Economic benefit of methane hydrate reformation management in transport pipeline by reducing thermodynamic hydrate inhibitor injection

Journal Pre-proof Economic benefit of methane hydrate reformation management in transport pipeline by reducing thermodynamic hydrate inhibitor injection Hyunho Kim, Jakyung Kim, Yutaek Seo PII:

S0920-4105(19)30919-2

DOI:

https://doi.org/10.1016/j.petrol.2019.106498

Reference:

PETROL 106498

To appear in:

Journal of Petroleum Science and Engineering

Received Date: 3 July 2019 Revised Date:

14 September 2019

Accepted Date: 16 September 2019

Please cite this article as: Kim, H., Kim, J., Seo, Y., Economic benefit of methane hydrate reformation management in transport pipeline by reducing thermodynamic hydrate inhibitor injection, Journal of Petroleum Science and Engineering (2019), doi: https://doi.org/10.1016/j.petrol.2019.106498. This is a PDF file of an article that has undergone enhancements after acceptance, such as the addition of a cover page and metadata, and formatting for readability, but it is not yet the definitive version of record. This version will undergo additional copyediting, typesetting and review before it is published in its final form, but we are providing this version to give early visibility of the article. 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. © 2019 Elsevier B.V. All rights reserved.

Gas Export

MEG injection

MEG injection

MEG regeneration and injection

Gas dehydration

MEG regeneration

Separation

Hydrate reformation in dissociated water induces high relative torques in early stage of hydrate formation. The addition of MEG maintains torque and reduces the capacity of the regeneration process

Economic benefit of methane hydrate reformation management in transport pipeline by reducing thermodynamic hydrate inhibitor injection

Hyunho Kima, Jakyung Kimb, Yutaek Seoa* a

Department of Naval Architecture and Ocean Engineering, Research Institute of Marine Systems Engineering, Seoul National University, Seoul, Republic of Korea, Seoul 151-744, Republic of Korea

b

Department of Mechanical Engineering, KAIST, 291 Daehak-ro, Yuseong-gu, Daejeon 305-701, Republic of Korea

*Corresponding author: Yutaek Seo E-mail: [email protected] (Yutaek Seo), Phone.: +82 2 880 7239; Fax: +82 2 880 9298 Key words: Gas hydrates; Methane production; Hydrate re-formation; Mono ethylene glycol; Life cycle cost; Regeneration process

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ABSTRACT Hydrate reformation is an inherent risk in a methane hydrate production system because dissociated water and methane flow through transport pipelines simultaneously. In this study, comprehensive experiments using high-pressure autoclaves are performed to investigate the effect of reducing the concentration of thermodynamic hydrate inhibitor on hydrate reformation characteristics. In order to simulate the production of methane hydrates, these are formed and dissociated in the preliminary stage; thereafter, the hydrates are reformed from dissociated water. Although dissociated water and fresh water are found to have similar formation rates, they have a major difference in relative torque, especially in the early stage of hydrate formation. The instant rise of the relative torque in the dissociated water is observed to be up to 2.5; thereafter, it remains approximately 3.0 until hydrate conversion reaches 28 vol%. This trend is different from that in fresh water, where a steady increase in relative torque with sudden spikes is observed. These results clearly show the high risk of hydrate blockage in the transport pipelines of a methane hydrate production system; hence, the injection of hydrate inhibitor is necessary. In this study, monoethylene glycol (MEG) is the selected thermodynamic hydrate inhibitor; its concentration is maintained below a specific value to completely avoid hydrate formation. Despite the occurrence of hydrate reformation, a stable relative torque is observed in a 20 wt% MEG concentration, where the conversion of water to hydrate approximately reaches 40 vol%. For a 10 wt% MEG concentration, a sudden increase in relative torque of up to 2.0 is observed in the early stage of hydrate formation; this suggests that the foregoing is less effective in sustaining flowability. Compared to fresh water, the addition of a 20 wt% MEG to dissociated water results in a more stable torque. Once the performance of the MEG concentration reduction is confirmed, the MEG regeneration process is designed accordingly by using both multiphase flow and process simulation models. A life cycle cost (LCC) analysis is conducted by estimating the capital expenditures (CAPEX) and operating expenditures (OPEX) of the MEG regeneration process. Considering both CAPEX and OPEX, the LCC for complete inhibition (35 wt%) is 90.15 MUSD (million United States dollars); however, for the 20 and 10 wt% MEG concentrations, the LCC values are reduced to 86.79 and 84.58 MUSD, respectively. The

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foregoing suggests the potential economic benefit of methane hydrate reformation management to avoid blockage.

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1. Introduction Gas hydrates encage gas molecules via hydrogen-bonded water molecules under high-pressure and low-temperature conditions. Methane and carbon dioxide are known to form gas hydrates; accordingly, numerous studies related to gas hydrates for storage and separation have been conducted (Dashti et al., 2015; Prasad and Chari, 2015; Zhang et al., 2015; Li et al., 2016; Narayanan and Ohmura, 2016; Tumba et al., 2016; Erfani et al., 2017; Abdi-Khanghah et al., 2018; Chaturvedi et al., 2018). In the late 1990s, significant amounts of methane were discovered as hydrates in marine sediments and permafrost regions (Sloan, 2003; Klauda and Sandler, 2005; Boswell and Collett, 2011). Recent studies have focused on ensuring an efficient recovery of methane from hydrate deposits. The permeabilty of hydrate-bearing sediment is a critical factor for evualting the efficiency of gas hydrate production. Gas injection has been used to measure the permeability of hydrate-bearing sediments (Delli and Grozic, 2014; Li et al., 2016). To improve the resolution, the new method via hydrate dissociation using temperature control was verified to determine the permeability the variation in water phase permeability (Chen et al., 2018). Field test results demonstrate that depressurization, rather than heat stimulation or inhibitor injection, is more effective for methane dissociation (Xu and Li, 2015; Kurihara et al., 2008; Oyama et al., 2009). However, the depressurization method has a limitation of methane production in the later period of hydrate exploitation and endothermic hydrate dissociation reaction may result in secondary hydrate formation (Chen et al., 2019c; Li et al., 2017). To improve the efficiency of gas hydrate exploitation, recent studies investiagated the comibitation method of depressurization with water flow erosion (Chen et al., 2019a; Chen et al., 2019b; Chen et al., 2019c). The methane hydrate decomposed maredly with higher simultanious cold water flow in ther later stages of depressurization through visualization sutdy (Chen et al., 2019a).

