New Gasoline AbsorptionâStabilization Process for Separation

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Process Systems Engineering

A new gasoline absorption-stabilization process for separation intensification and flowsheet simplification in refineries Jing Y. Chen, Ming Pan, Chang He, Bing J. Zhang, and Qinglin Chen Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b04477 • Publication Date (Web): 21 Sep 2018 Downloaded from http://pubs.acs.org on September 30, 2018

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Industrial & Engineering Chemistry Research

A new gasoline absorption-stabilization process for separation intensification and flowsheet simplification in refineries

Jing Y. Chen, Ming Pan, Chang He, Bing J. Zhang*, Qing L. Chen* School of Chemical Engineering and Technology, Guangdong Engineering Centre for Petrochemical Energy Conservation, Sun Yat-sen University, Zhuhai, 510275, China

Graphical Abstract

Abstract A new gasoline absorption-stabilization process (GASP) is presented to intensify the separation process and significantly simplify the conventional GASP flowsheet. The conventional GASP includes four columns and recycling streams between refinery units, 1

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makes the process very complex. In this study, the solubility and volatility of the C5-C11 hydrocarbons are discussed and a new indicator is presented to represent the performance of supplementary absorbent streams. Then the best suitable components are found. A new flash tank is introduced into the GASP to obtain the suitable absorbent components to intensify the absorption process. As a result, the original secondary absorption column and the large recycling streams between the units are cancelled. The conventional and new GASPs are simulated and optimized respectively. Heat exchanger networks are designed for the two processes. The new process can significantly reduce the total annual cost by 11.24% as to the conventional process.

Keywords: Absorption-stabilization process; Optimization; Heat exchanger network; Process intensification.

Highlights •

The most suitable supplementary absorbent is found for GASPs.

•

A new process is presented to intensify and simplify the conventional GASP.

•

The new process reduces columns from 4 to 3 and erases external recycling streams.

•

The new process reduces the total annual cost by 11.24%.

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1. Introduction Gasoline absorption-stabilization processes (GASPs) are important parts of crude oil refining and are widely applied. In 2016, China had 140 million tons of gasoline throughput.1 All the crude oil refining plants with product gasoline have to contain GASPs to separate out the gasoline from other refinery products. GASPs separate the mixture of crude gasoline and rich gas streams coming from the main fractionators in fluid catalytic cracking units (FCCUs) or delayed coking units (DCUs), and produce dry gas, liquefied petroleum gas (LPG) and stabilized gasoline. A conventional GASP includes four columns and two recycling streams. This makes the flowsheet very complex and results in a higher total annual cost (TAC). The question of how to intensify and simplify such a complex process is interesting and challenging. Many studies have concentrated on process intensification and simplification for better economic, environmental and energy performance.2 Process intensification in chemical engineering can be divided into four categories from macro- to micro-scale, including process synthesis3, unit operations4, new devices5 and new solvents6. Several publications have reviewed these topics.7,8 On the process synthesis and integration level, Babi et al. developed systematic, multi-stage framework for process synthesis-intensification that identifies more sustainable process designs.9 Baldea established a connection between process integration and process intensification, and proposed a novel avenue for discovering intensification opportunities at the process design stage.10 Chen and Grossmann reviewed recent developments in process synthesis and discussed some of the major challenges in theory and 3



practice in the area.11 In this study, we focus on the levels of process synthesis and unit operations for process intensification and simplification to obtain the minimum TAC. A number of methodologies have been investigated to design intensified and simplified chemical processes. Gopalakrishnan et al. presented two process simplification procedures based on “maximum extendible zones” and “graphical and mathematical programming techniques”.12 Lutze et al. presented a phenomena-based synthesis/design methodology incorporating process intensification, to reveal improved process options and break the barrier of just focusing on classical unit operations and mainly consider retrofitting.13 Holtbruegge et presented a systematic conceptual process design approach for reaction-separation processes applying process intensification and used it for an ethyl lactate process, from the insight of physicochemical and thermodynamic properties.14 Živković and Nikačević proposed a novel method connecting process intensification principles and process systems engineering techniques. The method includes three steps: reaction screening, reaction system and mathematical modelling, and optimization.15 Bertran et al. presented a systematic framework for novel and sustainable synthesis-design of processing routes along with the associated computer-aided methods and tools.16 Tula et al. presented an integrated computer-aided software-tool that searches the design space for hybrid/intensified process options, and embedded process synthesis and intensification methods that operate at multiple scales.17 Kuhlmann et al. presented an approach providing maximum flexibility for and modifying the phenomena building blocks involved.18 Demirel et al. used a building-block-based superstructure to incorporate many process intensification pathways and 4


