EP4655081A1

EP4655081A1 - Devolatilization of polyolefin polymers

Info

Publication number: EP4655081A1
Application number: EP24702011.8A
Authority: EP
Inventors: Abolfazl NOORJAHAN; Stephen Brown; Nafiseh Dadgostar; Eric Clavelle
Original assignee: Nova Chemicals International SA
Current assignee: Nova Chemicals International SA
Priority date: 2023-01-26
Filing date: 2024-01-17
Publication date: 2025-12-03
Also published as: MX2025008082A; WO2024157121A1

Abstract

A method and a system for devolatilizing a polymer are provided. An exemplary method includes flashing a polymer solution across a first pressure control valve into a first low-pressure separation vessel to form a more concentrated polymer solution from the first low-pressure separation vessel, and flashing the more concentrated polymer solution across a second pressure control valve into a second low-pressure separation vessel to form an outlet stream from the second low-pressure separation vessel.

Description

DEVOLATILIZATION OF POLYOLEFIN POLYMERS

TECHNICAL FIELD

The present disclosure is directed to the devolatilization of polyolefin plastomers and elastomers.

BACKGROUND ART

Solution polymerization processes are carried out at temperatures that are above the melting point of the product polymer. In a typical process, catalyst components, solvent, polymerizable monomers, and hydrogen are fed under pressure to one or more stirred reactors. Catalyst components may be fed to the reactor as a solution or as a slurry and the temperature of the reactor is controlled by the rate of catalyst addition, the temperature of the catalyst feed stream, and the use of heat transfer systems. Typical polymerizable monomers for solution phase polymerization processes include ethylene, styrene, propylene, and various dienes.

For ethylene polymerization, reactor temperatures can range from about 130°C to about 250°C, while pressures are generally in the range of from about 500 psi to about 4000 psi. Although catalyst residence times are generally short, for example, minutes, solution polymerization may be operated under a wide range of process conditions to allow tailoring of the product polymer as well as rapid product transitions.

Due to the high temperatures, the product polymer is dissolved in the solvent under reactor conditions, forming a polymer solution. After the catalyst residence time is completed, the polymer solution leaves the reactor as an effluent stream and the polymerization reaction is quenched, typically with coordinating polar compounds, to prevent further polymerization. Once quenched, the polymer solution is fed to a flash devolatilization system for solvent removal. Flash devolatilization also removes unreacted monomers from the polymer solution.

In a typical devolatilization process, the polymer solution is heated in a heat exchanger and then passed into a chamber of reduced pressure. Flashing of the solvent and unreacted monomers occurs, and the vapors are sent to a solvent and monomer recovery system to be recycled back to the reactor. Heating the polymer solution upstream of the devolatilization system increases the enthalpy of the product stream, providing high temperatures to the polymer melt after devolatilization. The high temperatures facilitate flow of the polymer melt by reducing its viscosity. The heat exchangers used are most commonly shell and tube type heat exchangers and can increase the temperature of the polymer solution to as high as about 280°C.

The molten polymer from the devolatilization system is then recovered as a bottom stream from the devolatilization system. The bottom stream is a viscous stream comprised essentially of molten ethylene polymer that contains a small amount of deactivated catalyst, residual solvent, and residual comonomer. The bottom stream can be conveyed to a pelletizer, for example, by gravity, gear pumps, single screw extruders, or twin-screw extruders, among others. In some cases, sub-atmospheric pressure, vacuum extruders with vents that allow residual solvent and a-olefm comonomers to be removed are used. Once pelletized, the ethylene polymer may be dried in one or more sequential drying bins before being transported to a product silo.

In PCT application 98/02471, fded by Dow Chemicals, a solution polymerization process is described in which a two stage devolatilization system is used to remove solvent and un-reacted monomers from a polymer solution of EPDM (ethylene-propylene-diene monomer). A dual reactor system is used in which the temperature of the second reactor is between 90°C and 120°C. For flash devolatilization, the temperature of the reactor effluent is raised to between 210°C and 250°C by passage through a heat exchanger prior to entering the flash chamber, which is at a lower pressure.

U.S. Pat. No. 5,691,445 assigned to Novacor Chemicals describes a polymer solution devolatilization process in which less than 150 ppm of residual volatiles is retained in the isolated polymer. In the process, the polymer solution leaves the reactor and travels through a pre-heat exchanger. The pre-heat exchanger heats the polymer solution to temperatures from about 200°C to 270°C to increase the vapor pressure of volatiles and to reduce the polymer solution viscosity. A super-critical fluid is added to the process at a point between the first and second devolatilization chambers to enhance polymer melt foaming.

SUMMARY OF INVENTION

An embodiment described herein provides a method for devolatilizing a polymer. The method includes flashing a polymer solution across a first pressure control valve into a first low-pressure separation vessel to form a more concentrated polymer solution from the first low-pressure separation vessel, and flashing the more concentrated polymer solution across a second pressure control valve into a second low-pressure separation vessel to form an outlet stream from the second low-pressure separation vessel. Another embodiment described herein provides a system to decrease volatile compounds in polymer pellets. The system includes a first pressure control valve on a line that is fluidically coupled to an inlet of a first low-pressure separation vessel, and a second pressure control valve on a line that fluidically couples the first low-pressure separation vessel to an inlet of a second low-pressure separation vessel.

BRIEF DESCRIPTION OF THE DRAWINGS

Figure 1 is a simplified process flow diagram of a system for devolatilizing a solution phase polymer.

Figure 2 is a simplified process flow diagram of another system for devolatilizing a solution phase polymer.

Figure 3 is a simplified process flow diagram of another system for devolatilizing a solution phase polymer.

Figure 4 is a simplified process flow diagram of another system for devolatilizing a solution phase polymer.

Figure 5 is a simplified process flow diagram of another system for devolatilizing a solution phase polymer.

Figure 6 is a simplified process flow diagram of another system for devolatilizing a solution phase polymer.

Figure 7 is a simplified process flow diagram of another system for devolatilizing a solution phase polymer.

Figure 8 is a simplified process flow diagram of another system for devolatilizing a solution phase polymer.

Figure 9 is a simplified process flow diagram of another system for devolatilizing a solution phase polymer.

Figure 10 is a simplified process flow diagram of another system for devolatilizing a solution phase polymer.

Figure 11 is a simplified process flow diagram of another system for devolatilizing a solution phase polymer.

Figure 12 is a plot comparing holdup time (HUT) for the systems of Figures 3 to 11.

Figure 13 is a process flow diagram of a method for devolatilizing a plastomer.

