Thermodynamic Fundamentals and Energy Efficiency for the

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Separations

Thermodynamics fundamentals and energy efficiency for separation and high-valued utilization of light naphtha from Fischer-Tropsch synthesis Xin Gao, Yue Zhao, Wei Yuan, Suli Liu, Xingang Li, Hong Li, Sixiao Wang, and Xueshan Lu Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.9b01002 • Publication Date (Web): 06 May 2019 Downloaded from http://pubs.acs.org on May 13, 2019

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Industrial & Engineering Chemistry Research

Thermodynamics fundamentals and energy efficiency for separation and high-valued utilization of light naphtha from FischerTropsch synthesis Xin Gao†, Yue Zhao†, Wei Yuan§, Suli Liu§, Xingang Li†, Hong Li,†, Sixiao Wang†, Xueshan Lu† †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 §Shenhua

Ningxia Coal Industry Group Co. Ltd, Yinchuan 750011, China

ABSTRACT: In order to improve the economic potential of Fischer-Tropsch (F-T) products, the low-cost separation of linear α-olefins is a flexible technology. However, the scarcity of thermodynamic data and high production cost were the key R&D issues essential for the commercialization. In this paper, the isobaric vapor-liquid equilibria (VLE) data of five binary systems composed of 1-hexene, 2-hexene, 2-methylpentane, 1-octene, 2-octene were measured and regressed by the activity coefficient models. With the binary interaction parameters regressed by experimental data, two separation processes for obtained α-olefins fraction products such as C6 (1-hexene) and C8 (1-octene) were simulated comparing with a traditional sequence distillation route and a dividing wall column (DWC) route. A practicable



Corresponding author. Tel: +86-022-27404701 (H.L.); Fax: +86-022-27404705 (H.L.). E-mail: [email protected] (Hong Li).

ACS Paragon Plus Environment

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

configuration of heat-integrated distillation process was proposed to reduce furthermore the energy consumption for separation of olefins and alkanes. The energy consumption and total annual cost (TAC) of the whole separation technology were estimated and compared, which indicated that advanced separation technology could reduce the energy consumption and TAC of the production of α-olefins from F-T products. KEYWORDS: F-T synthesis, α-olefins, DWC, heat-integrated distillation, process design 1. Introduction Fischer-Tropsch (F-T) synthesis process

1, 2,

i.e. the indirect coal liquefaction, is an

important route to product clean liquid fuel from synthetic gas which can be derived from nonpetroleum material, such as coal. The Fischer-Tropsch products are usually a broad range of hydrocarbons and oxygenate, the content of sulfur, nitrogen and aromatics is very low3-5. Furthermore, the oxygenate product is very low and a large proportion of the hydrocarbons are 1-alkenes which are of high added value. Linear α-olefins, especially 1-hexene and 1-octene, are key components for the production of Linear Low Density Polyethylene (LLDPE) and the demand for 1-hexene and 1-octene increased enormously in recent years 6-9. However, the F-T products are produced in the first place for utilization as fuel, generally used as gasoline or diesel after hydrogenation. Apparently, this utilization method of F-T products is unreasonable which results in the waste of valuable product. Moreover, economic benefit would be increased remarkably if the high added value substances, especially α-olefins could be separated from FT products and used for fine chemical. However, the technology for separate α-olefins from other F-T products is immature. There is only one company around the world, Sasol, could accomplish this process so far 10, whose


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process included traditional distillation, azeotropic distillation, super distillation and extractive distillation sections. For the preliminary separation of F-T synthesis oil, which is an energy intensive process, the major problem is that the total energy requirement and the capital cost are very high. Therefore, further energy-saving efforts are needed to reduce the production costs. In this respect, fully thermally coupled dividing wall columns (DWCs), in which two columns are integrated into one shell, represent a very promising technology allowing a significant energy and capital cost requirement reduction

11-13.

DWC can also be used in

reactive distillation14, 15, extractive distillation16-19 and azeotropic distillation20, 21 processes. The improvement of the thermal efficiency due to avoiding unnecessary mixing effects, leads to considerable energy savings of about 30% compared with the direct or indirect sequences 22. Since only one column along with one reboiler and one condenser are used, the capital costs are also reduced. The energy-saving effect of DWCs is remarkable when the multicomponent separation process is performed and F-T oil separation process is exactly a typical multicomponent separation process. What’s more, in order to obtain α-olefins from the narrow fractions, i.e. the products of preliminary separation process, in which the boiling point of the components are really close to each other, different pressure thermally coupled distillation technology23, 24 (DPTC) could be used. In the DPTC process, a traditional distillation column is divided into two individual columns, one is operated at high pressure as rectifying section, while the other is operated at low pressure as stripping section. So that the reboiler at the bottom of the stripping section can be heated by the vapor obtained from the top of rectifying section. Meanwhile the stream obtained from the bottom of stripping section could serve as the condensing agent for the condenser at the top of the rectifying section. Therefore the energy of



