Effect of heavy hydrocarbons on CO2 removal from natural gas by low

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Separations 2

Effect of heavy hydrocarbons on CO removal from natural gas by low-temperature distillation Giorgia De Guido, Mattia Riccardo Fogli, and Laura A. Pellegrini Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b00614 • Publication Date (Web): 03 May 2018 Downloaded from http://pubs.acs.org on May 5, 2018

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Industrial & Engineering Chemistry Research

Effect of heavy hydrocarbons on CO2 removal from natural gas by low-temperature distillation Giorgia De Guidoa,*, Mattia R. Foglia, Laura A. Pellegrinia aDipartimento

di Chimica, Materiali e Ingegneria Chimica “G. Natta”, Politecnico di Milano, Piazza Leonardo da Vinci 32, 20133 Milano, Italy Corresponding author: Giorgia De Guido ([email protected]; phone: +39 02 2399 3260; fax: +39 02 2399 3280)

Abstract Natural gas is expected to have the most important growth among fossil fuels being the cleanestburning one and also taking into account the ever-increasing role LNG is playing within the gas sector. To meet the increasing demand for natural gas, it will be necessary to develop both new unconventional gas fields and conventional gas fields that have not been developed yet because of the high acid gas content. The conventional acid gas removal technology (chemical absorption into aqueous alkanolamine solutions) results to be energy-intensive when applied to low-quality gas reserves. This is why, more recently, low-temperature processes have gained attention as an alternative to traditional CO2 removal technologies. An important advantage the use of such novel technologies poses consists in the separation of CO2 as a liquid at high pressure, which is very interesting for further applications, such as geo-sequestration or Enhanced Oil Recovery. In this scenario, this work investigates the use of a new low-temperature CO2 removal process that is based on dual pressure distillation. This simulation study shows how the presence of hydrocarbons heavier than methane can affect the energy requirements of the overall process, also including natural gas liquids fractionation.

Keywords: natural gas, carbon dioxide, acid gas removal, low-temperature distillation, natural gas liquids, energy saving

1.

Introduction

The energy market still depends heavily on fossil fuels which means that they will continue to be used in the future1 in order to meet the expected 37% increase in energy demand by 2040.2 ACS Paragon Plus Environment 1

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Compared with other fossil fuels, natural gas is generally considered as the best energy resource for the transition from high-carbon fossil fuel energy to renewable energy because of its low emissions and abundant global reserves. Raw natural gas from gas reservoirs contains compounds other than methane, including hydrocarbons, CO2 and/or H2S, N2, He, etc. Among them, H2S and CO2 are responsible for its sour or acidic nature and pose some challenges to natural gas processing, requiring special construction materials to be used. Focusing the attention on high CO2-content gases, currently there exist many undeveloped or underdeveloped gas reserves that are characterized by a significant fraction of CO2, ranging from 15% to 80%, which limits their economic and environmental viability.3 An example is given by the LaBarge field,4 located in Wyoming (USA): it was discovered in 1963 but its production was delayed until 1986 because of the challenging gas composition (i.e., 65% CO2, 21% CH4, 7% N2, 5% H2S, 0.6% He) that made it the lowest hydrocarbon content natural gas commercially produced in the world. An even higher CO2 content (70 %) can be found in the gas extracted from the Natuna field, located 225 km northeast of the Natuna Islands (Indonesia’s northernmost territory in the South China sea).5 To be considered for exploitation and development, the gas produced from these low-quality gas reserves must be processed by separating out the CO2 and/or H2S to give a natural gas product that meets commercial specifications. Current methods used to separate CH4/CO2 mixtures are chemical or physical absorption, adsorption, membrane permeation and low-temperature distillation. Chemical absorption processes rely on the use of amine-based solvents (e.g., MDEA,6 DGA® agent,7 proprietary/licensed MDEA): they permit to efficiently remove CO2 and/or H2S, but they require significant amount of heat for solvent regeneration, which is typically higher than that involved in physical absorption processes. The latter ones, based on the use of physical solvents (e.g., DEPG, PC, NMP, refrigerated methanol),8 allow the removal of more organic sulfur, but present a drawback due to the coabsorption of hydrocarbons. Historically, chemical and physical absorption systems have suffered a number of economic disadvantages when the feed stream contained high amounts of carbon dioxide, including high utility consumption, degradation of the solvents and high maintenance costs due to their corrosive nature.9 Adsorption-based techniques have been reported to show promise due to their inherent simplicity, low operational requirements, ease of control and high efficiency provided that an appropriate adsorbent is used, which must have a high adsorption capacity, large specific surface area and low heat for regeneration.10 However, adsorption by solids has generally been economically practical only when the feed gas contained relatively small amounts of CO2 and

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a complete removal was required.9 Membranes rely on the relative permeation rate of CO2 compared to the other components of natural gas. They are typically quoted in the literature because they do not require any phase transformation and the necessary process equipment is very simple.11 Membranes are usually installed for small size applications and remote locations12 even if, more recently, they have been also used on larger scales.13 An issue to be taken into account is related to their selectivity: early membrane plants had difficulty in meeting a low hydrocarbon loss target as in the standard CO2 removal technique (i.e., amine absorption)14 or in low-temperature natural gas purification processes. However, better membranes and process designs have recently improved the competitiveness of membranes. Due to their challenging composition, low-quality gas fields will benefit from cheaper processing options that may foster their future exploitation and development. In this scenario, new separation methods that perform a cryogenic or low-temperature CO2 removal have recently received an increasing attention. In addition to lower energy consumptions with respect to conventional amine scrubbing, these processes also offer another advantage since they can separate pure CO2 as a liquid under pressure rather than in the gaseous state at near ambient pressure, thus making it relatively easy to pump underground for storage or to be used for Enhanced Oil Recovery (EOR) applications. The distinctive solidification properties of pure CO2 have led to the development of different processes solutions operated at low-temperature conditions, at which dry ice can form. As a consequence, some processes allow the formation of solid CO2 to favor its separation from the gas stream, whereas others adopt a specific solution (e.g., the addition of a solid-preventing agent or the operation according to a thermodynamic cycle which by-passes the Solid-Liquid-Vapor (SLV) locus of the system CH4-CO2) to avoid that. Some of these low-temperature CO2 removal processes currently available in the literature are based on distillation. One of them is the extractive distillation process proposed by Ryan and Holmes:9, 15, 16 it comprises adding a solids-preventing agent (e.g., one or more C2-C5 alkanes or other nonpolar liquids which are miscible with methane) to avoid that carbon dioxide freezes out and plugs the distillation column as well as other equipment thereby. Another distillation-based process is the Controlled Freeze ZoneTM (CFZTM) technology: 4, 11, 17, 18