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Although several studies have been conducted to explore the best dissociation schemes, investigations regarding the transportation of dissociated methane and water to processing facilities are relatively few. The stoichiometric compositions of methane hydrate to form structure I hydrates are 14.8 mol% methane and 85.2 mol% water. Methane hydrate dissociation results in a two-phase flow (methane gas and liquid water); the amount of water may vary according to hydrate saturation. In the field test performed offshore in Japan, the gas flow rate is approximately 20,000 Sm3/day with a liquid water production of 200 m3/day when hydrate saturation ratios are larger than 80% (Yamamoto et al., 2014). For the methane hydrate deposits in the Ulleung Basin, reservoir simulation results suggest that the water and methane production rates vary with the water–gas ratio in the range 0.49–84.9 kg H2O/m3 CH4 as the hydrate saturation ratio ranges from 30% to 70% (Kim et al., 2017). Our previous work (Kim et al., 2017) indicated that there are two major flow assurance problems that arise during the production of marine hydrate deposits. First, the methane and water mixture should be in the dispersed bubble flow regime at high water and low gas flow rates; however, it becomes a slug flow at low water and high gas flow rates. A slug flow results in extreme methane and water flow fluctuations at the platform that may aggravate the operation complexity of the production system. Another problem is the hydrate reformation inside the production pipeline that is caused by the low temperature of seawater. The pressure and temperature profiles of methane and water mixture that traverse the transport pipeline coincide with the hydrate formation curves of most water and gas flow rates. The dissociated water from hydrate deposits is known to reform hydrates faster than fresh water; this can induce hydrate blockage inside the transport pipeline. Several theories pertaining to the fast hydrate reformation from dissociated water have been formulated. Chaouachi et al. (2015) found that gas hydrate nucleation and growth pattern from dissociated water is considerably different from that observed from fresh water using a synchrotron X-ray tomographic study. The formation rate of gas hydrates from dissociated water increases due to the enrichment of gases in the aqueous phase beyond equilibrium solubility. The impurities and/or container walls could cause more potentional subsequent heterogeneous nucleation site for gas hydrates (Chen and

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Hartman., 2018; Sowa and Maeda, 2015; Maeda, 2016). Through the microfluidic system with in situ Raman spectroscopy, the intrinsic kinetics dominates at the end of dissociation, which results in the residual of undissociated methane hydrate crystals in fluid (Chen and Hartman., 2018). Maeda (2016) inverstigated that only heterogeneous nucleation is possible at the guest-aqueous interface due to the low solubility of guest gases in host water and easily occurs in the presence of solid wall. The heterogenous nucleation rate of gas hydrates could be normalized to the unit length of the interface where aqueous phase, the guest phase and solid wall meet. Recent hypothesis was proposed that the interfacial nanobubbles or other interfacial gaseous states, which could form narrow solidaqueous interfacial area along the three-phase line, can provide the heterogeneous nucleation sites for gas hydrates formation (Maeda, 2018). The conventional strategy for mitigating hydrate formation involves the injection of alcohol or glycol solutions to shift hydrate equilibrium conditions. Methanol is widely used because its inhibition performance is better than that of glycols. However, it would be lost to the gas phase; this can possibly induce operational problems in gas processing systems (e.g., dehydration units). Therefore, monoethylene glycol (MEG) is preferred for gas field developments because its loss to the vapor phase is negligible. For the marine hydrate deposits in the Ulleung Basin, the required concentration of MEG to fully inhibit hydrate formation varies depending on the flow rates of water and gas; it is estimated as 35 wt% under the equilibrium conditions of methane hydrate in pure water and MEG solutions (Fig. S1). We investigated the feasibility of using Poly(N-vinylcaprolactam) (PVCap) to delay hydrate formation, with a delay time 76% longer than the residence time of methane and water mixture in the subsea transport pipeline. However, the transient operation of the methane hydrate production system may involve long periods of methane and water mixtures staying under the hydrate formation condition, where the amount of hydrates cannot be controlled upon hydrate onset. For methane hydrates formed in pure water, the water-to-hydrate conversion achieves as high as 80 vol%. Sohn et al. observed from their autoclave that if the hydrate volume fraction in the aqueous phase reached approximately 7 vol%, the relative torque started to increase and remained high; this is possible

6

because of the increased viscosity of hydrate slurry (Sohn et al., 2015). The Japan Oil, Gas, and Metals National Corporation (JOGMEC) conducted a flow loop study to investigate the flow characteristics of hydrate slurry in water-dominant systems; it was found that some hydrates could be carried by water flow when the amount of hydrates is less than the critical value required to block the flow loop (Sakurai et al., 2014). Joshi et al. (2013) also observed in the flow loop experiments that the volume fraction of hydrates in the aqueous phase for the transition from homogeneous to heterogeneous segregation was approximately 10 to 20 vol%. However, the addition of MEG to the water phase resulted in less amounts of hydrates; this is because the MEG molecules allow hydrogen to bond with water molecules, excluding them from the hydrate formation. Thus, hydrate fraction can be controlled by adjusting the MEG concentration, which should be lower than that required to completely prevent hydrate formation (Kim et al., 2014; Sohn et al., 2017). Several studies have been conducted to manage the hydrate risk by using lesser amounts of inhibitors instead of full inhibition (Hemmingsen et al., 2008; Boxall and May, 2011; Li et al., 2011; Lorenzo et al., 2014). However, the injection of MEG must be analyzed along with the design of the MEG regeneration process; this is because the hydrate inhibition system is composed of the MEG injection system and its regeneration process. Decreasing the MEG concentration at constant rich MEG flowrate leads to a high water fraction in the feed stream; this increases distillation costs during regeneration. In this study, comprehensive experiments are performed to investigate hydrate conversion and flowability during hydrate reformation from dissociated water when the MEG concentration is controlled. Considering the methane hydrate production scenario in the Ulleung Basin, the concentration of MEG is changed from 20 to 10 wt% accordingly. Thereafter, the MEG regeneration process is designed to distill the water off and increase the MEG concentration to 80 wt%; the unit for this process is installed onboard the production platform. A life cycle cost (LCC) analysis is conducted by estimating the CAPEX and OPEX of the MEG regeneration process. The obtained results suggest a holistic approach to manage the hydrate reformation risk of methane hydrate production.