distillation column could be internally coupled and little external energy is needed. Moreover, since the condenser and reboiler could be merged into one heat exchanger, the capital costs could be decreased. Nevertheless, the thermodynamic data, which is very important for process design, of the F-T system is deficient, which makes it difficult to study the F-T products and develop energysaving separation process for α-olefins. Therefore, it’s necessary to study the VLE of the F-T system, especially the VLE of key components. The property databases of simulation software have already contained several thermodynamic interaction parameters as well as the VLE experiment data for some binary systems in F-T synthesis oil. Considering that the composition of F-T synthesis oil was extremely complex, it’s infeasible to measure all the thermodynamic data. Thus we only measured the VLE data of some key components and took the existing thermodynamic data or adopted the idealized model for other components. In this paper, two F-T synthesis oil separation processes, including one traditional direct separation sequence (TDS) process and one diving wall column (DWC) process, as well as two narrow fraction separation processes, including one traditional column (TC) process and one different pressure thermally coupled distillation technology (DPTC) process, were designed and evaluated by the simulation software, Aspen Plus 8.4. Eventually, good purity olefins, i.e. 1-hexene and 1-octane could be obtained as high value product. The energy consumption and TAC of these two F-T oil separation processes were compared with each other, as well as the two narrow fraction separation process. To ensure the accuracy of process simulation, the VLE data of key components binary systems, i.e. 1-hexene + 2-hexene, 1-hexene + 2-methylpentane, 1-hexene + 3-methylpentane, 1-octene + 2-octene and 1+octane + 3-methylheptane, were


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measured by a double circulating vapor-liquid equilibrium glass still at 101.3 kPa. The main purpose to conduct those VLE experiments was for the narrow fraction separation processes so that the VLE data were concentrated on the C6 and C8 system. Van Ness test was employed to verify the thermodynamic consistency of the obtained VLE data. What’s more the VLE data were correlated with three activity coefficient models, namely Wilson, nonrandom two liquid (NRTL) and universal quasichemical (UNIQUAC) model. In addition, the binary VLE for 1hexene + n-hexane system was measured at 101.3 kPa to test the performance of the experimental apparatus. 2. Thermodynamic data acquisition 2.1 vapor-liquid equilibrium experiment. The specification of reagents were summarized in Table S1 in the Supporting Information. Those regents were detected by gas chromatography (PE, American) using a flame ionization detector (FID). All of the reagents were used without further purification. A double circulating vapor-liquid equilibrium glass still is used in the present equilibria study, which was used in previous study by our team

25-27.

The still and other apparatus are

shown in Figure 1. Both vapor and liquid are circulate in this still, which assures the vapor and liquid fully contact with each other, and then cuts down the equilibrium time evidently. The inner volume of the still is about 110cm3. The still was heated by a heat-up magnetic stirrer (SZCL-4B, Yu Hua Instrument, China). The apparatus was operated at atmospheric pressure which was measured by a barometer with an accuracy of ±0.2kPa (DYM3, Jiang Shan Instrument, China). For the temperature measurement, a digital thermometer (1552A-12-DL, Fluke, the United States) with an accuracy of ±0.01K was used.



Figure. 1 Scheme of the double circulating vapor-liquid equilibrium glass still 1- heat-up magnetic stirrer; 2- feed inlet; 3- boiling chamber; 4- condenser; 5- connecting port; 6- equilibrium chamber; 7- vacuum insulating layer; 8- digital thermometer; 9- thermocouple thermometer; 10- liquid-phase receiving section; 11- vapor-phase receiving section; 12sampling port; 13- discharge port To test the experimental apparatus’s accuracy, the binary vapor-liquid equilibrium data of the system of 1-hexene (1) + n-hexane (2) were measured at 101.3 kPa. The results are shown in Figure 2. The experimental data are in good agreement with those reported by Marrufo et al.28 and Suryanarayana et al.29, proving that the equilibrium still is reliable. The experiments were carried out by the following procedure. One of the pure components in the binary system was injected into the boiling chamber for cleaning until no impurities were detected in both condensed vapor and liquid samples. Then about 50ml pure component was injected into the boiling chamber and was heated up by the heat-up magnetic stirrer. Regulate the heating power to keep the vapor and liquid reflux rate at 2~3 drop/sec. When the temperature of the vapor-liquid interface was stable for more than 30 min, the vapor-liquid equilibrium was considered to be established. Record the temperature and pressure at this time.


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Vapor and liquid samples were taken out simultaneously from the sampling port for analysis. To change the composition in the equilibrium still, a certain amount of liquid was taken out from the boiling chamber using syringe and equal amount of another component in the binary system was injected to it. Repeat the above steps, then a series of VLE data could be obtained.