contrarily to the Ryan-Holmes process, it involves the formation of solid CO2 by means of a

single step separation of CO2 and other contaminants from a natural gas stream without the use of solvents or adsorbents. The solidification is allowed to take place within the distillation tower in the so-called Controlled-Freeze-Zone, which has been specially designed to provide a specific section for solid CO2 to form and melt.19 More recently, another low-temperature distillation process has been proposed 20, 21 as an alternative to conventional amine scrubbing to be considered for CO2 removal from acid gases characterized by



a high CO2 content. It is based on a dual pressure low-temperature distillation method (Figure 1), and it has been conceived to perform a bulk removal of the impurities in a distillation column operated at a pressure (typically 50 bar) such that CO2 freeze-out is avoided. Methane production at commercial grade is performed in a second column operated at a lower pressure (typically 40 bar, anyway lower than the critical pressure of methane). This operation mode allows to bypass the SLVE locus of the CH4-CO2 binary system by performing a bulk removal in the high-pressure section and to produce a methane-rich stream avoiding supercritical conditions by finishing the separation in the low-pressure section. [FIGURE 1 HERE]20

A further advantage of the Dual Pressure Low-Temperature distillation process (denoted by DPLT in the following) is the production of a very CO2-rich stream at 50 bar and in liquid phase: it can be used for Enhanced Oil Recovery (EOR) purposes with relatively low additional costs. In previous works, a thermodynamic investigation has been carried out (either experimentally22 or from a modelling point of view23) for a validation of the DPLT process, to properly investigate the possibility of solid phase formations, allowing the rigorous check on CO2 freeze-out. Moreover, an analysis on the operating costs of the DPLT process and of the conventional absorption one for different CO2 concentrations in the feed gas has shown that the DPLT process is more profitable for carbon dioxide contents above 10 mol %, the exact value depending on the geographic area where the gas reserve is located.24 Considering also different H2S concentrations, it has been observed that its presence favours the profitability of the DPLT process with respect to the conventional amine absorption one.24 In a previous work,25 the DPLT distillation process has been studied taking into account several possibilities to perform Natural Gas Liquids (NGLs) recovery, to establish if it is energetically profitable to recover NGLs before or after the purification unit on the basis of the content (0-25 mol%) of C2+ (where heavier hydrocarbons have been lumped as ethane) in the inlet gas stream. This work contributes to this discussion by investigating the effect of ethane, propane and n-butane on the energy consumptions of the DPLT distillation process in the case in which hydrocarbons heavier than methane are separated from methane and CO2 before the purification unit and further processed downstream. Energy requirements have been evaluated by adopting the net equivalent methane method concept,26-29 which provides a common basis to compare mechanical and thermal energies through the equivalent amount of methane that is needed to produce the required duties.


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2.

Process description

The natural gas separation train consists of a number of columns depending on the components. As shown in Figure 2, the process can be seen as composed of three different parts: the one (blue box in Figure 2) which separates hydrocarbons heavier than methane (i.e., ethane, propane and nbutane) from the CH4-CO2 mixture, the DPLT distillation (which performs the CH4-CO2 separation – red box in Figure 2) and the C2+ separation (which also includes n-butane recovery – green box in Figure 2).

[FIGURE 2 HERE]

The first column operates to separate the heavier compounds from the CH4-CO2 mixture as much as possible. In particular, a good separation between ethane and CO2 must be achieved since these two components are known to form an azeotrope, which represents a problem to be faced also in the CO2 recovery unit that is part of the DPLT distillation section. As reported in the literature, this challenge may be overcome by recycling some of the NGLs as solvent or entrainer,30 which helps breaking the CO2-ethane azeotrope by means of extractive distillation. In the process under study, n-butane is used as entrainer so that in the first column (T-100) CH4 and CO2 are recovered from the top (stream 4) while ethane and the other heavier hydrocarbons (propane and n-butane) are collected in the bottom product (stream 5). In all the case studies simulated in this work (as outlined in the following) a small amount of CO2 remains in this stream (of the order of 2e-5). The solvent recovery column (T-101) separates n-butane from the ethane-propane mixture. The entrainer recovered from the bottom (in which a negligible amount of CO2 is present) is, then, recycled back to the extractive distillation unit after withdrawing a certain fraction of the recovered stream as final product (stream 10). Before being fed to column T-100, the recovered n-butane is cooled down to 10°C since this turns out to improve the separation in column T-100, reducing the CO2 content in stream 6, methane losses in stream 5, and the duty required at the top condenser of column T-100. Ethane and propane remain in the top product (stream 6) of column T-101 and are further separated in column T-102. The CH4-CO2 stream recovered from the top of the extractive distillation unit (T-100) is sent to the DPLT distillation section, already shown in Figure 1 and described in the introductory section. The top product from column T-104 (stream 27) is used to cool down the CH4-CO2 mixture before it enters the DPLT distillation section.



3.

Methods

Data about gas fields compositions31-34 show that for associated gases the amount of the lighter species is higher, thus meaning that reasonable concentrations can be found according to the scale xethane > xpropane > xn-butane, with values ranging from 0 to 20% for ethane, from 0 to 15% for propane and from 0 to 5% for n-butane. The same also applies to non-associated gases, but the concentrations are far lower, with peaks of 5% for ethane and no more than 3% for propane. Gases with no ethane, propane and n-butane can be found as well. In this work, several simulations have been performed using Aspen Hysys® 35 as process simulator to take into account the aforementioned concentration values. The following section (3.1) contains a description of the simulations, while the next paragraph (3.2) outlines the equivalent methane method that has been used for their analysis and comparison on the basis of their energy requirements.

3.1.

