Modeling and simulation of a multistage absorption hydration hybrid process using equation oriented modeling environment

Modeling and simulation of a multistage absorption hydration hybrid process using equation oriented modeling environment

Accepted Manuscript Title: Modeling and simulation of a multistage absorption hydration hybrid process using equation oriented modeling environment Au...

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Accepted Manuscript Title: Modeling and simulation of a multistage absorption hydration hybrid process using equation oriented modeling environment Author: You Li Xingang Li Hong Li Luhong Zhang Feng Xin Jingyan Lian Yonghong Li PII: DOI: Reference:

S0098-1354(15)00385-3 http://dx.doi.org/doi:10.1016/j.compchemeng.2015.12.021 CACE 5340

To appear in:

Computers and Chemical Engineering

Received date: Revised date: Accepted date:

10-10-2015 21-12-2015 25-12-2015

Please cite this article as: Li, Y., Li, X., Li, H., Zhang, L., Xin, F., Lian, J., and Li, Y.,Modeling and simulation of a multistage absorption hydration hybrid process using equation oriented modeling environment, Computers and Chemical Engineering (2016), http://dx.doi.org/10.1016/j.compchemeng.2015.12.021 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.

Highlights:

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1. An absorption hydration method is established using of ACM® and Aspen Plus®.

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2. Key parameters of the absorption hydration process are analyzed.

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3. Based on optimization calculation, optimal operation conditions are found.

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4. Effect of hydrate structure on separation of dry gas is discussed.

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5. Based on economic evaluation, aspects need further investigation are pointed out.

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Modeling and simulation of a multistage absorption hydration hybrid process using equation

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oriented modeling environment

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You Li1, Xingang Li1, 2, Hong Li*1, 2, 3, Luhong Zhang1, 2, Feng Xin1, Jingyan Lian2, 4, Yonghong Li1, 2

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1. School of Chemical Engineering and Technology, Tianjin University, Tianjin 300072, China

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2. National Engineering Research Center of Distillation Technology, Tianjin 300072, China

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3. Collaborative Innovation Center of Chemical Science and Engineering, Tianjin 300072, China

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4. School of Chemical Engineering and Technology, Tianjin University of Technology, Tianjin 300072, China

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* Corresponding Author. E-mail:[email protected] Tel: +86-022-27404701

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Abstract: Separation of light hydrocarbon mixtures is one of the most important topics in chemical

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engineering research. With development of theories on hydrate equilibriums and kinetics,

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researchers are trying to apply hydration based separation technology to industrial applications. It

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is increasingly important to develop the corresponding simulation strategies for process design

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purposes. In this work we use an equation oriented modeling environment, named Aspen Custom

7

Modeler® (ACM®), which enables rapid model development and provides powerful simulation

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solvers. With the help of ACM®, a multistage absorption hydration hybrid process (AHHP) for

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refinery dry gas separation is modeled and simulated. Sensitivities of key parameters such as water

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content and absorbent flow rate, are analyzed. Features of the multistage AHHP are discussed. For

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comparison, based on an industrial data, a butane absorption process is established and simulated.

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Economic evaluation shows that the multistage AHHP is competitive compared to current

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absorption process.

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Key words: hydrate, absorption, water in oil emulsion, separation, process simulation, refinery

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dry gas, ethylene

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1. Introduction: Separation of light hydrocarbon mixtures is one of the most important topics in chemical

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engineering research. Conventional methods for separating low boiling point gas mixture include

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cryogenic distillation and alkane absorption. New technologies such as reactive absorption,

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membrane separation and adsorption are investigated.(Huang et al. 1999; Padin and Yang 2000;

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Adhikari and Fernando 2006) Recently, hydrate is proposed as a new tool for light hydrocarbon

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separations. (Eslamimanesh et al. 2012)

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1.1 Literature review on hydrate based separation experiment

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Formed at high pressure and low temperature, hydrates are cage like, non-stoichiometric

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compounds composed of water and hydrocarbons.(Sloan 2003) When a mixture of hydrocarbons

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forms hydrate, some components are more easily captured than others. Large amount of work is

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performed in hydrate based light hydrocarbon separation. Ma and coworkers investigated the

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equilibriums of CH4 and C2H4. (Ma et al. 2001, 2008) Englezos and coworkers studied the kinetics

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of hydrate formation. (Englezos et al. 1987a, 1987b)

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Recent research shows that for industrial purpose, forming hydrate with water as the continuous

phase is unviable, because hydrate formed would float on top of water and thereby impede further hydration.(Liu et al. 2013b) Chen and coworkers put forward a new material named water in oil emulsion (w/o emulsion) and recommended it as the new absorbent(Turner et al. 2009; H et al.

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2013; Liu et al. 2013a; Ma et al. 2013a). Their observation shows hydrate agglomeration can be

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prevented, since hydrate generated is dispersed as tiny particles. Meanwhile, due to the large

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contact area between oil and water, hydration rate is facilitated (Mu et al. 2014) .

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Though excellent work is carried out by the scholars, literature survey shows current study has

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been focused on hydrate formation kinetics and equilibriums in single reactors. And there is a

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growing demand for a process scale modeling and simulation method which can provide insights

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and guidance for the further application of hydration technology.

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1.2 Literature review on hydration simulation

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A number of theories are proposed for hydrate equilibrium calculation, among which

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Van-der-walls equations are the most famous. However, tough accurately derived, solving the

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van-der-walls equations requires global optimization algorithm that is time consuming. As an

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alternative, Chen-Guo equations have a simpler formulation and can provide a quick solution

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without losing much accuracy, and thereby more appealing for process simulation purposes (Chen

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and Guo 1998, Sloan 2007).

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Naeiji and coworkers calculated the performance of hydrate in separating methane and ethane

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with the help of a phase map. (Naeiji et al. 2014) The phase map can provide an intuitive

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understanding in the hydration separation mechanism. Yet, there are two problems ahead for this

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method. First, the phase map cannot provide a quick and accurate solution. Second, it can only

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calculate the hydrate composition formed from pure water. Therefore, more robust simulation methods are needed.