1.0

0.8

0.6

y1

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


0.4

0.2

0.0 0.0

0.2

0.4

0.6

0.8

1.0

x1

Figure 2. x–y phase diagram for the 1-hexene (1) + n-hexane (2) system at 101.33kPa ○, Marrufo et al.28; △, Suryanarayana et al.29; ×, experimental data The compositions of the condensed vapor and liquid samples were analyzed by gas chromatography (PE Auto System XL, PE, the United States), which was equipped with a flame ionization detector (FID) and a HP-5 capillary column (30m×0.32mm×0.25μm, Agilent). High-purity nitrogen (mass fraction 0.99999) was used as carrier gas with a rate of 1.0mL/min. The hydrogen and air flow rates were 45 and 450mL/min, respectively. Injector and detector temperatures were both 533.15 K for all systems. The column temperature for C6 and C8 systems were 318.15 K and 353.15 K, respectively. The injected volume of the samples was 0.2μL with a split ratio of 50:1. Good peak separation could be achieved under the conditions. The analysis was performed at least two times for each sample to ensure that the mass fraction



uncertainty was within ± 0.5%, and then the averaged data were used. The VLE data for the binary systems are shown in Tables S2-S6. The barometric pressure was recorded after each run and then the observed temperatures were corrected to 760 mmHg by eq.1 30: 𝑡𝑐 = 𝑡𝑜 +0.00012(𝑡𝑜 + 273.2)(760 ― 𝑃)

(1)

Where tc, to and P were corrected temperature, observed temperature and atmospheric pressure during experiment, respectively. The activity coefficients γi were calculated from eq.2 assuming non ideality of both liquid and vapor phase:

(

𝜑𝑖𝑦𝑖𝑝 = 𝑥𝑖𝛾𝑖𝜑𝑣𝑖𝑝𝑠𝑖𝑒𝑥𝑝

)

1 𝑝 𝑙 𝑅𝑇∫𝑐 𝑉𝑖𝑑𝑝

(2)

Where 𝑦𝑖 and 𝑥𝑖 were the vapor and liquid phase molar fraction of pure component i, 𝑝 and 𝑝𝑠𝑖 were total pressure of system and statured vapor pressure of pure liquid component i at system temperature T, 𝜑𝑖 and 𝜑𝑣𝑖 were the fugacity coefficient of component i in the system and in the pure state at system temperature, 𝛾𝑖 was the activity coefficient of the component i in liquid phase, 𝑉𝑙𝑖 was the molar volume of pure liquid component i, and R was the universal gas constant, respectively. The statured pressure 𝑝𝑠𝑖 could be calculated by the following extended Antoine equation eq.3: 𝐶2𝑖

ln 𝑝𝑠𝑖 = 𝐶1𝑖 + 𝑇 + 𝐶3𝑖 + 𝐶4𝑖𝑇 + 𝐶5𝑖ln 𝑇 + 𝐶6𝑖𝑇𝐶7𝑖

𝐶8𝑖 ≤ 𝑇 ≤ 𝐶9𝑖 (3)

Where C (1-7) i were the extend Antoine parameter of pure component i, C8i and C (1-7) i are the applicable temperature range. Those parameters, which were obtained from the Aspen property database, are shown in Table S7.


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Thermodynamic consistency test was performed to check the accuracy of experimental data, usually the Gibbs-Duhem equation 31 is employed. Herington consistency test

32, 33,

i.e. area

test which testifies thermodynamic consistency of the experimental data on the whole, is widely used. However that method cannot be applied to highly ideal systems (0.95 < γ < 1.10) 34, 35. As seen from Table S2-S6, the activity coefficients of the component were closed to 1.0 indicating these systems could not be evaluated using Herington method. The Van Ness test36 is used to test the reliability of each experimental point, which is expressed as eq. 4 and eq. 5: 1

1

𝑁

𝑁

∆𝑝 = 𝑁∑𝑖 = 1∆𝑝𝑖 = 𝑁∑𝑖 = 1100 1

1

𝑁

|

|

𝑝𝑒𝑥𝑝 ― 𝑝𝑐𝑎𝑙 𝑖 𝑖 𝑝𝑒𝑥𝑝 𝑖

(4)

𝑁

Δ𝑦 = 𝑁∑𝑖 = 1Δ𝑦𝑖 = 𝑁∑𝑖 = 1100|𝑦𝑒𝑥𝑝 ― 𝑦𝑐𝑎𝑙 𝑖 𝑖 |

(5)