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designed a process by optimizing a performance metric for given raw materials and product specifications, material properties and bounds on flow rates.19 Anantasarn et al. proposed a systematic, 3-stage synthesis-intensification framework and applied it to achieve a more sustainable design.20 Kuhlmann et al. further presented the extension of a phenomena-based process synthesis method by an additional building block for reactor network synthesis and reactive separations. The method facilitates the automatic generation of thermodynamically feasible phenomena-based flowsheet variants by means of superstructure optimization.21 Distillation is the most important separation method in the process industry and is energy-intensive. Its intensification is significantly investigated for better process Kiss and Olujić reviewed and proposed the design optimization, process intensification and operation issues for internally heat-integrated distillation columns.22 Gutiérrez-Antonio et al. proposed the intensification of the hydrotreating production process through the use of thermally coupled distillation for the purification stage and incorporated a turbine in order to generate electricity with the energy contained in the outlet stream of the reactor for energy savings.23 Anantasarn et al. proposed an approach to generate more sustainable intensified process designs for the production of important chemicals and used it on a production process for para-xylene.24 Wang et al. proposed a new sulfolane aromatic extractive distillation to separate aromatic and non-aromatic hydrocarbons and optimized it for better energy utilization.25 Contreras-Zarazúa et al. presented the simultaneous design and optimization of three alternative azeotropic distillation processes to purify furfural by the mathematical technique known as Differential Evolution with Tabu List (DETL). The results of the 5



simultaneous optimization show that the intensification processes DHI and DWC showed an important reduction in the cost and environmental impact, with respect to the conventional process.26 Meirelles et al. presented a general approach for calculating the number of ideal stages and the corresponding concentrations in absorption and stripping with parallel streams. The approach can be used for the optimization and intensification of complex chemical processes involving absorption or stripping steps.27 Columns and entrainers are further combined to intensify and optimize separation systems. Han et al. proposed to intensify the separation of benzene and trace thiophene through optimum selection of a suitable entrainer and improvement of the process. The improved process decreases energy consumption and equipment investment compared with the conventional six‐column process.28 Iyer et al. proposed a hybrid stochastic-deterministic optimization approach combined with a genetic algorithm to simultaneously find solvent molecular structures and the operating conditions of coupled absorption-desorption process.29 Some separation processes in petroleum refining systems have been studied for intensification and simplification. Hsu et al. found the C5 separation process difficult to due to a reactor being coupled with one of the separation columns, and thus found a wide of opportunities for process simplification and intensification.30,31 Liu et al. introduced a new flash tank into a toluene disproportionation process and reduced hot utility consumption and TAC according to heat integration and economic evaluation comparisons.32 Al-Shatri et al. discussed the control and optimization aspects of multivariable distillation processes in petroleum refineries for the separation of aromatic mixtures.33 Niu et al. developed a 6


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heuristic-based systematic methodology for process retrofitting via intensification/integration, and applied the methodology to retrofitting the conventional isopropyl alcohol (IPA) Li et al. proposed a process and built a hydrate formation reactor model for ethylene recovery from refinery dry gas. The process includes hydrate formation, absorption and cryogenic distillation.35 Process intensification and optimization can be combined to obtain better performance, including local and global process enhancement. Portha et al. introduced the concepts of local and global process intensification. They used several examples to illustrate that the classical approach of process intensification based on single-unit improvement (local intensification) presents several limitations when compared to holistic overall process-based intensification (global intensification).36 Onel et al. presented a multi-scale framework for the intensification of small scale gas-to-liquid (GTL) processes through CFD modelling, process synthesis, and global optimization.37 Carrasco and Lima introduced a framework for process design and intensification of modular systems. The framework combines non-linear programming concepts and process simulator Aspen Plus for obtaining the relationship between input and output spaces.38 Gong et al. developed a simulation-based process intensification method for the design and optimization of shale gas processing and NGLs recovery process systems uncertain feedstock compositions. The intensified design of the novel system shows a lower TAC than that of the conventional system.39 Pistikopoulos proposed multi-parametric optimization in process intensification and introduced a systematic framework and prototype software system for the representation, modelling and solution of integrated design, operation 7