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2. Description of hydrate inhibition system for marine methane hydrate production Fig. 1 shows a simplified marine hydrate production system incorporated with a MEG injection and regeneration process. Because of low seawater temperature, the produced water and methane mixture quickly cools to 0.25 °C, whereas the pressure decreases from 200 to 10 bar; consequently, there is a high risk of methane hydrate reformation in the transport pipeline. The required MEG concentration is 35.0 wt% to completely avoid hydrate reformation. As shown in Fig. 1, the lean MEG has a high MEG concentration (80 wt%) and is injected upstream of the transport pipeline. Once the lean MEG is mixed with produced water, it is termed as rich MEG, which indicates that it has more water content. In this work, the rich MEG concentration is 10 or 20 wt%; hence, it is necessary to manage hydrate reformation because the MEG concentration is less than 35 wt%. Although a certain amount of hydrate is present in the liquid phase, it flows up to the platform once the flowability is maintained. Therefore, the first objective of this work is to confirm the flowability of methane and rich MEG mixture with the presence of hydrate particles. The methane and rich MEG mixture flow into the separator located on the platform; thereafter, it is separated into methane gas and aqueous phase, i.e., the rich MEG. Methane gas is separated from the rich MEG, and then processed to reduce the water contents before it is transported through a gas export pipeline. Because there are no liquid hydrocarbons and acid gases, such as carbon dioxide, the gas processing system is sufficiently simple to satisfy the pipeline sales gas requirement. The rich MEG enters the MEG regeneration process to increase the MEG concentration up to 80 wt% through the distillation column, i.e., lean MEG; thereafter, the lean MEG is injected back into the downhole via the injection pump. The MEG stream cycles through this production system, as shown in Fig. 1 (Latta et al., 2013); details of the MEG regeneration process are described in a later section. For the MEG regeneration process, the rich MEG concentration is reduced to 20 wt% and thereafter to 10 wt% in this work to minimize the investment in the hydrate inhibition system. Reducing the MEG concentration is indeed beneficial in terms of storage and injection capacities. However, a reduced MEG concentration in the feed stream to the distillation column may demand an increased energy input. In our previous work (Kim et al., 2017), for a constant rich MEG flow rate,

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reducing the MEG concentration from 50 to 30 wt% resulted in an increased water fraction in rich MEG; consequently, sum of CAPEX and OPEX increased from 120 to 150 MMUSD due to increased amount of steam for the distillation column. The second objective of this work is to confirm the economic benefit of reducing MEG concentration by decreasing the rich MEG flow rate based on the LCC analysis of the hydrate inhibition system. The MEG regeneration is designed with the process simulation model using Aspen Plus. Thereafter, the economic benefit is quantitatively examined when reducing the MEG concentration from 35 wt% to 20 and 10 wt%.

Fig. 1. Simplified marine methane hydrate production system. Methane is processed for export, and MEG is injected into the downhole to avoid hydrate blockage formation. (Onset) Pressure– temperature changes in methane and water mixture along the transport pipeline under methane hydrate equilibrium conditions in the case of Ulleung Basin.

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3. Experimental materials and procedures 3.1. Materials and apparatus Methane gas (Daesung Industrial Gases Co., Ltd, Korea) with a stated purity of 99.95 mol% and deionized water with a purity of 99.99 mol% are used for the hydrate reformation experiments. Monoethylene glycol with a 99.9% purity is purchased from SAMCHUN Pure Chemical Co., Ltd. A high-pressure autoclave equipped with a magnet stirrer coupling and impeller (Fig. 2) is used to observe the methane hydrate reformation. The autoclave has an internal volume of 360 mL and a working pressure of 160 barg. A thermocouple probe is inserted into the cell to monitor the temperature of the liquid phase with a 0.05 oC uncertainty. The pressure is measured by a transducer with an uncertainty of 0.1 bar in the range 0–160 barg. A sensor is used to measure the torque of the continuously rotating shaft with an uncertainty of 0.3%. The cell is immersed in a temperaturecontrolled liquid bath connected to an external refrigerator.

Fig. 2. Schematic of experimental apparatus. The initial pressure and temperature are controlled using the high-pressure syringe pump and external refrigerator/heater, respectively. Every experiment is performed under an isochoric condition.

3.2. Procedures Similar to our previous work, the experiment sequence is divided into three stages (Kim et al., 2017). First, methane hydrates are formed to prepare the dissociated water. The autoclave cell is filled with a total of 80 mL of water; this cell is flushed with methane gas for at least three times to 10

remove the remaining air inside. Thereafter, with the desired amount of pressure (depending on the cooling temperature and MEG concentration), the cell is pressurized with methane gas with a rate of 1.0 bar/min. After stabilizing the pressure, the cell is cooled to the desired temperature (0.25
for pure water and MEG solutions) without stirring. When the temperature is stabilized to the desired value, stirring is initiated at 600 rpm and maintained for 12 h. The hydrate formation, which is indicated by the temperature spike and pressure drop, is soon observed. At the second stage, the methane hydrate formed in the previous step is dissociated. Once the hydrate formation is complete, the temperature is increased to 18 °C to dissociate the hydrate phase; the cell is maintained at this temperature for 5 h. When it is confirmed that the pressure inside the cell has returned to its initial pressure, hydrate dissociation is considered complete. The third step includes hydrate reformation using dissociated water under the same conditions and procedure in the first step. In this work, “fresh” is used to indicate that there is no hydrate formation in the liquid phase; otherwise, “dissociated” is used to indicate the liquid phase was experienced hydrate formation. Fig. 3 shows an example of the hydrate formation process observed in methane hydrate reformation in pure water. Based on our previous work (Kim et al., 2017), the largest subcooling is observed at approximately 12 °C when methane gas is produced at low gas and moderate water productions (0.15 and 6.5–18.5 kg/s, respectively). The fluid velocity in this production scenario is between 0.75 and 2.14 m/s; it is in the dispersed bubble flow regime. If hydrates are formed in the transport pipeline, then water becomes slurries that contain hydrate particles; these may induce pipeline blockage. In order to simulate the production scenario, the mixing rate of the autoclave is set to 600 rpm, where the flow velocity is 1.75 m/s; fluid flowability is investigated by measuring the relative torque exerted on the rotating shaft. As illustrated in Fig. 3, hydrate formation results in torque increase, which indicates the aggregation and deposition of hydrate particles. In this study, 12 experiments are performed using pure water, and 10 and 20 wt% MEG solutions to investigate the water-to-hydrate conversion (conversion to hydrate) and flowability of hydrate slurry.

