar; apply radio frequency power of 40W (below the scope of the claims) and process for 45 seconds (below the scope of the claims) to obtain an insufficiently fluorinated PTFE membrane with a liquid wetting pressure of 20.5 bar, a contact angle of 130°, a roll-off angle of 15°, and a membrane surface energy of 17.5 mN / m.

[0047] The ammonia-removed waste liquid (temperature 92℃, TDS 110,000 mg / L, pH 10.3) was sent to a sedimentation tank, and FLoCat™ PEI-100 polyethyleneimine derivative surfactant (manufactured by SNF Floerger, dosage 8 mg / L) was added. After clarification and sedimentation, a clear ammonia-removed liquid was obtained. The pressure delayed membrane distillation unit was replaced with a conventional hollow fiber or tubular membrane module structure. The clear ammonia-removed liquid flowed through the shell side via an external circulation pump, and the salt solution entered the tube side. The initial flux was 28 L / (m²). 2 The flux initially increased by 1.5 L / (m³), but after 480 hours it dropped to 15 L / (m³). 2 The decay rate reached 46%; the CaCO3 deposition thickness on the film surface exceeded 120μm after 720 hours, and local wetting caused the actual value of liquid wetting pressure to drop to 20 bar; the thermal-mechanical efficiency was only 2.0%. Comparative Example 4

[0048] (Pressure over-limit operation of the tube side of the pressure delayed membrane distillation unit) Membrane fluorination treatment: The GORE® Membrane Type GHPTFE base membrane (supplied by WLGore & Associates, average pore size 0.25 μm, thickness 150 μm) was placed in a vacuum plasma reaction chamber and evacuated to a pressure of 8 × 10⁻⁶. -5 bar; SF6 gas is introduced to maintain the pressure inside the reaction chamber at 60 × 10⁻⁶. -5 bar; apply RF power of 140W, process for 200 seconds, and obtain plasma fluorinated PTFE curtain membrane with liquid wetting pressure of 28 bar, contact angle of 144°, roll-off angle of 7.5°, and membrane surface energy of 14.0 mN / m.

[0049] The ammonia-removed waste liquid (temperature 92℃, TDS 110,000 mg / L, pH 10.3) was sent to a sedimentation tank, and ClearAmine™-PEI10 polyethyleneimine derivative surfactant (manufactured by Merck, Germany, dosage 8 mg / L) was added. After clarification and sedimentation, a clear ammonia-removed liquid was obtained. The remaining steps were the same as in Example 1, but the tube-side operating pressure was set to 28.5 bar, and the membrane liquid wetting pressure was 28 bar. After 90 minutes of operation, the tube-side effluent flow rate decreased. After 2 hours, the pressure-delayed membrane distillation unit completely lost its function and could not achieve heat and freshwater recovery. The subsequent nanofiltration membrane unit shut down within 24 hours due to insufficient influent flow rate. Comparative Example 5

[0050] (Pretreatment with polyethyleneimine derivative surfactant scavenger omitted) Membrane fluorination treatment: The same membrane preparation method as in Example 1 (GORE® Membrane Type GHPTFE base membrane) was used.

[0051] The ammonia-containing waste liquid underwent only sedimentation without the addition of polyethyleneimine derivative surfactants for removal, resulting in a surface tension of 50 mN / m (below the required 65 mN / m). During the initial operation (0-200 h), the system performed normally, but after 380 hours, the TDS of the tube-side effluent began to fluctuate drastically. After 600 hours, the TDS consistently exceeded 100,000 mg / L, causing the pressure delayed membrane distillation unit to fail. Analysis showed that residual ammonia and quaternary ammonium salt surfactants in the waste liquid significantly reduced the surface tension of the liquid, causing the actual liquid wetting pressure to drop from 27 bar to approximately 21 bar. Even though the tube-side pressure of 22 bar did not exceed the nominal liquid wetting pressure, "non-pressure wetting" still occurred. Comparative Example 6

[0052] (A standard PTFE pressure delayed membrane distillation apparatus membrane without plasma fluorination treatment) Fluorination treatment of membrane: No plasma fluorination treatment is performed. The GORE® Membrane Type GHPTFE base membrane (supplied by WLGore & Associates, with an average pore size of 0.25 μm, a thickness of 150 μm, a liquid wetting pressure of 22.5 bar, a contact angle of 128°, a roll-off angle of 18°, an F / C atomic ratio of 1.9, and a surface energy of 18 mN / m) is used directly.