Description of simulations

Simulations have been performed assuming the same molar flow rate (10000 kmol/h), temperature (25 °C) and pressure (35 bar) for the inlet gas stream. The inlet composition has been changed by increasing the content of a specific component at the expenses of the CO2 content. This criterion has been chosen considering that raw natural gas varies substantially in composition from source to source and it is not typically characterized by a constant CH4/CO2 ratio (contrarily, for example, to biogas). In particular, one component concentration at a time has been increased starting from the lowest reasonable value (2 mol% for ethane and propane and 1 mol% for n-butane) and reaching the highest one that is possible to find in real gas reserves. For each set of simulations, all the other components concentrations have remained unchanged (particularly, in all cases the methane inlet molar fraction is 45 mol%), except for the CO2 one. We have started to perform simulations considering a mixture containing 45 mol% CH4, 40 mol% CO2 and an equimolar composition for ethane, propane and n-butane. Then, we have varied the composition of the feed gas as previously explained. The first part of this work aims at analysing how a change in the concentration of ethane or propane in the raw feed gas affects the energy performances of the process under investigation. In all the simulations considered in this first part, the same n-butane recycle ratio (i.e., a fixed split ratio of 98% between stream 11 and stream 10 for TEE-100 in Figure 2) has been considered. The ethane and propane content has been modified by changing the mole fraction of one of them at a time (from 2 to 20 mol% for ethane, and from 2 to 15 mol% for propane), keeping the one of the others equal to 45 mol% for methane and to 5 mol% for the other two hydrocarbons.


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The feed tray in every column for the C2+ separation train has been optimized (by choosing the one that minimizes the column energy consumptions). The specifications for every column have been chosen so that the separation downstream of column T-100 is not compromised by thermodynamic issues (i.e., ethane-CO2 azeotrope in the DPLT section) and to obtain the highest allowable separation. Moreover, the optimum feed tray has been considered for all columns in all simulations, which has been determined by moving the feed tray location in order to find the one where the energy requirements are minimum. For column T-100 the number of theoretical trays has been fixed at 50: this allows to reduce the CO2 content in stream 6, requires a lower duty at the condenser of column T-100 and reduces methane losses from the bottom of column T-100. For this column the most restricting case has turned out to be the one having an inlet concentration of 5 mol% ethane, 2 mol% propane and 5 mol% butane (namely, the case study involving the lowest inlet propane concentration considered in this work): for this case, the recoveries of CO2 and ethane, respectively, in the top and bottom products are 99.99% and 99.2%. Indeed, higher propane inlet contents favour the CO2-ethane separation, allowing to reach a higher ethane recovery in the bottom product of column T-100. Therefore, these two specifications have been taken into account in the simulations of the first part of the work. For column T-101, which performs the entrainer (n-butane) regeneration, the number of theoretical trays has been fixed at 50 and a 99.9% recovery and a 99.5 mol% purity in the bottom product have been specified. The 99.9% recovery specified for n-butane together with the recycle ratio of 98%, leads to a massive n-butane flow coming from the bottom of column T-100. Column T-102 separates ethane from propane. The number of theoretical trays has been set equal to 50 in order to reduce the duty required at the top condenser and at the bottom reboiler. In addition to this, the recovery of ethane and propane has been specified and set equal to 99.99%. Good purity levels are, thus, achieved: >99.9% for ethane (with propane impurities) and from 95% to 97% for propane, due to n-butane losses from the top of column T-101. Columns T-103 and T-104 are equipped with, respectively, 20 and 15 theoretical trays and the given specifications for all the simulations refer to the methane molar fraction in the bottom product of column T-103 (set equal to 1e-4) and to the reflux ratio, R, in column T-104 (set equal to 1.4). In all simulations, a check has been performed on CO2 freeze-out in this section of the plant, even if the presence of some heavy hydrocarbons (mainly n-butane) in stream 17 enhances CO2 solubility in the liquid phase, as already stated in our previous work.25



For comparison purposes, all the specifications listed above are the same in every simulation performed to investigate the effect on the energy performances of the process due to an increase in the ethane or propane concentration in the raw feed gas. In the second part of the work, the n-butane recycle ratio has been considered different for each simulation. This choice has been made when performing the simulations for different n-butane inlet contents (in the range 1-5 mol%). Indeed, it was observed that the reduction from the starting value of 5 mol% keeping the recirculation ratio at 98% for all case studies significantly lowers the amount of n-butane recycled to column T-100 when the feed gas is characterized by low n-butane concentrations. Moreover, this leads to an ethane recovery not higher than 75% at the bottom of column T-100. This fact is not acceptable either for the loss of precious substance or for the azeotrope forming in column T-103 that prevents the CO2-methane separation. As a consequence, the recirculation ratio has been changed for each simulation in such a way that the entrainer molar flow in the bottom product of column T-100 is the same and equal to the minimum necessary to achieve the 99.2% ethane recovery (considered in all the previous simulations) in the most limiting case, that is the one for a 1 mol% butane concentration in the inlet raw natural gas stream. For this case, the molar flow rate of n-butane from the bottom of column T100 has resulted to be 9086 kmol/h and this value has been chosen and used in all the other simulations relative to higher n-butane inlet concentrations. The choice of keeping it constant, even if it could have been reduced when the raw feed gas contains higher n-butane concentrations, results from the decision not to introduce too many different parameters in every simulation and to be conservative later on with the energy assessments. For this set of simulations (characterized by a different split ratio for TEE-100 and by a fixed amount of n-butane in the bottom product of column T-100), the variation of the concentrations of ethane and of propane in the inlet raw gas has been also investigated. In particular, while parametrizing the ethane (10 mol% and 15 mol%) and nbutane (1, 3, and 5 mol%) inlet molar fraction, the one for propane has been varied from 2 mol% up to a value however lower than the one for ethane. The case characterized by a molar fraction of ethane equal to 5 mol% has not been taken into account because such a content is typical for nonassociated gases with very low propane content (no more than 1 mol%) and the process scheme can hardly be applied since the lack of heavier hydrocarbons would require the use of an external entrainer. The last set of simulations aims at estimating if another configuration is possible when the propane/n-butane separation is not necessary. Indeed, as proved in the following, it is one of the most expensive unit operations involved in the process scheme shown in Figure 2. The goal is the production of LPG (assumed to comprise propane and n-butane only) instead of two separate


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propane-rich and n-butane-rich streams and its direct use within the process as entrainer instead of high-purity n-butane.

3.2.

Equivalent methane calculation method

To evaluate the operating costs, the equivalent methane approach has been used. Its main features consist in the evaluation of the equivalent quantity of methane that must be burned or that can be recovered as a saving to allow a particular equipment to work. The amount of equivalent methane is determined as explained in the following, making use of its lower heating value (LHVCH4),36 of the efficiency of a methane-fired boiler,37 of the efficiency of a natural gas combined cycle power plant,38 and of the second principle efficiency39 (Table 1 summarizes the values used for these parameters).

