Understanding gas-phase breakout with high H2 content in CCS pipeline gathering networks

International Journal of Greenhouse Gas Control 90 (2019) 102816

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Understanding gas-phase breakout with high H2 content in CCS pipeline gathering networks Matthew Healeya, Ketan Mistryb, Thomas Jonesa, Eduardo Luna-Ortiza, a b

T

⁎

Pace Flow Assurance, 10 Lower Thames Street, London, EC3R 6EN, UK National Grid, 35 Homer Rd, Solihull, B91 3QJ, UK

A R T I C LE I N FO

A B S T R A C T

Keywords: CO2 transport CCS Flow assurance Hydrogen production Hydrogen breakout Impure CO2

An accurate understanding of the behaviour of impure carbon dioxide (CO2) during pipeline transport stage is required for commercial scale deployment of Carbon Capture and Storage (CCS) networks. Impurities in the CO2 stream modify phase behaviour and change the thermophysical and transport properties of the stream. CO2 streams containing hydrogen (H2) are a particular challenge due to the unique physical properties of H2. In a CCS gathering network where CO2 is sourced from processes such as pre-combustion capture or hydrogen production, H2 with a concentration of 2% mol can be expected to be present in the mixture. This paper describes foreseeable operating scenarios where gas breakout in single-phase CCS gathering networks might occur, leading to pockets of gas with high H2 concentration. Multiphase flow modelling of a CCS gathering network with impure CO2 containing H2 has been performed, demonstrating these operating scenarios and providing a basis for design. The H2 content of the gas breakout is shown to be > 20% mol when the conditions are close to bubble point. High concentrations of H2 due to gas breakout must be considered as part of a single-phase CCS gathering network design, whenever H2 is present. This paper provides practical guidelines for understanding, quantifying and, managing the worst design cases for H2 exposure due to gas breakout. Specific recommendations for what must be included in the project design basis are presented. Potential mitigating factors and engineering measures that can be taken to manage high H2 concentration in a CCS system are also discussed.

1. Introduction

existing high CO2 concentrations are stripped of the CO2 using solvent.

Carbon Capture and Storage (CCS) involves capturing carbon dioxide (CO2) emissions from industrial processes, transporting predominantly through pipelines, and storing safely in suitable geological formations deep underground. The four common capture processes are (Metz et al., 2005; Heltand et al., 2014):

• Post combustion: CO • • • ⁎

2 from fossil fuel combustion is extracted from the exhaust gas using a solvent. The CO2 is then separated and compressed for transportation while the solvent is recycled. Pre-combustion: Fossil fuel is combined with oxygen and steam through a series of reforming processes to produce CO2 and H2. The CO2 is disposed of whilst the H2 can be used as fuel (“blue hydrogen”). Oxyfuel: Fuel is burned in oxygen instead of air to produce a flue gas consisting mostly of CO2. High CO2 Hydrocarbon Stream: Hydrocarbon streams with pre-

Transportation of CO2 is possible in small quantities by lorry or ship, however over large distances and for large scale continuous operation it is more cost effective to use pipelines. For pipeline transportation it is more efficient for the CO2 to be compressed and transported as a liquid. Notwithstanding these advantages, CO2 transportation in pipelines is challenging in particular with respect to flow assurance, integrity management and, health and safety factors (Onyebuchi et al., 2018). The captured CO2 can be stored either onshore or offshore in geological formations such as depleted hydrocarbon fields (as part of Enhanced Oil Recovery (EOR)) or saline aquifers. The design of a CCS pipeline transportation network must consider all possible sources of CO2 and the impurities that can be collected from those sources. There are economic advantages to transport impure CO2 over further CO2 purification at the capture sites (Kolster et al., 2017). Yan et al. (2008) showed that costs associated to purification of CO2

Corresponding author. E-mail address: eduardo@paceflowassurance.co.uk (E. Luna-Ortiz).

https://doi.org/10.1016/j.ijggc.2019.102816 Received 10 April 2019; Received in revised form 2 August 2019; Accepted 19 August 2019 1750-5836/ © 2019 Elsevier Ltd. All rights reserved.