Where N was the number of experimental data points; the superscript exp and cal indicated experimental data and values calculated by the thermodynamic model, respectively. If Δp and Δy are less than 1, the data set passes the thermodynamic consistency test. The test results calculated by the Wilson, NRTL and UNIQUAC models were reported in Table S8. All values of each system were less than 1, indicating that the experimental data were reliable. 2.3 Data regression. The experimental data were correlated with the Wilson, NRTL and UNIQUAC models, then the binary interaction parameters of these models were estimated. In doing so, the following objective function eq. 6 was used: 𝑁

[(

OF = ∑𝑖 = 1

𝑒𝑥𝑝 2 𝑇𝑐𝑎𝑙 𝑖 ― 𝑇𝑖

𝜎𝑇

) +(

𝑒𝑥𝑝 2 𝑝𝑐𝑎𝑙 𝑖 ― 𝑝𝑖

𝜎𝑝

) +(

𝑒𝑥𝑝 2 𝑥𝑐𝑎𝑙 𝑖 ― 𝑥𝑖

𝜎𝑥

) +(

𝑒𝑥𝑝 2 𝑦𝑐𝑎𝑙 𝑖 ― 𝑦𝑖

𝜎𝑦

)]

(6)

Where σ was the standard deviation. The absolute deviation of temperature and vapor phase molar fraction between the calculated and experimental values, i.e. ΔT and Δ𝑦𝑖 for all models were shown in Table S2-S6.



The regressed binary interaction parameters of each model as well as the root-mean-square deviations (RMSD) of the temperature and vapor molar fraction between the experimental data and the calculated values were shown in Table 1. The RMSD (T) and RMSD (yi) were defined as eq. 7 and eq. 8 respectively: RMSD(T) = RMSD(y𝑖) =

1 𝑁

∑(𝑇𝑐𝑎𝑙 ― 𝑇𝑒𝑥𝑝)2

1 𝑁

𝑒𝑥𝑝 2 ∑(𝑦𝑐𝑎𝑙 𝑖 ― 𝑦𝑖 )

(7) (8)

As shown in Table 1, the RMSDs of the equilibrium temperature were no more than 0.30K, and the RMSDs of the vapor mole fraction are no more than 0.004 except parameters correlated by UNIQUAC model. Figs. 3-7 presented T-x-y diagram of the experimental data and values calculated by Wilson, NRTL and UNIQUAC models, from which we could see that the differences among the data calculated by the three activity coefficient models and the experimental data were slight. The results presented that the obtained binary VLE data can be well correlated by the three activity coefficient models, especially the Wilson model, indicating that the obtained thermodynamic data was reliable and could be introduced to the Aspen Plus property database. The deviation of the pure component may due to that the reagents were a little bit impure. What’s more, the boiling point difference between the components in binary system were pretty small causing the error seemed to be amplified. In fact, the root-meansquare deviations and the absolute deviations were pretty low. Specific absolute deviation data for Figure 3-7 could be seen in Table S2-S6, respectively.


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Table 1. Regressed binary interaction parameters of Wilson, NRTL and UNIQUAC models for the systems as well as RMSD between experimental data and values calculated by those models aAij

model

aAji

aBij

aBji

/K

/K

α

RMSD (T)

RMSD (yi)

——

0.16

0.0029

0.24

0.0036

0.17

0.0061

1-hexene(1)+2-hexene(2) Wilson

3.4533

NRTL

464.011

5.1217

-980.3391

-1998.4044

2924.0691

464.0112

-3573.8622

0

-251.9615

182.3990

0.3

2 UNIQUAC

0

——

2-methlypentane (1) + 1-hexene (2) Wilson

4.7843

4.0365

-1876.4208

-1174.4840

——

0.13

0.0034

NRTL

-3.5857

-0.7971

1208.5643

269.3622

0.3

0.05

0.0027

UNIQUAC

0

0

-24.9962

18.6284

——

0.13

0.0058

3-methlypentane (1) + 1-hexene (2) Wilson

0

0

-76.4314

49.2933

——

0.17

0.0024

NRTL

0

0

-88.4985

119.3108

0.3

0.19

0.0024

UNIQUAC

0

0

37.8352

-46.2190

——

0.19

0.0024

1-octene (1) + 2-octene (2) Wilson

-11.4582

-11.1410

4852.2019

3869.6377

——

0.10

0.0033

NRTL

10.2837

8.7835

-3564.8272

-3843.5564

0.3

0.10

0.0038

UNIQUAC

0

0

155.1835

-194.7469

——

0.11

0.0086

3-methlyheptane (1) + 1-octene (2) Wilson

-1.3668

7.8820

233.5713

-2904.8316

——

0.26

0.0028

NRTL

-0.5427

38.8345

240.8815

-4570.1265

0.3

0.25

0.0032

UNIQUAC

0

0

74.2793

-89.7225

——

0.26

0.0038

aThe

𝐵𝑖𝑗

interaction parameters of various models: for Wilson, ln Λ𝑖𝑗 = 𝐴𝑖𝑗 +

𝑇 ; for UNIQUAC, 𝜏𝑖𝑗 = exp (𝐴𝑖𝑗 +

𝐵𝑖𝑗

𝑇) .


