Simultaneous Extractive and Azeotropic Distillation Separation

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Separations

Simultaneous Extractive and Azeotropic Distillation Separation Process for Production of PODEn from Formaldehyde and Methylal Zhenwei Han, Yuanyang Ren, Hong Li, Xingang Li, and Xin Gao Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b06044 • Publication Date (Web): 26 Feb 2019 Downloaded from http://pubs.acs.org on March 6, 2019

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Industrial & Engineering Chemistry Research

Simultaneous Extractive and Azeotropic Distillation Separation Process for Production of PODEn from Formaldehyde and Methylal Zhenwei Han, Yuanyang Ren, Hong Li, Xingang Li, Xin Gao School of Chemical Engineering and Technology, National Engineering Research Center of Distillation Technology, Collaborative Innovation Center of Chemical Science and Engineering(Tianjin), Tianjin University, Tianjin 300072, China

ABSTRACT： Poly(oxymethylene) dimethyl ethers (PODEn) as eco-friendly diesel additives have been widely paid attention by engineers. Among different routes for production of PODEn, highly concentrated formaldehyde and methylal, both of them from methanol, is thought to be an economical way. However, unreacted formaldehyde solution mixed with PODE2 reaction products, resulting in serious separation difficulties. The present work first proposed a novel separation concept, simultaneous extractive and azeotropic distillation, to solve the key problem of separation process of PODE2, formaldehyde, water and methanol. The experimental results showed that formaldehyde could be totally separated from bottom of the column by adding supplementary water as both extractant and entrainer. And less than 0.05wt% formaldehyde distillated from the column condenser could avoid formaldehyde precipitation. Then a rigorous model based on the chemical equilibrium and phase equilibrium was established and validated by good agreement with the experimental data. The model was then applied to investigate the effect of the main variables on the simultaneous extractive and azeotropic distillation process. Finally, a complete conceptual eco-friendly production process of PODEn was presented. 1. INTRODUCTION Poly(oxymethylene) dimethyl ethers (PODEn) with the linear structure of CH3-O-



Corresponding author, Tel: +86-022-27404701(X.G.); Fax: +86-022-27404705(X.G.). E-mail: [email protected] (Xin Gao).

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(CH2O)n-CH3 (n≥2)1 are promising diesel additives. Among different chain lengths, PODEn of chain lengths n=3-5 (PODE3-5) are preferable to diesel additives as fuel due to their similar properties to diesel2. Their high cetane numbers ( ﹥63) and oxygen contents ( ﹥45%) could reduce the emissions of soot and NOx during the combustion process3, 4. Large scale green production of PODEn would be crucial if diesel-PODEn blends are applied to tackle the worldwide formidable air pollution by particulate matter5. Besides, PODEn are discussed as safe fuels to direct oxidation fuel cells and green solvents for the chemical industry6-8. Two kinds of educts are necessary for the synthesis of PODEn: a methyl end group (CH3) provider such as methylal (DMM), methanol (MeOH) or dimethyl ethers (DME) and a monomer unit formaldehyde (-CH2O) provider such as trioxane (TOX), formaldehyde (FA) solution or paraformaldehyde (PF). In published works, there are three common synthesis routes. The first is DMM plus TOX route9-11. The second is DMM plus PF route12, 13. The last is MeOH plus conventional FA solution route14-16. The reactions of TOX or PF over an acid catalyst produce a high yield of PODE3-59, 12, but the reactions have high costs of raw materials TOX and PF. Conventional FA solution plus MeOH route has lower yield of PODE3-5 comparing to the TOX and PF routes14, 17. DMM and PODE2 are main products in this route due to Schulz-Flory distribution law of PODEn products 5, 18, 19. When MeOH is replaced by DMM in the reactants, the yield of PODE3-5 can be improved. And high water in the reactants competes for FA, decreasing the FA concentration for chain elongation20-22. Besides, water leads to decomposition of PODE3-5. So concentrated FA is always used to improve the selectivity of PODE3-5. Thus, highly concentrated FA solution plus DMM is a promising route with relatively high yield and selectivity of PODE3-5 and low costs of materials23. PODEn synthesis reactions from DMM and highly concentrated FA solution are described according to reactions (1)-(8): FA + H2O ⇌ MG1 FA +MGn−1⇌ MGn