and advanced control problems.40 To the best of our knowledge, there are no publications which address intensification and simplification of GASPs for better cost-efficient performance from the view of column reduction and supplementary absorbent improvement. This study aims to develop a new GASP to simplify the complex flowsheet and intensify the separation process. First, we analyzed the influence of absorbent stream compositions (C5-C11 hydrocarbons) on the solubility and volatility, and then obtained the most suitable components for the supplementary absorbent stream to intensify the separation process. Second, a new flash tank was introduced into the GASP to generate the new absorbent stream. As a result, the original secondary absorption column and recycling material stream were cancelled and a new GASP was drawn out. Finally, parameter optimization, heat integration and economic analysis were investigated and compared for the new and conventional GASPs.

2. Conventional process description and analysis 2.1 Conventional GASP description Fig. 1 shows the conventional flowsheet of the GASP. The condensed rich gas from units is first mixed with the stripper gas from the stripper and the rich oil from the primary absorber. The mixture is separated into gas and liquid phase streams through a condensed oil tank after being cooled to 40°C. The gas stream is fed into the bottom of the primary absorber and the liquid stream into the top of the stripper. The crude gasoline enters the top of the 8


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primary absorber as the main absorbent stream. A portion of the stabilized gasoline from the bottom of the stabilizer is fed into the top of the primary absorber as the supplementary absorbent stream. The lean gas from the primary absorber is fed into the bottom of the secondary absorber to recover the gasoline components using the lean absorption oil. The lean absorption oil is the light diesel product from the main fractionators of the upstream units, which is then recycled to the upstream units after use. The liquid stream from the bottom of stripper enters the stabilizer and separates into the LPG and the stabilized gasoline products.

Figure 1. Flowsheet of a conventional GASP.

2.2 Process flowsheet analysis The conventional GASP has three feed streams: the rich gas, crude gasoline and lean 9



absorption oil; and has four product streams: the dry gas, LPG, stabilized gasoline and rich absorption oil. Two recycling material streams exist in the conventional GASP. One is the supplementary absorbent stream, which is used to absorb C3 and C4 hydrocarbons in the primary absorber. The other is the lean absorption oil, which comes from the main in the upstream units, enters the secondary absorber as an absorbent stream, and is recycled back to the upstream units after use. It is worth mentioning that the lean absorption oil and crude gasoline are separated at the main fractionators in the upstream units. The lean absorption oil enters the secondary absorber to absorb the light gasoline components escaping from the crude gasoline and becomes a rich absorption oil. The rich absorption oil is then sent back to the main fractionators to separate again. This repeated “separate-mix-separate” process of diesel and gasoline components results in lots of unnecessary exergy loss. Hence, reducing the recycling lean absorption oil can not only simplify the complex flowsheet of GASP, but also decrease the separation load and improve energy performance in the upstream units. In the primary absorber, the supplementary absorbent stream and the crude gasoline are used to absorb the LPG components in the rich gas. Apart from the temperature and pressure of the primary absorber, the composition and flow rate of the supplementary absorbent stream are the most important factors affecting the absorption performance. A suitable absorbent stream should not only have high solubility to absorb the solutes, but also low volatility to avoid being entrained by the gas. Hence, finding suitable components for the supplementary absorbent stream is important to intensify and simplify the absorption process. 10