DESCRIPTION OF EMBODIMENTS

In solution polymerization processes, the typical volatile organic compound (VOC) content of the molten polymer recovered from the devolatilization system is about 2 to about 4 wt. %. The polymer melt is then fed to a pelletizer, such as an underwater pelletizer, to form polymer pellets. The resulting pellets are then held in a stripping bin where a flow of an inert gas (e.g., air, steam, nitrogen, etc.) allows the VOCs to weather off until the VOC content reaches a specification, for example, less than about 150 ppm. The time until the pellets reach the target VOC level is measured as holdup time (HUT).

The required HUT for stripping pellets of higher density materials such as UUDPE and HPDE resins is generally not a problem as a steam flow at high temperatures, for example, greater than about 80-85°C, can be used. However, softer resins, such as polyolefin plastomers (POPs) which have a typical density in the range of from 0.886 g/cm³ to 0.912 g/cm³ and polyolefin elastomers (POEs) which have a typical density of less than 0.886 g/cm³, cannot be stripped with high temperature steam due to their soft nature (low VICAT softening temperature) and other operation complexities, such as degradation and slower diffusion rates of solvent out of the pellets, and pellets agglomeration and blockage of the stripping bin outlet nozzle. Therefore, a series of stripping bins utilizing tempered nitrogen have been designed to strip POPs and POEs. Stripping at low temperature without vacuum results in high HUTs, for example, greater than 40 hours at densities of less than 0.880 g/cc (termed an 880 density, herein).

Embodiments described herein provide a modified devolatilization system to reduce the HUT of POPs and POEs by reducing volatile organic compounds (VOCs) at the cutter of the pelletizer. Modeling is used to predict the values for VOCs at the cutter to predict the HUT for stripping of a resin with an 880 density and a 0.5 MI to achieve a final VOC level of 150 ppm in the stripping bins. This resin is selected as the limiting case for reaching the desired HUT in the stripping process.

In various embodiments, a single low-pressure devolatilization vessel is replaced with two low-pressure devolatilization vessels. The polymer solution is flashed across a valve into the first vessel, and then the more concentrated solution from the first vessel is flashed across a valve into the second vessel. The lowest HUT resulted from the injection of a devolatilization agent, for example, in a static mixer upstream of the valve that flashes into the second low-pressure devolatilization vessel. The devolatilization agent in this example was steam provided to the devolatilization system at a temperature of about 250°C.

In the present disclosure, solution polymerization reactor(s) outlet composition was modeled using heat and mass balance calculation based on the polymerization reaction operating conditions for the polymer grade in question. The modeling of the devolatilization system was performed using ASPEN Plus, which provided predictions of composition of the polymer solution for the selected resin in streams 208 and 134 based in part on heat and mass balance, the operating conditions in the phase separation area, and the phase behavior of the polymer solution for the grade in question. Ascertaining the phase behavior characteristics of the polymer solution was carried out off-line using phase equilibrium calculations with a suitable thermodynamic model, which was calibrated with phase behavior data collected using a multi-pass rheometer (MPR) on liquid-liquid phase transition conditions using polymer solutions of known composition. A non-limiting example of a suitable thermodynamic model is the PC-SAFT equation of state. Details of the multi-pass rheometer (MPR) experiments are provided in the Testing Procedures section.

For the prediction of the HUTs, a stripping bin model was developed and used to simulate and analyze the time required to reach the 150-ppm final VOC target in a commercial-scale batch stripping process in a silo-shaped bin with circular cross-sectional shape and with a capacity of 200 x 10³ kg (200 Mg) while an upflow of nitrogen was provided to the stripping bin at a constant nitrogen inlet temperature. This model was applied to the pellets of the above-described POE resin with an 880 density and a 0.5 MI using the mathematical model described in equations 1 through 8.

The stripping bin in this model comprised an upper cylindrical portion with a height of 24.0 m and a base diameter of 4.5 m. The stripping bin further comprised a truncated conical portion with a top diameter of 4.5 m, a height of 3.5 m and a bottom end wall diameter of 0.6 m. Nitrogen gas was continuously provided to an area proximal the bottom end wall using a plurality of feed nozzles as an upward flow at a mass flow rate from 2 weight % to 15 weight % per hour, based on the total weight of pellets in the stripping bin. To be specific, for example, a stripping bin that contains 200 x 10³ kg (200 Mg) of pellets can be provided with a nitrogen mass flow rate of 10⁴ kg per hour (10 Mg per hour) to provide a nitrogen flow rate of 5 weight % per hour, based on the total weight of pellets in the stripping bin.

In the present disclosure, for the sake of simplicity, for the purpose of computational model, it was assumed that pellets are spherical in shape with a single diameter value (Sauter Mean Diameter) characterizing the distribution of the pellets size and shape. The Sauter Mean Diameter (d_SM) is defined as the diameter of a sphere whose ratio of volume (Vp) to surface area (A_p) is equal to that of a pelletized particle with an arbitrary shape (i.e., (ISM bI^,/Ap). Polymer pellets volume and surface area were determined using an optical method according to the following steps. First approximately 200 pellets of the resin of interest laid flat on a black plate and photographed. To reduce inaccuracies due to transparency, the pellets were coated with icing sugar. An adaptive thresholding step was then applied to generate a binary image of the pellets. A clustering algorithm was used to extract each individual pellet allowing for the major and minor diameter of the pellets to be measured in pixels. The pixel to millimeter conversion factor was calculated by precisely measuring the base plate in mm and in pixels. The third dimension (or the thickness) of each pellet was then calculated based on the mass and density of the batch.

The stripping bin model included a multi-phase heat transfer model, which was developed and solved simultaneously with a mass transfer computational model. The multiphase heat transfer model was defined based on heat conduction within pellets, heat convection between pellets surface and the bulk nitrogen phase, VOCs latent heat of evaporation and heat advection representing the enthalpy transfer along the pellets bed height by nitrogen flow.

