Reversible reaction-assisted intensification process for separating the

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Article

Reversible reaction-assisted intensification process for separating the azeotropic mixture of ethanediol and 1,2butanediol: Vapor-liquid equilibrium and Economic evaluation Hong Li, Zhenyu Zhao, Jie Qin, Rui Wang, Xingang Li, and Xin Gao Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.7b04921 • Publication Date (Web): 21 Mar 2018 Downloaded from http://pubs.acs.org on March 22, 2018

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Industrial & Engineering Chemistry Research

1

Reversible reaction-assisted intensification process for

2

separating the azeotropic mixture of ethanediol and

3

1,2-butanediol: Vapor-liquid equilibrium and Economic

4

evaluation

5

Hong Li, Zhenyu Zhao, Jie Qin, Rui Wang, Xingang Li, Xin Gao∗

6

School of Chemical Engineering and Technology, National Engineering Research Center of Distillation

7

Technology, Collaborative Innovation Center of Chemical Science and Engineering (Tianjin), Tianjin University,

8

Tianjin 300072, China

9

Abstract: Vapor-liquid equilibrium data of six binary systems composed of ethylene

10

glycol

11

2,3-butanediol (2,3-BD) were measured and regressed by NRTL model. With the

12

binary interaction parameters regressed by experimental data, four separation

13

processes to obtain polymer grade EG from the synthesis products were simulated,

14

including conventional distillation route, azeotropic distillation route, distillation

15

coupled with liquid-liquid extraction route and reaction-assisted distillation route. The

16

total annual costs and profits of the four routes were estimated and compared, which

17

indicates that the annual profit was mainly determined by the yield of EG. The results

18

also demonstrate the reaction-assisted distillation route was proved to be the best

19

strategy currently among those separation methods.

20

Keywords: Ethylene glycol, Vapor-liquid equilibrium, Process simulation, Economic

21

evaluation, Reactive separation ∗

(EG),

1,2-propylene

glycol

(1,2-PG),

1,2-butanediol

Corresponding author: Xin Gao E-mail: [email protected]. Tel: +86-022-27404701. 1

ACS Paragon Plus Environment

(1,2-BD)

and

Industrial & Engineering Chemistry Research 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

1

1. Introduction

2

Ethylene glycol (EG) is widely used in petrochemical industry and other fields, such

3

as polymer ester manufacturing and anti-freezer. It is mainly produced by ethylene

4

epoxidation and ethylene epoxide hydrolysis. However, the worldwide oil shortage

5

and the increasing demand of EG call for alternative non-oil routes to produce

6

glycols.1 Among those routes, C1 chemical process based on syngas mainly from coal

7

is proved to be promising.2 The liquid product of indirect syngas-to-EG reaction

8

contains 48.0wt%-51.0wt% methanol, 39.0wt%-43.0wt% EG, 1.5wt-2.5wt% water,

9

2.6wt-3.5wt% ethanol (EtOH), 2.0wt%-2.5wt% glycols including 1,2-butanediol

10

(1,2-BD), 2,3-butanediol (2,3-BD), 1,2-propylene glycol (1,2-PG), 1.5wt%-2.1wt%

11

unreacted methyl nitrite (MN), dimethyl oxatate (DMO), methyl glycolate (MG),

12

dimethyl carbonate (DMC) and 0.3wt-0.5wt% heavy glycols, including diethylene

13

glycol (DEG), triethylene glycol (TEG) and others.3

14

Currently, the traditional mode of separating the mixture consumes excessive amount

15

of energy because of the application of distillation column with large reflux ratio and

16

tray number to remove other glycols from EG, resulting in poor competitive

17

advantage of coal-based glycol production. Therefore, several methods have been

18

proposed to enhance the separation process, for example, azeotropic distillation.

19

Much work4,

20

reports on industrial application have rarely been seen. Besides, extraction is expected

21

to be used in separation process but the recovery and consumption of extracting

22

solvent should be taken into account.6 Moreover, reaction-assisted separation for the

5

has been devoted to selecting proper entrainer for the system but

2


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1

glycols developed in recent years was proved to be of great potential.7 In addition,

2

other novel separation ideas have been reported, such as extractive distillation2, 8 and

3

adsorption9, but they didn’t quite come off due to high toxicity or large wasting and

4

were dismissed in this paper.

5

Those ideas are all great but it is still not clear which route is more applicable to the

6

glycol system. Therefore, in this paper, four separation processes were designed and

7

evaluated based on process simulation by Aspen Plus. To ensure the accuracy of

8

process simulation, the vapor-liquid equilibrium data of glycol systems should be

9

firstly measured to retrieve the parameters of thermodynamic model.

10

2. Vapor-liquid equilibrium data

11

2.1 Measurement of thermodynamic data

12

The isobaric vapor-liquid equilibrium (VLE) data of the glycols has already been

13

measured by many researchers but it is still not systematic, especially the data of

14

binary systems composed of 1,2-BD, 2,3-BD, 1,2-PG and EG.10-15 Therefore, the

15

double circulating vapor-liquid equilibrium still was used to take measurements. The

16

temperature was measured with a precise digital thermometer (1552A-12-DL model,

17

provided by Fluke Calibration) with the accuracy ±0.01K. The pressure was regulated

18

by a pressure control system, including one vacuum pump, one digital manometer

19

(AOB-20 model) with the accuracy ±0.1 kPa and an aneroid barometer (DYM3 model)

20

with the accuracy ±0.2 kPa.

21

The VLE measurements for the systems (EG +1,2-BD), (EG +1,2-PG), (EG +2,3-BD),

22

(1,2-BD +1,2-PG), (1,2-BD +2,3-BD), (1.2-PG +2,3-BD) at pressure of 10.0kPa, 3



Page 4 of 32

1

30.0kPa and 101.3 kPa were performed in the still with 110 cm3 capacity. Liquid

2

phase and condensed vapor phase were continuously circulated to ensure rapid

3

establishment of vapor-liquid equilibrium. The system was thoroughly cleaned by the

4

standard sample before experiment to eliminate contaminants. During the experiments,

5

the equilibrium was assumed to be reached when the pressure and temperature

6

maintained for 45 minutes.