Table 1. Parameters used for the calculation of the equivalent amount of methane produced or consumed by the process. Parameter

Value

LHVCH4 [MJ/kg]

50

ηboiler [-]

0.80

ηCC [-]

0.55

ηII,R [-]

0.60

Some energy requirements have been not included in the equivalent methane analysis: cooler E-100 and heater E-104, considering the possibility of energy integration within the process itself; cooler E-101b, which is not suitable for steam generation due to the low-temperature levels; cooler E-103 and T-101 condenser, in which cooling water can be used as coolant; T-103 reboiler, in which water can be used as heating medium. On the contrary, all the other energy requirements have been considered and can be classified in one of the following categories, depending whether they can be conceived as methane-consuming or methane-producing processes. The main methane-consuming processes are listed in the following. -

Use of low-pressure (LP) steam to supply heat to the process at temperatures higher than the

ambient one. This applies to the reboilers of columns T-100 and T-101 (where a temperature of ca. 130-140°C and 146°C is, respectively, reached), and to the reboiler of column T-102 (where a temperature of ca. 87°C is reached). The required steam is assumed to be produced by a CH4-fired ACS Paragon Plus Environment 9


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boiler. The amount of equivalent methane can be computed according to Eq. (1), given the boiler efficiency (ηboiler) and the fuel lower heating value reported in Table 1, and the thermal duty to be provided to the process (resulting from simulations).

m& CH 4 = -

Q& ηboiler LHVCH 4

(1)

Cooling duty produced by a proper refrigeration cycle to cool a process stream down at

temperatures lower than the ambient one. This applies to the heat exchangers E-102 (which cools the recovered n-butane stream from 25°C down to 10°C), E-105 (which condensates part of the top product from column T-103), and to the top condensers of columns T-100, T-102 and T-104 (where temperatures of ca. -20°C, 16°C and -87°C are, respectively, reached). These refrigeration cycles require mechanical work to be operated, which is considered to be provided by a combined-cycle plant powered by natural gas (methane). To calculate the amount of methane equivalent to the required mechanical duty, the coefficient of performance of the refrigeration cycle (COPR) can be calculated from Eq. (2), where ηII,R denotes the second principle efficiency (Table 1) – a measure of the system efficiency compared to the theoretical maximum possible one - and COPR,id stands for the coefficient of performance of an ideal Carnot cycle. It is given by Eq. (3) as a function of the temperatures of the hot and cold reservoirs, respectively TH and TL. The temperature of the hot reservoir has been set equal to 25°C, whereas that of the cold reservoir has been considered equal to the temperature (resulting from simulations) reached at the outlet of each heat exchanger used to cool a process stream down at temperatures below the ambient one.

COPR = COPR ,id ⋅η II , R COPR ,id =

(2)

1 TH −1 TL

(3)

By definition of the COPR, the mechanical duty (ܹሶ el) can be computed given the cooling duty to be removed from the system (Eq. (4)), resulting from simulations. W&el =

Q& COPR

(4)

Then, considering the definition of the efficiency of a combined-cycle (ηCC) (Eq. (5)), which is given by the ratio between the electric power produced by the gas turbine and the steam turbine and the heat received by the gas turbine (given by the product of the fuel flow rate and its LHV), the amount of equivalent methane can be determined according to Eq. (6).

ηCC =

W&el m& CH 4 LHVCH 4

(5)


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m& CH 4 = -

W&el Q& = ηCC LHVCH 4 COPRηCC LHVCH 4

(6)

Electric energy obtained by means of an equivalent CH4-fired combined-cycle power plant.

This applies to the P-100 pump (which pumps the bottom stream from column T-104 to the operating pressure of column T-103) and to the compressor K-100 that brings the pressure of the top product from column T-100 to the operating pressure of column T-103. The amount of methane equivalent to the required electric energy (ܹሶ௘௟ ) can be determined knowing the efficiency of the combined cycle and the fuel lower heating value (Eq. (6)). The main methane-producing processes are listed in the following. -

Heat recovered as energy available to produce LP steam (or, also, MP or HP steam,

depending on the temperature level) that can be exploited for further uses. This applies to the heat exchanger E-101a, which cools the bottom product from column T-101 from ca. 146°C down to 135°C. Once the thermal duty to be removed from the process is known from simulations, the amount of methane equivalent to it is determined according to Eq. (1), the same considered when heat has to be provided to the process by means of LP steam. The difference between the two cases lies in the fact that when heat is produced by the process under study, the equivalent amount of methane is considered to be produced rather than consumed by the process. The different process configurations have been compared on the basis of the net equivalent methane that must be consumed in order to operate the whole process. This is given by the difference between the amount of methane produced by the process (given by stream 30 in the process flow diagram illustrated in Figure 2) and by the methane-producing processes described above, and the amount of methane consumed by the methane-consuming processes listed above, Eq. (7). m& CH 4 , net = n&30,tot + m& CH 4 , produced − m& CH 4 ,consumed

(7)

Equivalently, the energy requirements will be illustrated in terms of the percentage of net equivalent methane in the raw natural gas stream that is required to run the process, Eq. (8).28

CH 4 % = 100 ⋅

4.

m& CH 4 ,consumed − m& CH 4 , produced m& CH 4 , feed NG

(8)

Results and discussion

In this section, the effect due to a change in the content of hydrocarbons heavier than methane in the raw feed gas will be discussed considering the trends of the net produced equivalent methane (for a fixed quantity in the inlet stream of 72193 kg/h). The impact on the duties required at columns condensers/reboilers will be taken into account, while thermal duties related to heat exchangers will not be reported because they are mainly affected by the molar flow rate passing through them, ACS Paragon Plus Environment 11


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having specified the temperature or associated variables (i.e., the dew point temperature for a stream with almost the same composition).

4.1.

Variation of the ethane and propane concentration in the feed gas at constant n-

butane recycle ratio Simulations for different ethane contents in the raw feed gas have shown that a higher content decreases the energy consumption at fixed ethane recovery in the bottom product of column T-100 and at fixed CO2 recovery in the top product of column T-100, as shown in Figure 3a and in Figure 3b (solid lines), respectively, for the condenser and reboiler duties. This can be explained considering the decreasing amount of CO2 fed to column T-100, having fixed its recovery in the top product from the same column. For the same reason, the same decreasing trend for the duties required at the condenser and reboiler of column T-100 is caused by the increase of the propane molar fraction in the feed gas, as also shown in Figure 3a and in Figure 3b (dotted lines), respectively.