Thermal conduction in each pellet was described by the Fourier’s law in spherical coordinates:

(eq. 1)

In Eq. 1, p, C_p and k are the density, specific heat-capacity and thermal conductivity of the POE resin in the solid state. The partial differential equation in eq. 1 was solved using appropriate initial and boundary conditions to yield temporal variation of temperature within pellets along their radius — i.e., T = T(r, t). A constant thermal conductivity k of 0.2 W/m.K was implemented for the purpose of solving the partial differential equation in eq. 1. A linear relationship with temperature T (in °C) and resin density p (in g/cm³) was used to estimate the specific heat-capacity C_p (in J/kg«°C) of the POE resin in eq. 1 as follows:

C_p = 1000 X [(0.2648/? - 0.2172) X T + (-12.2172/? + 12.9296)]

(eq. 2)

The pellets were assumed to have an initial temperature (T_o) before entering the stripping bin, and thus their initial temperature at t = 0 was T_o for all values of r. For symmetry reasons, the first boundary condition was defined as dT/dr = 0 at the pellet center for all values of t. The second boundary condition was used to connect the pellet surface temperature (i.e., the temperature at r = r_p where r_p = d_SM/ 2) to the bulk nitrogen temperature (T_N) and evaporation of VOCs as defined below:

(eq. 3) and h_h are pellet surface area, the i-th VOC evaporation mass flowrate at the pellet surface, the i-th VOC latent heat of evaporation and heat convection coefficient between pellets surface and bulk of nitrogen, respectively. Each VOC component latent heat of evaporation could be estimated using known equations of state in the art. One example of an appropriate equation of state is the PC-SAFT equation of state with the proper parameters. Close to atmospheric pressures, the heat of evaporation is not a strong function of pressure. Thus, in the present disclosure, a temperature dependent equation has been fitted to the values predicted by PC-SAFT for numerical purposes.

The latent heat of evaporation (in kJ/kg) of the VOC components, e.g., process solvent (2 -methylpentane) and the comonomer (1 -octene), using the PC-SAFT model predictions can be described as a function of temperature (in °C) using the following linear relationships:

^1-octene = -0.5715T + 376.77

^2-methylpentane = -0.6412T + 358.67

(eq. 4)

The convective heat transfer coefficient between pellets surface and bulk of nitrogen (h_h) was calculated according to generally known correlations in chemical engineering literature as follows:

(eq. 5)

In Eq. 5, Pr, Re, j_h and St_h are Prandtl’s number, Reynold’s number, Chilton and Colburn J-factor for heat transfer in packed beds and Stanton’s number for heat transfer in packed beds, respectively. In eq. 5, p_mt_x, C_pmix, k_mix and p_mt_x are vapor phase viscosity, specific heat capacity, thermal conductivity and density and the terms D_p and V represent pellet diameter and the superficial velocity defined as vapor phase volumetric flowrate divided by bin cross sectional.

Knowing the composition of the vapor phase (mixture of nitrogen and VOCs), temperature, and pressure, the Peng-Robinson equation of state could be used to predict the transport properties of the vapor phase. The vapor phase comprises a mixture of varying percentages of nitrogen, at least one a-olefinic comonomer and the process solvent. However, due to low pressure (~1 atm) and temperature of the stripping process (which was limited by the VI CAT softening temperature of the resin), the ideal gas law can be used for prediction of the vapor mixture density of the vapor phase mixture at a temperature range of 40-85°C and a pressure range of 100-120 kPa. This assumption was checked and confirmed by comparing the ideal gas predictions to those of Peng-Robinson equation of state (EOS) in ASPEN Properties V10.

In this context, by assuming ideal gas properties, the specific heat capacity C_pmix and thermal conductivity k_mix (denoted as B_mix in eq. 6) were estimated by applying a simple mixing rule as follows:

(eq. 6)

In Eq. 6, A, is the ideal gas property of the i-th component in the vapor phase, x_t is the mass fraction of the i-th component. Using a simplified version of the Chapman-Enskog-Brokaw- Wilke mixing rule, the viscosity of the vapor mixture ( _mix) was estimated as follows:

(eq. 7)

In Eq. 7, is the mole fraction of the i-th VOC component in the mixture, is the viscosity of the i-th VOC component and is the in which MWj and Mw_t were i-th and j-th components molecular weights. Accuracy of above estimations were evaluated against Peng-Robinson equation of state and transport properties models in ASPEN Properties V10. It was determined that the specific heat capacity, thermal conductivity and viscosity of the vapor mixture could be estimated by applying the simple mixing rules described in Eqs. 6 and 7 within an average error of 0.47%, 1.40% and 2.24%, respectively.

The local and temporal variation of the bulk nitrogen temperature (T_w) was determined based on the heat advection due to nitrogen flow along the bed height as

(eq. 8)

In Eq. 8, p_N, C_pN, A_b, e, m_N and z are nitrogen density, nitrogen heat capacity, pellets bed cross-sectional surface area, bed void fraction, nitrogen mass flow rate and spatial coordinates along the pellets bed height, respectively. As can be seen, heat conduction in the nitrogen phase along the bed height has been neglected due to high superficial velocity of nitrogen. In the computational model, the bed void fraction e had a value of 0.35.

For smooth initialization of the problem, the initial temperature of nitrogen was assumed to be at equilibrium with pellets, i.e., T_N = T_o at t = 0 for all values of z. This did not introduce any error in the calculations, as the initial nitrogen content in the bed will leave the process within seconds. The boundary condition required to solve this equation was the inlet nitrogen temperature, i.e., T_N = T_Niniet at z = 0 for all values of t. The inlet nitrogen temperature T_Niniet is known as nitrogen stream is heated to a specific temperature before entering the bin.

At time zero there was no pellets in the bin and the bin was filled at a rate of 53.4x 10³ kg/h taking about 4 h to reach capacity of the bin at 200 x 10³ kg. An algorithm according to the following steps was applied to incorporate the filling stage into the mathematical model: (a) time required to fill one discrete node t_node in the bin was calculated based on the node capacity and production rate, (b) first node is loaded with pellets and is simulated for duration of t_node, (c) second node is loaded with pellets and both nodes are simulated for duration of t_node. and (d) steps (a) and (b) were continued until all the nodes are filled and bin is full.

Assuming the total time for filling the bin is t^m, the pellets stripping in the model was only simulated during the time period of t^m — t_node. This sequence was intentionally designed to underestimate the level of stripping happening during the filling period, therefore the required HUT for stripping was always overestimated by a time ranging from 0 to t_node. A similar approach was taken during the interval where the stripping bin was emptied.

During the stripping process, pellets were recirculated from the bottom to the top of the bed at a certain frequency and rate. A recirculation rate of 5.0 x 10³ kg/h was applied for 10 minutes in every hour in the current disclosure. Thereby, the Hourly Recirculation Rate (HRR) is 1.67 x 10³ kg/h. An algorithm according to the following steps was applied at the end of each hour during simulation to incorporate the recirculation process by: (a) continuing the stripping process for a period of time in hours defined by node mass (mass of pellets in each node) divided by HRR, (b) removing the first node at the bottom of the bin, (c) shifting all other nodes down by one node, and (d) placing the first node at the top.