7

The composition of the sample was analyzed by gas chromatograph (GC7890Ⅱ,

8

Techcomp (China) Ltd. ) equipped with a hydrogen flame ionization detector (FID)

9

and a DB-WAX capillary column (30m × 0.25 mm × 0.25 µm, Agilent). The

10

temperature of injector, oven and detector were kept at 523.15K, 393.15K and

11

543.15K, respectively. The carrier gas was high purity nitrogen (mass fraction 0.9999)

12

and the flow rate was 1 mL•min−1. Each VLE data was measured at least three times

13

to ensure accuracy.

14

results by using the external standard curve to convert the peak areas into mole

15

fractions of each sample. Taken the measurement error into consideration, the

16

uncertainties of the vapor and liquid mole fractions were both calculated to be within

17

0.001. The isobaric VLE data (T, xi, yi) for the six binary mixtures were measured and

18

shown in Table S1-S6 in Supporting Information. The activity coefficients γi of the

19

liquid phase could be calculated from the experiment data by Eq.(1), as the vapor

20

phase was assumed as ideal gas under low experimental pressures.

The external standard was chosen to obtain quantitative analysis

(1) 21

where

was the saturated vapor pressure of pure substance i at the system 4


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1

temperature T, which was calculated by Extended Antoine equation based on Aspen

2

Plus database. yi is the vapor phase mole fraction, xi is the liquid phase mole fraction,

3

and P is the total pressure of the system. The VLE data of EG +1,2-PG and 1,2-PG

4

+1,2-BD at 10.0kPa and EG +1,2-BD at 101.3kPa has already been measured by other

5

researchers10,13 and was not measured repeatedly.

6

2.2 Consistency test and data regression

7

To verify the measured quality of VLE values, all experimental results were checked

8

with Herington test and van Ness test proposed by Kang et.al16. As Herington17

9

suggested, if

, the pertinent experimental isobaric VLE data can be

10

considered thermodynamically consistent. In the process, D and J were calculated as

11

Eq.(2) and Eq.(3).

(2)

(3)

12

where Tmax and Tmin were the highest and lowest temperature in the system,

13

respectively. The γ values of the systems (2,3-BD +1,2-PG, 2,3-BD +1,2-BD, EG

14

+1,2-PG, 1,2-PG +1,2-BD) were all between 0.95 and 1.10 at 10.0kPa, 30.0kPa and

15

101.3kPa, hence the area test was not employed. The values of D -Jfor the system

16

{EG +1,2-BD} were calculated to be 1.48, 0.09, 1.05 at 10.0kPa, 30.0kPa, 101.3kPa,

17

respectively, while the values of the binary system {2,3-BD +EG} were 3.71, 2.27

18

and 1.88 at 10.0kPa, 30.0kPa and 101.3kPa, which meant the experimental data

19

passed the Herington test. 5



Page 6 of 32

1

Van Ness point test described by Fredenslund18 was applied to confirm the accuracy

2

of the experimental data. The index average absolute deviation (AAD) was defined by

3

following equation: (4)

4

where Ci is the independent variable (bubble point temperature, vapor phase mole

5

fraction) of the pure component i, N is the number of experimental data points, the

6

superscript exp denotes data from experiments and cal denotes data calculated with

7

thermodynamic models.

8

The experimental data were correlated by minimizing object function Q, which was

9

defined by the sum of squared differences between calculated and experimental data

10

as Eq.(5): N

Q = ∑[( i =1

Te ,i - Tm ,i σ T ,i

)2 + (

Pe,i - Pm ,i σ P ,i

)2 + (

xe ,i - xm ,i σ x ,i

)2 + (

y e ,i - y m ,i σ y ,i

)2 ]

(5)

11

where y, x represent mole fraction of pure component in the vapor phase and in the

12

liquid phase. T and P are temperature and pressure. i represents each pure component

13

in the binary mixtures, N is the number of experimental data points, and σ represents

14

the estimated standard deviation of measured experimental values. e, m represent

15

estimated value and measured value, respectively.

16

Correlated interaction parameters, together with the average absolute deviation (AAD)

17

of equilibrium temperature, vapor phase mole fraction for the binary systems between

18

experimental values and calculated values (obtained from Wilson, NRTL and

19

UNIQUAC models) were shown in Table 1.

20

Table 1 Correlated Interaction Parameters and the Average Absolute Deviations (AAD) between 6


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Experimental and Calculated Values