1.3 Current process flow chart for dry gas separation Ethylene is a fundamental a raw material in petroleum industry. Refinery dry gas usually

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contains a significant amount of ethylene (about 10~20 mol %), which, however, is often flared

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for energy.(Li and Luo 2014) Moreover, the development in catalytic cracking, which primarily

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aimed to improve the light oil generation, results in ethylene content increment in dry gas.(Rahimi

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and Karimzadeh 2011; Sadeghbeigi 2012) Therefore, it becomes increasingly important to recover

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ethylene from dry gas. A list of dry gas composition can be observed in Table 1.(Lei Si 2013) Admittedly, cryogenic distillation system is known to produce high quality products. However,

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at the same time, it also demands huge initial investment. As one of the most expensive units in

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cryogenic process, cold box is usually used to chill gas at cryogenic temperatures. To save

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investment in cold box, before cryogenic distillation, dry gas should be preprocessed, as is the

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case of Sheng Li refinery.(Lei Si 2013)

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The C4 absorption flow chart for dry gas pretreatment can be observed in Figure 1. The

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absorption flow chart can be divided into three parts: absorption (B-1), stripping (B-2) and

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desorption (B-3). Feed gas is absorbed in the absorption section. In stripping section, part of

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methane and nitrogen are driven away from the absorbent. Desorption column extracts enriched

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dry gas from butane. According to industrial data(Lei Si 2013), the operating conditions are listed

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in Table 2. Pressure of enriched dry gas is 2.6MPa. Before getting into cryogenic distillation, the

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enriched dry gas should be compressed up to 3.3MPa, which is a typical condition for

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methane/ethylene separation.(Salerno et al. 2011)

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1.4 Contribution of this work

In this work, we build the models in Aspen Custom Modeler® (ACM), an integrated develop

environment for chemical process modeling.(Aspen Technology, 2012) ACM can greatly reduce the amount of labor in coding by providing with the built in flash functions and property

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calculation procedures. The models are then exported as unit operation blocks to Aspen Plus® for

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integrated simulation and optimization.

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Chen and Naeji proposed a multistage separation strategy for hydrate based separation.

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Following their footprints, we propose and simulate a multistage absorption hydration hybrid

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process (AHHP) for dry gas separation applications. Sensitivity of the key parameters are studied.

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Then we point out important features and evaluate some economic issues. For comparison, a C4

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absorption process is also simulated and evaluated. It is found the proposed AHHP is more

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economical. Reasons for the savings are discussed and analyzed. Nevertheless, we found there is

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still room for improvement provided that the refrigeration is reused.

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To summarize, most previous work on hydration experiment and simulation is on unit operation

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scale. In this article we put forward a powerful tool, which is especially suitable for large scale

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process calculation. Then the suggested approach is applied to model a hydration based dry gas

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separation system. This article is organized as follows. The next section introduces the multistage

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schematic. Third section explains the overall solution strategy including hydration modeling,

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mathematical formulation of optimization problem as well as the corresponding optimization

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algorithm. The fourth section shows, explains and discusses the calculation results. Eventually, the

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last section states the conclusions.

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2. Multistage AHHP schematic

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The conceptual design of multistage schematic can be found in the work of Liu (Liu et al.

2013b) and Naeji (Naeiji et al. 2014). As is shown in Figure 2a, feed gas and water forms hydrate in the reactor. After hydration, equilibrium gas phase is treated as off gas, while hydrate is controlled to dissociate. Gas released from hydrate then forms hydrate again. There is a defect in

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the design: though composition can be updated trough pressure and gas/water ratio control,

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recovery of desired component is low. This is because gas that failed to enter hydrate is all treated

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as off-gas.

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To tackle with the above issue, we redesigned the flow chart. The new design is shown in

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Figure 2b and is named multistage AHHP. The operation condition can be found in Table 3.

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Refinery dry gas is feed to Stage-n after chilling in HX-2 (n is the stage number). Gas flows from

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Stage-n to Stage-(n-1). The absorbent, which in this case is the w/o emulsion, goes the opposite

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way. The final liquid phase out of Stage-n is transported to R-1. The cryogenic process that

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follows AHHP operates at 3.3 MPa. Hence, R-1 operates at 3.3 MPa. In R-1, the temperature is

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raised to 40℃ for hydrate dissociation. The liquid phase out of R-1 is flashed at atmosphere

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pressure in R-2. Gas released from R-2 should be recompressed to 3.3MPa. The emulsion is

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reused for continuous operation.

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The stages in Figure 2b stand for stirred tanks in which the phases are at equilibrium. In unit

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operation text books, these connected stages are represented as columns.(Seider et al. 2009)

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However, to avoid hydrate agglomeration, stirring is required in each tank. For this reason, and

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because refrigeration device is needed for temperature control, making inner structure differ a lot

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from column trays, in this work the stages cannot be simply represented by a column.

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3. Solution strategy

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3.1 Modeling and simulation environment Instead of FORTRAN or phase map, in this work we build the hydration models in an equation

oriented modelling environment, which has user friendly interface and whose code is easy to implement and maintain. There are a number of options, such as gPROMS® and SystemModeler®.

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And we choose Aspen Custom Modeler® due to the following reasons. First, ACM® has powerful

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built in functions, property packages which greatly lighten the burden in coding. Second, the

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models built in ACM® can be exported to Aspen Plus® for process integration purposes. Last but

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not least, for complex flow charts, we must tune the parameters through optimization, which is

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perfectly handled by the built-in optimization modular in Aspen Plus®.

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3.2 Multiphase equilibrium modeling In the proposed AHHP, each of the stage in Figure 2b is assumed to be at multiphase

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equilibrium. We will introduce the calculation procedure in four aspects including structure

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determination, mass and energy balance equations, equilibrium correlations and phase

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determination.

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3.2.1 Hydrate structure determination

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As is known, three types of hydrate structures have been recognized, namely SH, SI and

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SII.(Subramanian et al. 2000; Murshed et al. 2010) An important aspect in hydration modeling is

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structure determination. Structure of hydrate is influence by a number of factors including pressure,

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temperature, additives as well as composition if gas mixture is concerned. With the existence of

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propane and since composition of ethane and ethylene in the gas phase is about 3~25mol%, it is

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found that the SII type is preferred for the dry gas separation system (Sun et al. 2007; Watanabe et

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al. 2011; Kondo et al. 2014).

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3.2.2 Equilibrium correlations

Temperature and pressure are controlled at certain value to ensure hydrate formation.

T  T0

(1)

P  P0

(2)

The vapor phase and liquid phase are considered at equilibrium and thereby we have identical fugacity for vapor and liquid.

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f jL  f jV , j  C

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Peng-Robinson EOS (Equation of State) is employed to calculate the fugacity of components in

(3)

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Chen-Guo equations are used for hydrate-vapor equilibrium. (Chen and Guo 1998; Ma et al. 2013b). The occupancy of small and large hydrate cages are calculated separately.

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oil and vapor phase.