Figure 1 is a simplified process flow diagram of a system 100 for devolatilizing a polymer solution 102. The polymer solution 102 is generated in a reactor (not shown), and fed into the system 100 through a pressure control valve 104. An intermediate pressure separation feed heater (IPSHTR) 106 heats the polymer solution 102 to the flashing temperatures, for example, greater than about 250°C, forming a heated polymer solution 108.

A deactivator solution 110 may be injected into the heated polymer solution to kill any remaining catalyst. The heated polymer solution 108 is then flashed across an intermediate pressure separation (IPS) control valve 112 into an intermediate pressure separation (IPS) vessel 114. The vapor 116 flashed off the heated polymer solution 108 is removed through an overhead outlet, and sent to a condenser, and then a distillation column for purification and reuse. The concentrated polymer solution 118 is removed through a bottom outlet.

Polymer additives 120, such as antioxidants, among others, can be added to the concentrated polymer solution 118. The concentrated polymer solution 118 is then flashed across a low-pressure separation (LPS) control valve 122 into an upper chamber (LPS1) 124 of a LPS vessel. The vapor 126 flashed from the concentrated polymer solution 118 is then removed through an overhead line and sent to the distillation column.

The resulting polymer solution flows through a sieve plate 128 into a lower chamber (LPS2) 130 of the LPS vessel. More vapor 132 flashes from the polymer solution as it flows through the sieve plate 128 and is removed from an outlet positioned under the sieve plate 128. The final polymer melt 134 is removed from the lower chamber 130 through a bottom outlet. In embodiments described herein, the LPS system 136 is modified to increase the extent of devolatilization to reduce the VOC content in stream 134, and consequently to lower the HUT required to achieve a final VOC level of 150 ppm in the finished pellets. The modifications are shown in Figures 2-11. In Figure 2, modifications to the LPS system 136 to lower the VOC content in stream 134, as described herein, are shown. The modifications may be used individually or in any combinations, including the complete set, as shown in Figure 2, to provide further decreases in the VOC content in stream 134. Figures 3-11 describe results for modeling of the individual elements of the LPS system 136.

Figure 2 is a simplified process flow diagram of a generic LPS system 200 for devolatilizing a solution phase polymer. Like numbered items are as described with respect to Figure 1. In the generic low-pressure separation system 200, the low-pressure separation vessel of Figure 1 is divided into two vessels. In various embodiments, different process units are placed between the two vessels to enhance the removal of solvent.

In the generic LPS system 200, the concentrated polymer solution 118 from the IPS vessel 114 of Figure 1 is flashed through the LPS control valve 122 into a first low-pressure separation vessel, LPS1 204. LPS1 204 corresponds to the upper chamber 124 of the single low-pressure separation vessel described with respect to Figure 1. The vapor 206 flashed from the polymer solution exits through a top outlet of LPS 1 204, and a more concentrated polymer solution 208 exits through a bottom outlet of LPS 1 204. As described with respect to Figure 1, the vapor 206 is directed to a condenser and a distillation column for purification and reuse.

In some embodiments, the more concentrated polymer solution 208 is passed through a heat exchanger 210 to increase the temperature, for example, by heat exchange with steam. In some embodiments, the more concentrated polymer solution 208 is passed from the heat exchanger 210 into a mixer 212.

The mixer 212 may be an in-line type static mixer. A devolatilization agent 214 can be introduced into the more concentrated polymer solution 208 immediately before the mixer 212, between sections of the mixer 212, after the mixer 212, or in other locations along the line, as described herein. The devolatilization agent 214 can include the polymerization solvent, hot ethylene, steam, or any combinations thereof.

The mixed polymer solution 216 is then flashed across a second LPS control valve 218 into a second low-pressure separation vessel, LPS2 220. In some embodiments, a vacuum pump 222 is used to pull a vacuum on LPS2 220 to increase the amount of vapor 224 removed from the mixed polymer solution 216 through a top outlet of LPS2 220. The vapor 224 is then directed to the condenser and distillation column for purification and recycling.

The final polymer melt 134 exits LPS2 220 through a bottom outlet. As described herein, the final polymer melt 134 may then be further processed, for example, through a devolatilizing extruder, and then pelletized. The pellets are directed to one or more stripping bins in which further solvent is removed, for example, by tempered nitrogen prior to the pellets being directed to a final storage bin.

Figure 3 is a simplified process flow diagram of another system 300 for devolatilizing a solution phase polymer. As described herein, the test polymer is the 880 density, 0.5 MI POE resin. The heat and mass balance were modeled in ASPEN Plus to determine the expected composition of streams 208 and 134.

The system 300 of Figure 3 is the base case for dividing the LPS vessel into two vessels, LPS1 204 and LPS2 220. In the system 300, the more concentrated polymer solution 208 from LPS1 204 is directly flashed across the second LPS control valve 218 into a second low-pressure separation vessel, LPS2 220.

Using this system 300 as the input to the ASPEN Plus model, the residual VOCs in the more concentrated polymer solution 208 coming from LPS1 204 and the final polymer melt 134 coming from LPS2 220 has been simulated. TABLE 1 summarizes the results for these two streams, labeled by the reference numbers 208 and 134, for each stream. As used herein, SM is the process solvent, which is generally methyl pentane, and FC is the comonomer, which is generally octene. Further, FE is ethylene and RA is the polymer phase.

TABLE 1: LPS1 and LPS2 Stream Condition in System 300

The model predicts that the VOCs in the final polymer melt 134 from LPS2 220 includes about 1.6 wt. % of SM, and about 0.8 wt. % of FC. From the simulated VOCs content for stream 134 for the base case in TABLE 1, the required HUT for stripping of this resin has been estimated using the modeling scheme described in Eqs. 1 through 8 to be 44.2 hours. In this context, for the base case and all other cases modeled, as a conservative assumption, the VOCs content of polymer melt stream 134, the VOCs content of pellets at the cutter location, and the initial VOCs content of pellets when entering the stripping bin were assumed to be equal. For the purpose of this simulation, the normal operating conditions for the stripping bins are assumed to be at a flow of 10 Mg per hour ( 10⁴ kg per hour) of nitrogen, stripping at a nitrogen inlet temperature of about 5°C below the VICAT softening point, or about 48°C, performed in a 200 Mg (200 x 10³ kg) bin. The conditions of the stripping bin have been assumed to be constant for all other cases modeled.

The modeling in the following figures was performed with the LPS1 tail conditions, e.g., the VOCs, of the more concentrated polymer solution 208 kept constant and modifications applied only to downstream of LPS1 204. Accordingly, the modeled HUTs can be directly compared between the different modifications.