1

AAD (T)/K

AAD kPa

11.2114 1025.1681 -481.7063

1,2-PG(1) +EG(2) -890.5198 0.0059 -164.0382 0.3 0.0060 102.8876 0.0061

0.0285 0.0402 0.0404

6.7907e-06 9.3288e-06 9.3784e-06

30.0

Wilson NRTL UNIQUAC

-303.8596 379.5311 -149.2076

-435.0841 407.0284 -160.0364

0.3

0.0086 0.0087 0.0087

0.0290 0.0367 0.0363

2.2166e-05 2.9864e-05 2.9381e-05

101.3

Wilson NRTL UNIQUAC

-303.8596 306.2324 -59.9569

-502.8533 598.3744 -301.4157

0.3

0.0068 0.0081 0.0080

0.0898 0.0942 0.0932

2.9896e-04 3.0985-04 3.0700e-04

10.0

Wilson NRTL UNIQUAC

-365.3776 -274.0664 76.2143

1,2-BD(1) +EG(2) 177.9303 0.0074 471.9168 0.3 0.0074 -137.4402 0.0074

0.0055 0.0094 0.0057

1.3376e-06 2.2593e-06 1.3666e-06

30.0

Wilson NRTL UNIQUAC

-553.1914 -344.9806 137.4240

218.5276 688.6855 -253.4244

0.3

0.0039 0.0046 0.0040

0.0137 0.0225 0.0132

1.1115e-05 1.8333e-05 1.0805e-05

101.3

Wilson NRTL UNIQUAC

-357.4847 -234.7664 12.3188

129.9490 476.8356 -87.5923

0.3

0.0030 0.0032 0.0030

0.0099 0.0110 0.0098

3.3322e-05 3.6732e-05 3.2597e-05

10.0

Wilson NRTL UNIQUAC

-466.6218 -343.8058 154.1807

2,3-BD(1) +EG(2) 251.7104 0.0085 550.6957 0.3 0.0086 -233.8679 0.0084

0.0468 0.0313 0.0450

1.0411e-05 6.8939e-06 1.0131e-05

30.0

Wilson NRTL UNIQUAC

-345.4576 -293.5719 72.6219

182.9973 477.0140 -126.3332

0.3

0.0039 0.0038 0.0039

0.0152 0.0206 0.0149

1.2047e-05 1.6138e-05 1.2037e-05

101.3

Wilson NRTL UNIQUAC

-174.4196 -94.9450 -91.3732

44.6094 227.5011 31.7564

0.3

0.0082 0.0082 0.0082

0.0118 0.0094 0.0128

4.0027e-05 3.3662e-05 4.2203e-05

10.0

Wilson NRTL UNIQUAC

214.5182 385.0055 -192.5574

1,2-PG(1) +1,2-BD(2) -309.1564 0.0039 -290.5894 0.3 0.0040 150.5187 0.0039

0.0494 0.0510 0.0492

1.1516e-05 1.1953e-05 1.1472e-05

30.0

Wilson NRTL UNIQUAC

-104.6842 -139.0940 98.9830

82.7958 167.2831 -119.0317

0.3

0.0066 0.0066 0.0066

0.0197 0.0193 0.0186

1.6738e-05 1.6403e-05 1.5832e-05

101.3

Wilson NRTL UNIQUAC

242.3635 408.5374 -225.1837

-369.2433 -300.8118 171.5125

0.3

0.0047 0.0050 0.0047

0.0122 0.0130 0.0122

3.8772e-05 4.0436e-05 3.8509e-05

model

B12/K

B21/K

10.0

Wilson NRTL UNIQUAC

p/kPa

α

AAD (y)

7


(P)/


Page 8 of 32

10.0

Wilson NRTL UNIQUAC

-304.6439 -296.5335 149.0168

2,3-BD(1) +1,2-PG(2) 215.6571 0.0043 389.0352 0.3 0.0042 -188.3749 0.0043

30.0

Wilson NRTL UNIQUAC

-346.6179 -316.4019 165.1422

240.3541 416.5423 -210.8901

101.3

Wilson NRTL UNIQUAC

219.9860 333.0377 -29.0645

-299.5595 -268.0018 24.5504

0.0093 0.0092 0.0087

0.0030 0.0042 0.0226

1.1982e-05 1.716e-05 7.0655e-05

Wilson NRTL UNIQUAC

277.1605 418.4868 -259.2443

2,3-BD(1) +1,2-BD(2) -454.6703 0.0054 -310.2117 0.3 0.0061 192.6849 0.0056

0.0931 0.0792 0.0864

2.0524e-05 1.7419e-05 1.9031e-05

Wilson

361.2147

-817.1036

0.0053

0.0081

6.3589e-06

NRTL

799.0103

-464.4162

0.0056

0.0427

3.3472e-05

UNIQUAC

-457.6764

282.4736

0.0052

0.0251

1.9698e-05

Wilson NRTL UNIQUAC

-412.5052 -366.2347 198.3000

290.0799 469.8231 -249.2973

0.0041 0.0042 0.0041

0.0547 0.0548 0.0561

1.7443e-04 1.74244-04 1.7866e-04

10.0

30.0

101.3

0.0210 0.0246 0.0217

4.692e-06 5.4823e-06 4.8438e-06

0.3

0.0058 0.0056 0.0058

0.0032 0.0053 0.0032

3.1805e-06 4.3514e-06 3.1468e-06

0.3

0.3

0.3

1

The AAD of bubble point temperature, pressure and vapor phase mole fraction are

2

less than 0.1K, 1Pa and 0.01 respectively in all of cases, which indicate that the three

3

models agree well with experimental data and can be used in industrial application.

4

Among those models, non-random two liquid (NRTL) model, as a common tool for

5

fitting in many literatures19-21, has been reported for VLE and LLE in multicomponent

6

systems consisting alcohols, water, aromatic hydrocarbons, e.g. toluene and glycols

7

that are all involved in the simulation while the other two models have rarely been

8

reported. Therefore, NRTL thermodynamic model was applied to simulate the

9

separation process. The activity coefficient calculated in Aspen Plus is shown as

10

Eq.(6). (6)

8


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1

where

2 3 4

In order to retireve the binary interaction parameters correspongding the models, the

5

data regression module of Aspen Plus was employed based on the least-squares

6

method subject to the maximum likehood principle following Britt-Luecke

7

algorithm22.The model parameters regressed by experimental data were shown in

8

Table 2. Table 2 Model parameters fitted by experimental data

9 component i

1,2-PG

1,2-BD

2,3-BD

1,2-PG

2,3-BD

2,3-BD

component j

EG

EG

EG

1,2-BD

1,2-PG

1,2-BD

0 0 306.2324 598.3744 0.3

0 0 -234.7664 476.8356 0.3

0 0 -94.9450 227.5011 0.3

0 0 408.5374 -300.8118 0.3

0 0 333.0377 -268.0018 0.3

0 0 -366.2347 469.8231 0.3

(K) (K)

10

The NRTL parameters of the binary systems composed of methanol, ethanol, glycols,

11

DMC, DMO and MG were adjusted based on the reported literatures1, 23-25.

12

According to VLE results correlated by NRTL model, the key to obtain EG with high

13

quality is removal of other glycols, including 1,2-PG, 2,3-BD as these components

14

form close boiling point mixtures at the entire range of their compositions, especially

15

1,2-BD which forms azeotrope with EG shown in Figure 1. The difficulty to separate

16

the mixture of glycols only by conventional distillation requires application of other

17

special distillation technology.