[0053] The ammonia-evaporized wastewater (temperature 92℃, TDS 110,000 mg / L, pH 10.3) was sent to a sedimentation tank, and ClearAmine™-PEI10 polyethyleneimine derivative surfactant (manufactured by Merck, Germany, dosage 8 mg / L) was added. After clarification and sedimentation, a clear ammonia-evaporized liquid was obtained. The remaining steps were exactly the same as in Example 1, except that the membrane of the pressure delayed membrane distillation device was replaced with the above-mentioned unfluorinated commercial PTFE membrane. The system was initially stable, but after 1,200 hours of operation, local wetting spots appeared on the membrane surface. After 3,500 hours, the wetting area expanded, and the TDS of the tube-side effluent rose to 50,000 mg / L. The cumulative flux decline rate reached 66.4% after 8,400 hours, and the thermal-mechanical efficiency dropped to 1.65%. The system required membrane cleaning and physical drying after 3,500 hours, but only 62% of the initial performance was restored, indicating that the wetting was partially irreversible.

[0054] The key performance parameters of the embodiments and comparative examples are summarized in Table 1.

[0055]

[0056] Comparative Example 1 (ORC power generation driven NF-RO) reveals the fundamental shortcomings of traditional heat-to-work conversion technology in highly polluted wastewater. The ammonia-containing wastewater contains a high concentration of Ca. 2+ Mg 2+Ammonia nitrogen and tar-like organic matter cause rapid scaling in the ORC evaporator (a 45% decrease in heat transfer coefficient within 72 hours) and alkaline stress corrosion cracking (leaking within 15 days). The actual thermoelectric efficiency is only 1.2%, far below the theoretical value, and the system reliability is extremely poor (over 900 hours of downtime within 8,400 hours). In contrast, the PRMD of this invention uses a hydrophobic film as the mass transfer interface to physically isolate contaminants, eliminating the need for metal heat exchange surfaces and fundamentally avoiding the risks of corrosion and scaling.

[0057] Comparative Example 2 (316L stainless steel tube heat exchanger) only achieves heat transfer and cannot generate driving pressure, thus having a single function. More seriously, under the synergistic effect of "high temperature-high pH-high Cl⁻", the heat exchanger perforates and fails after 500 hours, with the heat transfer efficiency plummeting by 60%. In contrast, the PRMD unit of this invention uses a corrosion-resistant PTFE membrane + immersion structure + microbubble flushing to achieve >8,000 hours of trouble-free operation, while simultaneously achieving the dual goals of heat energy to mechanical energy conversion and freshwater recovery.

[0058] Comparative Example 3 (conventional non-immersion PRMD) suffered from indirect contact between the membrane and waste liquid, dead zones in the flow channel, and a lack of dynamic anti-scaling measures, resulting in rapid CaCO3 deposition (scale thickness >120 μm after 720 h), flux attenuation exceeding 57%, and a significant decrease in LEP leading to wetting. This invention, through immersion-based zero-distance heat transfer and uniform microbubble aeration (50–200 μm), forms interfacial microturbulence, improving heat transfer efficiency by more than 2 times and controlling flux attenuation to within 5%.

[0059] Comparative Example 4 (operating pressure exceeding LEP) verified the thermodynamic boundary conditions for PRMD operation. When the operating pressure (28.5 bar) exceeded the membrane liquid wetting pressure (28 bar), hydraulic pressure overcame capillary resistance, irreversible wetting occurred within 2 hours, and the system completely failed. This invention strictly controls the operating pressure below LEP, maximizing energy output while ensuring a safety margin.

[0060] Comparative Example 5 (surfactant scavenger omitted) demonstrates from an interfacial thermodynamic perspective that even if the operating pressure does not exceed the nominal LEP value, if the surface tension of the waste liquid drops to 50 mN / m (<65 mN / m requirement) due to residual surfactant, the actual LEP value will decrease significantly, leading to "non-pressure wetting." This invention, by adding a polyethyleneimine derivative scavenger, increases the surface tension to >65 mN / m, which is a key prerequisite for the long-term stable operation of the system.

[0061] Comparative Example 6 (unfluorinated PTFE membrane) shows that ordinary PTFE membranes have insufficient hydrophobicity (contact angle 128°, roll-off angle 18°), and under harsh operating conditions, the flux attenuation rate reaches 66.4% after 8,400 hours, with irreversible wetting. In contrast, this invention employs plasma fluorination treatment (contact angle ≥135°, roll-off angle ≤10°, LEP >25 bar), significantly improving the membrane's anti-wetting and anti-fouling capabilities, providing a core material guarantee for achieving long-term stable operation.