For small cages,

j 

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f jV C j

(4)

1   f iV Ci

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Cj is the Langmuir constant. For large cages,

z j 

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f jV

Yj T Zj

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C j  X j  exp[

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1

j

(5)

(6)

(7)

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j

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fTj  exp[ 0

  Aij j j

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f j 0  f Tj 0  exp[

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T

zj 

P ]  a w 1/ 2 T

]  [ A ' j exp(

B'j T C'j

  1 / 2 z j *   j

(z

* j

  j )

(8)

)]

(9)

(10)

(11)

j

zj* denotes mole fraction of hydrocarbon component j in large cavities. zj stands for the mole

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fraction of component j in hydrate phase. fjV is the fugacity of component j in the gas phase. fj0

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stands for the hydrate-phase fugacity of component j when the small cavities are not occupied by

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hydrocarbon molecules. f Tj0 denotes the effect of temperature on fj0. For SI hydrate, α=1/3. For SII

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hydrate, α=2. aw represents the activity of water. A’i, B’i, C’i, Xi, Yi, Zi and binary interaction

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coefficients.(Chen and Guo 1998; Sun et al. 2007) Aij are constants. The parameters values can be

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found in the work of Ma. (Ma et al. 2013)

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3.2.3 Mass and energy balance equations The mass balance equations are shown as follows,

j j j Vin yinj  Lin xinj  HYin zinj  Winj  Vout yout  Lout xout  HYout zout  Woutj

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(12)

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Hydrocarbons dissolved in water phase are so trivial that they can be neglected in the

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calculation.(Chen and Guo 1998)

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Woutj  0, j  C and j≠ w

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A hydrate unit is a fundamental brick with which a hydrate crystal particle is build. The

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additional mass balance equation correlates water molecules and number of units (NOU) in

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hydrate.

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NOU 

w (W inw  HYinw  W out ) NOM

(14)

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The superscript w denotes water flow rate in each phase. NOM stands for the H2O molecule

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λ1 and λ2 are the number of small and large cages in a hydrate unit. For SI hydrate, λ1 is 2 and λ2 is

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6. For SII hydrate, λ1 is 16 and λ2 is 8.

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number in a unit. According to Sloan (Sloan 2003) NOM is determined by the type of hydrate structure. Since large cages are always considered fully filled, therefore we can calculate the total amount of light hydrocarbons in hydrate using the occupancy of small cages.

HYout  NOU  (1  j  2 ) j

(15)

Energy balance is as follows:

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Vin H inv  Lin H inl  HYin H inhy  Win H inw  Q

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Heat generated from hydrate formation described by the following equation:



H total   H j  n j cpt

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j



(17)

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(16)

v v hy w  Vout H out  Lout H out  HYout H out  Wout H out

where njcpt is the amount of component j captured in hydrate. In dry gas separation system,

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composition of hydrogen and propane in hydrate is less than 1%, hence they are not incorporated

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in the energy balance calculation. For simplicity, oil in water oil emulsion is calculated as

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n-dodecane.

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3.2.4 Phase determination

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An issue that deserves attention is that there are two categories for the multiphase calculation.

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First, when sufficient water is supplied, the system will have 4 phases, namely vapor, oil, water

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and hydrate. And if we perform calculation using the equations (1-15) we will have a positive

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value for Wout. However, when water is entirely transformed into hydrate, the system will contain

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3 phases namely, vapor, oil and hydrate. In this case, after the calculation using equations (1-15),

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we will obtain a negative value for Wout, which means that all water is transformed into hydrate. In

this situation, the calculation procedure should be carried out again replacing equation (6) and (7) with the following equations

W out  0 

zi 

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f jV f j0

(18)

 [ i

fiV 1 ] fi 0

(19)

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As shown in equation (19), the large cage occupancy is replaced by a competition mechanism

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(Zhang et al. 2004a, 2004b; Ma et al. 2008, 2013b).

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3.3 Economic evaluation methods

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In this work, we evaluate the economic performance of different process in terms of total annual

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cost (TAC). TAC is comprised of annualized utility cost and equipment investment. Utility price is

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referred to Seider (2004). Formulas for the equipment investment can be found in appendix. We

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choose a payback period of 10 years with annualized interest rate of 10%.

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3.4 Problem formulation and solution

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Economic evaluations in this work are aimed to compare economic behavior between the

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proposed AHHP and the C4 absorption process. Since current dry gas recovery system operates at

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conditions where ethylene recovery is 90%, we choose total annual cost (TAC) as the target

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function with the constraint that ethylene recovery equals to 90%. However, the enriched dry gas

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composition from different process may not be the same. Therefore, we have to add cost functions

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for different product gas composition. The functions are obtained through evaluation of a

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simplified cryogenic process, which separates methane and nitrogen from enriched dry gas. The

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flow chart and operating conditions can be found in Appendix. Eventually, an optimization

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problem is formulated as follows for the multistage AHHP.

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Min TAC s.t. Unit operation models;

Specifications on yield and product purity; Variable ranges.

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In the above formulation, TAC incorporates the investment and operating cost of a dry gas

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pre-separation process and a simplified cryogenic process. The constraints include unit operation

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models, parameter ranges and design specifications. Specifications and parameter ranges are listed

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in Table 4. Temperature should be above 0℃, which is the ice formation temperature. Pressure is

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set to be the same with C4 absorption. Water content is supposed to be greater than 10v% and less

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than 30v%. When water content is too low, the absorbent cannot reflect the effect of hydration.

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When the water content is over 30v%, the viscosity becomes too large for efficient stirring if

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hydrate is formed. Feed gas composition is listed in Table 1. Dry gas flow rate is 7.5t/h at 10

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℃,4.25MPa.

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The optimization problem can be categorized as nonlinear programming problem, which in this

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work is solved by the DMO solver. The DMO solver implements a variant of the successive

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quadratic programming (SQP) algorithm. In Aspen Plus®, the optimization algorithms are

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conveniently embedded in the software. Equation oriented mode (EO mode) is used, which is

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known to facilitate the convergence of recycled flow charts. Specifications on yield and product

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purity is added as equations in Calculator Blocks (Aspen Technology, 2012).

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4. Results and discussion

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4.1 Optimized profiles

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The optimized profile of AHHP is listed in Table 5a. As is shown, optimal water content is

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30v%, which is the upper bound. Temperature is also optimized to the lower bound, 0oC.

Dodecane flow rate is 105.7 kmol/hr. In the next section, we will look into the sensitivities of the key parameters one by one. For comparison, profile of C4 absorption process is listed in Table 5b. The parameters in C4 absorption is also tuned trough optimization. Sensitivities in next section are obtained through fixing the investigated parameters and optimize other parameters.

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The process simulated is the combination of two parts. One is the absorption pre-separation part.