(1) n≥2

FA + MeOH ⇌ HF1


(2) (3)

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FA + HFn−1⇌ HFn

n≥2

(4)

DMM + FA ⇌ PODE2

(5)

H+

H+

PODEn−1 + FA ⇌ PODEn

n≥3

(6)

H+

DMM + H2O ⇌ HF1 +MeOH H+

PODEn + H2O ⇌ HFn + MeOH

(7) n≥3

(8)

In aqueous FA solution, FA is almost completely bond in the oligomers poly(oxymethylene) glycols (MGn, HO-(CH2O)n-H), described in reactions (1) and (2)24. MeOH is a byproduct of the process according to reactions (7) and (8). Once MeOH appears in FA solution, high degree of polymerization of MGn depolymerize to low degree of polymerization of MGn and FA monomer in the reactions (1) and (2), then MeOH reacts with FA monomer to form poly(oxymethylene) hemiformals (HFn, HO-(CH2O)n-CH3), described in reactions (3)-(4)25, 26. The reactions (1)-(4) occur at all pH levels even without addition of any catalyst. By contrast, the following reactions (5)-(8) occur only in acid environment. The chain elongation from DMM to PODE2 and PODEn, FA is added in reactions (5)-(6). Due to existence of water in the reactants, decomposition of PODEn with water occur in the process described in reactions (7)(8)15. Complex reactants result in a big trouble in the distillation step especially when reactions (1)-(4) have to be taken explicitly into account. Burger et.al had shown PODEn with n ≥ 3 can be separated as bottom product from mixtures of FA, water, MeOH, DMM and PODEn with n ≥ 2 by simulations and continuous experiments27, 28. DMM can be distillated with a portion of MeOH to form DMM-MeOH binary azeotrope due to the lowest boiling point in the system. But no work investigates the separation of remaining system (PODE2 + FA + H2O + MeOH). There are two serious problems in the separation process. The first is ternary azeotropes in the system. PODE2-FA-H2O, PODE2-FA-MeOH can form two ternary reactive azeotropes28. Besides, FA is easily polymerized to solid PF in high concentration and low temperature. So a distillation of the mixture (PODE2 + FA + H2O + MeOH) will lead to FA polymerized to solid PF in the condenser when the ternary azeotropes are



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distillated from the overhead of the column. Solid PF can gradually blocks the condenser and greatly affects the continuity of the production process. Aiming at the separation problems of this route, we present simultaneous extractive and azeotropic distillation to separate FA from system (PODE2 + FA + H2O + MeOH) to break down the ternary reactive azeotropes and avoid FA precipitation. The feasibility of this process is verified by lab-scale experiments, and the process is also simulated by Aspen Plus simulation software. Important operation parameters of the column were studied by sensitivity analysis. Finally, we develop a novel production process of PODEn to enhance the efficient use of each componnet.

2. EXPERIMENTS 2.1 Chemicals PODEn products were supplied by Chengdu Organic Chemicals Co., Ltd. Chinese Academy of Sciences. The components were summarized in Table S1 in the Supporting Information by analyzing the composition. Pure water was provided by Yong Qingyuan Distilled Water Co., Ltd. The mixture feed (PODE2 + FA + H2O + MeOH) for simultaneous extractive and azeotropic distillation were obtained by distillation separation of PODE3-10 heavy components first and then DMM-MeOH azeotrope light components from the raw PODEn products. The composition of MeOH, FA, H2O and PODE2 are shown in Table 1. Table 1. Components of simultaneous extractive and azeotropic feed Components Mass fraction (%)

Boiling points (℃)