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Fig. 3. Pressure, temperature, and torque profiles during hydrate formation, dissociation, and regeneration in methane and pure water. In the first hydrate formation, stable relative torques are marked with red circles. However, in the early stage of hydrate reformation, the relative torque suddenly increases followed by spikes.

The relative torque is calculated as the ratio of the torque recorded during the experiment at a certain time (τ ) to the torque measured prior to nucleation (τ ) using Eq. (1). 

     

(1)



The consumed amounts of gas are calculated using Eq. (2) (Kim et al., 2014): ∆,  

   !

" # 

$%   !

"



(2)

where n denotes the consumed gas moles for hydrate formation at a certain time; Pcal is the calculation pressure with the postulation of no hydrate formation; Pexp is the experimental pressure;

12

Vcell is the volume of gas; z is the compressibility factor calculated from the Cubic-Plus-Association equation of state; R is the ideal gas constant; T is the gas temperature. The conversion to hydrates in an aqueous phase is calculated using the amount of gas consumed, hydration number, and Eq. (3) (Kim et al., 2014): Conversion to hydrate (∅5.789: , %) =

∆=> × hydration number × 100 5>CD

(3)

where nwater denotes the moles of water in the system at the initial condition, and ∆ngas represents the moles of gas consumed in hydrate formation at a given time calculated from the difference between the experimental pressure and estimated equilibrium pressure without hydrate formation. When the hydrate cavities are fully occupied by CH4 molecules, n approaches the ideal value of 5.75 (Seo et al., 2002). The calculated amount of hydrates is compared to the remaining amount of water to determine the conversion to hydrate. The hydrate growth rate is calculated based on the slope of the conversion-to-hydrate curve within a certain time interval to elucidate the hydrate formation characteristics using Eq. (4). GH G

H

LH

=  IJK L I " IJK

I

MNOP

(4)

Each experiment is repeated three times, and the mean value and standard deviation for the final conversion-to-hydrate value, maximum growth rate, and maximum relative torque are calculated.

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4. Result and discussion 4.1. Fast agglomeration of hydrates in dissociated water Table 1 summarizes the final conversion-to-hydrate (vol%) value, maximum growth rate, and maximum relative torque of both fresh and dissociated water. Fig. 4a shows that the conversion-tohydrate value changes with time; Figs. 4b and 4c show that the relative torque changes as a function of conversion-to-hydrate for both the fresh and dissociated water. For both fresh and dissociated water, hydrates quickly nucleate and reach the final conversion-to-hydrate value within 1 h; the final conversion is 94.1 vol% for fresh water and 90.4 vol% for dissociated water. The maximum growth rate in fresh water is 0.21 min−1, which is faster than the maximum growth rate in the dissociated water, 0.16 min−1. During the hydrate formation in fresh water (Fig. 4b), the relative torque remains at approximately 1.5 at the early stage, i.e., a conversion of less than 15%; however, it starts to increase monotonically when the conversion-to-hydrate value is increased from 15% to 50%, and several spikes are observed at a 40% conversion. The relative torque is maintained at approximately 1.7 between a conversion of 50% and 85%; however, it sharply increases to 3.6 when the conversion reaches 85%. It is presumed that the flowability is maintained if the conversion-to-hydrate value is less than 15%; thereafter, the viscosity increases followed by hydrate segregation, which seems to reduce the flowability. As shown in Fig. 4c, in the case of the hydrate reformation from dissociated water, the relative torque increases upon hydrate formation and remains at approximately 3.0 until the conversion reaches 30%; thereafter, it decreases to 1.5 and is maintained until the conversion reaches 85%. However, when the conversion exceeds 85%, the relative torque sharply increases with fluctuation; this is observed in all the three repeat cycles. During the initial stage of hydrate formation, compared to that in dissociated water, the relative torque in fresh water is stable. Although the growth rate in dissociated water is slightly lower than that in the fresh water, the relative torque at the early stage of hydrate formation is two times higher. The increase in relative torque is attributed to the increase in viscosity of hydrate slurry induced by fast-growing hydrate particles and their agglomeration. These results suggest that the hydrate blockage risk is

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considerably higher in dissociated water. Hydrate formation in fresh water exhibits a stepwise decreasing flowability; on the other hand, in dissociated water, the flowability drops at the early stage of hydrate formation until the aggregated chunks are distributed in the aqueous phase because of the high shear from fluid flow. Previous results suggest that the micromechanical force among hydrate particles increases and induces the increase in viscosity of the aqueous phase (Lee et al., 2014; Lee and Sum, 2015; Wang et al., 2017). In this work, the hydrate formation rate and torque are found to be higher in dissociated water than those in fresh water; this indicates high cohesion among hydrate particles. As previously discussed, dissociated water may contain residual hydrate structures or nanobubbles that can facilitate hydrate formation. In fresh water, hydrate formation usually occurs at the gas-water interface; on the other hand, the hydrate formation in dissociated water would occur heterogeneously from the interfacial nanobubbles (Maeda, 2018). Although the growth rate was similar between the fresh water and the dissociated water, the relative torque during the initial hydrate formation was much higher in the dissociated water. This indicates that the hydrate nucleation was faster in the dissociated water, but the growing hydrate particles were agglomerated quickly, offsetting the overall growth. Kim et al. (2014) developed a work process for evaluating the feasibility of hydrate inhibition strategy by incorporating a multiphase flow simulation tool, OLGA, and hydrate kinetics experiments that measure hydrate delay time and hydrate fraction in liquid phase (Fig. S2). Defining field conditions and multiphase flow simulation was carried out in our previous work.. In our previous work, PVCap and MEG were injected to avoid hydrate formation during the residence time of methane and water mixture in the transport pipeline; the experimental studies focused on hydrate delay time as a function of inhibitor concentrations. This present work extensively performs experimental studies to investigate the hydrate fraction in the aqueous phase and the resulting flowability, as suggested in Fig. S2. The technical feasibility and economic benefit of the adopted inhibition system are discussed in the following section.