[0062] The above are merely specific embodiments of the present invention, but the scope of protection of the present invention is not limited thereto. Any changes or substitutions that cannot be easily conceived by those skilled in the art within the technical scope disclosed in the present invention should be included within the scope of protection of the present invention. In the absence of conflict, the embodiments of the present invention and the features in the embodiments can be combined with each other. Therefore, the scope of protection of the present invention should be determined by the scope of the claims.

Claims

1. A method for salt separation of a salt solution driven by the thermal energy of ammonia stripping waste liquid, characterized in that, Includes the following steps: (1) Add polyethyleneimine derivative surfactant to the ammonia stripping waste liquid, clarify and settle to obtain ammonia stripping clear liquid; the temperature of the ammonia stripping clear liquid is 92~95℃, the surface tension is >65mN / m, and the SS value is <500mg / L; (2) The ammonia distillate is fed into the shell side of the submerged pressure delayed membrane distillation device, and a salt solution with a pressure of 20-24 bar flows in the tube side of the submerged pressure delayed membrane distillation device. Driven by the temperature difference and concentration difference on both sides of the pressure delayed membrane, water molecules in the ammonia distillate migrate to the salt solution in the form of vapor to obtain a dilute salt solution and ammonia distillate concentrate. When the temperature of the ammonia distillate concentrate drops to 70-85°C, the ammonia distillate concentrate in the shell side is discharged. (3) Pressurize the dilute salt solution to 30~35 bar and send it into the nanofiltration membrane device to separate the salts and obtain monovalent salt permeate and divalent salt concentrate; (4) Pressurize the monovalent salt permeate to 60~70 bar and send it into the reverse osmosis device to separate salts to obtain concentrated brine and fresh water.

2. The salt separation method for ammonia waste liquid driven by thermal energy according to claim 1, characterized in that, The pressure delay membrane is a plasma-fluorinated PTFE curtain membrane with a liquid wetting pressure >25 bar, a temperature resistance ≥90℃, a strong alkali corrosion resistance pH>10, a membrane surface energy ≤15 mN / m, a contact angle ≥135°, a roll-off angle ≤10°, an average pore size of 0.1~0.35 μm, and a membrane thickness of 90~180 μm.

3. The salt separation method for ammonia waste liquid driven by thermal energy according to claim 2, characterized in that, The plasma-fluorinated PTFE curtain membrane was prepared by the following steps: (a) Place the PTFE base film in a vacuum plasma reaction chamber and evacuate to a pressure ≤10×10 -5 bar; (b) Fluoride gas is introduced, and the pressure inside the reaction chamber is maintained at 20 × 10⁻⁶. -5 ~100×10 -5 bar; (c) Plasma treatment is applied to the PTFE base membrane, with a radio frequency power of 50~200W and a treatment time of 60~300 seconds, to obtain plasma-fluorinated PTFE curtain membrane.

4. The salt separation method for ammonia waste liquid driven by thermal energy according to claim 3, characterized in that, The PTFE base membrane is a porous hydrophobic membrane with an average pore size of 0.1~0.35μm, a thickness of 90~180μm, a porosity of 75~85%, and an initial contact angle of 120~125°.

5. The salt separation method for ammonia waste liquid driven by thermal energy according to claim 3, characterized in that, The fluorinated gas is selected from at least one of CF4, C2F6 or SF6.

6. The salt separation method for ammonia waste liquid driven by thermal energy according to claim 1, characterized in that, The concentrated brine and salt solution obtained in step (4) are sent to a pressure exchange device for pressure exchange, and the salt solution is pressurized to 20~24 bar and then sent to the tube side of an immersion pressure delayed membrane distillation device.

7. The salt separation method for ammonia waste liquid driven by thermal energy according to claim 1, characterized in that, In step (4), pressurizing the monovalent salt permeate to 60-70 bar means first sending the monovalent salt permeate and the divalent salt concentrate into a pressure exchange device to exchange pressure and increase the pressure of the monovalent salt permeate, and then using a booster pump to pressurize the monovalent salt permeate to 60-70 bar.

8. The salt separation method for ammonia waste liquid driven by thermal energy according to claim 1, characterized in that, The pressure-delayed permeation membrane has uniformly distributed aeration tubes on its surface, which provide microbubbles to flush the membrane surface. The microbubbles have a size of 50~200μm and an aeration intensity of 0.5~1.5Nm. 3 / (m 2 ·h).

9. The salt separation method for ammonia waste liquid driven by thermal energy according to claim 1, characterized in that, The polyethyleneimine derivative surfactant scavenger is a branched polyethyleneimine with a molecular weight range of 60,000 to 100,000 Daltons, a nitrogen content of 25 to 30 wt%, and a high cationic charge density.

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