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2.3 Supplementary absorbent and energy analysis The properties of the supplementary absorbent significantly affect the performance of the separation system. In the conventional GASP, the components in the supplementary absorbent include C5-C11 hydrocarbons. The solubility of C5-C11 hydrocarbons to C3-C4 hydrocarbons and the volatility of C5-C11 hydrocarbons are discussed below. We assume the mixture of C3-C4 hydrocarbons (Pseudo-component A) and the supplementary absorbent (Pseudo-component B) to be a binary system. The phase equilibrium curves are drawn in Fig. 2. The green line x = y means that the compositions in the liquid and vapour phases are exactly identical. In the primary absorber, Pseudo-component B in the liquid phase should solve Pseudo-component A in the vapour phase as much as possible. With the same mole fraction of Pseudo-component A in the vapour phase, the higher mole fraction of Pseudo-component A in the liquid phase suggests that Pseudo-component B has a better solubility. Hence the maximum distance dmax between the curve and the line x = y represents the solubility of Pseudo-component B into Pseudo-component A. The larger dmax, the smaller the solubility. Low volatility can reduce the loss of the supplementary absorbent caused by the entrainment of the gas flow, thus decreasing the load of the second absorber. We simulated and obtained dmax and the equilibrium vapour pressure for each component from C5-C11 hydrocarbons. The average dmax of the components with the same carbon number is marked as d, and the average equilibrium vapour pressure of the components with the same carbon number is marked as p. The results are shown in Fig. 3. As shown in Fig. 3, with the carbon number increasing, the volatility decreases, while the 11



solubility also decreases. However, the equilibrium vapour pressure drops rapidly from components C5 to C7, while the maximum vertical distance increases steadily. Hence, the heavier components as the supplementary absorbent can intensify the absorption process.

Figure 2. Phase equilibrium curves of a pseudo-binary system.

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Figure 3. Maximum distance dmax and equilibrium vapour pressure.

In order to further investigate the effect of different supplementary absorbent components on their flowrate, absorption performance and energy requirement, the GASP is simulated using the process simulation software PRO/II V9.0.41 The SRK thermodynamic method is used to describe the phase equilibrium accurately. The screening of thermodynamic methods was proposed by Chen and Mathias.42 The binary interactions and the physical properties of components are the default values. The gasoline consists of n-paraffins, i-paraffins, c-paraffins, n-olefins, i-olefins, c-olefins and aromatics.43,44 The quality specification of the dry gas is set so that the mole fraction of C3+ components in the dry gas must not be higher than 3%. Each component from C5 to C11 is respectively set as the supplementary absorbent and simulated. The flowrate of the supplementary absorbent stream, absorption effects and the heat duty of the process system are obtained and shown in Fig. 4. It is shown that when the supplementary absorbent is C5 or C6 hydrocarbon, the flowrate of the supplementary absorbent stream is higher, the lean absorption oil must be introduced and the secondary absorber must be installed in order to satisfy the quality of the dry gas. When using heavier hydrocarbons as the supplementary absorbent, including C7, C8, C9, C10 or C11, the gas from the top of the primary absorber can satisfy the specification of the dry gas through changing the flowrate of the supplementary absorbent streams. Thus, the secondary absorber can be cancelled. It is worth noting that the C9 hydrocarbon shows the best absorption effect, with the minimum flowrate and heat duty of the stripper and stabilizer. 13



However, the supplementary absorbent with the same carbon number is hard to obtain in real practice. For example, it requires a great amount of energy to separate C9 hydrocarbons from the stabilized gasoline. Whereas obtaining the hydrocarbons with a few continuous carbon numbers is much easier. Similarly, the mixtures with continuous carbon numbers, such as hydrocarbons C5, C5-C6, …, C5-C11; C6, C6-C7, …, C6-C11; …; C10, C10-C11, C11, are set as the supplementary absorbent streams, respectively. The simulation results are shown in Fig. 5. When using the supplementary absorbent streams containing C5 hydrocarbons, the C6 hydrocarbon stream or the C6-C7 hydrocarbons stream, the flowrate must be greater than 40.0 t/h, which is higher than other supplementary absorbent streams. More importantly, the secondary absorber must be used to control the quality of the dry gas. The rest of the supplementary absorbent streams absorb well enough to cancel the secondary absorber, and the heat duty of the stripper and stabilizer varies between 15.2 to 16.1 MW. The C9 hydrocarbon supplement stream still achieves the minimum flowrate and results in the minimum heat duty of the stripper and stabilizer. According to the analysis above, the heavy components in the supplementary absorbent can intensify the absorption process more than the light components. The heavy components, like C11 hydrocarbons, have poorer solubility to LPG components, which is a disadvantage for absorption. However, the light components, like C5 hydrocarbons, have higher volatility and thus are easily entrained by the rising gas phase in the primary absorber. Therefore, if the supplementary absorbent contains large amounts of light components, the secondary absorber must exist to absorb the light components in the lean gas. Conversely, using heavier 14


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supplementary absorbent can lead to the cancelation of the secondary absorber.