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The other is the added cryogenic de-methane part. To illustrate the composition of enriched dry

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gas produced from absorption part, C1/C2 ratio is used as an indicator, which is the ratio of

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methane to ethane and ethylene. Higher C1/C2 ratio means greater methane content in enriched

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dry gas.

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4.2 Sensitivity analysis

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4.2.1 Ethylene recovery

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Optimized profiles of different ethylene recovery are shown in Table 6. As ethylene recovery

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steps from 0.88 to 0.92, optimized absorption temperature and water content are 0℃ and 30v%,

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respectively. The oil flow rate in emulsion grows almost linearly from 103.4 kmol/h to 108.0

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kmol/h. Yet, variation in C1/C2 ratio is found negligible.

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4.2.2 Water content

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Optimized profiles for different water content are listed in Table 7. Influence of water content is

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reflected on two aspects. First, with the increase of water content from 10v% to 30v%, oil flow

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rate in emulsion decreases rapidly from 164.2 kmol/h to 105.7 kmol/h. Second, as water content

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increase, C1/C2 ratio has grown from 0.350 to 0.511, which indicates greater methane content in

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the enriched dry gas.

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It is interesting to find the change in TAC not significant for different water content (less than

0.01 M$/yr). With larger water content, cost of compressor duty and equipment investment can be reduced, because gas released from hydrate have a higher pressure which enables this portion of gas to bypass the compressor. However, on the other hand, as more hydrate is formed,

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corresponding refrigeration duty consumed by hydrate formation will go up. Compromise of the

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two cost sources mentioned above has resulted in the small variance of TAC.

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Since water content has little effect on TAC, we now try to determine water content considering

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other aspects. First, for the sake of easy operation, less water content is better, which makes the

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hydrate/oil slurry behave more like liquid and will alleviate the risk of forming plugs. However, as

2

is known(Sloan Jr and Koh 2007), hydrate dissociation can provide a significant amount of

3

refrigeration. It is found phase change materials (PCM) are promising refrigeration recovery

4

methods. During hydration, PCM transforms from solid to liquid, providing cold duty and the

5

other way around, stores the refrigeration when hydrate dissociates.(Song et al. 2015) Yet,

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currently no successful application of the above method is reported. But if the energy recovery

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method were utilized, more water content would no doubt become more favorable.

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In sum, dispute exists in the choice of water content. First, the TAC variance with respect to

9

water content is not significant. Second, even if minor water content makes operation easier,

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however, on the other hand, development in refrigeration recovery technology may turn larger

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water content to the advantage. Consequently, we cannot state a definite conclusion on which is

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preferable temporarily.

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4.2.3 Temperature

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As is shown in Table 8, shift in temperature has little effect on TAC and other parameters. The

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insignificance can provide some design space for other concerns. For example, for the sake of process intensification, 0℃ is preferred in that lower temperature can boost hydrate formation. 4.2.4 Number of stages

The number of stages has always been an important factor for absorption process design.

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(McCabe et al. 1993) The effect of increasing stage number becomes less noticeable as the number

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grows larger. A heuristic is to choose a stage number which is so large that when we add another

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stage the recovery of key component does not evidently change. Therefore, we fix the emulsion

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flow rate and investigate the change of ethylene recovery with respect to the number of stages. As

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can be seen from Figure 3, ethylene recovery becomes larger with the increment of stage number.

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When stage number is larger than 10, the change of recovery becomes less than 0.1%.

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4.3 Composition distribution

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Figure 4a, b show the composition distribution of some key components along the stages in oil

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and hydrate phase, respectively. As is shown in Figure 4a, ethane and ethylene content in oil has

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grown along the direction of emulsion flow, whereas the absorption of methane and nitrogen is not

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significant. Figure 4b shows the composition variance of hydrate, it is found that the increase of

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ethane and ethylene content is not as obvious as that of oil. For better illustration, Figure 5 is

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drawn to illustrate the C1/C2 ratio in both oil and hydrate. It is noticed that in stage-1, the C1/C2

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ratio in oil is slightly higher than hydrate. However, as the slurry flows to downward stages,

11

C1/C2 ratio in oil drops rapidly from 1.1 to 0.2, while in hydrate the C1/C2 ratio remains around

12

1.0.

te

d

M

an

us

cr

4

This phenomena can be explained by the special structure in hydrate. As is known, SII hydrate

14

has two cage types. One is large enough to capture ethane or ethylene. The other is comparatively

15 16 17 18 19

Ac ce p

13

small, which can accommodate methane and nitrogen, but unable to hold ethylene or ethane.(Sloan 2003) We can see from Figure 7 that the large cages are already fully occupied even in the first stage. Thus there is no more room for more ethylene in the hydrate. Consequently, though effective for the oil phase, the multistage configuration fails to improve composition of ethylene in hydrate.

20

The analysis into composition distribution informs us that due to the special structure, hydrate is

21

suitable to recover certain components from its low concentration mixture but may be not as

22

effective in improving composition profile as oil.

Page 17 of 48

1

4.4 Alternative configuration An alternative flow chart configuration is considered, which enables the process to control

3

enriched dry gas C1/C2 ratio. As is shown in Figure 8, in R-3 hydrate is dissociated at 4.25MPa by

4

raising temperature. The gas released is split by SP into two streams. One stream is returned to

5

stage-15 (stage number is increased to 15) just like a stripping gas for an absorber. The other goes

6

to mix with gas produced by R-1 and R-2. The dry gas still feed at stage-11. The optimized profile

7

for different split fraction is shown in Table 9. As can be seen, the C1/C2 ratio has fallen from

8

0.511 to 0.178. However, the TAC also raises significantly from 0.604 to 0.894 M$/yr. To

9

summarize, the alternative process can offer a much lower C1/C2 ratio but at the expense of higher TAC.

11

4.5 Cost distribution

M

10

an

us

cr

ip t

2

Here we analyze the cost distribution of the proposed AHHP and compare it to the C4

13

absorption process. As we can see from Table 10 a, in C4 absorption, re-boiler duty of B-2, 3 takes

14

43% of the TAC.

16 17 18

te

Ac ce p

15

d

12

For AHHP, refrigeration and compressor duty add up to 50% of the TAC, as is shown in Table

10 b. It is found refrigeration cost originates from removal of the heat generated during hydrate formation, as shown in Table 11. On the other hand, the process has a gas stream produced at a lower pressure, which requires recompression before further treatment. For this reason and

19

because electricity is comparatively dear (0.04$/kWh), therefore, compressor duty becomes the

20

other expense consuming item. It is also noticed that equipment cost for AHHP is higher. The

21

main reason is the compressor (0.95M$) that recompress a product stream.