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Table 1 Experimental results for pure water including final conversion-to-hydrate value, maximum growth rate, and maximum relative torque. The standard deviations observed in the three repeat cycles are enclosed in parentheses. Final conversion-to hydrate (vol%)

Maximum growth rate (min-1)

Maximum relative torque

Fresh pure water

94.1 (2.2)

0.21 (0.02)

3.6 (0.3)

Dissociated pure water

90.4 (1.6)

0.16 (0.01)

3.8 (0.3)

16

(a)

(b)

(c)

Fig. 4. Experimental results for pure water. (a) Conversion-to-hydrate changes; (b) Relative torque as a function of conversion-to-hydrate in fresh water and (c) in dissociated water.

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4.2. Improved flowability by injecting MEG Table 2 summarizes the final conversion-to-hydrate (vol%) value, maximum growth rate, and maximum relative torque observed during hydrate formation when MEG is added to both fresh and dissociated water. Figs. 5 and 6 show the change in relative torque as a function of conversion-tohydrate value in the fresh and dissociated MEG solutions, respectively. When MEG is added to fresh and dissociated water, both the conversion-to-hydrate value and maximum relative torque significantly decrease. In the fresh MEG solution, the final conversion values are 50.8 vol.% for a 10 wt% MEG and 40.7 vol.% for a 20 wt% MEG, which is relatively smaller than that of pure water. The hydrate growth mostly occurred within 2 h to 39.8 vol.% for a 10wt% MEG and 38.3 vol.% for a 20wt% MEG. Since then, the conversion of a 20 wt% MEG solution is maintained at approximately 40.7 vol% for 12 h; on the other hand, the conversions of a 10 wt% MEG solution continuously develop for 12 h and finally reach 50.8 vol% (Fig. S3a). The maximum relative torques are 2.2 and 1.8 in 10 and 20 wt% MEG solutions, respectively. As shown in Fig 5b, the relative torque sharply increases to 2.2, and several spikes are observed with the increase in the conversionto-hydrate value of a fresh 10 wt% MEG solution. When the MEG concentration is increased to 20 wt%, the relative torque is managed to be lower than 1.5; however, a sharp spike is observed at a conversion of 40% (Fig. 5c). Assuming that the MEG molecule does not participate in the hydrate cage when hydrates form, the MEG concentration that remains in the aqueous phase increases as the water molecules are converted to hydrates. If the increasing concentration of MEG is sufficient to inhibit hydrate formation thermodynamically, then hydrate formation ceases to occur. The change in MEG concentration is measured from the mass of the remaining amount of water and loaded MEG. For both fresh and dissociated MEG solutions, the MEG concentrations increase to 18.5 and 28.5 wt% in 10 and 20 wt% MEG solutions, respectively. The MEG concentration required to completely prevent hydrate formation is approximately 35 wt% for MEG solutions. During the hydrate formation under an isochoric condition, the pressure decreases, and the MEG concentration increases; this results in the reduction of thermal driving force for hydrate formation. For the 10 wt%

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MEG solution, the subcooling temperature is 3.7 oC when the hydrate formation stops; on the other hand, it becomes 4.4 oC for the 20 wt% MEG solution.

Table 2 Experimental results for MEG including final conversion-to-hydrate value, maximum growth rate, and maximum relative torque. The standard deviations observed in the three repeat cycles are enclosed in parentheses.

Fresh MEG

MEG concentration (wt%) 10

Final conversionto-hydrate (vol%) 50.8 (2.66)

Maximum growth rate (min-1) 0.10 (0.03)

Maximum relative torque 2.2 (0.4)

solution

20

40.7 (1.03)

0.15 (0.01)

1.8 (0.3)

Dissociated

10

53.3 (2.80)

0.13 (0.04)

1.8 (0.4)

MEG solution

20

37.2 (0.94)

0.11 (0.01)

1.5 (0.3)

The addition of MEG to dissociated water induces a slow hydrate growth because of the decrease in thermal driving force; a stable torque is observed during the hydrate formation process for the 10 and 20 wt% MEG solutions (Fig. 6). The inhibition performance evidently increases with the increase in MEG concentration. Hydrates continuously form until the conversion reaches 53.3 vol% with a maximum growth rate of 0.13 min−1 in a 10 wt% MEG solution. In a 20 wt% MEG solution, the hydrate grows to a final conversion of 37.2 vol% with a maximum rate of 0.11 min−1; the relative torque remains constant during the entire formation period. For dissociated water, a high relative torque is observed at the early and later stages of hydrate formation. Fig. 6b shows a sharp increase in relative torque at the moment of hydrate growth; however, it thereafter decreased to approximately 1.5 in the dissociated 10 wt% MEG solution. By increasing the MEG concentration to 20 wt%, a stable relative torque is observed throughout the formation process, as illustrated in Fig. 6c. These results suggest that the MEG performs an important function during the hydrate formation in the aqueous phase. As the MEG homogeneously dissolves in the aqueous phase, its hydrogen bonding with water molecules competes with the hydrate formation among water molecules, thus limiting hydrate growth. Based on visual observation and obtained torque changes, brittle hydrates are formed and dispersed in the aqueous phase.