Figure 4. The flowrate of the supplement and the heat duty of the system when using components with the same carbon number as the supplement.

Figure 5. The flowrate of the supplement and the heat duty of the system when using components with continuous carbon numbers as supplement.

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3. A new indicator guiding process intensification According to the above analysis, the absorption process in the absorber has two aspects. The heavier supplementary absorbent can reduce the volatility, while having lower solubility. The lighter supplementary absorbent can increase the solubility to C3-C4 hydrocarbons, while having larger volatility, and a secondary absorber must be installed to further recover the volatile components. Hence, we must make a trade-off between the solubility and volatility of supplementary absorbent. In this study, we present a new indicator to evaluate the performance of the supplementary absorbent, as shown in Equation (1). = log

/( − )

(1)

where R is the index, SLK is the flowrate of light key component absorbed by the supplementary absorbent, VSA is the flowrate of heavy key component volatilizing from the supplementary absorbent, FSA is the flowrate of supplementary absorbent, and c is a constant. The indicator represents the ability of mass transfer in the absorber. The higher the indicator, the more the mass transfer process intensified. We calculated the indicator R for different supplementary absorbent streams. The results are shown in Fig. 6. R value increases sharply when we change the composition of the supplementary absorbent. The lower value of R is 0.03 where the composition of supplementary absorbent is C5, and the largest value of R is 0.80 where the composition of supplementary absorbent is C9. Hence, appropriate composition of supplementary absorbent can intensify the separation process in the absorber.

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Figure 6. Various trends of R with the composition of supplementary absorbent streams

4. New process A new process flowsheet with a flash tank and without the secondary absorber was introduced to improve the conventional GASP. The new flowsheet is shown in Fig. 7. The solid black line represents the same parts as the conventional GASP, the solid red line represents the new parts, and the dotted grey line represents the deleted parts compared to the conventional GASP.

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Figure 7. Flowsheet of the new process.

The purpose of the new flash tank is to obtain suitable hydrocarbon components used as the supplementary absorbent stream to intensify the absorption process. As shown in Figs. 4 and 5, the secondary absorber can be cancelled when the supplementary absorbent is heavy enough and the gas phase from the top of the primary absorber can satisfy the specification of the dry gas. The specifications of the dry gas, LPG and stabilized gasoline remain the same. The columns, except the secondary absorber, remain and function as before. The columns in the conventional and new processes are operated at the same pressures. Nevertheless, the stream directions are varied. The bottom product of the stabilizer is fed into the new flash tank that operates at approximately atmospheric pressure. The liquid phase of the flash tank is partly sent to the primary absorber as the supplementary absorbent stream, the rest of the 18


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liquid phase and the gas phase from the flash tank are sent to other units for further processing. The gas phase from the top of the primary absorber is incorporated into the fuel gas network. The lean absorption oil, which is the diesel product from the main fractionators in the upstream units, and the rich absorption oil, which is sent back to the main fractionators, are all cancelled. The process comparison between the conventional and new GASPs is shown in Table 1.

Table 1. Comparison of the conventional and new GASPs Items

Conventional process

New process

Number of columns

4

3

Number of absorption streams

2

1

Number of recycling streams between units

1

0

Process simplification in the chemical process industry has a significant impact on efficiency and sustainability.12 The new GASP reduces the original 4 columns to 3, and 2 absorption streams to 1, compared with the conventional GASP. The reduction of the secondary absorber process can significantly decrease the investment of the equipment costs and avoid the introduction of the lean absorption oil, a material stream outside the GASP. The lean absorption oil is separated from the crude gasoline at the main fractionators in the upstream units and the rich absorption oil, consisting of diesel components and a small amount of gasoline components, is sent back to the main fractionators to separate again. 19



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Hence, the cancellation of the recycled lean absorption oil can also decrease the equipment cost and energy requirement at the upstream units. In this case, the high-temperature heat removal from the main fractionator in the up-stream unit is increased from 35.09 to 38.02 MW when using the new GASP. The high-temperature can be used to generate 3.5 MPa steam, as shown in Fig. 8.