18

3. Processes descriptions and simulations

19

To obtain polymer grade EG, three separation processes were proposed and simulated 9



1

by Aspen Plus to calculate equipment sizes and energy cost, including azeotropic

2

distillation route, distillation coupled with liquid-liquid extraction and distillation

3

coupled with acetalization reaction. Here separation process currently applied in

4

conventional industry was also simulated as an investment for comparison. The

5

distillaton columns were firstly designed by DSTWU modules to determine tray

6

number, reflux ratio and distillate rate. The columns were then rated by RadFrac

7

modules. The operational conditions of the equipments were optimized by sensitive

8

analysis with the object to decrease energy cost, e.g. reduce reflux ratio and reboiler

9

duty under the prerequisite of product quality.

10

3.1 Conventional distillation route (CD route)

11

The four-column distillation process widely used in coal-based ethylene glycol

12

industry was shown in Figure 2. The operational conditions of columns and main

13

information of streams were also demonstrated in the figure while some details were

14

omitted due to space reasons.

15

Note that methanol could be served as raw materials to produce MN and DMO in

16

previous CO coupling reaction.3 Therefore, the majority of methanol was recovered

17

by Col-11 with 57 stages, where the reflux ratio was set at 1.9 to ensure the methanol

18

purity according to simulation results. The recovery of methanol was determined

19

according to sensitive analysis of Col-11. From Figure 3, when the recovery of

20

methanol is too high a sharp increase of reboiler duty occurs due to increasing

21

difficulty of separation according to the tendency of minimum reflux ratio Rmin. Thus

22

the methanol recovery was reasonably set at 95.0%. B11 from the bottom stream of 10


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1

Col-11 was separated by Col-12 to remove the remaining methanol, by-product

2

ethanol and other light components. D12 was assumed to be used as fuel and was

3

unnecessary to be purified. Then unreacted MG and DMO was recovered by Col-13.

4

For separation of glycols, it is widely used as expediency in industry to distillate a

5

part of EG carrying almost all other glycols to the top of column, of which the

6

concentration of EG should be up to 95wt% so that it could be still sold as qualified

7

product at relative lower price. Then EG (> 99.9wt%) could be gained from the

8

sidestream while other heavy glycols (DEG, TEG) were removed from the system at

9

the bottom of Col-14. Obviously, it is necessary to apply large tray number and large

10

reflux ratio to Col-14 to ensure the purity of EG at the sidestream. The concentration

11

profile of EG in Col-14 (at RR=24) was shown in Figure 4. As it will keep up to

12

99.9wt% below the 15th plate in vapor phase, the drew-position was accordingly set at

13

the 16th plate considering tower performance. Col-14 was operated at 0.05bar to

14

decrease reboiler temperature so that cheaper low pressure steam could be used.

15

3.2 Azeotropic distillation route(AD route)

16

The form of close boiling point mixture and even azeotropic mixture discounts EG

17

quality and yield in CD route and hence causes large diseconomy in manufacture.

18

Therefore, many researches have been devoted to development of azeotropic

19

distillation technology for glycols system.

20

The application of azeotropic distillation to separating glycols system was firstly

21

proposed by Berg26 in 1990. He listed about 30 potential entrainers to enhance the

22

separation via heterogeneous azeotropic distillation. The past two decades has 11



1

witnessed developing performance of azeotropic distillation technology in separation

2

of EG and 1,2-BD. Song5 has used homotopy-Newton united method to predict

3

azeotropic component and proved ethylene benzene was the best entrainer by fuzz

4

mathematics. Wang27 simulated separation process for EG and 1,2-BD through

5

azeotropic distillation and compared it with conventional distillation. The result

6

showed that azeotropic distillation is superior to conventional distillation on the

7

operating cost and investment. Li28 conducted several experiments to separate EG and

8

1,2-BD and the purity of EG can be up to 99.7wt% and recovery reached 80.32%.

9

Furthermore, the feasibility of application of azeotropic distillation to separating

10

glycols has been studied by experiments.4 Based on those researches, a feasible

11

azeotropic distillation process with the optimized results was shown in Figure 5.

12

Similar to CD route, methanol and DMO should be recovered firstly and meanwhile

13

light impurities, including ethanol and water were removed from the system by

14

Col-21, Col-22 and Col-23 which were similar with Col-11, Col-12 and Col-13 and

15

not shown repeatedly in Figure 5. Then the mixture of glycols was pumped into

16

Col-24, where the mixture of EG, 1,2-BD, 1,2-PG, 2,3-BD was distillated out. The

17

majority of EG mixed with heavy glycols from the bottom was separated by Col-25.

18

1,2-BD in the mixture from the top of Col-24 was expected be removed by azeotropic

19

distillation using ethylbenzene (EB) as entrainer. The model parameters of binary

20

mixture of EG and ethylbenzene were adjusted according to Yang’s work29 to ensure

21

the accuracy of simulation. According to the result of property analysis in Aspen Plus,

22

1,2-PG is bad for the process of azeotropic distillation. Therefore, the mixture should 12


Page 12 of 32

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1

be pretreated to remove a part of 1,2-PG from the system by Col-26 with 69 trays to

2

ensure quality of the final product. The bottom product of Col-26 was pumped into

3

Col-27, together with EB, where 1,2-BD was carried by EB to the top and thus EG

4

(>99.9wt%) could be obtained at the bottom. Then ethylbenzene-rich phase from the

5

decanter flowed back to Col-27 while glycol-rich phase was treated by Col-28 to

6

recover EB.

7

The mass ratio of entrainer to feed was 0.5-0.8 according to the result of sensive

8

analysis shown in Figure 6. Under the goal that the recovery of 1,2-BD in Col-27 was

9

set at 0.995, the addition of entrainer makes it easier to separate the two glycols,

10

therefore the minimum tray number and input energy decrease with the increase of EB.

11

However, too much entrainer will cause the increase of column capacity and thus the

12

increase of energy cost.