Figure 4 is a simplified process flow diagram of another system 400 for devolatilizing a solution phase polymer. In this system 400, the more concentrated polymer solution 208 is heated to 230°C in the heat exchanger 210. This provides a higher temperature in LPS2 220 and increases the devolatilization of both SM and FC. The resulting VOCs and conditions in the final polymer melt 134 are shown in TABLE 2. As for TABLE 1, the output streams are labeled with the reference numbers 208 and 134, for each stream.

TABLE 2: LPS2 Tail Conditions for System 400

Increasing the temperature of the more concentrated polymer solution 208 to 230°C before flashing the more concentrated polymer solution 208 across the second LPS control valve 218 into LPS2 220 decreases the VOCs of the final polymer melt 134 by about 50 wt. % from the more concentrated polymer solution 208. The calculated HUT for stripping of the resin with this level of cutter VOCs is 36.2 hours.

Heating the more concentrated polymer solution 208 to high temperatures reduces the cutter VOCs, as both SM and FC are more volatile at higher temperature. However, this may create black specks as a higher concentration polymer solution is being heated. Further, it may cause problems at the cutter of the pelletizer due to the high melt temperatures that are fed to the extruder.

However, since more heat is provided in the phase separation, the outlet temperature of the IPSHTR 106 (Figure 1) can be reduced to compensate. Reducing the outlet temperature may increase the VOC content of the more concentrated polymer solution 108, without affecting the VOC content of the final polymer melt 134.

Figure 5 is a simplified process flow diagram of another system 500 for devolatilizing a solution phase polymer. In system 500, a static mixer 212 on the more concentrated polymer solution 208 is mixed with a devolatilization agent 214, which is, in this example, 5 Mg/h of SM at 250°C, resulting in a combined stream 226. The temperature of the devolatilization agent 214 may range from an equivalent temperature to the more concentrated polymer solution 208 to about 100°C above the temperature of the more concentrated polymer solution 208. This will increase the temperature of the more concentrated polymer solution 208.

While not wishing to be limited by theory, it was believed that the increase in SM in the vapor space of LPS2 220 reduces the octene vapor pressure and promotes more removal of the FC. The resulting VOCs and conditions for the final polymer melt from LPS2 220 are shown in TABLE 3.

TABLE 3: LPS2 Tail Conditions for System 500

As expected, relative to the stream 134 in the base case in system 300, the SM content of the final polymer melt 134 from LPS2 220 in system 500 has not changed. However, the FC content has dropped by about 60% from the more concentrated polymer solution 208. The calculated HUT for stripping of the resin with this level of cutter VOCs is 38.9 hours. Generally, the HUT for stripping of 880 density is mostly determined by the FC content of the pellets. Therefore, 60% less FC in the pellets, even with the same SM content, is effective in reducing the HUT.

As shown, the addition of the hot solvent to the more concentrated polymer solution 208 coming from UPS1 204 can be performed through a static mixer before UPS2 220. However, this requires the addition of a solvent heater, and a change in the diameter of UPS2 220, since more solid is now removed from the melt, increasing the risk of polymer carryover. The solvent (SM) used can be sourced through makeup solvent. The vaporized solvent will be sent to the distillation column with the vapor 206 from the overhead of UPS 1 204 and the vapor 224 from the overhead of UPS2 220.

In some embodiments, the devolatilization agent 214 can be injected downstream of the second UPS control valve 218 or directly injected into UPS2 220. Either of these cases will eliminate the need for the static mixer 212. Thus, these modifications could be performed with fewer equipment changes. Further, the improvement in the VOCs for these implementations are close to those shown in TAB EE 3.

Figure 6 is a simplified process flow diagram of another system 600 for devolatilizing a solution phase polymer. In system 600, the devolatilization agent 214 added to the mixer 212 is 3 Mg/h of ethylene (FE) at 200°C. Compared to the use of hot SM, less mass is used, as the molar flow of the devolatilization agent (FE or SM) is the determining factor. The addition of hot FE to the solution has same effect on the VOCs removal in LPS2 220, except that both FC and SM are reduced compared to the base case. The VOCs and conditions of the final polymer melt 134 from LPS2 220 are shown in TABLE 4.

TABLE 4: LPS2 Tail Conditions for System 600

The calculated HUT for stripping of pellets with this level of cutter VOCs is 34.4 hours. The addition of the hot ethylene as the devolatilization agent 214 can be performed through the static mixer 212 by the addition of an ethylene heater. As for the system 500 of Figure 5, the diameter of LPS2 220 may be increased, since more ethylene is now removed from the melt, which increases the risk of polymer carryover. Further, the extra ethylene will affect the performance of the LPS overhead condenser and increase the ethylene load in distillation.

The ethylene used can be sourced from the top product of the solvent column and returned to the solvent column after use. This option will provide a constant flow of ethylene to the solvent column, which can help with drying of the solvent column overhead condenser and diminish the corrosion issue there.

Figure 7 is a simplified process flow diagram of another system 700 for devolatilizing a solution phase polymer. Ethylene is more volatile than SM and cannot be liquefied at high temperatures easily. Mixing the ethylene in vapor phase with polymer melt can be challenging. To reduce these problems, in system 700, a mixture of SM and FE is used. Accordingly, the devolatilization agent 214 is 5 Mg/h of a 60%/40% (by weight) mixture of SM/FE. Temperature of the mixture has been set to 230°C. The VOCs and tail conditions for the final polymer melt 134 from the LPS2 220 are shown in TABLE 5. The calculated HUT for stripping of the pellets with this level of cutter VOCs is 34.9 hours.

TABLE 5: LPS2 Tail Conditions for System 700

Figure 8 is a simplified process flow diagram of another system 800 for devolatilizing a solution phase polymer. In system 800, both of the methods described with respect to the previous figures are combined. Specifically, the more concentrated polymer solution 208 from LPS1 204 is heated to 230°C, and then mixed with 3 Mg/h of a 60%/40% (by weight) mixture of SM/FE as devolatilization agent 214. The heating and mixing can be performed in any order or at the same time. The VOCs and conditions in the final polymer melt 134 from LPS2 220 are shown in TABLE 6. The calculated HUT for stripping of the pellets with this level of cutter VOCs is 30.7 hours. TABLE 6: LPS2 Tail Conditions for System 800

Figure 9 is a simplified process flow diagram of another system 900 for devolatilizing a solution phase polymer. The pressure in LPS2 220 controls the partial pressure of both SM and FC in the vapor space. Accordingly, as the pressure is lowered the VOCs in the final polymer melt 134 is also decreased. In system 900, the pressure of LPS2 220 is reduced to 80 kPa(a) compared to 140 kPa(a) in the base case. The VOCs and tail conditions for the final polymer melt from LPS2 220 are shown in TABLE 7.