(a)

(b)

Figure 8. Pinch analysis of GASP and upstream units (a) conventional process; (b) new process.

5. Process simulation and optimization 5.1 Process simulation A real industrial GASP in a Chinese refinery is investigated. The composition of the rich gas is listed in Table S1. The basic data and product specifications are shown in Table S2. The existing GASP is simulated using the process simulation software PRO/II V9.0. We can also 20


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use other process simulators to simulate the process; which does not affect the simulation results. To assess the accuracy of the simulation, the results obtained from the simulation software are compared with the operating data extracted from the real process, as shown in Table S3. The difference between the calibration value and simulation value of the parameters is less than 5%, indicating that the model is reliable for engineering simulation.

5.2 Optimization of parameters for the new flash tank A new flash tank was introduced in the new GASP. As a result, the secondary absorber was cancelled. The energy requirement of the new process is significantly affected by the composition and flowrate of the supplementary absorbent when the feed is given and the columns have been optimized. The operating pressure and temperature of the new flash tank governs the flowrate and composition of the supplementary absorbent. Hence, the operating pressure and temperature of the flash tank affect the energy consumption of the GASP independently. For the given new process, the input variables are the operating pressure and temperature of the new flash tank and the output variable is the energy requirement. We can use Equation (2) to represent the optimization, with the constraints of mass and energy balance. = (, )

(2)

where Q is the energy requirement, T and P are the operating pressure and temperature of the new flash tank, respectively. The mathematical model based on first principles for minimum energy requirement is very 21



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complex and presents difficulties in optimization, especially in global optimization.45 In this study, we use response surface methodology to optimize the two parameters on the basis of high-throughput process simulation. Through initial process simulation, we set the upper and lower bounds of operating pressure at 0.25 and 0.1 MPa, and the upper and lower bounds of the operating temperature were set at 185 and 140°C. The lowest bound of operating pressure is set at 0.1 MPa to avoid using vacuum environment. At the lowest operating pressure, when the operating temperature is lower than 140°C, the light components in the supplementary absorbent are high. As a result, we cannot cancel the secondary absorber. Hence, we determine the lower bound of operating temperature at 140°C, then the optimization problem can be further represented by Equation (3). = (, ) . . 0.1 ≤ ≤ 0.25 140 ≤ ≤ 185

(3)

In the process simulation, the product specifications are kept the same. The response surface of energy requirement is shown in Fig. 9. Q reaches the minimum value of 21.00 MW when P is at 1.0MPa and T is at 178.5°C. This temperature equals the bottom temperature of the stabilizer. Hence, the bottom product of the stabilizer is fed to the new flash tank directly, neither heating nor cooling.

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Figure 9. Utility requirement varying with operating pressure and temperature of the flash tank.

5.3 Absorbent comparison The optimal supplementary absorbent’s compositions are shown in Table 2. In the conventional GASP, C5 hydrocarbons in the supplementary absorbent have the highest content, which is not good for the absorption process. In the new GASP, the mass fractions of the C4 and C5 hydrocarbons in the supplementary absorbent drop dramatically, and the mass fraction of the C9 hydrocarbons increases from 11.30 to 26.76%, which can significantly intensify the absorption process, as mentioned above. When using the new supplementary absorbent, the R value is 0.13; this is 1.63 times of the conventional supplementary absorbent.

Table 2. Composition of supplementary absorbent (mass fraction/%) 23



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Hydrocarbons

New supplementary absorbent

Conventional supplementary absorbent

C4

0.17

2.00

C5

4.98

36.79

C6

4.86

16.07

C7

7.12

12.16

C8

10.94

8.86

C9

26.76

11.30

C10

24.58

8.04

C11

20.59

4.78

Total

100

100

6. Heat integration and HEN design In order to compare the new and conventional GASPs, heat exchanger networks (HENs) were designed for the two processes. Table 3 shows the flowrates of the supplementary absorbent and the lean absorption oil, and the heat duties of the stripper and stabilizer in the conventional and new processes. The mass flowrate of the supplementary absorbent stream and the total heat duty in the new process are similar to those in the conventional process. While the lean absorption oil is reduced to 0 in the new process.