19

60

Conversion to hydrate (vol% )

(a)

50

40

30

20

10

MEG 10wt% MEG 20wt%

0 0.0

0.5

1.0

1.5

2.0

Time after nucleation (hr)

(b)

(c)

Fig. 5. Experimental results for fresh MEG solution. (a) Conversion-to-hydrate value changes; (b) relative torque as a function of conversion-to-hydrate for 10 wt% MEG and (c) 20 wt% MEG.

20

(a)

(b)

(c)

Fig. 6. Experimental results for dissociated MEG solution. (a) Conversion-to-hydrate value changes; (b) relative torque as a function of conversion-to-hydrate for 10 wt% MEG and (c) 20 wt% MEG. Time 0 indicates the initiation of hydrate formation.

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Compared with dissociated water, a moderate trend of hydrate growth and an improved transportability of fluid are observed with the addition of MEG. It should be noted that the addition of MEG removed the high relative torque at the early stage of hydrate formation that is observed in dissociated water. The experimental results suggest that the MEG concentration should not be reduced to less than 20 wt%; this is because hydrate blockage may occur at a low MEG concentration. For the scenario of gas production with 0.15 kg/s and water production rates of 6.5– 18.5 kg/s, a dispersed bubble flow regime with a superficial velocity of 0.75–2.1 m/s is expected to occur in the transport pipeline. The multiphase flow model applied in this work calculates the pressure drop along the pipeline using the momentum balance equation in Eq. (5), which is a steadystate expression of the pressure gradient: ST

P S

#QR  S "  # U S (VQR WR RX < + QR WR Z # [= 

\

\ O]

> # [C N + [G ^ +

_` a  Vb

(5)

where QR and Q^ are the liquid–film and liquid–droplet volume fractions, respectively; WR and W=

denote the densities of the liquid and gas phases, respectively; A is the pipe cross-sectional area; [=

is the mass transfer rate between the two phases; [C and [G are the entrainment and deposition rates, respectively (Bendiksen et al., 1991). The frictional pressure drop results from the friction loss between the pipeline wall and the liquid phase and the interaction between the gas and liquid phases, as expressed in Eq. (6). The wall friction factor, λL, is expressed in Eq. (7); Reynolds number is applied in Eq. (8): cdefghfij UR

P

kR  0.0055 p1 +  {

C



m

P

m

 #kR X WR | R | R nU\ # kN X W= | D | D nUI

|\ b }ℎ \

X∗Prs ∈ ^u

(6)

K

Prv y

+w " z x

(7)

(8)

where, SL and Si are the wetted perimeter liquid and interface, respectively; R and D are the

liquid and relative velocity, respectively; > is the velocity between the two phases; d is the droplet deposition; kR and kN are the applied wall friction factor of the liquid phase and interfacial friction

factor, respectively; ∈ is the absolute roughness; Dh is the hydraulic diameter; €R is the viscosity of the liquid phase. 22

Considering that the upward subsea pipeline is long, the hydrate deposition on the wall reduces the hydraulic diameter as well as the Reynolds number; the wall friction factor increases, which can eventually increase the pressure drop along the pipeline. The increase in fluid viscosity also reduces the Reynolds number, resulting in an increased pressure drop. Because of the high water production of hydrate deposits and accompanied high hydrostatic head, it is essential to maintain the flowability by minimizing the frictional pressure loss along the pipeline; a high energy input is required for the subsea multiphase pump. The experimental studies in this work clearly suggest that the rich MEG concentration can be reduced from 35 to 20 wt% to avoid hydrate deposition and secure the flowability of methane and rich MEG mixture with the presence of hydrate particles. Accordingly, the first objective of this research is achieved.

23

4.3. LCC analysis for MEG regeneration process Although the reduction of MEG concentration is applicable to manage hydrate reformation, it cannot be adopted if the economic penalty becomes inevitable. Thus, a MEG regeneration process for producing marine methane hydrates is designed; subsequently, the feasibility of the MEG regeneration process is evaluated from the perspective of CAPEX and OPEX of the MEG regeneration system with an injection pump. The scope of the simulation model, developed with Aspen Plus v9.0, encompasses the downstream of the methane/rich MEG separator to the MEG injection pump. The flow diagram of the brief process is shown in Fig. 7a (Brustad et al., 2005; Laborie et al., 2007; Billington, 2009; Baraka-Lokmane et al., 2012; Lee and Sams, 2016; Kim et al., 2017). The stream from the rich MEG tank enters the pre-treatment vessel (S101) to remove the divalent salt. After the pre-treatment, the rich MEG is heated in the distillation column (C101) to increase the MEG concentration to 80 wt%. A part of the MEG stream is diverted into the flash separator (S102) to remove monovalent salt; this is operated under vacuum conditions (15 kPa). The obtained lean MEG is reinjected into the downhole through the injection pump (P106). The MEG concentration along the transport pipeline is analyzed using the MEG tracking module in the multiphase simulation software, OLGA v2017.2.0, which models transport pipelines, as shown in Fig. 7b. (Kim et al., 2017). Fig. 8 shows the simulation results of the developed MEG process simulation model and the subsea transport pipeline model. Increasing the MEG concentration by increasing the temperature in the distillation column is modeled by applying an electrolyte NRTL in the liquid phase and RK equation of state in the gas phase. The target MEG concentration of 80 wt% is achieved at a temperature of 137 °C at 1.5 bar, as shown in Fig. 8(a). Because MEG is mixed with dissociated water, the MEG concentration reaches the target concentration downstream of the MEG injection point (Fig. 8 (b)); thereafter, it becomes an input variable for the MEG regeneration process simulation model (Fig. 8 (c)).