Analysis method

MeOH

15.57

64.7

GC

H2O

2.99

100

PODE2

58.26

105

Karl-Fisher titration titration GC

FA

23.18

19.9

Sodium sulfite method

2.2 Apparatus Simultaneous extractive and azeotropic distillation experiments are performed in a lab-


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scale glass distillation column (diameter 30mm, 2.1m triangular screw packing) as shown in Figure 1. It is equipped with a total condenser and an electrically heated thermosiphon reboiler. Temperatures are measured at the bottom and top of the column by mercurial thermometer and digital thermometer with an accuracy of±0.3K. The pressure is measured at the top of the column using a pressure gage with an accuracy of ±0.5 mbar. The mixture and supplementary water feed locations are in the 1/3 and 2/3 of the column. Mass flows of the mixture and supplementary water are measured by coriolis flow meters with an accuracy of ±0.5%.

Figure 1. Simultaneous extractive and azeotropic distillation apparatus. 2.3 Procedure The simultaneous extractive and azeotropic distillation procedure is like this: for the start-up, the distillation flask of the column is filled with the feed mixture and some supplementary water. The column is operated on total reflux after heat duty of the electrical heating jacket is set. When reflux liquids emerge at the condenser, the mixture



and supplementary water are pumped to the column in desired values. Reflux ratio is also set to a desired value. Establishment of the steady state takes up several hours. It is determined by monitoring the temperature profile and consecutive sampling at both the overhead and bottom of the column. After steady state is reached, the column is left undisturbed for one hour. Then overhead product is collected from collecting bottle while bottom product is obtained from distillation flask. For vacuum distillation, the vacuum pump is operated from the beginning and keeps constant until the distillation is over.

2.4 Analysis The overall mass fraction of

MeOH, DMM and PODEn are determined by gas

chromatography (GC, Agilent 7890B with HP-5 capillary column) with ethanol as internal standard and a flame ionization detector. The exact contents of HFn and MGn are hardly determinable by GC29. So the sodium sulfite method is applied to determine the overall FA content of the samples and the Karl-Fisher titration is used for analyzing the overall water content. Each sample is analyzed at least twice. The relative error for each of the three methods is typically below 2% and the sum of mass fraction is between 97% and 103% in all tests.

3. MODEL BUILDING 3.1 Physico-chemical model In the absence of acid catalyst, only reactions (1)-(4) have to be considered. The vaporliquid equilibrium is calculated from the extended Raoult’s law. The vapor phase is assumed to be a mixture of ideal gases of FA, H2O, MeOH, PODE2, MG1 and HF1. MGn and HFn with n≥2 are ignored in the vapor phase for their low saturation vapor pressure. The liquid-liquid equilibrium is calculated from the isoactivity criterion. In the calculations, the maximal chain length of MGn and HFn is limited to n=10. Increasing the chain length does not significant affect the model results, as the corresponding components are presented only in negligible amounts28. Their chemical equilibrium constants of reactions (1)-(4) are given in Table S2 in the Supporting


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Information. The nonideality of the liquid phase is considered using an UNIFAC based activity coefficient model. The structure groups in the system (FA + water + MeOH +PODE2) are given in Table S3 in the Supporting Information, which also contains the size and surface parameters R and Q of the groups. The UNIFAC group assignment is given in Table S4 in the Supporting Information. Monomeric FA, water and MeOH are modeled as individual groups. PODE2 consists of one H3C-O-CH2O-CH3 group (methylal) and one (CH2O)OME group. The UNIFAC group interaction parameters are given in Table S5 in the Supporting Information. Interaction parameters between group 1-9 were adopted from Kuhnert30 et al, which were originally fitted to VLE data in the system (FA + water + MeOH + DMM). Interaction parameters of group 10, namely (CH2O)OME, with group 1-9 were reported by Schmitz et al31 to consider PODEn. The parametrization, which was established using LLE data, does not perfectly predict the VLE in the system. Schmitz et al27 had modified parameters of group 10 with group 8 and 9 to fit experimental vapor-liquid equilibrium data. The remaining parameters of group (CH2O)OME were adopted from Schmitz et al. without any change31. 3.2 Process model The simultaneous extractive and azeotropic distillation process for separation of FA from the system (FA + water + MeOH + PODE2) using supplementary water as extractant/entrainer is depicted in Figure 2. The process has one column. The mixture and supplementary water are fed to the column, in which the components boil separated allowing the less volatile components to collect in the bottom and in the overhead of the column the light key components. Reliable model is important to design the process. In this section, a rigorous EQ stage model with reliable chemical equilibrium of chemical reactions (1)-(4) and thermodynamic method of UNIFAC is established by using RADFRAC module of commercial Aspen Plus software (version 8.4).