13

3.3 Distillation coupled with liquid-liquid extraction route(ED route)

14

Apart from azeotropic distillation, there are some researches on application of

15

liquid-liquid extraction on separating EG and 1,2-BD. It has been proved that toluene

16

was the best agent among aromatics, ethers, alkanes as well as ketones.6, 9, 28 Sun23

17

simulated several extraction processes for separating EG and 1,2-BD using toluene as

18

extracting solvent. Based on those researches, a detailed conceptual design and

19

operation of separation process was shown in Figure 7. Col-31-Col-33 were similar

20

with Col-11-Col-13 in Figure 2 and not shown.

21

After the recovery of methanol, MG and DMO, the stream B33 was pumped to the

22

extraction section. The extraction column was operated at 313K and 1 bar based on 13



Page 14 of 32

1

the experimental results in the literature28, in which the feed is counter-currently

2

brought into contact with toluene i.e. extracting solvent. The extract stream leaving

3

the extraction column contains 1,2-BD, coextracted EG and extracting solution.

4

The process designs included extraction and distillation, which were simulated by

5

using the Extract and RadFrac modules of Aspen Plus, respectively while the

6

equilibrium stage model was used, of which the distribution coefficient of toluene was

7

estimated by UNIFAC group-contribution method with the help of LLE data20. To

8

ensure the quality of EG product, the mass fraction of 1,2-BD in raffinate phase

9

should be below 0.001, then the minimum solvent to feed mole ratio was calculated to

10

be 0.1 using Eq. (7). (7)

11

where S is flow rate of the solvent, F is flow rate of the feed, Xin is mass fraction of

12

the solute i.e. 1,2-BD in the feed, Xout is mass fraction of the solute in the raffinate

13

phase, Yin is the mass fraction of the solute in the entering solvent, and Yout is the mass

14

fraction of the solute in the extractant phase, which is in equilibrium with Xin. As

15

suggested by King30, an appropriate estimation of the optimal solvent to feed ratio

16

should be 1.15 ~2 times of (S/F)min to ensure product quality and in this work it was

17

set to be 1.5. Therefore, the flow rate of solvent was calcultated to be about 5000kg/h.

18

The extraction phase (D34) including the majority of toluene and 1,2-BD flowing out

19

from the top of extraction tower was separated by Col-36 to recover the solvent while

20

the raffinate phase (B34) from the bottom was separated by Col-35, where the mixture

21

of 1,2-PG and toluene was distillated from the system and treated as fuel. The bottom 14


Page 15 of 32 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


1

liquid of Col-35 was separated by Col-38, where EG (>99.9wt%) was obtained at the

2

top. Col-36-Col-38 were operated under lower pressure to reduce operational

3

temperatures to keep products/solvent from spoilling and save energy.

4

3.4 Reaction-assisted distillation route (RD route)

5

The novel idea has been developed in recent years that chemical reactions could be

6

used to enhance separation process. Dhale’s work31 has shown excellent performance

7

in the recovery of EG, 1,2-PG and 1,2-BD from aqueous solution through

8

acetalization reaction. The acetals formed by acetaldehyde (ALK) and glycols were

9

easier to be separated and then additional reactive distillation columns were needed

10

for hydrolysis of acetals back to EG and 1,2-PG, which showed great advantage for

11

EG with low concentration. As for manufacturing of EG, Huang et al.32 studied

12

reaction kinetics and developed a reactive distillation process for separation of EG

13

and 1,2-BD. Based on those previous researches, a conceptual design was established

14

shown in Figure 8.

15

Similar to AD route in Figure 5, light components were removed by Col-41-Col-43

16

and the majority of EG was obtained after distillation by Col-44 and Col-45. Here

17

they were omitted as they were similar with AD route. The brief introduction of

18

Reaction-assisted distillation (RD) route begins with the stream D44 from Col-44.

19

Unlike AD route, the distillates from Col-44 (similar with Col-24) was pumped,

20

together with acetaldehyde into the reactor R41, which was operated under 1.5 bar,

21

333.15 K. A 2:1 mole ratio of acetaldehyde to glycols was set to ensure the conversion

22

of glycols based on the previous study32. Besides, the conversions of EG and 1,2-BD 15



Page 16 of 32

1

were fixed to 0.61 and 0.70, respectively, based on the previously published

2

experimental results7. The conversions of 1,2-PG and 2,3-BD have not been reported

3

and they were assumed to be 0.70, due to their chemical properties were similar with

4

1,2-BD and their concentrations were too little to have effect on energy cost. The

5

acetalization reactions of glycols following Eq(8)-(11) occurred in the reactor. OH H

O

-H2O

+ O

OH

(ALK)

(EG)

(8) O

(2-MD)

OH H

O

-H2O

+

(9) O

OH

O

(ALK)

(1,2-PG)

(2,4-DMD)

OH H

O

(10)

-H2O

+ O

OH

O

(ALK)

(1,2-BD)

(4-EMD)

OH H

O

-H2O

+ OH

(2,3-BD)

O

(11) O

(2,4,5-TMD)

(ALK)

6

The distillation tower Col-46 was to recover the unreacted acetaldehyde and glylcols.

7

The distillation tower Col-47 and Col-48 were aimed to separate the acetals.Then EG

8

could be obtained by acetals hydrolysis in the continuous reactive distillation column

9

R42, where the conversions of hydrolysis reaction were set at 100% according to the

10

experimental results in a pilot scale column as mentioned by Atul D. Dhale33. The

11

hydrolysis of 2,4-DMD, 4-EMD and 2,4,5-TMD occurred in R43, where acetaldehyde

12

was recovered. Then water in S42 and S43 was removed by Col-48 and Col-49,

13

respectively. 16


Page 17 of 32 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


1

The process included reactor and distillation tower, which were simulated using the

2

Rstoic and RadFrac modules of Aspen Plus, respectively, while the reactive

3

distillation columns R42-R44 were simulated by using RadFrac modules. The binary

4

interaction parameters of acetals and water were obtained from regression of the VLE

5

data measured by Chopade et al.34,

6

equilibrium (LLE) data of the binary system 2MD +4EMD were measured in the

7

preliminary work7.