𝐵𝑖𝑗

𝑇 ; for NRTL, 𝜏𝑖𝑗 = 𝐴𝑖𝑗 +


70 69 68

T/℃

67 66 65 64 63 62 0.0

0.2

0.4

0.6

0.8

1.0

x1, y1

Figure 3. T-x–y phase diagram for the 1-hexene (1) + 2-hexene (2) system at 101.33kPa: □,△, (𝑥1, 𝑦1) experimental data in this work; —, (𝑥1, 𝑦1) correlated results by the Wilson model; - -, (𝑥1, 𝑦1) correlated results by the NRTL model; -·-·-, (𝑥1, 𝑦1) correlated results by the UNIQUAC model. 64 63 62

T/℃

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

61 60 59 0.0

0.2

0.4

0.6

0.8

1.0

x1, y1

Figure 4. T-x–y phase diagram for the 2-methlypentane (1) + 1-hexene (2) system at 101.33kPa: □, △ , (𝑥1, 𝑦1) experimental data in this work; —, (𝑥1, 𝑦1) correlated results by the Wilson model; - - -, (𝑥1, 𝑦1) correlated results by the NRTL model; -·-·-, (𝑥1, 𝑦1) correlated results by the UNIQUAC model.


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64.0

T/℃

63.5

63.0

62.5

62.0 0.0

0.2

0.4

0.6

0.8

1.0

x1, y1

Figure 5. T-x–y phase diagram for the 3-methlypentane (1) + 1-hexene (2) system at 101.33kPa: □, △ , (𝑥1, 𝑦1) experimental data in this work; —, (𝑥1, 𝑦1) correlated results by the Wilson model; - - -, (𝑥1, 𝑦1) correlated results by the NRTL model; -·-·-, (𝑥1, 𝑦1) correlated results by the UNIQUAC model.

127 126 125 124

T/℃

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


123 122 121 120 0.0

0.2

0.4

0.6

0.8

1.0

x1, y1

Figure 6. T-x–y phase diagram for the 1-octene (1) + 2-octene (2) system at 101.33kPa: □,△, (𝑥1, 𝑦1) experimental data in this work; —, (𝑥1, 𝑦1) correlated results by the Wilson model; - -, (𝑥1, 𝑦1) correlated results by the NRTL model; -·-·-, (𝑥1, 𝑦1) correlated results by the UNIQUAC model.



123 122 121

T/℃

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

120 119 118 117 0.0

0.2

0.4

0.6

0.8

1.0

x1, y1

Figure 7. T-x–y phase diagram for the 3-methlyheptane (1) + 1-octene (2) system at 101.33kPa: □,△, (𝑥1, 𝑦1) experimental data in this work; —, (𝑥1, 𝑦1) correlated results by the Wilson model; - - -, (𝑥1, 𝑦1) correlated results by the NRTL model; -·-·-, (𝑥1, 𝑦1) correlated results by the UNIQUAC model. 3. Process design and simulation In this section, the simulation of F-T synthesis oil separation processes and even carbon number narrow fractions (i.e. C6 and C8) separation processes were rigorously simulated using Aspen Plus 8.4. Two processes for F-T synthesis oil separation were studied, including one TDS process and one DWC process. Moreover two processes for α-olefins were studied, including one TC process and one DPTC process. Eventually, we could obtain the high purity α-olefins, i.e. 1-hexene and 1-octene. The thermodynamic data obtained in previous section were introduced to the aspen property database and the existing Wilson interaction parameters of some other components involved in this paper were shown in Table S9.All processes were simulated with the Wilson property method. 3.1 F-T synthesis oil separation process simulation. The F-T synthesis oil feedstock included hydrocarbons of C5-C10, mainly n-hydrocarbons, i.e. n-alkanes and 1-olefins, as well


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as a small amount of internal olefins and iso-hydrocarbons. The specific composition of the F-T synthesis oil was shown in Table S10. The crude products of F-T synthesis oil preliminary separation process were C6, C7, C8 and C9 narrow fraction. The feed temperature and pressure were 40 ℃ and 2 bar respectively with a feed rate at 46,975 kg/h. The overhead temperature of each column, which can be changed by adjusting the column operating pressure, should be no less than 40 ℃, in order to ensure that the ascending vapor could be condensed with normal cooling water. The design specifications were that the purity and yield of products were greater than 90 wt.％ and 80 wt.％, respectively. The product means that the hydrocarbons with the same carbon number in the corresponding narrow fraction .The flowsheet of F-T synthesis oil separation process were shown in Figure 8.