[FIGURE 3 HERE]

Focusing on column T-101, when the ethane content is varied in the feed gas from 2 to 20 mol%, the duty at the condenser (Figure 4a, solid line) decreases monotonically, whereas that at the reboiler (Figure 4b, solid line) exhibits a minimum against the ethane content. The latter fact can be explained considering the duty at the reboiler to be proportional to the amount of product withdrawn from the bottom of column T-101, i.e. F7. This is given by Eq. (9), where RecT-101(n-C4) and x7(n-C4) denote, respectively, the recovery of n-butane and its molar fraction in the bottom stream 7. At fixed RecT-101(n-C4) and x7(n-C4), F7 depends upon the product of the total flow rate of stream 5 and the n-butane molar fraction in it (i.e., the product F5·x5(n-C4)), which exhibits a minimum trend as well, as a result of the fact that F5 and x5(n-C4), respectively increases and decreases at increasing ethane molar fraction in the raw feed gas. F7 =

RecT −101 ( n − C4 ) F5 ⋅ x5 (n − C4 ) x7 ( n − C4 )

(9)

A different trend has been found varying the content of propane in the feed gas (dotted lines in Figure 4a, b). Indeed, a higher amount of propane in the feed gas results in a higher amount of propane entering column T-101, which leads to higher energy consumptions to reach the desired purity level for n-butane in the bottom product stream. This result may be taken into account for further scopes, such as LPG production exploiting the beneficial role played by propane as entrainer


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in column T-100, which adds up to that of n-butane, thus eliminating a column (i.e., column T-102) in the process scheme shown in Figure 2.

[FIGURE 4 HERE] e

For T-102, which performs the separation between ethane and propane, both the condenser (Figure 5a, solid line) and reboiler (Figure 5b, solid line) duties increase as the ethane molar fraction in the feed stream increases. This also occurs by varying the propane molar fraction from 2 to 15 mol% (dotted lines in Figure 5a, b).

[FIGURE 5 HERE]

A linear decreasing trend is observed for the duty of the reboiler of column T-103 (Figure 6), which results from the reduced traffic in the column given by the lower CO2 content in the raw natural gas (at increasing ethane or propane content in it) and, thus, in stream 17 entering column T-103. However, this energy requirement is not considered as an expense for running the process since the temperature is such that water can be used as heating medium. [FIGURE 6 HERE]

Contrarily, no significant effect is observed on the condenser duty of column T-104 (Figure 7) having specified the same value for the reflux ratio for this column in all simulation cases. [FIGURE 7 HERE]

The overall effect due to a variation of the ethane content in the raw feed gas on the process energy consumption is illustrated in terms of net methane production, given by the difference between the amount of methane produced (as sum of the amount present in the raw feed gas to be treated and that produced by the methane-producing processes previously mentioned) and the amount of methane consumed by the process. As shown in Figure 8a, the net methane production increases at increasing ethane content (solid line) in the raw feed gas, since less energy is required to operate the process scheme shown in Figure 2. The same result can be also expressed in terms of the percentage of methane available in the raw feed gas, which must be used to run the process (Eq. (8)): as



illustrated in Figure 8b (solid line), it decreases from 34 to 31% as the ethane inlet concentration increases from 2 to 20 mol%. On the contrary, an increase in the propane content in the raw feed gas leads to higher energy consumptions and, thus, to a lower production of net equivalent methane (Figure 8a, dotted line) or, equivalently, to a higher percentage of methane in the raw feed gas to be used for running the process (Figure 8b, dotted line). This result can be explained considering that a higher amount of propane in the feed gas affects the recovery of the entrainer in column T-101, requiring a higher amount of energy for separation of propane (as overhead product, together with ethane) from nbutane (as bottom product).

[FIGURE 8 HERE]

4.2.

Variation of n-butane concentration in the feed gas

For the sensitivity analysis on the n-butane content in the feed gas in order to determine the value of the flow rate of n-butane in stream 5 to be used in all the simulations at different n-butane inlet concentrations, the worst case scenario has been taken into account, which is the one containing 1 mol% n-butane molar fraction in the inlet stream 2. In this way, a constant molar flow rate of 9086 kmol/h has been considered regulating the split ratio for TEE-100. The choice for the flow rate of nbutane in the bottom product of column T-100 is based on the results obtained by feeding the inlet gas on the 29th tray, due to the impossibility of simultaneously optimizing the feed tray of column T-100 and T-101 and the split ratio for TEE-100. Of course, as the amount of n-butane in the raw feed gas increases a lower split ratio for TEE-100 (defined as the flow rate of stream 11 with respect to that of stream 10 in Figure 2) permits to ensure the same amount of n-butane in the bottom product of column T-100 (Table 2).

Table 2. Split ratio (F11/F10) for TEE-100 in the scheme illustrated in Figure 2 as a function of the n-butane molar fraction in the feed stream determined to have the same n-butane flow rate in the bottom product of column T-100.

mol% n-butane feed gas

Split ratio TEE-100

1

0.99680

2

0.98592

3

0.97487

4

0.96377

5

0.95264 ACS Paragon Plus Environment 14

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Results show that the variation of the n-butane content in the feed gas mainly affects the energy performances of the first column (T-100), where the increasing n-butane content in the inlet stream favors the separation. As a result, the net methane production increases as the content of n-butane in the feed gas increases (Figure 9a) or, equivalently, the percentage of inlet methane to be used for running the process decreases (Figure 9b).

[FIGURE 9 HERE]

Therefore, it is possible to sum up the results discussed so far as follows: 1)

an increasing ethane content in the raw feed gas reduces the energy consumptions required

to perform the desired separation; 2)

an increasing propane content in the raw feed gas, though favoring the ethane recovery from

the initial mixture in column T-100, strongly affects the performances of the entrainer regeneration section determining a decrease in the net equivalent methane production; 3)

an increasing n-butane content in the raw feed gas allows to perform the desired separation

at lower energy expenses.

For the second set of simulations (at different split ratios for TEE-100) the ethane and propane contents in the feed gas have been also varied in order to investigate the combined effect due to the variation of the inlet concentration of more than one heavier hydrocarbon. As shown in Figure 10, both the favoring effect of ethane (the curves in Figure 10b are shifted upwards with respect to the ones in Figure 10a) and n-butane (the curves for higher n-butane concentrations lie above the ones for lower n-butane contents) is kept together with the propane issues concerning solvent regeneration (all curves illustrated in Figure 10a,b exhibit a decreasing trend as the propane inlet concentration increases).