Table 3. Flowrates and reboiler heat duties for the two processes 24


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Item

New Process

Conventional Process

Supplement flowrate/t·h-1

33.67

33.35

Lean absorption oil flowrate/t·h-1

0

7.16

Stripper heat duty/MW

6.75

6.94

Stabilizer heat duty/MW

14.25

14.02

Total heat duty/MW

21.00

20.96

The software Aspen Energy Analyzer V7.2 was used to carry out the pinch analysis and economic evaluation for the HENs. The stream information is extracted from the simulation files. Table S4 shows, in detail, the extracted information for the new and conventional GASPs. The cooling water from 20 to 25°C is used as the cold utility stream. The middle-pressure steam (MP steam) from 175 to 174°C and the high-pressure steam (HP steam) from 250 to 249°C are used as the hot utility streams. The utility streams are listed in Table S5. The default cost and heat transfer coefficient (HTC) for utility streams are retrieved from the database in the software. The minimum approach temperature (∆Tmin) of HEN is a key parameter in balancing the capital cost and operating cost. The optimal ∆Tmin must minimize the total annual cost (TAC) of HEN. The TAC of HEN varying ∆Tmin are calculated using the Aspen Energy Analyzer, the rate of return (ROR) and plant life (PL) are set as default at 10 and 5, respectively. The optimal ∆Tmin is 19°C in the conventional process and 15°C in the new process. The utility duty is shown in Table 4. 25



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Table 4. Utility duty of two GASPs Duty

New GASP

Conventional GASP

Heat duty/MW

21.00

19.31

Cooling duty/MW

23.15

24.56

The Aspen Energy Analyzer can automatically generate HEN design with the objective function of minimum TAC. Table 5 gives four possible optimal designs for the two GASPs. Due to the objective of minimum TAC, Design 2 is the best design option for the conventional process, and Design 3 is the best design option for the new process. The HEN diagrams of the two optimization designs are shown in Figs. 10 and 11. In the conventional GASP, the supplementary absorbent stream provides part of the heat to the stripper’s reboiler, while all the heating and cooling duties are provided by utilities in the new GASP. According to HENs design, the new process increases the hot utility by 8.75%, and decreases the cold utility by 5.74%.

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Table 5. Recommend designs for the conventional and new GASPs Process

Schemes

Operating cost/($103·yr-1)

Capital cost/($103)

TAC/($103·yr-1)

Conventional

Design 1

220.0

232.9

288.6

Design 2

213.5

237.0

283.3

Design 3

227.9

183.1

281.8

Design 4

236.2

181.8

288.6

New

Figure 10. HEN diagram of the conventional GASP for the least TAC design.

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Figure 11. HEN diagram of the new GASP for the least TAC design.

7. Economic Evaluation 7.1 TAC comparison Four columns including the primary absorber, stripper, secondary absorber and stabilizer, and the two flash tanks are counted into the TAC for the conventional and new processes. The capital cost data are often expressed as a power law of capacity, as in Equation (4). +

-

() = (* +,

(4)

where CE is the equipment cost with capacity, CB is the known base cost for equipment with capacity, M is constant depending on equipment, and fM is the correction factor for materials of construction. Data for the equipment items is on the basis of January 2000 costs (CE Index of Equipment = 435.8) and listed in Table S6.46 The parameters of the equipment are shown in Table S7. The material of column Q245R is a carbon steel, the same as the base material. The material 28


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of the valve trays is 06Cr13Al, a low-grade stainless steel. Thus the - is equal to 2.4.47 The material of the tanks is Q245R. The capital costs are shown in Table S8. The data shown in Table S8 is old, but such data can be brought up to date and put on a common basis using cost indexes as in Equation (5). ./

.0

=

123)4/

(5)