(a) Traditional direct separation sequence process



(b) Dividing wall column process Figure 8. Scheme and simulation results of the F-T oil separation process For the TDS process (Figure 8 a), the overhead outcome of distillation column Col-11, i.e. C5-C6 systems, was feed to distillation column Col-12 and the bottom outcome, i.e. C7-C10 systems, was feed to distillation column Col-13. Later the C6 system product was obtained from the bottom of distillation column Col-12, while C7, C8 and C9 system products were obtained from the top of distillation column Col-13, Col-14 and Col-15, respectively. Column Col-12 was operated at 1.34 bar while other columns were operated at atmosphere. These distillation columns were optimized individually by adopting the appropriate stage number, reflux ratio and feed position. The separation process in this paper is commonly used in industry. Since the boiling point of C5 components are pretty low (e.g. n-pentane 36.07℃, 1pentene 30.07 ℃ ), the overhead temperature of the Col-11 would below 40 ℃ if only C5 system was took out from the overhead. Then in order to ensure that the ascending vapor could be condensed by normal cooling water, Col-11 should be operated at pressurized condition


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which would reduce the relative volatility of substances and increase the heat duty. Therefore we took out both C5 and C6 from the overhead so that Col-11 could be operated under atmosphere condition. Even though the Col-12 needed be operated at pressurized condition in this case, the operation was relatively easy due to its much smaller size. What’s more, the relative volatility between n-pentane-1-hexene are larger than n-hexane-1-heptene and the percentage decrease during pressurized operation is less. If indirect sequence was adopted, the reboiler duty and column diameter of Col-11 would increase significantly (the reboiler duty of Col-11 could reach up to 10 Gcal/h while the total reboiler duty of the process used in the paper was only 10.9739 Gcal/h). For the sake of economic efficiency and operational simplicity, we adopted this TDS process. For the DWC process (Figure 8 b), the C5-C6 systems, C7-C9 systems and C10 system were obtained from the top, side and bottom of Col-21, respectively. The C5-C6 systems were feed to distillation column COL-22 and then the C6 system product was obtained from the bottom of Col-22. Meanwhile the C7-C9 systems were feed to Col-23, the C7, C8 and C9 system products were obtained from the top, side and bottom of Col-23, respectively. This DWC process was similar to the TDS one, but replace those four traditional distillation column, i.e. column Col-11, Col-12, Col-13 and Col-14, with two dividing wall column, i.e. Col-21 and Col-23. We adopted this DWC process since it’s convenient to revamp on the basic of industrial process. We may continue to study other DWC processes systematically in the future. The theoretical stage number (NS), reboiler duty (QR), condenser duty (QC), product purity and yield of these process were summarized in Table S11. As we can see, the total reboiler duty of the DWC process could be reduced compared with the TDS process, the percentage



reduction was 11.7% and the absolute reduction was 1.2879 Gcal/h. Moreover the total theoretical stage number of both DWC processes were decreased significantly which could be cut down 25.4% (i.e. 45 theoretical trays) compared with the TDS process. Accordingly, the diving wall column is a very promising technology allowing a significant energy and capital cost requirement reduction. 3.2 Narrow fractions separation process simulation. The C6 and C8 fraction feedstock were obtained from the F-T synthesis oil separation process, then the C5 component in C6 fraction and C7 component in C8 fraction were removed respectively. The composition of the C6 and C8 fraction feedstock were shown in Table S12 and Table S13, respectively. The design specifications were that the purity and yield of products were both greater than 95 wt.％. The product means 1-hexene and 1-octene respectively in C6 and C8 fraction. The flowsheet of C6 and C8 fraction separation process were shown in Figure 9 and Figure 10, respectively. For the DPTC processes, the high pressure columns were operated under atmosphere condition, in order for these processes to be thermally coupled, the temperature difference between the top of high pressure column and the bottom of low pressure column was controlled to be greater than 10℃ by adjusting the operating pressure of the low pressure column. For the C6 fraction TC process shown in Figure 9 a (or C8 fraction TC process shown in Figure 10 a), the 1-hexene (or 1-octene) product and other C6 (or C8) component, mainly nhexane (or n-octane), were obtained from the top and bottom of Col-31(or Col-41), respectively. For the C6 fraction DPTC process shown in Figure 9 b, the rectifying and stripping section of Col-31 were separated into Col-32 and Col-33 respectively. The Col-32 was operated under


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atmosphere condition while Col-33 was operated under decompression condition, so that the condenser of Col-32 could perform heat transfer with the reboiler of Col-33. The bottom outcome of Col-32 was feed into the top of Col-33, meanwhile the overhead outcome of Col33 was feed into the bottom of Col-32 after pressurized by the compressor Com-31. Since the condenser duty of Col-32 was larger than the reboiler duty of Col-33, the reboiler of Col-33 could be offset, thus only an auxiliary condenser in Col-32 was needed. The C8 fraction DPTC process shown in Figure 10 b worked in the same way.