22

From the analysis above, we find steam utility is the main cost in C4 absorption and therefore

Page 18 of 48

we may try to integrate the reboilers with other units in the refinery. For AHHP, as refrigeration

2

and compression is the main cost, refrigeration recovery methods and pressure equalization

3

methods should be adopted for further improvement.

4

5. Conclusions

ip t

1

In this work, we put forward ACM® for the modeling and simulation of complex hydration

6

process. This approach is found to be robust and convenient. With the help of ACM®, a multistage

7

AHHP for ethylene recovery from dry gas is modeled and simulated. We found the optimal

8

operating conditions for AHHP. Influence of the key parameters are discussed. It is found

9

emulsion flow rate is controlling the ethylene recovery. Influence of temperature and water

10

content on TAC is found insignificant. As a consequence, the two parameters should be

11

determined by other concerns. By analyzing composition distribution along the stages, we can see

12

that ethylene and ethane composition in hydrate cannot improve as effectively as that of oil, which

13

is attributed to the special structure in hydrate. Economic evaluations on AHHP and C4 absorption

14

show the former has a lower TAC. The cost distribution indicates most cost of AHHP comes from

16 17 18

us

an

M

d

te

Ac ce p

15

cr

5

refrigeration and compression. Adoption of refrigeration recovery methods is recommended for AHHP’s further improvement. Acknowledgments:

Financial support for this investigation was received from the National Key Basic Research

19

Program of China (No. 2012CB215005), the 2013 China−Europe Small-Sized and Medium-Sized

20

Enterprises Energy Saving and Carbon Reduction Research Project (No. SQ2013ZOA100002),

21

and the National Natural Science Foundation of China (No. 21336007).

22 23

Appendix

Page 19 of 48

1

Introduction to cryogenic process: In order to compare different process, we are supposed to make different process at the same

3

end point. As the two process introduced above may have different enriched dry gas composition,

4

a simplified cryogenic process is presented in Figure 8. By adding such a cryogenic process the

5

two flow charts can now share the same finish point, at which methane/ethylene ratio satisfies the

6

requirement of polymerization (methane/ethylene less than 0.05mol%). Cost of the cryogenic

7

distillation part is also incorporated in our evaluation. The flow chart of cryogenic process is a

8

simplified one, in which we ignored some unit operations that are considered identical and

9

therefore make little difference in economic comparison, such as alkaline wash tower and de-NOx

cr

us

an

tower.

M

10

ip t

2

As is shown in Figure 8, the enriched dry gas is chilled through several stages. After each stage,

12

the mixtures are flashed. The liquid phase is lined to DM and the vapor phase is further cooled in

13

the next stage. This stage by stage refrigeration pattern is devised for energy saving. Finally, after

14

the final flash operation, the vapor phase is mixed with gas out of DM condenser. The mixture

16 17 18 19 20 21 22

te

Ac ce p

15

d

11

then goes to expander to provide refrigeration. Operating conditions are listed in Table 14.

Equipment cost ($):

Vessel investment= 17640·D1.066L0.802

(20)

Heat Exchanger investment= 7296·A0.65

(21)

Compressor= 1293·517.3·3.11· (0.746Wcmp)0.82/280

Cryogenic Heat Exchange investment = 30000+750·A0.81

(22) (23)

Space between trays are calculated as 0.64m. As is shown in equation (23), cryogenic (-60~-101

23

℃) heat exchangers are more expensive than ordinary heat exchangers.

24 25 26

Chen-Guo equation parameters

Page 20 of 48

te

d

M

an

us

cr

ip t

Parameters for Chen-Guo equation can be found in Table 15.

Ac ce p

1 2

Page 21 of 48

Nomenclature Meaning (unit of measurement)

A

Heat exchange area (m2)

C

Collection of components

Cj

Langmuir constants for component j

D

Vessel diameter (m)

fjL

Fugacity of j in liquid (bar)

fjV

Fugacity of j in vapor (bar)

Hhy

Molar enthalpy of hydrate (GJ/kmol)

HL

Molar enthalpy of liquid (GJ/kmol)

HV

Molar enthalpy of vapor (GJ/kmol)

HYin

Hydrate mole flow rate going in (kmol/h)

HYout

Hydrate mole leaving (kmol/h)

L

Vessel length (m)

Lin

Oil mole flow rate going in (kmol/h)

Lout

Oil mole flow rate leaving (kmol/h)

P

Pressure (bar)

Q

cr

us

an

M

d

te

P0

ip t

Symbol

Operating pressure (bar) Heat removed from stage (GJ)

Ac ce p

1 2

T

Temperature (℃)

T0

Operating temperature (℃)

Vin

Vapor mole flow rate going in (kmol/h)

Vout

Vapor mole flow rate leaving (kmol/h)

Wcmp

Working load of compressor (kw)

Wj

Mole flow rate of j in water (kmol/h)

xjin

Mole composition of oil liquid going in

xjout

Mole composition of oil liquid leaving

yjin

Mole composition of vapor going in

yjout

Mole composition of vapor leaving

zjin

Mole composition of hydrate going in

Page 22 of 48

zjout

Mole composition of hydrate leaving

1 Reference:

3

Adhikari S, Fernando S. Hydrogen membrane separation techniques. (2006) Ind Eng Chem Res.

4

ACS Publications; 45(3):875–81.

5

Chen GJ, Guo TM. (1998) A new approach to gas hydrate modelling. Chem Eng J.

6

71(March):145–51.

7

Englezos P, Kalogerakis N, Dholabhai PD, Bishnoi PR. (1987a) Kinetics of formation of methane

8

and ethane gas hydrates. Chem Eng Sci. 42(11):2647–58.

9

Englezos P, Kalogerakis N, Dholabhai PD, Bishnoi PR. (1987b) Kinetics of gas hydrate formation from mixtures of methane and ethane. Chem Eng Sci. 42(11):2659–66.

11

Eslamimanesh A, Mohammadi AH, Richon D, Naidoo P, Ramjugernath D. (2012) Application of

12

gas hydrate formation in separation processes: A review of experimental studies. J Chem

13

Thermodyn. 46:62–71.

14

Huang HY, Padin J, Yang RT. Comparison of π-complexations of ethylene and carbon monoxide

15

with Cu+ and Ag+. (1999) Ind Eng Chem Res. ACS Publications; 38(7):2720–5.