123)40

where Ca is the equipment cost in year a, Cb is the equipment cost in year b, INDEXa is the cost index in year a, and INDEXb is the cost index in year b. Capital costs can be expressed on an annual basis if it is assumed that the capital has been borrowed over a fixed period at a fixed rate of interest, in which case the capital cost can be annualized according to Equation (6). C(DEC)F

566789:;:?89 @A? = 8>:?89 @A? × (DEC)F GD

(6)

where i is the fractional interest rate per year, and n is the number of years. The index in 2000 was 435.8, and in 2017 was 541.7. The capital cost is to be annualized over a five-year period at a fixed rate of interest of 5%. Thus, the total annualized equipment cost is calculated and shown in Table S9. The TACs of the two GASPs, including the columns, tanks and HEN equipment costs, and operating costs, are shown in Table 6. The conventional process was counted without the new flash tank, and the new process was counted without the secondary absorber. According to the analysis above, including the considerations of the HENs and equipment, the new GASP can reduce annual costs by 11.24%, compared to the conventional GASP.

29



Page 30 of 38

Table 6. TAC of the conventional and new GASPs Processes

HEN TAC / ($103·yr-1)

Equipment Cost / ($106)

Total Cost / ($106·yr-1)

Conventional

283.3

1.747

2.030

New

281.8

1.527

1.809

7.2 Sensitivity analysis The new process has a lower TAC and decreases the cold utility by 5.74%, but it increases the hot utility by 8.75%. Energy prices have varied over a relatively large range in recent years. We changed the prices of hot and cold utility streams in the same ratio, resulting in the variation of TAC, as shown in Figure 12. We varied the prices of utility streams from 0.5 to 20 times. As a result, the TAC of the conventional process increased from $1.923 to $6.088 million, and the TAC of the new process increased from $1.695 to $6.139 million. When the prices of utility streams are lower, the TAC of the new process is lower than the conventional process. When the prices of utility streams are increased to 16 times, the TACs of the conventional and new processes are the same. When the prices of utility streams are increased over 16 times, the TAC of the conventional process is lower. When the prices of utility streams are increasing, the ratio of TACs between the new and conventional processes rises more and more slowly, as shown by the purple curve in Figure 12.

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Figure 12. Sensitivity analysis for utility price.

8. Conclusions This study presents a new GASP to intensify the absorption process and simplify the flowsheet of the conventional GASP. The solubility and volatility of the C5-C11 hydrocarbons were discussed. A new indicator is presented and the software PRO/II was used to aid finding the most suitable components for the supplementary absorbent. The conventional and new processes were fully simulated under the same feedstock and product specifications. For further comparison, the process data was extracted to perform the HEN design and economic evaluation. Three conclusions can be drawn from this study: (1) With the increase of carbon number in the supplementary absorbent’s components, the solubility and volatility all decrease. The supplementary absorbent with C9 hydrocarbons gets the minimum supplementary absorbent flowrate and heat duty. 31



(2) The new GASP need not install the secondary absorber because the new flash tank can obtain a more efficient supplementary absorbent stream to intensify the separation process in the primary absorber. Thus, the number of columns is reduced from 4 to 3, and the external recycling stream is also cancelled. The process has been significantly simplified. (3) Compared to the conventional process, the new process can significantly decrease TAC by 11.24%. Furthermore, the new process has a better economic performance when the energy price does not rise over 16 times.

ASSOCIATED CONTENT Supporting Information The Supporting Information is available free of charge on the ACS Publications website. Table S1. The composition of the rich gas; Table S2. Operating pressure and product specifications; Table S3. Comparison between operating and simulation parameters; Table S4. Stream data of two GASPs; Table S5. Utility data for heat integration; Table S6. Typical equipment capacity delivered capital cost correlations; Table S7. Parameters of equipment; Table S8. Capital cost of items; Table S9. Annualized equipment cost.

AUTHOR INFORMATION Corresponding Authors *(B. J. Zhang) Tel.: +86 20 84113731, Email: [email protected]; (Q. L. Chen) Tel.: +86 20 84113659, Email: [email protected].

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ACKNOWLEDGMENTS This research is supported by the National Natural Science Foundation of China (Nos. 21776323 and U1462113), the project of Guangdong Provincial Natural Science Foundation of China (No. 2015A030313112), and the Fundamental Research Funds for the Central Universities.

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