(a) Traditional column process

(b) Different pressure thermally coupled distillation technology process Figure 9. Scheme and simulation results of the C6 fraction separation process



(a) Traditional column process

(b) Different pressure thermally coupled distillation technology process Figure 10. Scheme and simulation results of the C8 fraction separation process The theoretical stage number (NS), reboiler duty (QR), condenser duty (QC), product purity and yield of these process were summarized in Table S14. In both DPTC processes no extra reboiler duty was needed which means that the DPTC processes could save 4.3090 Gcal/h and 5.7104 Gcal/h reboiler duty compared with corresponding traditional ones. Accordingly, the different pressure thermally coupled distillation is a very promising technology allowing a significant energy cost requirement reduction, which could be used in α-olefin purification process.


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4. Economic estimation In this section, the total annual costs (TAC) of the three separation processes were estimation and compared with each other. The TAC including the capital investigation and operational cost could be evaluated by eq. 9. 𝑐𝑎𝑝𝑖𝑡𝑎𝑙 𝑐𝑜𝑠𝑡

TAC = 𝑜𝑝𝑒𝑟𝑎𝑡𝑖𝑜𝑛𝑎𝑙 𝑐𝑜𝑠𝑡 + 𝑝𝑎𝑦𝑏𝑎𝑐𝑘 𝑝𝑒𝑟𝑖𝑜𝑑

(9)

The operational cost mainly included electricity cost and energy cost such as heating steam and condensing water, while the capital investigation included columns, condensers and reboilers cost, the purchase costs of pumps and other equipment were too small to be taken into account. The payback period was usually set at 3 years in industry. The energy cost was calculated on the basis of the heat duty of heat exchangers37. For reboilers the price of high-pressure (41 barg, 251 °C) steam was 9.87 $/GJ while the prices of moderate-pressure (10 barg, 184 °C) and low-pressure (5 barg, 160 °C) steam were 8.22 and 7.78 $/GJ, respectively25. In these separation processes, the condensers were all cooled by condensing water, whose price was 0.354 $/GJ25. Moreover the price of electricity was 10.99 $/GJ39. Stainless steel was speciﬁed as the material in column, condensers, reboilers and internals. The cost of condensers and reboilers was directly estimated by the Aspen Exchanger Design and Rating module. The cost of column was calculated by eq. 1025 on the basic of the height of the column L and the diameter of the column D, which were estimated by the Tray Sizing module in Aspen Plus. For the DWC, its column diameter were estimated by combining the diameter of the main column and side stripper column following eq. 1138. The cost of compressor was calculated by eq. 1239, where 𝑊𝑐𝑜𝑚 was the work required by the compressor.



The size of the distillation columns adopted in this paper was normal so that the parameters involved in eq.10-12 were considered valid in this case. 𝑐𝑜𝑠𝑡𝑐𝑜𝑙𝑢𝑚𝑛

𝑑

$ = 17640( 𝜋

1.066

0.802 𝑙 ( 𝑚)

𝑚)

𝜋

𝑑𝐷𝑊𝐶 = 2 [4𝑑2𝑚𝑎𝑖𝑛 + 4𝑑2𝑠𝑖𝑑𝑒] 𝜋 𝑐𝑜𝑠𝑡𝑐𝑜𝑚𝑝𝑟𝑒𝑠𝑠𝑜𝑟

(11) 0.67

𝑊𝑐𝑜𝑚

$ = 91562(

𝑘𝑊 445

(10)

)

(12)

Details for columns and heat exchangers price of those processes were shown in Table S15. 4.1 F-T synthesis oil separation processes estimation. The economic estimation results of the two F-T synthesis oil separation processes were shown in Table 2. Since the TDS process consisted of 5 columns while there’re only 3 columns in DWC process, the column price could be reduced in the TDS process. The column cost of the DWC process was 7.9% (i.e. 1.45×105 $) lower than the TDS process as expected. However the heat exchanger cost of the DWC process was a little higher than the TDS process result in that the capital cost reduction was only 3.8% (i.e. 7.97×104 $) on the whole. As for the energy cost, the DWC process could cut down about 11.5% (i.e. 3.58×105 $/a) compared with the TDS process since the total reboiler duty of the former process was lower which brought a considerable reduction in heat steam consumption. Moreover the energy cost accounted for a large proportion of TAC in both process, the percentages were 81.5% and 80.2%, respectively. Consequently, the TAC of the DWC process was decreased by around 10.1% (i.e. 3.86×105 $/a) compared with the TDS process.