16

Kondo W, Ohtsuka K, Ohmura R, Takeya S, Mori YH. (2014) Clathrate-hydrate formation from a

17

hydrocarbon gas mixture: Compositional evolution of formed hydrate during an isobaric

18

23 24

Condition Measurements for CH4+C2H4+C2H6+N2 Gas Mixtures in Diesel Oil + Water +

25

Anti-agglomerant Emulsion System, Proceedings of the 13th International Conference on

26

Properties and Phase Equilibria for Product and Process Design.

27

Liu H, Mu L, Wang B, Liu B, Wang J, Zhang X, et al. (2013) Separation of ethylene from refinery

28

dry gas via forming hydrate in w/o dispersion system. Sep Purif Technol. 116:342–50.

29

Ma C, Chen G, Wang F, Sun C, Guo T. (2001) Hydrate formation of (CH4 + C2H4) and (CH4 +

19 20 21 22

te

d

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an

10

Ac ce p

us

cr

ip t

2

semi-batch hydrate-forming operation. Appl Energy. 113:864–71. Lei Si WZ. Summary on running of 110kt/a catalytic dry gas ethylene recovery unit. (2013) Qilu petrochemical Technol. 4(41):274–7. Li Y, Luo H. Li Y, Luo H. (2015) Integration of light hydrocarbons cryogenic separation process in refinery based on LNG cold energy utilization. Chem Eng Res Des. 93:632-9. Liu B, Liu H, Mu L, Sun C, Chen G. (2013) Compositional Analysis and Hydrate Dissociation

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C3H6) gas mixtures. 191:41–7.

2

Ma Q, Chen G, Zhang L. (2008) Vapor-hydrate phases equilibrium of (CH4+C2H6) and (CH4+C2H4)

3

systems. Pet Sci. 5:359–66.

4

Ma Q, Huang Q, Chen G, Wang X, Sun C, Yang L. (2013) Kinetic and phase behaviors of catalytic

5

cracking dry gas hydrate in water-in-oil emulsion. Chinese J Chem Eng. 21(3):295–300.

6

Ma QL, Chen GJ, Sun CY. (2013b) Vapor-liquid-liquid-hydrate phase equilibrium calculation for

7

multicomponent systems containing hydrogen. Fluid Phase Equilib. 338:87–94.

8

Ma Q-L, Chen G-J, Sun C-Y. (2013c) Vapor–liquid–liquid–hydrate phase equilibrium calculation

9

for multicomponent systems containing hydrogen. Fluid Phase Equilib. Elsevier; 338:87–94. McCabe WL, Smith JC, Harriott P. (1993) Unit operations of chemical engineering. McGraw-Hill

11

New York;

12

Mu L, Li S, Ma Q-L, Zhang K, Sun C-Y, Chen G-J, et al. (2014) Experimental and modeling

13

investigation of kinetics of methane gas hydrate formation in water-in-oil emulsion. Fluid Phase

14

Equilib. Elsevier; 362:28–34.

15

Murshed MM, Schmidt BC, Kuhs WF. (2010) Kinetics of methane-ethane gas replacement in

16

clathrate-hydrates studied by time-resolved neutron diffraction and Raman spectroscopy. J Phys

17

Chem A. 114:247–55.

18

Naeiji P, Varaminian F. (2013) Experimental study and kinetics modeling of gas hydrate formation

19

24 25

Rahimi N, Karimzadeh R. (2011) Catalytic cracking of hydrocarbons over modified ZSM-5

26

zeolites to produce light olefins: A review. Appl Catal A Gen. 398(1):1–17.

27

Sadeghbeigi R. (2012) Fluid catalytic cracking handbook: An expert guide to the practical

28

operation, design, and optimization of FCC units.

29

Salerno D, Arellano-Garcia H, Wozny G. (2011) Ethylene separation by feed-splitting from light

30

gases. Energy. 36(7):4518–23.

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d

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an

10

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ip t

1

of methane–ethane mixture. Journal of Non-Equilibrium Thermodynamics. 38(3):273-86. Naeiji P, Mottahedin M, Varaminian F. (2014) Separation of methane–ethane gas mixtures via gas hydrate formation. Separation and Purification Technology. 123:139-44. Padin J, Yang RT. (2000) New sorbents for olefin/paraffin separations by adsorption via π-complexation: synthesis and effects of substrates. Chem Eng Sci. Elsevier; 55(14):2607–16.

Aspen Technology Inc. http://www.aspentech.com. Cambridge, MA, USA. 2012.

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Seider WD, Seader JD, Lewin DR. (2009) Product & Process Design Principles: Synthesis,

2

Analysis and Evaluation. John Wiley & Sons;

3

Sloan ED. (2003) Fundamental principles and applications of natural gas hydrates. Nature.

4

426(November):353–63.

5

Sloan Jr ED, Koh C. (2007) Clathrate hydrates of natural gases. CRC press;

6

Song X, Xin F, Yan H, Li X, Jia H. (2015) Intensification and kinetics of methane hydrate

7

formation under heat removal by phase change of n‐tetradecane. AIChE J. 61(10): 3441-3450.

8

Subramanian S, Kini R a., Dec SF, Sloan ED. (2000) Evidence of structure II hydrate formation

9

from methane+ethane mixtures. Chem Eng Sci. 55:1981–99.

Sun C-Y, Ma C-F, Chen G-J, Zhang S-X. Sun CY, Ma CF, Chen GJ, Zhang SX. (2007)

11

Experimental and simulation of single equilibrium stage separation of (methane+ hydrogen)

12

mixtures via forming hydrate. Fluid Phase Equilibria. 261(1):85-91.

13

Turner DJ, Miller KT, Sloan ED. (2009) Methane hydrate formation and an inward growing shell

14

model in water-in-oil dispersions. Chem Eng Sci. Elsevier; 64(18):3996–4004.

15

Watanabe S, Saito K, Ohmura R. (2011) Crystal growth of clathrate hydrate in liquid water

16

saturated with a simulated natural gas. Cryst Growth Des. ACS Publications; 11(7):3235–42.

17

Zhang LW, Chen GJ, Guo XQ, Sun CY, Yang LY. (2004) The partition coefficients of ethane

18

between vapor and hydrate phase for methane +ethane +water and methane +ethane +THF +water

19 20

te

d

M

an

10

Ac ce p

us

cr

ip t

1

systems. Fluid Phase Equilib. 225:141–4.