Figure 2. Process flowsheet for separation of FA from the system (FA + water + MeOH + PODE2) by simultaneous extractive and azeotropic distillation. To establish the operating conditions for simultaneous extractive and azeotropic distillation process, sensitivity analysis was done. The design of simultaneous extractive and azeotropic distillation column operation was shown in Figure 2. As can be seen, the mixture contains MeOH, PODE2, FA and some water. The supplementary water fed above the mixture feed stage to extract FA from the mixture. The distillate component was MeOH, H2O and PODE2. The bottom stream was mainly conventional FA solution with some MeOH.

4. RESULTS AND DISCUSSION 4.1 Lab-scale distillation experiments To verify the feasibility of separation of FA from the system (FA + water + MeOH + PODE2) by simultaneous extractive and azeotropic distillation, five experiments were performed typically. The effects of reflux molar ratio (R) and mass feed ratio of the supplementary water as extractant/entrainer to mixture (E/M) on the composition profile in the overhead and bottom stream were investigated. The operation parameters and mass composition in the overhead and bottom stream of these five runs are given in Table 2. Table 2. Operation parameters and overhead and bottom composition of the experiment.


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Experiments

Run 1

Run 2

Run 3

Run 4

Run 5

Reflux Ratio Mixture feed rate (g/h) Supplementary water feed rate (g/h) E/M mass feed ratio Overhead rate (g/h) Bottom rate (g/h) Overhead stream composition (wt%) MeOH FA H2O PODE2 Bottom stream composition (wt%) MeOH FA H2O PODE2

1 100 50 0.50 85.9 64.1

2 100 45 0.45 85.1 59.9

2 100 50 0.50 86.4 63.6

2 100 55 0.55 87.8 67.2

3 100 50 0.50 86.9 63.1

16.321 16.589 16.520 16.751 16.448 0 0.051 0.021 0 0.024 15.937 15.035 16.090 16.879 16.274 67.730 68.354 67.386 66.355 67.254 2.417 2.426 2.039 1.284 1.657 36.153 38.666 36.442 34.474 36.515 61.309 58.657 61.359 64.242 61.653 0.121 0.251 0.160 0 0.175

From the experiment, we can see that nearly no formaldehyde presents in the overhead stream and nearly no PODE2 presents in the bottom stream. These results indicate that supplementary water was a suitable extractant/entrainer to separate FA from the system (FA + water + MeOH + PODE2). In the quaternary mixture before supplementary water added, due to vast FA and small numbers of water and methanol, water and methanol are nearly represented by MG and HF according to reactions (1)-(4). Although PODE2water binary azeotrope is the lowest azeotrope in the system, there is little true water component actually. After vast supplementary water added, there is true water component existing in the system. Then, PODE2-water binary azeotrope can be distillated from the overhead of the column. And vast supplementary water reacts with FA to exchange some methanol that is bond in HF, so some methanol are released in the column and distillated from the overhead of the bottom. The process is an azeotropic distillation. Supplementary water was an entrainer in the azeotropic process. Also, to break down PODE2-FA-water and PODE2-FA-MeOH ternary azeotropes, supplementary water was an extractant in the extractive distillation process. For water has a good solvability with FA mainly because of its quick reaction with FA to form