8

4. Economic estimation

9

After the initial steady state design of each section, optimization was performed to

10

find a set of optimal design and operating variables with minimizing total annual cost

11

(TAC), including adjusting tray number and reflux ratio, choosing the optimal

12

position to feed and draw-out36. According to Luyben’s work37, the four processes

13

were evaluated by annual cost and profit. Total annual cost (TAC) contains capital

14

investigation and operational cost.

35

. The VLE data as well as liquid-liquid

(12) 15

Payback period was usually set at 3 years in industry. The capital investigation

16

included columns and heat exchangers and the purchase costs of pumps and other

17

equipments were too small to be taken in account38.

18

Stainless steel was specified as the material in coolers, reboilers, tanks, and internals

19

to avoid iron contamination in the final product. Therefore, the cost of columns was

20

calculated by Eq.(13) based on the height of column L and the diameter of column D

21

esitimated by Tray Sizing module in Aspen Plus following Eq.(14). 17



Page 18 of 32

(13) (14) 1

where NT is tray number of each column.

2

The estimation of reactor cost followed Eq. (13), either. Under the assumption that the

3

reactor could be considered as CSTR(continuous stirred tank reactor) and the ratio of

4

D and L was generally set at 3. D and L were estimated by reactor volume, which was

5

calculated by Eq. (15). (15)

6

where Q is the flow rate of input stream, t is space time depending on optimal time of

7

reaction at different operating conditions. α is the ratio of effective volume to actual

8

volume, which is always set at 1.5~2 in industrial design to take the cost of additional

9

volumes and internals into account. The cost of catalyst is too low to be considered.

10

The cost of heat exchangers was calculated following Eq.16 based on the heat transfer

11

area of each heat exchanger. The latter was estimated by Aspen Exchanger Design and

12

Rating following Eq.(16). (16)

13

Energy cost was calcuated based on the heat duty of heat exchangers. For reboilers

14

and heaters, the price of high pressure (41barg, 251℃) steam was 9.87$/GJ while the

15

prices of middle pressure (10barg, 184℃) and low pressure (5barg, 160℃) steam

16

were 8.22$/GJ and 7.78GJ$, respectively. In the four separation processes, the

17

condensers were all cooled by cooling water, whose price was 0.354$/GJ.

18

Apart from input energy, those separations need additional materials for supplement, 18


Page 19 of 32 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


1

including ethylbenzene in AD route, toluene in ED route as well as acetaldehyde in

2

RD route. The output value was determined by production of EG. The prices of

3

chemicals were inquired in Molbase database. The ecomomic estimation results of the

4

four separation processses were shown in Table 3. Table 3 Economic estimation of separation processes

5 Price

6

Column&Reactor/10 $ heat exchanger/106$ energy cost/106$/a additional material//106$ TAC/106$ output/(106$/a) profit/(106$/a)

CD

AD

ED

RD

9.72 4.36 44.00 0.00 48.69 296.28 247.59

8.99 6.25 17.94 0.40 23.42 317.75 294.33

7.88 5.63 25.10 3.74 33.34 318.42 285.08

9.19 6.08 23.76 0.01 28.86 336.76 307.89

6

Table 3 reveals that TAC is mainly determined by energy cost and the annual profit

7

depends mainly on the yield of EG (output). In CD route, the existence of 1,2-PG,

8

1,2-BD and 2,3-BD will result in decrease of EG output and increase of imported

9

energy and hence cause reduction of annual profit. Compared with the CD route, the

10

application of azeotropic distillation technology improves the yield of EG and

11

reduced the energy cost of separation, which demonstrated great potential. However,

12

ethylbenzene, as is proved to be the best entrainer though, has danger of serious

13

damage to health by prolonged exposure through inhalation and if swallowed.

14

Therefore, the application of azeotropic distillation technology on separation of

15

glycols system in industry requires development of non-toxic entrainer, e.g.

16

ionic-liquid. The application of liquid-liquid extraction on separating EG and 1,2-BD

17

is able to increase the productivity of EG but the process to remove 1,2-PG and

18

recover extracting agent still consumes large quantity of energy. In addition, owing to

19

great quantity and high toxicity of extracting solvent, ED route shows no advantage 19



1

whether from economic angle or from the perspective of the environment. By

2

comparison, reaction-assisted separation with acetaldehyde is an effectively way to

3

solve the puzzle of separating glycols mixture due to low energy consumption and

4

large EG output while additional products such as 1,2-PG and 1,2-BD could be

5

obtained. Thus, the RD route was calculated to be the best strategy for separation of

6

liquid product of ethylene glycol synthesis from syngas. Furthermore, as an idea

7

developed in recent years, reverse reacion-assisted separation needs more researches,

8

such as, developing new suitable reverse reactions e.g. reaction with butanone and

9

propionaldehyde39, application of reactive distillation to promote reaction conversion

10

and save energy, etc.

11

5. Conclusion

12

In this work, the rigorous steady state design and optimization work of four separation

13

processes for liquid products of EG synthesis from syngas were investigated. The

14

thermodynamic data of glycol system was measured to ensure the accuracy of process

15

simulation by Aspen Plus.

16

The result of economic analysis indicated that azeotropic distillation owned great

17

potential due to lower energy cost compared to CD route but it still needed promotion,

18

for example, development of entrainer. Among those separation routes, the distillation

19

process coupling with acetalization reaction was currently proved to be the best

20

strategy for the glycols system in industry due to high yield of EG and low energy

21

cost.

22

Supporting Information 20


Page 20 of 32

Page 21 of 32 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


1

Tables of vapor-liquid equilibrium data measured for the binary mixture composed of

2

EG, 1,2-PG, 1,2-BD and 2,3-BD are available in the Supporting Information.

3

Acknowledgements

4

The authors are grateful for the financial support from National Natural Science

5

Foundation of China (Nos. 21690084, 21336007).