[FIGURE 10 HERE]

4.3 Improvements to the process scheme From the results presented above, it turns out that a high amount of energy is required by columns T-100 and T-101, as shown in the pie charts reported in Figure 11. They illustrate which percentage of the methane equivalent to the condenser duties (Figure 11a) and reboiler duties (Figure 11b) is related to each column considering the reference case study (i.e., the one characterized by a molar ACS Paragon Plus Environment 15


fraction of methane, ethane, propane, n-butane and CO2 in the raw feed gas equal to 45, 5, 5, 5, and 40 mol%, respectively).

[FIGURE 11 HERE]

Thus, in addition to column T-100, which performs the separation between the CH4-CO2 mixture and hydrocarbons heavier than methane, column T-101 is the most expensive one (the reader can refer to the values reported on the y-axis in Figure 4, which are at least one order of magnitude larger than the ones for the duties in the other distillation columns). For energy-saving purposes, an alternative process scheme can be considered, which differs from the one illustrated in Figure 2 because of the elimination of column T-102. In the process scheme in Figure 2, a part of n-butane slips from the top of column T-101 and is found into the bottom product of column T-102 affecting the purity of the propane-rich product (on the contrary, ethane purity is always high). Thus, one may wonder whether the separation between propane and n-butane is really required or can be avoided. To answer this question the following observations concerning the nature of the product streams are helpful. Indeed, while ethane is generally sold as a pure compound, it is more likely to find propane not as a pure component but as LPG, whose composition is not defined, but generally indicated as a propane/n-butane mixture with propane as main component. Moreover, propane reduces separation costs in T-100, giving the possibility to achieve the same separation using a lower amount of entrainer. Thus, for the reference case (45 mol% methane, 40 mol% CO2 and 5 mol% for each heavier hydrocarbon) simulations have been performed for the alternative process scheme in which column T-102 is not included anymore and the specifications for column T-101 are the ethane purity and recovery (set at 99 mol% and 99.5%, respectively). The simulations for this case study are characterized by a different split ratio for TEE-100 regarding the amount of LPG to be recycled to column TEE-100 as entrainer. For a split ratio equal to 94.5% it results that the percentage of the methane available in the raw feed gas that must be used to run the process is 25%. This value can be compared with that one (33%) obtained for the simulation carried out for the same inlet composition but including column T-102 in the analysed process scheme and for a split ratio for TEE-100 equal to 98%. Thus, the elimination of column T-102 turns out to be advantageous from an energy point of view. This occurs even if in this case the possibility of LP steam production by using heat from the heat exchanger E-101 is not considered (due to the lower temperature of the bottom product of column T-101, now containing some propane in addition to nbutane), and even if an additional consumption of equivalent methane has been considered to account for the energy requirement at the condenser of column T-101, where lower temperatures ACS Paragon Plus Environment 16

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are reached that make the use of cooling water as refrigerant not possible. Therefore, despite that, the proposed alternative process configuration proves the advantage of considering the possibility to produce an LPG stream and using part of it as entrainer for the separation of the heavier hydrocarbons from the methane-CO2 mixture to be further processed in the acid gas removal unit by low-temperature distillation.

5.

Conclusions

This work investigates a possible process scheme to perform acid gas removal from natural gas by low-temperature distillation together with NGLs recovery. The impact of the upstream natural gas composition (in particular, of the content of ethane, propane and n-butane) has been studied to understand the process energy performances (accounted for by means of the net equivalent methane method) and to suggest further improvements. The following conclusions can be drawn about the energy performances of the process: -

they improve for an increasing ethane content in the raw feed gas;

-

they decrease as the propane content in the raw feed gas increases, even if this favors the

ethane recovery from the initial mixture in column T-100; -

they improve for an increasing n-butane content in the raw feed gas.

The fact that the overall energy consumptions increase as the propane content in the raw feed gas increases can be explained considering that a higher amount of propane in the feed gas affects the recovery of the entrainer (n-butane) in column T-101, requiring a higher amount of energy for separation of propane (as overhead product, together with ethane) from n-butane (as bottom product). It should be pointed out that these conclusions have been drawn thinking of every process unit as independent from one another and a pinch technology analysis could actually increase the obtained results allowing to exploit thermal recoveries and to significantly reduce energy consumptions. Moreover, considering the beneficial role played by propane to perform NGLs recovery by means of extractive distillation, a less energy-demanding process scheme has been also proposed. This aims at the production of LPG, to be used as entrainer instead of pure n-butane, so to avoid the separation of propane from n-butane, which resulted to be quite energy intensive. This process scheme could be taken into account whenever the market demand may profitably absorb the production of LPG with consequent increased revenues due to lower investment (a column can be avoided) and operating costs.

Nomenclature



Abbreviations CFZ

Controlled-Freeze-Zone

COP

Coefficient of Performance

DEPG

Dimethyl Ether of Polyethylene Glycol

DGA

DiGlycolAmine

DPLT

Dual Pressure Low-Temperature

EOR

Enhanced Oil Recovery

HP

High Pressure

LHV

Lower Heating Value, MJ/kg

LP

Low Pressure

LPG

Liquefied Petroleum Gas

MDEA

Methyl DiEthanolAmine

MP

Medium Pressure

NGLs

Natural Gas Liquids

NMP

N-Methyl-2-Pyrrolidone

PC

Propylene Carbonate

SLV

Solid-Liquid-Vapor

SLVE

Solid-Liquid-Vapor Equilibrium

Symbols Fj

Total Molar Flow Rate of Stream j, kmol/h

݉ሶ

Mass Flow Rate, kg/h

ܳሶ

Duty, kW

ܳሶc

Condenser Duty, kW

ܳሶr

Reboiler Duty, kW

R

Reflux Ratio

RecT-101(i)

Recovery of i-th Component in the Bottom Product of T-101 Column

TH

Temperature of the Hot Reservoir in the Refrigeration Cycle, K

TL

Temperature of the Cold Reservoir in the Refrigeration Cycle, K

ܹሶ el

Electric Power, kW

xj(i)