MG1. Supplementary water extracts FA from PODE2-FA-water, PODE2-FA-MeOH ternary azeotropes to form an analogous FA-water binary azeotrope in the bottom of the column. The influence of E/M mass feed ratio is presented by comparing run 2, 3 and 4. Run 2, 3 and 4 were operated at the same reflux ratio. When E/M mass ratio increased from 0.45 to 0.55, the distillate FA mass fraction decreased from 0.051 to 0 and bottom PODE2 mass fraction decreased from 0.251 to 0. When more supplementary water was fed to the system, PODE2 is easier to contact water to form PODE2-H2O binary azeotrope to the overhead, and monomer formaldehyde could also reacts with it to form MG1 and be extracted to the bottom. But more supplementary water leads to larger energy consumption in the distillate step. The appropriate amount of supplementary water was also determined by bottom FA mass fraction. Run 3 indicated that when E/M mass ratio was 0.50, bottom FA mass fraction was about 36%. This conventional FA solution could be directly used in the industrial process. The influence of reflux ratio is illustrated by comparing 1, 3 and 5. When reflux ratio increased, both distillate FA and bottom PODE2 increased. When increasing the reflux ratio from 1 to 3, the change of distillate FA mass composition was from 0 to 0.024, bottom PODE2 mass fraction from 0.121 to 0.175. Because when R increased, more components containing FA and PODE2 were recycled back. Thus, the extractive effects decreased. In simultaneous extractive and azeotropic distillation process, there was a strange phenomenon that MeOH presented in the bottom product even the temperature of the bottom is 99℃. Hasse et.al carried out laboratory distillation experiments with mixtures of FA, water and MeOH. They also found that small number for the MeOH concentration in the bottom product. This phonomenon could be seen as a result of reaction kinetics of FA-water-MeOH32. Due to the comparatively slow oligomerization reactions of HFn, MeOH bound in HFn was not readily released in the stripping section. In this work, we investigated influences of E/M mass ratio and reflux ratio on the mass fraction of MeOH in the bottom product as shown in Table 2. When supplementary water increased, the bottom MeOH mass fraction decreased. Because supplementary


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water competes for FA to form MGn as reactions (1) and (2), higher degree of HFn degrated to lower HFn and FA as reactions (3) and (4), the total amount of higher HFn decreased. Thus, less MeOH was bound in the stream products. When reflux ratio increased, the bottom MeOH mass fraction decreased too. For R increased, the residence time of vapor and liquid increased in the column, HFn had comparatively longer time to release MeOH. More MeOH were released in the distillation process. Based on the experimental results, it is concluded that supplementary water is a fesaible extractant/entrainer to separate FA from the system (FA + MeOH + PODE2 + H2O). The overhead stream is mainly MeOH, H2O and PODE2 while the bottom stream is mainly 36wt% FA solution and small amount of MeOH. Supplementary water breaks down the PODE2-FA-H2O and PODE2-FA-MeOH reactive azeotropes and 36% FA solution extracted from bottom could avoid FA precipitation in the condenser and pipelines. 4.2 Model validation According to the lab-scale distillation experiments, model parameters such as column pressure, reflux ratio, supplementary water as extractant/entrainer to mixture mass feed ratio were input previously. Column pressure was set at 1bar, reflux ratio was 2, mixture feed mass rate was 100Kg/h. Furthermore, to ensure the accuracy of the established model, theoretical plate numbers were set reasonably to meet the characteristics of the column internal, i.e., the number of stages per meter is about 20-25/m for triangular screw packing. The condenser is the first stage and the reboiler is the last stage. Therefore, theoretical plate numbers of the column were set to 42. The supplementary water feed stage and mixture feed stage were set to the 14th and 28th theoretical stage, respectively. To validate the accuracy of the model, the simulation results were compared with some experimental data. The three data were derived from E/M mass fraction 0.45, 0.50 and 0.55, respectively. Data of experiment and simulations are given in Table S6 in the Supporting Information. Figure 3 displays the comparison of the distillate and bottom composition profile between the experimental data and the simulated results. From Figure 3, we could find the three experimental data and simulated results fitted well.