6

Corresponding Author

7

*E-mail: [email protected] (Xin Gao) Tel: +86-022-27404701

8

Notes

9

The authors declare no competing financial interest.

10

References

11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34

(1) Sun, L.; Fu, J.; Li, W.; Zhan, X. L.; Xiao, W. D. Binary vapor-liquid equilibrium of methyl glycolate and ethylene glycol. Fluid Phase Equilib. 2006, 250, 33-36. (2) Chen, Y.; Hung, S.; Lee, H.; Chien, I. Energy-Saving Designs for Separation of a Close-Boiling 1,2-Propanediol and Ethylene Glycol Mixture. Ind. Eng. Chem. Res. 2015, 54, 3828-3843. (3) Yu, B.; Chien, I. Chemical Engineering Research and Design Design and optimization of dimethyl oxalate ( DMO ) hydrogenation process to produce ethylene glycol ( EG ). Chem. Eng. Res. Des. 2017, 121, 173-190. (4) Niu, Y.; Liu, Z.; Qiao, K.; Chen, M. Study on Separation of Glycol and 1,2-Propanediol with Heterogeneous Azeotropic Distillation. Contemp. Chem. Ind. (China) 2011, 40, 560-561. (5) Song, G. Study on the Separation of Glycol-Propanediol-Butanediol System Using Azeotropic Distillation. Thesis, Tianjin University, Tianjin, 2006. (6) Dai, C.; Pan, G.; Liu, Y.; Li, Q. A novel method to separate ethylene and 1,2-butanediol with azeotropic distillation coupled with extraction technology. China Patent CN 103396290-A, 2013. (7) Li, H.; Huang, W.; Li, X.; Gao, X. Application of the Aldolization Reaction in Separating the Mixture of Ethylene Glycol and 1,2-Butanediol: Thermodynamics and New Separation Process. Ind. Eng. Chem. Res. 2016, 55, 9994-10003. (8) Figueirêdo, M. F.; Brito, K. D.; Ramos, W. B.; Vasconcelos, L. G. S.; Brito, R. P.; Brito, K. D.; Ramos, W. B.; Vasconcelos, L. G. S.; Brito, R. P., Optimization of the Design and Operation of Extractive Distillation Processes Optimization of the Design and Operation of Extractive Distillation Processes. Sep. Sci. Technol. 2015, 50, 2238-2247. (9) Xiao, J.; Zhong, L.; Guo, Y. Separation method for glycol, propylene glycol and butylene glycol. China Patent CN 102372600 A, 2012. (10) Zhu, L. Yan, J.; Xiao, W. Determination and correlation of VLE data for ethylene glycol and 21



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

1,2-butanedioi system. Chem. Eng. (China) 2012, 40, 34-37. (11) Chiavone-filho, O.; Proustj, P., Vapor-Liquid Equilibria for Glycol Ether + Water Systems. J. Chem. Eng. Data 1993, 38, 128-131. (12) Yue, H.; Zhao, Y.; Ma, X.; Gong, J. Ethylene glycol: properties, synthesis, and applications. Chem. Soc. Rev. 2012, 41, 4218. (13) Yang, Q.; Jin, F.; Feng, X.; Zhi, J.; Yang, C., Isobaric Vapor–Liquid Equilibrium for Two Binary Systems ( n -Butanol + 1,4-Butanediol and γ-Butyrolactone + 1,4-Butanediol) at p = (30.0, 50.0, and 70.0) kPa. J. Chem. Eng. Data 2016, 61, 3034-3040. (14) Zhong, Y.; Wu, Y.; Zhu, J.; Chen, K.; Wu, B.; Ji, L. Thermodynamics in separation for the ternary system 1,2-ethanediol + 1,2-propanediol + 2,3-butanediol. Ind. Eng. Chem. Res. 2014, 53, 12143-12148. (15) Yang, C.; Feng, X.; Sun, Y.; Yang, Q.; Zhi, J. Isobaric Vapor–Liquid Equilibrium for Two Binary Systems{Propane-1,2-diol + Ethane-1,2-diol and Propane-1,2-diol + Butane-1,2-diol} at p = (10.0, 20.0, and 40.0) kPa. J. Chem. Eng. Data 2015, 60, 1126-1133. (16) J.W., Kang; V., Diky; R.D. Chirico; J.W. Magee; C.D. Muzny; I. Abdulagatov; A.F. Kazakov; Michael. F. Quality assessment algorithm for vapour-liquid equilibrium data. J. Chem. Eng. Data 2010, 55, 3631-3640. (17) Herington, E. F. G. Tests for the consistency of experimental isobaric vapour-liquid equilibrium data. J. Pet. Ins. 1951, 37, 457-470. (18) Hack C. W. Winkle M. V. Vapor-Liquid Equilibria of the Diacetone Alcohol–Water System at Subatmospheric Pressures. Ind. Eng. Chem. Res., 1954, 46, 2392-2395. (19) Esteban J.; Ladero M.; García-Ochoa F. Liquid–liquid equilibria for the systems ethylene carbonate+ethylene glycol+glycerol; ethylene carbonate+glycerol carbonate+glycerol and ethylene carbonate+ethylene glycol+glycerol carbonate+glycerol at catalytic reacting temperatures. Chem. Eng.g Res. Des., 2014, 94, 440-448. (20) Mohsen-Nia M.; Doulabi F. S. M.; Manousiouthakis V. I. (Liquid + liquid) equilibria for ternary mixtures of (ethylene glycol + toluene + n-octane). J. Chem. Thermodyn. 2008, 40, 1269-1273. (21) Cháfer, A.; Burguet, M. C.; Monton, J. B.; Lladosa, E., Liquid–liquid equilibria of the systems dipropyl ether+n-propanol+water and dipropyl ether+n-propanol+ethylene glycol at different temperatures. Fluid Phase Equilib. 2007, 262, 76-81. (22) Britt, H. I.; Luecke, R. H., The Estimation of Parameters in Nonlinear, Implicit Models. Technometrics 1973, 15, 233-247. (23) Sun, L. Binary Vapor-liquid Equilibrium of Methyl Glycolate and Ethylene glycol and Separation Process for Hydrogenation of Dimethyl Oxalate to Ethylene Glycol. Thesis, East China University of Science and Technology, Shanghai, 2007. (24) Matsuda, H.; Takahara, H.; Fujino, S.; Constantinescu, D.; Kurihara, K.; Tochigi, K.; Ochi, K.; Gmehling, J., Fluid Phase Equilibria Selection of entrainers for the separation of the binary azeotropic system methanol + dimethyl carbonate by extractive distillation. Fluid Phase Equilib. 2011, 310, 166-181. (25) Ma, X.; Liu, X.; Li, Z.; Xu, G. Vapor–liquid equilibria for the ternary system methanol + dimethyl carbonate + dimethyl oxalate and constituent binary systems at different temperatures. Fluid Phase Equilib. 2004, 221, 51-56. (26) Berg, L. Recovery of ethylene glycol from butanediol isomers by azeotropic distillation. U.S. 22