24

(a)

(b)

Fig. 7. (a) Process diagram of the slip stream process with Aspen Plus v9.0, (b) Multiphase flow simulation for the transport pipeline with OLGA v2017.2.0

Table 3 Basic information for process simulations

Rich MEG flowrate into regeneration process (kg/s) MEG concentration in transport pipeline (wt%) Injection amount of MEG into downhole (kg/s) Injection pressure at topside (bar)

Case 1:

Case 2:

Case 3:

35 wt% MEG

20 wt% MEG

10 wt% MEG

12.6

9.3

7.6

38.8

23.9

11.7

6.1

2.8

1.1

213.0

245.9

259.5

25

(a)

(b)

(c)

Fig. 8. (a) Vapor-liquid equilibrium temperature (T) against the liquid phase mass concentration of MEG for MEG and water system (‚ƒ„ ) at 1.5 bar with eNRTL–RK model using Aspen Plus; (b) MEG concentration and (c) Mass flowrate in aqueous phase against pipeline length using OLGA.

26

As discussed above, a 20 wt% MEG solution exhibits a better hydrate inhibition performance than a 10 wt% MEG solution because the former has a stable flowability. Herein, a process simulation for three different cases is performed. In Case 1, the MEG concentration in the flowline is 35 wt%; full inhibition is considered to avoid hydrate formation. In Cases 2 and 3, the MEG concentrations are reduced to 20 and 10 wt%, respectively. The rich MEG concentration and its flow rate during the regeneration process are calculated through simulations; the foregoing are summarized in Table 3. The boundary of the injection system in the OLGA simulation is from the downstream of the topside injection pump to the downhole where the MEG is injected. The 80 wt% MEG solution from the topside is injected into the downhole of the production tube through a 1.5-in injection line. The amount of MEG to be injected is theoretically calculated based on the water production rate (6.5 kg/s) and target MEG concentration in the flowline. The safety factor amount of 20% is considered for the lean MEG injection. After calculating the mass to be injected to the topside, the required pressure to inject the MEG is thereafter measured; the pipeline pressure at the downhole and the geometry of the injection line are taken into consideration. The concentration of the MEG that arrives on the platform is slightly higher than the target concentration; this is attributed to the applied safety factor. As the MEG flow rate decreases, the injection pressure increases. In the present study, the obtained results in which the MEG injection line has a 1.5-in inner diameter show that the MEG injection amount decreases as the injection pressure at the topside increases; this indicates a gravitydominated flow (Guo et al., 2005). However, the rich MEG flow rate also decreases from 12.6 to 7.6 kg/s when the rich MEG concentration is reduced. A process simulation to regenerate the lean MEG to a desired rich MEG is conducted; the resulting operating conditions of each equipment are summarized in Table 4. By reducing the MEG concentration, both the concentration and flow rate of the rich MEG, which flows into the regeneration process, decrease. The heater in the pre-treatment stage (H101) heats the rich MEG stream of up to 92 °C. After the pre-treatment stage (S101), the stream enters the distillation column to evaporate the water. The distillation column (C101) is operated at 137°C (regardless of the

27

composition of the inlet rich MEG feed) because the required concentration of lean MEG (80 wt% in this work) is determined by the operating temperature in the distillation column. The streams from the flash separator (S102) and lean MEG cooler are combined and reinjected through the injection pump (P106).

Table 4. Description of equipment type and operating conditions for MEG regeneration process.

Vessel

Column

Name

Equipment type

Operating condition

S101

Pre-treatment vessel

80 °C, 1.5 bar

S102

Reclamation vessel

134 °C, 0.15 bar

V101

Vacuum receiver

50 °C, 0.15bar

T101

Rich MEG storage tank

50 °C, 1 bar

T102

Lean MEG storage tank

50 °C, 1 bar

C101

Distillation column

137 °C, 1.5 bar

P101

Booster pump after pretreatment

4 bar

P102

Booster pump after column

3.98 bar

P103

Booster pump after reclamation 1

1 bar

P104

Booster pump after reclamation 2

3.48 bar

P105

Vacuum pump

0.15 bar

P106

Injection pump

H101

Pre-treatment recycle heater

92 °C, 3 bar

H102

Lean MEG cooler

50 °C

H103

Cooler before vacuum receiver

50 °C

H104

Reclamation recycle heater

137 °C

F101

Filter after pretreatment

F102

Filter after reclamation unit

Pump

Heat exchange

Filter

28

1.0

Cost (MUSD)

0.8

Case 1: 35 wt% Case 2: 20 wt% Case 3: 10 wt%

0.6

0.4

0.2

0.0 Column

Injection pump

MEG storage tank

Fig. 9. Equipment cost of column and injection pump

Once the simulation model for the MEG regeneration process is validated, an economic evaluation is performed using the ASPEN Process Economic Analyzer. The size and cost of each equipment are also calculated based on the developed process simulation model coupled with an economic analyzer in ASPEN PLUS. The three cases are evaluated based on the LCC (i.e., the sum of capital expenditure and operating expenditure). The equipment cost of the distillation column, injection pump, and MEG storage tank, which are the major equipment used in the regeneration process are shown in Fig. 9 and summarized in Table S1. The “MEG storage tank” includes both rich and lean MEG solutions. Table 4 lists the costs of purchased equipment, which are calculated based on equipment size; the direct and indirect costs included in the CAPEX are calculated using the ratio indicated in literature (Peters et al., 1968; Javanmardi and Moshfeghian, 2003; Farid, 2007; Wittholz et al., 2008). Columns, pumps, and storage tanks account for approximately 87% of the