The largest deviation was not exceeding 5%. There were some reasons for the explanation of the deviation. First, UNIFAC interaction parameters of group (CH2O)OME with group 1-7 were established using LLE data, it did not perfectly predict the VLE in the system. Second, some measurement deviation occurred during the sample analysis. Besides, the EQ model assumed steadystate operation is reached in the experiment, however, actual thermodynamic and thermodynamic behavior were very complex in the column, and hence deviations were unavoidable.

Mass fraction of simulated values (%)

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

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100 90 80 70 60

+5

50

% -5

40

%

30 20 10 0 0

10

20

30

40

50

60

70

80

90

100

Mass fraction of experimental values (%)

Figure 3. Comparison between the calculated and experimental mass fraction of distillate H2O, MeOH and PODE2 and bottom H2O and FA.

4.3 Sensitivity analysis of simultaneous extractive and azeotropic column The following operating parameters were analysed by sensitivity analysis: number of total stages of the column (N), reflux ratio of the column (R), mass ratio of the supplementary water as extractant/entrainer to the mixture feed (E/M), mixture feed stage (MFS) and supplementary water as extractant/entrainer feed stage (EFS). The constraints for the designed specification are: (1) the overall mass fraction of PODE2, MeOH and water is excepted to be over 99% in the distillate, in other words,


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distillate FA mass fraction must be less than 1%; (2) the overall mass fraction of FA, MeOH and water is excepted to be over 99% in the bottom, so bottom PODE2 mass fraction must be less than 1%; (3) bottom FA mass fraction is excepted to be near 37%; (4) the least energy consumption is preferable. 4.3.1 Number of stages Figure 4 shows the influence of number of stages on bottom FA mass fraction and Figure 5 shows the influence of number of stages on distillate FA, bottom PODE2 mass fraction and reboiler duty. As can be seen in Figure 4, number of stages has little influence on bottom FA mass fraction. Combined with the experimental results, we could draw a conclusion that bottom FA mass fraction is mainly determined by E/M mass fraction. Other parameters such as N, R, MFS and EFS have little influence on bottom FA mass fraction and would not be discussed in these parameters analysis. It is observed that an increase of number of stages cause a decrease of bottom PODE2 mass fraction and reboiler duty from Figure 5. According to the constraints, it is observed among number of stages from 10 to 26, distillate FA and bottom PODE2 mass fraction constraints can be achieved. And when the number of stages is greater than 24, the reboiler duty would not decrease and then remained nearly unchanged. Therefore, number of stages of 24 is considered reasonable.



Mass fraction of bottom FA (%)

35.10

35.09

35.08

35.07

35.06

35.05

35.04 8

10

12

14

16

18

20

22

24

26

Number of stages

Figure 4. Influence of number of stages on bottom FA mass fraction. Operational conditions: D=84Kg/h, R=2, E/M=0.50, MFS=9, EMS=3.

57.910

1.0

Bottom PODE2 Distillate FA Reboiler duty

57.905

0.6

57.900

0.4

57.895

0.2

57.890

Reboiler duty (Kw)

0.8

Mass Fraction (%)

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

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57.885

0.0 8

10

12

14

16

18

20

22

24

26

Number of stages

Figure 5. Influence of number of stages on bottom PODE2 and distillate FA mass fraction and reboiler duty. Operational conditions: D=84Kg/h, R=2, E/M=0.50, MFS=9, EMS=3.


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4.3.2 Reflux ratio The effect of reflux ratio on distillate FA, bottom PODE2 mass fraction and reboiler duty are shown in Figure 6. It can be observed distillate FA mass fraction can be easily achieved and remained nearly unchanged among reflux ratio from 1 to 5. But increased reflux ratio results in increase in distillate FA mass fraction and reboiler duty. For more components are recycled back to the column, the extractive effects became worse. And when reflux ratio is larger than 3, bottom PODE2 mass fraction would be over 1%. From the sensitivity analysis, it is possible to conclude that at reflux ratio of 1, the energy consumption requirements and distillate FA and bottom PODE2 mass fraction are satisfied.