TABLE 7 : LPS2 Tail Conditions for System 900

The calculated HUT for stripping of the pellets with this level of cutter VOCs is 38.7 hours. The moderate vacuum level can be achieved with a vacuum pump 222. The diameter of LPS2 220 may be increased to lower the possible carryover of polymer due to the higher vapor velocity. Further, the discharge pressure of the vacuum pump 222 is set to at least 140 kPa(a), so that the vapor 224 can tie into the LPS overhead condenser.

Figure 10 is a simplified process flow diagram of another system 1000 for devolatilizing a solution phase polymer. The configuration of system 1000 is the same as the system 300 of the base case of Figure 3. However, the 1 -octene purity is increased from 70% to 85% (or higher) in the polymerization reaction. Accordingly, the reactor outlet composition will have a lower amount of octene isomers. This translates to a lower amount of FC in the final polymer melt 134 from LPS2 220.

The modeling indicates that the effect of this lower FC would be about the same has system 900 of Figure 9, resulting in an HUT of about 39.2 hours. However, increasing the octene purity to 85% in the distillation process requires a significant amount of purging from the FC column, which will add significantly to costs. This is the easiest option to manage the HUT for 880 density POEs, as there would be no change in the system configuration.

Figure 11 is a simplified process flow diagram of another system 1100 for devolatilizing a solution phase polymer. In system 1100, the devolatilization agent 214 is 3 Mg/h of steam at 250°C and saturated pressure (39.8 bar). The resulting VOCs and conditions for the final polymer melt 134 from the EPS2 220 are shown in TABEE 8.

TABLE 8: LPS2 Tail Conditions for System 1100

The calculated HUT for stripping of the pellets with this level of cutter VOCs is 25.4 hours. It should be noted that 3 Mg/h of steam is more effective than FE as the Mw of water (18) is lower than FE (28). The addition of the steam into the more concentrated polymer solution 208 can easily be done through the static mixer 212. However, this option will necessitate drying of the vapor 224 from the overhead of LPS2 220, as the hydrocarbons will be mixed with steam. Further, there will be some trace of water in the final polymer melt 134 to the extruder. Accordingly, additives may be used to prevent polymer oxidation due to the presence of the water.

Figure 12 is a plot comparing holdup time (HUT) for the systems of Figures 3 to 11. In addition, the HUTs calculated for each case is summarized in TABLE 9. All of the options discussed herein are effective in reducing the required HUT for stripping of POEs. Accordingly, anyone, or any combination, of the options may be chosen, for example, based on costs. As for the previous tables, the column labeled 208 are the VOCs of the more concentrated polymer solution in Figures 3 to 11, while the columns labeled 134 are the VOCs of the final polymer melt 134 in Figures 3 to 11.

TABLE 9: Summary of Systems

Here we have shown the effect of reduced cutter VOCs on the required HUT for stripping. It is also possible to use these options and lower costs further by reducing the nitrogen circulation rates and keeping the number of stripping bins the same. The reduced nitrogen circulation translates to lower operating expenses (OPEX) as compressor power goes down. Further, reduced nitrogen circulation reduces the compressor size, and thus, the capital expenses (OPEX).

Costs could also be lowered further by reducing the number of stripping bins. Fewer stripping bins lowers the amount of equipment used, and thus the CAPEX. Further, fewer stripping bins reduces the total nitrogen circulation which has benefits for both OPEX and CAPEX.

The fluid at the top of the stripping bins is a mixture of nitrogen and VOCs that has been stripped from the pellets. This fluid may be directed to a nitrogen purification system for the removal of VOCs. By transferring less VOCs from the phase separation area to the finishing area, the load on the nitrogen purification system would be decreased. A nonlimiting example of technologies suitable for the removal of VOCs is adsorption and regeneration operations, which include pressure swing adsorption (PSA), temperature swing adsorption (TSA), and vacuum swing adsorption (VSA). Thus, by decreasing the VOCs content at the cutter location, the absorbent beds would be swung less frequently between absorbing and regenerating, which can contribute to further savings in both CAPEX and OPEX.

Figure 13 is a process flow diagram of a method 1300 for devolatilizing a plastomer. The method 1300 begins at block 1302, when a polymer solution is flashed across a first pressure control valve into a first low-pressure separation vessel to form a more concentrated polymer stream from the first low-pressure separation vessel.

In various embodiments, the more concentrated polymer stream can then be heated, or mixed with a devolatilization agent, or both. The devolatilization agent can be polymerization solvent (SM), or ethylene (FE), or both. In some embodiments, the devolatilization agent is steam.

At block 1304, the more concentrated polymer solution is flashed across a second pressure control valve into a second low-pressure separation vessel to form a final polymer melt. In some embodiments, a vacuum is pulled on the second low-pressure separation vessel to increase the amount of solvent flashing off the more concentrated polymer solution.

The final polymer melt is then conveyed to a pelletizer, such as an underwater pelletizer. The pellets are conveyed to one or more stripping bins to allow any remaining solvent to vaporize, and be removed, for example, by a nitrogen stream.

Definition of Terms

Other than where otherwise indicated, all numbers referring to process conditions (temperature, pressure, etc.), quantities of ingredients, etc., used in the specification and claims are to be understood as modified in all instances by the term "about." Accordingly, unless indicated to the contrary, the numerical parameters set forth in the following specification and attached claims are approximations that can vary significantly depending upon the raw materials used or the desired ethylene polymer produced. At the very least, and not as an attempt to limit the application of the doctrine of equivalents to the scope of the claims, each numerical parameter should at least be construed in light of the number of reported significant digits and by applying ordinary rounding techniques.

It should be understood that any numerical range recited herein is intended to include all sub-ranges subsumed therein. For example, a range of “1 to 10” is intended to include all sub-ranges between and including the recited minimum value of 1 and the recited maximum value of 10; that is, having a minimum value equal to or greater than 1 and a maximum value of equal to or less than 10. Because the disclosed numerical ranges are continuous, they include every value between the minimum and maximum values. Unless expressly indicated otherwise, the various numerical ranges specified in this application are approximations. Similarly, a range of 0% to 100% is intended to include all subranges between and including the recited minimum value of 0% and the recited maximum value of 100%; that is, having a minimum value equal to or greater than 0% and a maximum value of equal to or less than 100%.