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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

Patent, US4966658. 1990. (27) Wang, Z. Separation Process for the Product of Ethylene Glycol Synthesis from Syngas. Thesis, East China University of Science and Technology, Shanghai, 2011. (28) Li, Q. Process Separation of ethylene glycol from the polyols mixture. Thesis, Jilin University, Jilin, 2014. (29) Yang, Y., Huang, M., Ma, Xiaoxun. Isobaric vapor-liquid equilibrium(VLE) data of the binary system 1,2-Ethanediol+Ethylbenzene at 3.55kPa. J. Northwest University(China), 2017. (30) King, C. J. Separaton Processes. McGraw-Hill: New York, 1971. (31) Dhale A. D.; Myrant L K; Chopade S. P. Propylene glycol and ethylene glycol recovery from aqueous solution via reactive distillation. Chem. Eng. Sci. 2004, 59, 2881-2890. (32) Huang, W.; Li, H.; Wang, R.; Li, X.; Gao, X. Chemical Engineering & Processing : Process Intensification Application of the aldolization reaction in separating the mixture of ethylene glycol and 1,2-butanediol: Kinetics and reactive distillation. Chem. Eng. Process.: Process Intensification 2017, 120, 173-183. (33) Dhale, A. D.; Myrant, L. K.; Chopade, S. P.; Jackson, J. E.; Miller, D. J. Propylene glycol and ethylene glycol recovery from aqueous solution via reactive distillation. Chem. Eng. Sci. 2004, 59, 2881-2890. (34) Chopade, S. P.; Dhale, A. D.; Clark, A. M.; Kiesling, C. W.; Myrant, L. K.; Jackson, J. E.; Miller, D. J. Vapor-Liquid-Liquid Equilibrium ( VLLE ) and Vapor Pressure Data for the Systems 2-Methyl-1,3-dioxolane ( 2MD ) + Water and 2,4-Dimethyl-1,3-dioxolane ( 24DMD )+ Water. J. Chem. Eng. Data 2003, 46, 44-47. (35) Shubham P. Chopade; Atul D. Dhale; Angela M. Clark; Chris W. Kiesling; Laurie K. Myrant; James E. Jackson.; Dennis J. Miller. Vapor−Liquid−Liquid Equilibrium (VLLE) and Vapor Pressure Data for the Systems 2-Methyl-1,3-dioxolane (2MD) + Water and 2,4 -Dimethyl-1,3-dioxolane (24DMD) + Water. J. Chem. Eng. Data 2003, 1, 44-47. (36) Chien, I Lung, B. Y. Yu, and Z. J. Ai. Design of Azeotropic Distillation Systems. Chemical Engineering Process Simulation. 2017. (37) Luyben, W. L., Comparison of extractive distillation and pressure-swing distillation for acetone/chloroform separation. Comput. Chem. Eng. 2013, 50, 1-7. (38) Dye, R. F. Ethylene glycols technology. Korean J. Chem. Eng. 2001, 18, 571-579. (39) Li, X.; Wang, R.; Na, J.; Li, H.; Gao, X., Reversible Reaction-Assisted Intensification Process for Separating the Azeotropic Mixture of Ethanediol and 1,2-Butanediol: Reactants Screening. Ind. Eng. Chem. Res. 2018, 57, 710-717.

23



470.5

Page 24 of 32

liquid phase vapor phase

470.0 469.5 469.0 468.5 468.0 467.5 467.0 0.0

0.2

0.4

0.6

0.8

1.0

Mass fraction of 1,2-BD Figure 1. Binary phase diagram of the binary system {EG +1,2-BD}.


Page 25 of 32 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


Figure 2. Flowsheet and simulation results of CD route.



Page 26 of 32

1000

80

Rmin

70

Reboiler Duty 800

60 50

600

40 400

30 20

200 10 0 0.90

0.92

0.94

0.96

0.98

0

Recovery of Methanol Figure 3. Minimum reflux ratio and reboiler heat duty of Col-11 at different methanol recovery.


Page 27 of 32 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


1.00 0.98 0.96 0.94 0.92 0.90

Vapor phase Liquid phase

0.88 0.86

0

5

10

15

20

76

77

78

79

80

81

82

Stage number Figure 4. The concentration profile of EG in vapor and liquid phase of Col-14.



)LJXUH

)ORZVKHHW DQG VLPXODWLRQ UHVXOWV RI $' URXWH


Page 28 of 32

Page 29 of 32 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


12

70

NT,min Reboiler duty

60

50

10

8

40 6 30 4 20 2 10

0.2

0.4

0.6

0.8

1.0

EB/Feed Figure 6. Minimum stage number and reboiler heat duty of Col-27 at different ratio of EB to feed.



Figure 7. Flowsheet and simulation results of ED route.


Page 30 of 32

Page 31 of 32 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


Figure 8. Flowsheet and simulation results of RD route.



Table of Content Graph


Page 32 of 32

Reversible reaction-assisted intensification process for separating the

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