120 1.8 110


1.4

100

1.2

90

1.0

80

0.8

70

0.6

60

0.4

Reboiler duty (Kw)

1.6

Mass fraction （%）

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


50

0.2

40

0.0 30 1

2

3

4

5

Reflux ratio

Figure 6. Reflux ratio influence on bottom PODE2 and distillate FA mass fraction and reboiler duty. Operational conditions: D=84Kg/h, N=24, E/M=0.50, MFS=9, EFS=3.

4.3.3 E/M mass fraction In simultaneous extractive and azeotropic distillation process, E/M mass fraction is a very important operation parameter. In this part, the influence of E/M mass fraction was investigated at various E/M mass fraction.



The bottom FA mass fraction over E/M mass fraction is shown in Figure 7. Distillate FA and bottom PODE2 mass fraction and reboiler duty over E/M mass fraction is shown in Figure 8. It is noticed that increased E/M mass fraction result in a decrease in bottom FA mass fraction. Due to the bottom stream contains MeOH, it can not directly recycled back to PODEn synthesis reactor as reactant of highly concentrated FA solution. And highly concentrated FA solution is not easy to preserve. When E/M mass fraction is over 0.60, the FA mass fraction is below 30% and when E/M mass fraction is below 0.40, the FA mass fraction is over 40%. Therefore, according to constraint (3), the E/M mass fraction may be 0.45 or 0.50. Also, the simulation results show that water is distributed in both overhead and bottom. But it is more inclined to be with PODE2 to form binary azeotrope, when there is no sufficient water, FA will bond more MeOH and PODE2 in the bottom. Figure 8 shows bottom PODE2 and distillate FA mass fraction can be easily achieved for all of the evaluated E/M mass fraction range. Also, a rise in E/M mass fraction increased reboiler duty. For least energy consumption, at E/M mass fraction 0.45 is a feasible choice.

70

Mass fraction of bottom FA (%)

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

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65 60 55 50 45 40 35 30 25 20 0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0

E/M mass fraction

Figure 7. E/M mass fraction influence on bottom FA mass fraction. Operational conditions: D=84Kg/h, N=24, R=1, MFS=9, EFS=3.


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0.8

46 45


0.6

44 43

0.5

42 0.4 41 0.3 40 0.2

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Mass fraction (%)

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39

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37 0.1

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E/M mass fraction

Figure 8. E/M mass fraction influence on bottom PODE2 and distillate FA mass fraction and reboiler duty. Operational conditions: D=84Kg/h, N=24, R=1, MFS=9, EFS=3. 4.3.4 Mixture feed stage Figure 9 presents the influence of mixture feed stage on the distillate FA, bottom PODE2 mass fraction and reboiler duty. For the given separation system, the separation objective can be attained when MFS exceeds 4. However, more energy consumption is required for higher MFS. Thus, an optimal value could be 4 because with this value it was possible to obtain high product purity and lower energy consumption.



7 40

39

5 Bottom PODE2 Dstillate FA Reboiler duty

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38

3

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2 36

Reboiler duty (Kw)

6

Mass fraction (%)

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1 35 0 2

4

6

8

10

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16

18

20

Mixture feed stage

Figure 9. Mixture feed stage influence on bottom PODE2 and distillate FA mass fraction and reboiler duty. Operational conditions: D=84Kg/h, N=24, R=1, E/M=0.45, EFS=3. 4.3.5 Supplementary water as extractant/entrainer feed stage Figure 10 shows the supplementary water as extractant/entrainer feed stage analysis. As can be seen, bottom PODE2 and distillate FA mass fraction increases when supplementary water feed stage is close to the mixture feed stage because the shorter contact time between the mixture and supplementary water. However, bottom PODE2 and distillate FA mass fraction can be both below 1% among EFS from 2 to 4. And for a least energy consumption, an optimal value of 4 is preferable.


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1.0 40.8 0.9 Bottom PODE2 Distillate FA Reboiler duty

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39.0 2

3

4

Entrainer feed stage

Figure 10. Mixture feed stage influence on bottom PODE2 and distillate FA mass fraction and reboiler duty. Operational conditions: D=84Kg/h, N=24, R=1, E/M=0.45, MFS=4.