Molar Fraction of Component i in Stream j

Subscripts CH4

Referred to Methane ACS Paragon Plus Environment 18

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Ethane

Referred to Ethane

n-butane

Referred to n-butane

n-C4

Referred to n-butane

propane

Referred to Propane

R

Referred to the Refrigeration Cycle

R,id

Referred to the Ideal Carnot Refrigeration Cycle

References

1. Erans, M.; Manovic, V.; Anthony, E. J., Calcium looping sorbents for CO2 capture. Applied Energy 2016, 180, 722-742. 2. IEA, World Energy Outlook 2014. IEA: 2014. 3. Burgers, W. F. J.; Northrop, P. S.; Kheshgi, H. S.; Valencia, J. A., Worldwide development potential for sour gas. Energy Procedia 2011, 4, (0), 2178-2184. 4. Parker, P. M. E.; Northrop, S.; Valencia, J. A.; Foglesong, R. E.; Duncan, W. T., CO2 management at ExxonMobil’s LaBarge field, Wyoming, USA. Energy Procedia 2011, 4, 54555470. 5. Sumartono, S., Pengolahan Gas Limbah Proyek Gas Natuna. Jurnal Teknologi Lingkungan 2011, 1, (1), 10-16. 6. Kohl, A. L.; Nielsen, R., Gas purification. Gulf Professional Publishing: 1997. 7. Blohm, C. L.; Riesenfeld, F. C. Amino-ether gas treating process. 1955. 8. Burr, B.; Lyddon, L. In A comparison of physical solvents for acid gas removal, 87th Annual Gas Processors Association Convention, Grapevine, TX, March, 2008. 9. Holmes, A. S.; Ryan, J. M. Cryogenic distillative separation of acid gases from methane. U.S. Patent No. 4318723 A, 9 March, 1982. 10. Chen, S.; Fu, Y.; Huang, Y.; Tao, Z.; Zhu, M., Experimental investigation of CO2 separation by adsorption methods in natural gas purification. Applied Energy 2016, 179, 329-337. 11. Shimekit, B.; Mukhtar, H., Natural gas purification technologies-major advances for CO2 separation and future directions. In Advances in Natural Gas Technology, InTech: 2012. 12. Bernardo, P.; Drioli, E.; Golemme, G., Membrane gas separation: a review/state of the art. Industrial & Engineering Chemistry Research 2009, 48, (10), 4638-4663. 13. Cnop, T.; Dortmundt, D.; Schott, M. In Continued development of gas separation membrane for highly sour service, Sour oil and gas advanced technology conference, Abu Dhabi, April 29th May 3rd, 2007. 14. Baker, R. W.; Lokhandwala, K., Natural Gas Processing with Membranes: An Overview. Industrial & engineering chemistry research 2008, 47, (7), 2109-2121. 15. Holmes, A.; Price, B.; Ryan, J.; Styring, R., Pilot tests prove out cryogenic acidgas/hydrocarbon separation processes. Oil Gas Journal 1983, 81, (26). 16. Holmes, A. S.; Ryan, J. M. Distillative separation of carbon dioxide from light hydrocarbons. U.S. Patent No. 4350511 A, 21 September, 1982. 17. Kelley, B.; Valencia, J.; Northrop, P.; Mart, C., Controlled Freeze Zone™ for developing sour gas reserves. Energy Procedia 2011, 4, 824-829. 18. Valencia, J. A.; Denton, R. D. Method and apparatus for separating carbon dioxide and other acid gases from methane by the use of distillation and a controlled freezing zone. U.S. Patent No. 4,533,372, August 6, 1985.



19. Valencia, J. A.; Kelman, S. D.; Nagavarapu, A. K.; Maher, D. W. In The controlled freeze zone technology for the commercialization of sour gas resources, IPTC 2014: International Petroleum Technology Conference, Doha, Qatar, January 19-22, 2014. 20. Pellegrini, L. A. Process for the removal of CO2 from acid gas. US Patent 20150276308 A1, October 1, 2015. 21. Pellegrini, L. A.; De Guido, G.; Langè, S.; Moioli, S.; Picutti, B.; Vergani, P.; Franzoni, G.; Brignoli, F., The Potential of a New Distillation Process for the Upgrading of Acid Gas. In Abu Dhabi International Petroleum Exhibition & Conference, Society of Petroleum Engineers: Abu Dhabi, UAE, 2016. 22. Langé, S.; Pellegrini, L.; Coquelet, C.; Stringari, P. In Experimental Determination of the Solid-Liquid-Vapor Locus for the CH4-CO2-H2S System and Application to the Design of a New Low-Temperature Distillation Process for the Purification of Natural Gas, GPA convention, 2015; pp 1-32. 23. De Guido, G.; Langè, S.; Moioli, S.; Pellegrini, L. A., Thermodynamic method for the prediction of solid CO2 formation from multicomponent mixtures. Process Safety and Environmental Protection 2014, 92, (1), 70-79. 24. Langè, S.; Pellegrini, L. A.; Vergani, P.; Lo Savio, M., Energy and Economic Analysis of a New Low-Temperature Distillation Process for the Upgrading of High-CO2 Content Natural Gas Streams. Industrial & Engineering Chemistry Research 2015, 54, (40), 9770-9782. 25. Langé, S.; Pellegrini, L. A., Energy Analysis of the New Dual-Pressure Low-Temperature Distillation Process for Natural Gas Purification Integrated with Natural Gas Liquids Recovery. Industrial & Engineering Chemistry Research 2016, 55, (28), 7742-7767. 26. Pellegrini, L.; Langè, S.; Baccanelli, M.; De Guido, G. In Techno-Economic Analysis of LNG Production Using Cryogenic Vs Conventional Techniques for Natural Gas Purification, Offshore Mediterranean Conference and Exhibition, Ravenna, Italy, 25-27 March, 2015. 27. Baccanelli, M.; Langé, S.; Rocco, M. V.; Pellegrini, L. A.; Colombo, E., Low temperature techniques for natural gas purification and LNG production: An energy and exergy analysis. Applied Energy 2016, 180, 546-559. 28. Pellegrini, L. A.; De Guido, G.; Langé, S., Biogas to liquefied biomethane via cryogenic upgrading technologies. Renewable Energy 2017. 29. Pellegrini, L. A.; De Guido, G.; Lodi, G.; Mokhatab, S. In CO2 Capture from Natural Gas in LNG Production. Comparison of Low-Temperature Purification Processes and Conventional Amine Scrubbing, Cutting-Edge Technology for Carbon Capture, Utilization And Storage (CETCCUS), Clermont-Ferrand, France, 24-27 September, 2017. 30. Lastari, F.; Pareek, V.; Trebble, M.; Tade, M. O.; Chinn, D.; Tsai, N. C.; Chan, K. I., Extractive distillation for CO2–ethane azeotrope separation. Chemical Engineering and Processing: Process Intensification 2012, 52, 155-161. 31. Jukić, A. Petroleum Refining and Petrochemical Processes - Natural Gas Composition, Classification, Processing. https://www.fkit.unizg.hr/_download/repository/PRPP_2013_Natural_gas.pdf 32. Luyben, W. L., NGL demethanizer control. Industrial & Engineering Chemistry Research 2013, 52, (33), 11626-11638. 33. Society of Petroleum Engineers Importance of Domestic Production Gas Field Molve –The Contribution of Croatian Energy Sector. http://oslo.spe.org/HigherLogic/System/DownloadDocumentFile.ashx?DocumentFileKey=a7e7682 b-74b0-4fdb-9211-e56a0d326109 34. Burruss, R. A.; Ryder, R. T. Composition of Natural Gas and Crude Oil Produced From 14 Wells in the Lower Silurian “Clinton” Sandstone and Medina Group Sandstones, Northeastern Ohio and Northwestern Pennsylvania; 2330-7102; US Geological Survey: 2014. 35. AspenTech Aspen Hysys®; AspenTech: Burlington, MA, United States, 2012.