Page 25 of 48

Table 1 FCC dry gas composition Composition (mol %)

H2

31.79

N2

11.88

O2

2.97

CO2

2.01

CO

0.77

CH4

24.65

C 2 H6

12.14

C 2 H4

12.56

C 3 H8

0.15

C 3 H6

1.05

cr

Ac ce p

te

d

M

an

2 3

ip t

Component

us

1

Page 26 of 48

1 2

Table 2 Operating condition of C4 absorption Pressure MPa

Temperature ℃

Tray Number

B-1 B-2 B-3 HX-1 HX-2 HX-3

4.25 4.25 2.6 2.6 2.6 2.6

18 18~118 18~120 in 120 out 40 in 40 out 18 in 40 out 18

15 20 40 -

Ac ce p

te

d

M

an

us

cr

3 4

ip t

Unit operation

Page 27 of 48

1 2 3

Table 3 Operating condition of AHHP Temperature (℃)

Tray Number

RCT

4.25

0

-

Absorber

4.25

0~10

11

R-1

3.3

40

-

R-2

0.1

40

-

3.3

-

-

HX-1

4.25

In 40 Out 0

-

HX-2

4.25

In 40 Out 0

-

Ac ce p

te

d

M

an

us

4 5

cr

CMP outlet

ip t

Pressure (MPa)

Unit operation

Page 28 of 48

1 2 3

Table 4 Design specifications and parameter ranges Parameters

Specifications and ranges = 0.9 ≤0.05 = 4.25 MPa >0 ≥0℃ 10-30 v%

cr

ip t

Ethylene recovery Final Product Methane/ (methane + ethylene) Absorption pressure Emulsion flow rate Absorption temperature Water volume fraction

Ac ce p

te

d

M

an

us

4 5 6

Page 29 of 48

Table 5a. Optimized profile for AHHP (ethylene recovery=90%) 0

Water content v%

30

Oil flow rate kmol/h

105.7

Ethylene recovery in AHHP

0.919

C1/C2 ratio

0.511

Equipment Cost M$/yr

0.149

Utility Cost M$/yr

0.455

TAC Cost M$/yr

0.604

us

3 4 5

Absorption Temperature

Table 5b. Profile of C4 absorption (ethylene recovery = 90%) Value

638.6

B2 reboiler duty GJ/h

10.6

B3 reboiler duty GJ/h

5.67

B3 condenser duty GJ/h

-0.69

Ethylene recovery in absorption

0.911

C1/C2 ratio

0.401

Equipment Cost M$/yr

0.195

Utility Cost M$/yr

0.590

TAC Cost M$/yr

0.786

Ac ce p

te

M

n-butane flow rate kmol/h

d

an

Parameter

6 7

ip t

Optimized value

Parameter

cr

1 2

Page 30 of 48

1 2 3 4

Table 6 Optimized Profile for different ethylene recovery 0.89

0.9

0.91

0.92

Absorption Temperature ℃

0

0

0

0

0

Water content v%

30

30

30

30

30

Oil flow rate kmol/h

103.4

104.5

105.7

106.8

Ethylene recovery in enriched dry gas

0.899

0.909

0.919

0.930

C1/C2 ratio

0.511

0.511

0.511

0.511

Equipment Cost M$/yr

0.147

0.148

0.149

0.150

0.151

Utility Cost M$/yr

0.445

0.450

0.455

0.460

0.465

TAC M$/yr

0.592

0.598

0.604

0.610

0.616

108.0 0.940

cr

0.511

Ac ce p

te

d

M

an

5 6

ip t

0.88

us

Ethylene recovery

Page 31 of 48

Table 7 Optimized profile for different water content (ethylene recovery = 0.9) 10

20

30

Absorption Temperature ℃ Oil flow rate kmol/h

0

0

0

164.2

130.7

105.7

Ethylene recovery in enriched dry gas

0.917

0.918

0.919

C1/C2 ratio

0.350

0.430

0.511

Equipment Cost M$/yr

0.184

0.166

0.149

Utility Cost M$/yr

0.424

0.439

TAC M$/yr

0.608

0.605

0.455

0.604

Ac ce p

te

d

M

an

us

4 5

ip t

Water content v%

cr

1 2 3

Page 32 of 48

Table 8 Optimized profile for different absorption temperature (ethylene recovery = 0.9) 0

1

2

3

4

Water content v%

30

30

30

30

30

Oil flow rate kmol/h

105.7

106.9

108.1

109.3

110.5

DM condenser duty GJ/h

-0.159

-0.159

-0.158

-0.157

-0.157

Ethylene recovery in enriched dry gas

0.919

0.919

0.919

0.919

0.919

C1/C2 ratio

0.511

0.511

0.510

0.509

0.508

Equipment Cost M$/yr

0.149

0.150

0.150

0.151

0.152

Utility Cost M$/yr

0.455

0.456

0.457

0.458

0.459

TAC M$/yr

0.604

0.606

0.609

0.610

0.608

cr

Ac ce p

te

d

M

an

5 6

ip t

Absorption Temperature ℃

us

1 2 3 4

Page 33 of 48

1 2

Table 9 Optimized profile for different split fraction

C1/C2 ratio Equipment Cost M$/yr Utility Cost M$/yr TAC M$/yr

0.2

0.4

0.6

0.8

1

0 30 113.2 0.917 0.472 0.158 0.475 0.633

0 30 122.5 0.913 0.425 0.169 0.502 0.671

0 30 134.5 0.910 0.366 0.183 0.539 0.723

0 30 151.1 0.906 0.288 0.203 0.590 0.793

0 30 176.9 0.901 0.178 0.233 0.661 0.894

Ac ce p

te

d

M

an

us

3 4

ip t

Ethylene recovery in enriched dry gas

0 0 30 105.7 0.919 0.511 0.149 0.455 0.604

cr

Split fraction Temperature ℃ Water content v% Oil flow rate kmol/h

Page 34 of 48

Cost(M$/yr)/ Percentage

Absorbent refrigeration

0.129/ 16%

B-2,3 reboiler utility

0.335/ 43%

Cryogenic refrigeration

0.112/ 14%

Annualized investment

0.195/ 25%

cr

ip t

Item

Table 10 b Cost distribution analysis of AHHP

us

4 5 6 7

Table 10 a Cost distribution analysis of C4 absorption

Item

Cost(M$/yr)/ Percentage

Absorption refrigeration

0.220/ 35%

Compressor utility

M

Cryogenic refrigeration

an

1 2 3

Annual investment

0.147/ 24% 0.149/ 25%

Ac ce p

te

d

8 9

0.087/ 14%

Page 35 of 48

Table 11 Enthalpy change of hydrate formation Component

ΔH kJ/mol

N2 CH4 C2H6 C2H4

65 54 74 74

ip t

1 2

Ac ce p

te

d

M

an

us

cr

3 4

Page 36 of 48

1 2

Cost 6.67 $/GJ 4.37 $/GJ 2.56 $/GJ 3.3 $/GJ 4.6 $/GJ 6.6 $/GJ 8.5 $/GJ 10.5 $/GJ 0.04 $/kWh