29

total equipment cost. The cost and size of the column slightly decrease as the rich MEG concentration decreases. This is because the column size and the operating temperature are the same in all cases to obtain an 80 wt% lean MEG solution; this is achieved when the rich MEG concentration varies between 10 and 35 wt%. The cost of the injection pump used to inject the MEG into the downhole significantly increases from 0.53 to 0.77 MUSD as the target concentration in the flowline decreases from 35 to 10 wt%. The operating pressure is a factor that significantly affects the specification of the pump; hence, the higher the operating pressure, the higher the purchasing cost. The last major equipment is the rich MEG and lean MEG storage tanks; these tanks are assumed to store the MEG solutions for 24 h during preparation. When full inhibition (35 wt%) is applied, the flow rate of MEG is the largest; this results in rich MEG flow rates that are 1.6 times higher than that of the 10 wt% MEG solution. According to literature (Misra, 2015; Lundal and Festøy, 2017; Kim et al., 2018), the average storage tank volume is 1000–1500 m3. Thus, for each case, one tank is required, and the purchase cost is calculated based on its storage capacity; as opposed to the injection pump, the cost of the storage tank decreases from 0.69 to 0.41 MUSD as the target rich MEG concentration decreases from 35 to 10 wt%. Taking these changes into consideration, CAPEX is calculated to be 11.10 MUSD for full inhibition with a 35 wt% MEG solution (Case 1); it becomes 10.86 and 10.65 MUSD for a 20 and 10 wt% MEG solutions (Cases 2 and 3, respectively), respectively. The CAPEX decreases when the MEG concentration decreases because of the reduced rich MEG flow rate. The utility costs (e.g., electricity, cooling water, and LP steam), which are calculated with an interest rate of 8% (Arias et al., 2016) during the operation period, are illustrated in Fig. 10 and summarized in Table S2. In particular, the cost of the used steam accounts for more than 70% of the utility cost; thus, it becomes a significant factor in the OPEX. The full inhibition with a 35 wt% MEG (Case 1) requires the highest utility cost for the distillation column compared to other cases. The cost for maintenance, repair, and depreciation is another major portion of OPEX. The cost of maintenance and depreciation is calculated at a constant rate of fixed capital investment. Therefore, Case 1, which is the largest CAPEX, requires the highest maintenance and depreciation costs (5.55

30

and 9.25 MUSD, respectively). The overall OPEX is calculated to be 79.05 MUSD for the full inhibition with a 35 wt% MEG solution (Case 1), and 75.93 and 73.93 MUSD for 20 and 10 wt% MEG solutions (Cases 2 and 3, respectively). Similar to CAPEX, OPEX can be reduced by diminishing the MEG concentration. Figure 11 shows the LCC analysis, which indicates that the OPEX occupies a larger proportion than CAPEX. Reducing the MEG concentration increases the cost of the MEG injection pump; this results in a reduction in the CAPEX and OPEX because of the decrease in rich MEG flow rate for the MEG regeneration process. The resulting LCCs are calculated as 90.15, 86.79, and 84.58 MUSD (Cases 1–3, respectively); this suggests the economic benefit of reducing the MEG concentration to 20 wt%. Accordingly, the second objective of this research is achieved. 25

Cost (MUSD)

20

Case 1: 35 wt% Case 2: 20 wt% Case 3: 10 wt%

15

10

5

0 Electricity

Cooling water

LP steam

Fig. 10. Utility cost during operating years (20 years)

31

Fig. 11. Life cycle cost including CAPEX and OPEX over 20 years 5. Conclusions Methane production from marine hydrate deposits is quite challenging because of the deep water depth and low-temperature seawater. Although methane hydrate is dissociated by depressurization in its reservoir temperature, a mixture of methane and dissociated water flows through the upward transport pipeline to the platform, where it cools down as a result of the heat exchange with cold seawater. The hydrate reformation from dissociated water results in a high resistance to flow in the early stage of hydrate formation; thus, there is practically no response time to avoid hydrate blockage upon the detection of hydrate formation. Similar to those employed in the conventional gas production system, this work suggests the injection and regeneration of MEG to prevent hydrate risk; moreover, the reduction of the concentration of MEG in the aqueous phase is also suggested to obtain economic benefits. As demonstrated by the experimental results, the addition of MEG in the aqueous phase significantly retards hydrate formation and growth. Moreover, a stable flowability is observed in the dissociated water because of the presence of homogeneously dispersed MEG in the aqueous phase. By reducing the MEG concentration, the flow rates of rich MEG streams are also reduced. The obtained process simulation results indicate that the size of the regeneration process can be decreased, which can thereby diminish the CAPEX and OPEX. Both the rich MEG concentration

32

and flow rate are reduced; the foregoing is beneficial for designing smaller MEG regeneration processes with less operating costs. In this work, the experimental results of the hydrate formation characteristics are incorporated with the process simulation to analyze the feasibility of using MEG as a hydrate inhibitor in marine methane hydrates. The obtained results support the good hydrate inhibition performance of MEG and the economic design of the MEG regeneration process by controlling the MEG concentration. In extracting marine hydrate deposits, both technical and economic problems arise. This work aims to provide insights for researchers and engineers involved in hydrate investigations related to marine hydrate deposits.

33

Acknowledgments This work was respectfully supported by the Technology Innovation Program (10060099) funded by the Ministry of Trade, Industry & Energy (MI, Korea). This work was also partially supported by the Korea Institute of Geoscience and Mineral Resources.

Nomenclature A CAPEX CH4 C_MT C_OL D FCI LCC MUSD n OPEX Pcal Pexp {C Sp Vcell R Z t T

VF Vw,conv VP z ‡ ∈ k € W ֎G [ ‘ 

Pipe cross-sectional area, m2 Capital expenditures methane Maintenance and repair cost Operating labor cost Diameter, m Fixed capital investment Life cycle cost Million USD Consumed gas moles for hydrate formation at a certain time, mol Operating expenditures; Total manufacturing cost Calculation pressure with postulation of no hydrate formation, barg Experimental pressure in the autoclave cell, barg Reynolds number Wetted perimeter at phase P, m Volume of gas, cm3 Ideal gas constant gravitational constant, m/s X Time Temperature at gas phase in the autoclave cell, K Velocity, m ∙ s LP Volume, cm3, (F=hyd, w) Volume of the water converted to hydrate, cm3 Volumetric fractions, (P=L, D) Compressibility factor Angle with graivity vector, rad Absolute roughness, m Friction coefficient Visocity, kg/m ∙ Š Density, kg/‹Œ Conversion volume fraction of water to hydrate, conversion-to-hydrate, vol% Mass-transfer term, kg/‹Œ ∙ Š Growth rate, min-1

Subscripts d D e F

Droplet deposition Droplet Droplet entrainment Friction

34

g h hyd w i L r

Gas phase hydrulic hydrate water Interfacial Liquid phase Relative

35

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39

Highlights
This research studied the hydrate re-formation in methane hydrate production system.
MEG solution prevented the hydrate agglomeration in early stage of hydrate growth.
Total cost for MEG regeneration was less when controlling MEG concentrations.