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Table 2. Economic estimation of the F-T oil separation processes price

TDS

DWC

Column/ 103 $

1845.45

1699.35

heat exchanger/ 103 $

269.81

336.20

energy cost/ (103 $/a)

3103.44

2745.33

TAC/ (103 $/a)

3809.53

3423.85

4.2 Narrow fractions separation processes estimation. The economic estimation results of the C6 and C8 fraction separation processes were shown in Table 3. As we can see, for the DPTC processes since the reboiler duty of Col-33 (or Col-43) could be completely offset by the condenser duty of Col-32 (or Col-42), no extra heat steam was required so that only a little bit condensing water consumption was needed. Therefore there’re almost no energy consumption cost in these DPTC processes while the energy consumption cost accounted for 61.5% and 63.3% respectively in corresponding TC processes. Nevertheless, the column costs of the DPTC processes was a little bit higher since the column diameter of the low pressure columns, i.e. Col-33 and Col-43, was enlarged slightly by decompression operation. Moreover owing to a compressor was added to both DPTC processes, corresponding capital cost and electricity cost were required while they were very small compared with energy costs required by the TC processes. Consequently, the TAC have been reduced by about 45.6 % (i.e. 8.70×105 $/a) and 45.0 % (i.e. 1.11×106 $/a) respectively for these DPTC processes. Accordingly the different pressure thermally coupled distillation technology could create great economic benefit during narrow fraction separation process.



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Table 3. Economic estimation of the narrow fraction separation processes price

C6 TC

C6 DPTC

C8 TC

C8 DPTC

column/ 103 $

2102.68

2413.49

2572.67

3022.65

heat exchanger/ 103 $

101.63

103.54

134.51

195.18

compressor / 103 $

0

109.70

0

126.72

energy cost/ (103 $/a)

1173.21

6.04

1555.46

8.14

Electricity cost/ (103 $/a)

0

184.38

0

236.82

TAC/ (103 $/a)

1907.98

1038.41

2457.85

1351.68

5. Conclusion The isobaric vapor-liquid equilibria (VLE) data of 1-hexene + 2-hexene, 1-hexene + 2methylpentane, 1-hexene + 3-methylpentane, 1-octene + 2-octene and 1+octane + 3methylheptane binary systems were measured and correlated. The results illustrated that all of the activity models, especially Wilson model, agreed well with the experimental data,indicating that the obtained thermodynamic data was reliable and could be introduced to the Aspen Plus property database. By doing so, the simulation accuracy of the F-T oil separation processes and narrow fractions separation processes could be improved. Then the rigorous simulations of these separation processes were conducted by Aspen Plus 8.4 .The results presented that when purity and yield of the narrow fraction products were similar, the DWC process could reduce about 11.7 % energy consumption and 10.1% TAC compared with the TDS one. What’s more, when purity and yield of the α-olefin products were the same, the DPTC processes had almost no energy consumption and the TAC could be cut down about 45 % compared with the corresponding TC ones. Accordingly, diving wall column and different pressure thermally coupled distillation are very promising technologies in F-T oil separation and narrow fraction separation process allowing a significant energy and capital cost requirement reduction.


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ASSOCIATED CONTENT Supporting Information Tables of vapor-liquid equilibrium experimental data, some other thermodynamic data involved in this work, the feedstock composition of each simulation process, simulation results of each process as well as details of capital cost. AUTHOR INFORMATION Corresponding Author *E-mail: [email protected]. Tel: +86-022-27404701 (H.L.) ORCID Hong Li: 0000-0002-7954-6107 Notes The authors declare no competing ﬁnancial interest. ACKNOWLEDGEMENTS The authors are grateful for the financial support from the National Key R&D Program of China (2018YFB0604903), and the Key Research and Development Program of Ningxia (2018BDE02057). REFERENCES (1) Fischer, F.; Tropsch, H., The preparation of synthetic oil mixtures (synthol) from carbon monoxide and hydrogen. Brennstoff-Chem 1923, 4, 276-285. (2) Leckel, D., Diesel Production from Fischer−Tropsch: The Past, the Present, and New Concepts. Energy & Fuels 2009, 23, (5), 2342-2358. (3) Egaña, A.; Sanz, O.; Merino, D.; Moriones, X.; Montes, M., Fischer–Tropsch Synthesis Intensification in Foam Structures. Industrial & Engineering Chemistry Research 2018, 57, (31), 1018710197. (4) Nakhaei Pour, A.; Housaindokht, M. R.; Kamali Shahri, S. M., Fischer–Tropsch Synthesis over Cobalt/CNTs Catalysts: Functionalized Support and Catalyst Preparation Effect on Activity and Kinetic Parameters. Industrial & Engineering Chemistry Research 2018, 57, (41), 13639-13649. (5) Muleja, A. A.; Yao, Y.; Glasser, D.; Hildebrandt, D., Variation of the Short-Chain Paraffin and



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Table of Contents (TOC) Graphic


Thermodynamic Fundamentals and Energy Efficiency for the

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