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36. Perry, R. H.; Green, D. W., Perry's chemical engineers' handbook. 7th ed.; McGraw-Hill: New York, USA, 2008. 37. CleaverBrooks Boiler Efficiency Guide-Facts about Firetube Boilers and Boiler Efficiency; Cleaver-Brooks, Inc.: Thomasville, GA, USA, 2011. 38. Kehlhofer, R.; Hannemann, F.; Rukes, B.; Stirnimann, F., Combined-cycle gas & steam turbine power plants. 3rd ed.; Pennwell Books: Tulsa, Oklahoma, USA, 2009. 39. DITEC Università degli Studi di Genova Elementi di Analisi Exergetica. http://www.ditec.unige.it/users/administrator/documents/FT2_DITEC_GG_PARTE1.pdf (14.07.2017).



Graphical Abstract

22 ACS Paragon Plus Environment

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Figure 1

Figure 1. Process flow diagram of the Dual Pressure Low-Temperature distillation unit.



Figure 2

Figure 2. Scheme of the simulated natural gas separation train.


Page 24 of 33

Page 25 of 33

Figure 3

19000

130000

18000

ethane

17000

propane

ethane

125000

propane

120000

16000

Qr [kW]

15000

Qc [kW]

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


14000 13000 12000

115000 110000 105000

11000 100000

10000 9000

95000 0

5 10 15 20 mol% ethane or propane feed gas

25

0

a)


25

b)

Figure 3. Thermal duty [kW] for the condenser (a) and reboiler (b) of column T-100 as a function of the ethane (solid line) or propane (dotted line) content (mol%) in the raw feed gas.



Figure 4

120000

160000

118000

150000

116000

140000

122000

170000 160000

120000 ethane

114000

150000

130000

propane

112000

120000

110000 110000

108000

100000

106000 104000

90000

102000

80000

100000

70000

140000

118000

Qr [kW]

Qc [kW]

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

Page 26 of 33

130000 116000 120000 114000

110000 100000

112000

0


ethane propane

110000

25

90000 80000

0

a)

5 10 15 20 25 mol% ethane or propane feed gas

b)



Page 27 of 33

Figure 5

16000

20000

14000

18000 16000

12000

14000

Qr [kW]

10000

Qc [kW]

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


8000 6000

12000 10000 8000 6000

4000 2000

ethane

4000

propane

2000

0

ethane propane

0 0


25

0

a)


25

b)




Figure 6

11000 ethane

10500

propane

10000

Qr [kW]

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

9500 9000 8500 8000 7500 0


25

Figure 6. Thermal duty [kW] for the reboiler of column T-103 as a function of the ethane (solid line) or propane (dotted line) content (mol%) in the raw feed gas.


Page 28 of 33

Page 29 of 33

Figure 7

4630 ethane

4629

propane

4628

Qc [kW]

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


4627 4626 4625 4624 4623 0


25

Figure 7. Thermal duty [kW] for the condenser of column T-104 as a function of the ethane (solid line) or propane (dotted line) content (mol%) in the raw feed gas.



Figure 8

51000

net CH4,eq [kg/h]

50000 49000 ethane

48000

propane

47000 46000 45000 0


25

a)

38 37 36 % CH4 [-]

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

35 ethane

34

propane

33 32 31 30 0


25

b)

Figure 8. Net equivalent methane production [kg/h] (a) and percentage of methane to be consumed for running the process (b) as a function of the ethane (solid line) or propane (dotted line) content (mol%) in the raw feed gas.


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Figure 9

52800 52600

net CH4,eq [kg/h]

52400 52200 52000 51800 51600 51400 51200 51000 0

1

2 3 4 mol% n-butane feed gas

5

2 3 4 mol% n-butane feed gas

5

6

a)

0.325 0.320 0.315

% CH4 [-]

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


0.310 0.305 0.300 0.295 0

1

6

b)

Figure 9. Net equivalent methane production [kg/h] (a) and percentage of methane to be consumed for running the process (b) as a function of the n-butane content (mol%) in the raw feed gas.



Figure 10

58500

58500 10 % ethane; 1% n-butane 10 % ethane; 3% n-butane 10 % ethane; 5% n-butane

58000

15 % ethane; 1% n-butane 15 % ethane; 3% n-butane 15 % ethane; 5% n-butane

58000 57500

net CH4,eq [kg/h]

57500

net CH4,eq [kg/h]

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

Page 32 of 33

57000 56500 56000 55500

57000 56500 56000 55500

55000

55000

54500

54500

54000

54000 0

1

2 3 4 5 6 mol% propane feed gas

7

8

0

1

2

a)

3 4 5 6 7 8 mol% propane feed gas

9

10 11

b)

Figure 10. Net methane production as a function of the propane content (mol%) in the raw feed gas for different n-butane contents (1, 3, 5 mol%, as indicated in the figure legend) at constant ethane content: a) 10 mol%; b) 15 mol%.


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Figure 11

a)

b)

Figure 11. Fraction of the net methane equivalent to (a) condenser duties and (b) reboiler duties of the columns in the process scheme illustrated in Figure 2 (reference case study).


Effect of heavy hydrocarbons on CO2 removal from natural gas by low

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