In order to form a more complete understanding of the invention, the following terms are defined and should be used with the accompanying figures, the detailed description of the various embodiments and the claims.

As used herein, the term “monomer” refers to a small molecule that chemically reacts to become chemically bonded with itself or other monomers to form a polymer. Nonlimiting examples of monomers include ethylene (ethene), propylene (propene), and C4 to C12 a-olefins.

As used herein, the term “polymer” refers to a macromolecule composed of one or more monomers connected together by covalent chemical bonds. The term polymer is meant to encompass, without limitation, homopolymers (containing one type of monomer), copolymers (containing two monomer types), terpolymers (containing three monomer types), and quatropolymers (containing four monomers types), etc.

As used herein, the term “ethylene polymer”, refers to polymers produced from the ethylene monomer and optionally one or more additional monomers. The term ethylene polymer is meant to encompass, ethylene homopolymers, ethylene copolymers, ethylene terpolymers and ethylene quatropolymers, etc. Other commonly used terms to describe ethylene polymers include, but are not limited to, high-density polyethylene (HDPE), medium density polyethylene (MDPE), linear low -density polyethylene (LLDPE), very low-density polyethylene (VLDPE), ultralow density polyethylene (ULDPE), plastomer, and elastomers.

The term “heterogeneously branched ethylene polymer” or “heterogeneous ethylene polymer” refers to a subset of the ethylene polymer group that are produced using Ziegler- Natta or chromium catalysts.

The term “homogeneously branched ethylene polymer” or “homogeneous ethylene polymer” refers to a subset of the ethylene polymer group that are produced using a single site catalyst or metallocene catalyst. It is well known to those skilled in the art, that the homogeneous ethylene polymer group is frequently further subdivided into “linear homogeneous ethylene polymer” and “substantially linear homogeneous ethylene polymer”. These two subgroups differ in the amount of long chain branching. More specifically, linear homogeneous ethylene polymers have an undetectable amount of long chain branching; while substantially linear ethylene polymers have a small amount of long chain branching, typically from 0.01 long chain branches/1000 carbons to 3 long chain branches/ 1000. A long chain branch is defined as a branch having a chain length that is macromolecular in nature, i.e., the length of the long chain branch can be similar to the length of the polymer back-bone to which it is attached. In this disclosure, the term homogeneous ethylene polymer includes both linear homogeneous ethylene polymers and substantially linear homogeneous ethylene polymers.

As used herein, the term “oligomers” refers to an ethylene polymer of low molecular weight, e.g., an ethylene polymer with a weight average molecular weight (Mw) of about 2000 to 3000 daltons. Other commonly used terms for oligomers include “wax” or “grease”. In a solution polymerization process the presence of oligomers in the process solvent can be problematic, e.g., oligomers may deposit on and foul heat transfer surfaces.

As used herein, the term “V/L” refers to a vapor/ liquid separator, wherein a process stream enters the V/L separator (vessel or tank) and is separated into two streams, wherein one stream is ethylene polymer rich, and the other stream is solvent rich. As described herein, this is generally performed by flashing solvent from a polymer solution.

As used herein, the term “light-end impurities” refers to chemical compounds with relatively low boiling points that may be present in the various vessels and process streams within a continuous solution polymerization process; non-limiting examples include, methane, ethane, propane, butane, nitrogen, CO2 , chloroethane, HC1, etc.

As used herein the term “oxygenated impurities”, refers to trace amounts of water, fatty acids, alcohols, ketones, aldehydes, etc.; such impurities are potential catalyst deactivating poisons.

As used herein, the term “heavy impurities” refers to linear or branched, saturated or unsaturated, C8 to C30 hydrocarbons.

Testing Procedures

The multi-pass rheometer is a capillary rheometer in which several capillaries of different lengths and diameters are enclosed within a high temperature/pressure cell that is capable of confining a solution. The vertical cell has pistons at both ends of the cell, and during operation, the solution is sheared back and forth through the capillaries. Under conditions in which a steady shear is achieved, the pressure drop across the capillaries, A/' or IN- OUT, is measured and, the apparent viscosity of the fluid is determined as a function of the shear rate in each capillary. Before the shearing is initiated, the pistons are moved with respect to each other in order to obtain a desired static pressure for the system.

The “apparent viscosity”, //a of a polymer solution is defined for a given shear rate as the pressure drop, AP or PIN-POUT across the capillaries when a polymer solution is forced through a capillary of constant diameter, at a constant static pressure, temperature and polymer solution composition. The apparent viscosity, //a is equal to the ratio of “shear wall stress”, T_W, to “wall shear rate”, (dy/dt)_w ^N, for a Newtonian fluid: where V is the average fluid velocity in the tube in m/s (i.e., the velocity at which the pistons are moved within the capillaries), AP is the pressure drop across the capillaries in MPa (or Pa), L is the length of the capillaries in meters, R is the radius of the capillaries in meters and r_w is defined as above.

A person skilled in the art will recognize, that use of a multi-pass rheometer in the prescribed manner, allows for the measurement of the apparent viscosity of an off-line polymer solutions as a function of temperature, pressure and shear rate.

The apparent viscosity of a polymer solution undergoes a dramatic change at the cloud point pressure, due to the formation of two-liquid phases. Measurement of the cloud point pressure at different temperatures yields a cloud point curve which demarcates the two-liquid region from the single liquid region of the phase diagram for a given polymer solution. Modeling the phase behavior of a polymer solution of interest can be carried out using phase equilibrium calculations with a suitable thermodynamic model calibrated with liquid-liquid equilibrium data collected using the multi-pass rheometer on cloud point conditions.

DESCRIPTION OF EMBODIMENTS

An embodiment described herein provides a method for devolatilizing a polymer. The method includes flashing a polymer solution across a first pressure control valve into a first low-pressure separation vessel to form a more concentrated polymer solution from the first low-pressure separation vessel, and flashing the more concentrated polymer solution across a second pressure control valve into a second low-pressure separation vessel to form an outlet stream from the second low-pressure separation vessel.

In an aspect, the method includes heating the more concentrated polymer solution prior to flashing the more concentrated polymer solution across the second pressure control valve.

In an aspect, the method includes injecting a devolatilization agent into the more concentrated polymer solution prior to flashing the more concentrated polymer solution across the second pressure control valve. In an aspect, the devolatilization agent comprises a heated ethylene stream. In an aspect, the devolatilization agent comprises a steam stream.

In an aspect, the devolatilization agent comprises a heated polymerization solvent stream. In an aspect, the devolatilization agent comprises a mixture comprising at a heated ethylene stream and a polymerization solvent stream. In an aspect, the polymerization solvent stream comprises mixed hexanes. In an aspect, the polymerization solvent stream comprises methyl pentane.