cr

Utilities Steam 234 oC Steam 180 oC Steam 138 oC Refrigeration 4.4 oC Refrigeration -12.2 oC Refrigeration -34.5 oC Refrigeration -67.8 oC Refrigeration -101 oC Electricity

ip t

Table 12 Utility cost

Ac ce p

te

d

M

an

us

3 4

Page 37 of 48

1 Table 13 Heat transfer coefficient in economic evaluation kWm-2K-1

Gas-Gas

0.17

Gas-Condensing vapor

0.28

Gas-vaporizing liquid

0.28

Liquid-Liquid

0.57

Liquid-Condensing vapor

0.85

Liquid-Vaporizing liquid

0.85

Ac ce p

te

d

M

an

us

3 4

ip t

Type

cr

2

Page 38 of 48

1 Table 14 Operating condition for cryogenic process CR-HX1 CR-HX2 CR-HX3 DM Turbine Refrigerant 1 Refrigerant 2 Refrigerant 3

Pressure

Temperature ℃

3.3MPa 3.3MPa 3.3MPa 3.3MPa 0.3MPa -

Tray Number

Dry gas target -29.5 Dry gas target -62.5 Dry gas target -96 -98~-5 -130 -34.5 -68.5 -101

Ac ce p

te

d

M

an

us

3 4

60 -

ip t

Unit operation

cr

2

Page 39 of 48

1 2

Table 15 Parameter value for Chen-Guo equations Antoine constant

SII structure

Gas

X (bar)

Y (K)

Z (K)

A′ (bar)

B′ (K)

H2

5.64E−06

120.775

253.1

1.00E+23

0

O2

9.50E−06

2452.29

1.03

4.32E+23

−12505.00

−0.35

N2

4.32E−06

2472.37

0.64

6.82E+23

−12770.00

−1.10

CH4

2.30E−06

2752.29

23.01

5.26E+23

−12955.00

4.08

1.65E−06

2799.66

3.45E+23

−12570.00

6.79

3.77E+21

−13841.00

0.55

C2H6b

3.99E+21

−11491.00

30.4

C3H8b

4.10E+23

−12312.00

39

3

a.

4 5

an

ip t

us

cr

0

M te

15.9

Ac ce p

C2H4b

d

COa

CO2

C′ (K)

There is currently no CO parameter for Chen-Guo equation in literature. The parameters of CO2 is used for

CO. b.

For the blank spaces, the parameter is considered to be 0.

6 7

Page 40 of 48

1

Figures Off gas

Feed dry gas

Enriched dry gas

MX-1

cr

HX-3

B-3

us

B-2

ip t

B-1

2

Figure 1. Absorption process flow chart

d

M

(B-1: absorber; B-2: stripper; B-3: desorption column; HX-1, 2, 3: heat exchanger 1, 2, 3; MX-1: mixer tank)

te

4 5

HX-2

Ac ce p

3

Pump

an

HX-1

Page 41 of 48

1 2 Off Gas 2

Off Gas 1

Hydrate

Water

Hydration P1,T1

Dissociation

Dissociation

Water

Water

us

3 4

Product Gas

ip t

Product Gas

Hydration P0,T0

cr

Feed gas and water

Stage-n

Feed dry gas

Stage-(n-1)

an

Figure 2a Conceptual Design of hydration process

Stage-(n-2)

Gas

Off-gas

...

Slurry

M

HX-2

Stage-1

5 6 7 8

Ac ce p

te

d

Hydrate in oil Slurry

Enriched dry gas

R-1

CMP

R-2 Emulsion reuse

Pump

HX-1

Figure 2b AHHP flow chart

(CMP: Compressor; HX: Heat exchanger; R-1, 2: Releaser 1, 2)

Page 42 of 48

1 2

0.930

ip t

0.925

0.915

cr

0.910 0.905

us

Ethylene recovery

0.920

0.900 0.895

5

6

an

0.890 7

8

9

10

11

Stage numbers

3

te

d

M

Figure 3 Effect of stage number on ethylene recovery

Ac ce p

4 5 6

Page 43 of 48

1 2

0.18

ethylene nitrogen

0.16

methane ethane

ip t

0.12 0.10 0.08

cr

Mole fraction in oil

0.14

0.06

0.02 0.00 0

1

2

3

4

5

6

7

8

9

10

11

12

11

12

an

Stage

us

0.04

3 4

M

Figure 4 a Composition Distribution in oil

0.6

ethylene nitrogen

d

0.4

te

0.3

0.2

Ac ce p

Mole composition in hydrate

0.5

methane ethane

0.1

0.0

0

5 6 7

1

2

3

4

5

6

7

8

9

10

Stage

Figure 4 b Composition in hydrate (water free basis)

Page 44 of 48

Oil Hydrate

1.4 1.2

ip t

0.8 0.6

cr

C1/C2 ratio

1.0

0.4

0

1

2

3

4

5

6

7

Stage

10

11

12

te

d

M

Figure 5 C1/C2 ratio of oil and hydrate

Ac ce p

2 3

9

an

1

8

us

0.2

Page 45 of 48

1

ip t cr

0.95

0.90

us

Cage occupancy

1.00

0.85

0.80 1

2

3

4

5

an

Sum of ethane and ethylene

6

7

8

9

10

11

M

Stage 2

te

d

Figure 6 Sum of ethane and ethylene occupancy in large cage

Ac ce p

3 4 5 6

Page 46 of 48

Feed dry gas HX-2 Stage-(n-1)

Stage-1 ...

Off-gas

...

ip t

Stage-n

Stage-i

cr

SP

Enriched dry gas

us

R-3 CMP

R-1

HX-1

M

Figure 7. An alternative configuration for AHHP

d

(R-3: hydrate dissociation tank, 4.25MPa, 40 ℃; SP: gas splitter)

te

3 4

Pump

Ac ce p

1 2

an

R-2

Page 47 of 48

1 2

Off gas

CR-HX2

Refrigerant 1

CR-HX3

Refrigerant 2

Turbine

Refrigerant 3

Enriched dry gas feed

F-3

DM

3

d

7

(CR-HX1, 2, 3: cryogenic heat exchanger 1, 2, 3; F-1, 2, 3: flash tank 1, 2, 3; DM: de-methane column)

te

6

Figure 8. Simplified cryogenic process

Ac ce p

4 5

Product

M

an

us

cr

F-2

F-1

ip t

CR-HX1

Page 48 of 48