5. CONCEPTUAL PROCESS DESCRIPTION In the production process, DMM and PODE2 are considered to be recycled back to the PODEn synthesis reactor while water and MeOH are not. Water and MeOH compete for FA according to reactions (1)-(4). The reactions decrease the FA concentration for chain elongation in reactions (5)-(6)20-22. Besides, water leads to the depolymerization of PODEn according to reactions (7)-(8). Thus, water and MeOH should be removed from DMM and PODE2 as far as possible before DMM and PODE2 are recycled back to PODEn synthesis reactor. A flowsheet of a novel PODEn production process is depicted in Figure 11. The process starts with the production of DMM. Pure DMM can be obtained through conventional FA solution reacts with MeOH in a reactive distillation column33 (DMM RDC) and then purified by pressure sewing distillation in PSD1 and PSD2 due to pressuresensitive of DMM-MeOH azeotropic mixtures34,35. Pure MeOH was recycled back to DMM RDC. Pure DMM was fed to PODEn synthesis reactor to react with highly



concentrated FA solution to produce PODEn. Then PODEn products and unreacted DMM and FA solution feedstocks were fed to a rectifying sequence to separate main products of PODE3-5 and recycle other components. The feed stream comprises MeOH, DMM, FA, water and PODEn of different chain lengths. Column C1 can avoid FA polymerization when all other light components are distilled together as overhead product. The bottom product PODE3-10 of column C1 is fed to the vacuum column C2. The aim of column C2 is to obtain PODE3-5 as product. The overhead product of column C1 containing water and MeOH needs further separations. DMM-MeOH azeotrope can be separated by column C3 as overhead product and then fed to PSD1 and PSD2 for further separation. Simultaneous extractive and azeotropic column C4 can separate FA from the system (PODE2 + FA + H2O + MeOH) with supplementary water as an extractant/entrainer. The overhead products MeOH, water and PODE2 was fed to membrane unit to dehydrate water first35, this unit can also be replaced by adsorption process using zeolites as adsorber28. Column C5 is to separate MeOH and PODE2 from the overhead and bottom, respectively, because of the large difference between their boiling points. Our lab-scale experimental data of column C1, C2, C3, C4 and C5 are presented in Table S7 in the Supporting Information to support the feasibility of the conceptual process.

Figure 11. Flowsheet of the novel PODEn production process


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As can be seen in Figure 11, through the novel separation process, MeOH can be recycled back to DMM RDC, also conventional FA solution with slight MeOH can be recycled back to DMM RDC. DMM, PODE2, PODE6-10 can be recycled back to PODEn synthesis reactor. This integration process can make efficient use of every component of PODEn products.

6. CONCLUSION In order to solve the separation problems of PODEn system containing FA solution, a simultaneous extractive and azeotropic process for separation of FA from the system (PODE2 + FA + H2O + MeOH) with supplementary water as extractant/entrainer has been proposed. This work focuses on the experiment and simulations of the simultaneous extractive and azeotropic distillation process. Experiment results show FA can be totally separated from the system.Then the influence of various parameters such as number of stages, reflux ratio, supplementary water as extractant/entrainer to mixture mass ratio, mixture feed stage and extractant/entrainer feed stage were studied in the distillation column. Finally, combined with pressure sewing distillation, reactive distillation and membrane separation technique, we propose a compelete PODEn production process from DMM production to realize a more economical manufacturing process.

Supporting Information This information is available free of charge via the Internet at http://pubs.acs.org/. PODEn products composition. Chemical equilibrium constants of reactions (1)(4). UNIFAC structure groups with size and surface parameters. UNIFAC group parameters assignment. UNIFAC interaction parameters. Comparison of the experimental data and simulation results. Distillation experimental data.

Acknowledgements



The authors are grateful for the financial support from the National Key Research and Development Program of China (2018YFB0604903), National Natural Science Foundation of China (No. 21336007).

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Simultaneous Extractive and Azeotropic Distillation Separation

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