In an aspect, the method includes injecting the devolatilization agent into a static mixer on the more concentrated polymer solution prior to the second pressure control valve.

In an aspect, the method includes injecting a devolatilization agent into the second low-pressure separation vessel.

In an aspect, the method includes injecting a devolatilization agent added downstream stream of a second pressure control valve.

In an aspect, the method includes heating the more concentrated polymer stream, and injecting the devolatilization agent into a static mixer on the more concentrated polymer solution prior to the second pressure control valve.

In an aspect, the method includes pulling a vacuum on the second low-pressure separation vessel.

In an aspect, the method includes increasing a purity of a comonomer used in the polymerization.

In an aspect, the method includes flashing a reactor effluent from a polymerization reactor into an intermediate pressure separator to form the polymer solution.

In an aspect, the method includes feeding the outlet stream to an extruder system to form an extruded stream.

In an aspect, the method includes feeding the extruded stream to a pelletizer to form polymer pellets.

Another embodiment described herein provides a system to decrease volatile compounds in polymer pellets. The system includes a first pressure control valve on a line that is fluidically coupled to an inlet of a first low-pressure separation vessel, and a second pressure control valve on a line that fluidically couples the first low-pressure separation vessel to an inlet of a second low-pressure separation vessel.

In an aspect, the system includes a heater fluidically coupled between an outlet of the first low-pressure separation vessel and the inlet of the second pressure control valve. In an aspect, the system includes a static mixer fluidically coupled between an outlet of the first low-pressure separation vessel and the second pressure control valve, wherein the static mixer has an inlet for a devolatilization agent. In an aspect, the devolatilization agent comprises ethylene. In an aspect, the devolatilization agent comprises a polymerization solvent.

In an aspect, the devolatilization agent comprises a mixture comprising a heated ethylene stream and a polymerization solvent stream. In an aspect, the polymerization solvent comprises an isomer of hexane.

In an aspect, the system includes a heater fluidically coupled between an outlet of the first low-pressure separation vessel and the inlet of a static mixer, and the static mixer fluidically coupled between the outlet of the heater and the inlet of the second pressure control valve, wherein the static mixer has an inlet for a devolatilization agent.

In an aspect, the system includes a vacuum pump fluidically coupled to a vapor outlet of the second low-pressure separation vessel.

In an aspect, the system includes an inlet for a devolatilization agent fluidically coupled to the second low-pressure separation vessel.

In an aspect, the system includes an inlet for a devolatilization agent fluidically coupled to a line upstream of the second pressure control valve.

Other implementations are also within the scope of the following claims.

Claims

1. A method for devolatilizing a polymer, comprising: flashing a polymer solution across a first pressure control valve into a first low- pressure separation vessel to form a more concentrated polymer solution from the first low- pressure separation vessel; and flashing the more concentrated polymer solution across a second pressure control valve into a second low-pressure separation vessel to form an outlet stream from the second low-pressure separation vessel.

2. The method of claim 1, comprising heating the more concentrated polymer solution prior to flashing the more concentrated polymer solution across the second pressure control valve.

3. The method of claim 1, comprising injecting a devolatilization agent into the more concentrated polymer solution prior to flashing the more concentrated polymer solution across the second pressure control valve.

4. The method of claim 3, wherein the devolatilization agent comprises a heated ethylene stream.

5. The method of claim 3, wherein the devolatilization agent comprises a steam stream.

6. The method of claim 3, wherein the devolatilization agent comprises a heated polymerization solvent stream.

7. The method of claim 3, wherein the devolatilization agent comprises a mixture comprising a heated ethylene stream, and a heated polymerization solvent stream.

8. The method of claim 6, wherein the polymerization solvent stream comprises mixed hexanes.

9. The method of claim 6, wherein the polymerization solvent stream comprises methyl pentane.

10. The method of claim 3, comprising injecting the devolatilization agent into a static mixer on the more concentrated polymer solution prior to the second pressure control valve.

11. The method of claim 1, comprising injecting a devolatilization agent into the second low-pressure separation vessel.

12. The method of claim 1, comprising injecting a devolatilization agent added downstream of the second pressure control valve.

13. The method of claim 3, comprising: heating the more concentrated polymer solution; and injecting the devolatilization agent into a static mixer on the more concentrated polymer solution prior to the second pressure control valve.

14. The method of claim 1, comprising pulling a vacuum on the second low-pressure separation vessel.

15. The method of claim 1, comprising increasing a purity of a comonomer used in the polymerization.

16. The method of claim 1, comprising flashing a reactor effluent from a polymerization reactor into an intermediate pressure separator to form the polymer solution.

17. The method of claim 1, comprising feeding the outlet stream to an extruder system to form an extruded stream.

18. The method of claim 17, comprising feeding the extruded stream to a pelletizer to form polymer pellets.

19. A system to decrease volatile compounds in polymer pellets, comprising: a first pressure control valve on a line that is fluidically coupled to an inlet of a first low-pressure separation vessel; and a second pressure control valve on a line that fluidically couples the first low- pressure separation vessel to an inlet of a second low-pressure separation vessel.

20. The system of claim 19, comprising a heater fluidically coupled between an outlet of the first low-pressure separation vessel and the inlet of the second pressure control valve.

21. The system of claim 19, comprising a static mixer fluidically coupled between an outlet of the first low-pressure separation vessel and the second pressure control valve, wherein the static mixer has an inlet for a devolatilization agent.

22. The system of claim 21, wherein the devolatilization agent comprises ethylene.

23. The system of claim 21, wherein the devolatilization agent comprises a polymerization solvent.

24. The system of claim 21, wherein the devolatilization agent comprises a mixture comprising a heated ethylene stream, and a heated polymerization solvent stream.

25. The system of claim 23, wherein the polymerization solvent comprises an isomer of hexane.

26. The system of claim 19, comprising: a heater fluidically coupled between an outlet of the first low-pressure separation vessel and the inlet of a static mixer; and the static mixer fluidically coupled between the outlet of the heater and the inlet of the second pressure control valve, wherein the static mixer has an inlet for a devolatilization agent.

27. The system of claim 19, comprising a vacuum pump fluidically coupled to a vapor outlet of the second low-pressure separation vessel.

28. The system of claim 19, comprising an inlet for a devolatilization agent fluidically coupled the second low-pressure separation vessel.

29. The system of claim 19, comprising an inlet for a devolatilization agent fluidically coupled to a line downstream of the second pressure control valve.