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required for the pipeline pressure to fall within the two-phase region and risk gas breakout.

(capital and operating costs) in a CCS system are dominant in comparison to transportation costs. However, it has been suggested a maximum content of 4% mol of non-condensable impurities (e.g. N2, H2 or Ar) for economical transportation over long distances (e.g. de Visser et al., 2008). H2 is a contaminant in the waste CO2 stream from pre-combustion capture processes as well as from hydrogen manufacturing. There is limited experience worldwide in the pipeline transportation of CO2 with impurities and, particularly, there is a blind spot in the design of pipelines in the presence of H2 as impurity, which, due to its low critical temperature and pressure, it has the highest impact on the phase behavior of the CO2-mixture. For CO2 streams from pre-combustion capture CCS, the H2 content is 0.8 to 2.0% mol (Metz et al., 2005). In general, these low concentrations in a single-phase fluid present minimal challenges for design. However, there are foreseeable operating scenarios where the system may be at risk of long-term exposure to significant higher local concentration of H2 than that of the bulk stream. Such scenarios need to be considered in the design of the CCS transportation networks. This is particularly important for material selection and integrity management of the pipelines, where exposure of high concentrations H2 can present metallurgical issues such as embrittlement and hydrogen induced cracking (Ohaeri et al., 2018). The presence of gas in a liquid CCS transportation network is undesirable during normal operation. Single phase flow, as either a liquid or dense phase fluid, is preferred. However, gas breakout may occur during some planned and unplanned operating scenarios:

• Scenario A - Pressurization and depressurization (CO

2

• Scenario C - Low pressure, downstream of the injection choke at wellhead

It is possible that pressure downstream of the injection choke will be below the bubble point of the CO2 stream. This is particularly likely early in field life, when the storage pressure is at its minimum. (Storage pressure will increase as CO2 is injected over time, whether the target is a depleted hydrocarbon reservoir, a saline aquifer, a salt cavern, or other geological formation.) Also, it is likely that early operation coincides with low injection rates as the CCS system is brought online. This means that there is low frictional pressure drop in the well and across the perforations into the storage reservoir, which also reduces the pressure immediately downstream of the injection choke. When the pressure downstream of the injection choke is below the bubble point, gas will evolve across the choke during injection. A gas cap at the top of the well may also be in place when injection is not taking place (for example during well shut down or abandonment). While design guidelines of CO2 pipelines with various impurities have been considered elsewhere (i.e. DNV, 2010; Mohitpour et al., 2012 and ISO, 2016), there are not specific considerations reported regarding H2 gas breakout occurring in operating scenarios as described above. ISO 27913 (ISO, 2016) provides standards for the design, construction, operation, environmental and risk management, and related activities in CO2 capture, transportation, and geological storage. It provides a wide range of specifications that ensure a CO2 pipeline would be fit for purpose; however, it does not mention any risks associated with localized high concentrations of H2 gas. The practice DNV-RP-J202 (DNV, 2010) provides guidance and sets out criteria for the development, design, construction, and operation of steel pipelines for the transportation of CO2. The recommended practice provides an overview of the factors that need to be considered for designing and operating CO2 transportation pipelines, but high local concentrations of H2 are not covered. Several research consortia such as CO2PIPEHAZ (Woolley et al., 2014), COOLTRANS (Cooper and Barnett, 2014), CO2QUEST (Porter et al., 2016) and IMPACTS (Brunsvold et al., 2016) have investigated CO2 transportation with impurities (pipeline flow simulation, pipeline rupture, gas dispersion, etc.) but the issue of long-term persistent H2-rich gas breakout in transportation networks has not been addressed. The aim of this paper is to introduce and define the issue of high H2 content gas breakout in single-phase CCS gathering networks. We investigate potential operating scenarios that could lead to localized H2 long-term exposure in the CO2 transportation pipeline and may represent safety and integrity risks. The rest of the paper is organised as follows. Firstly, we present a review of the thermodynamic and flow modelling of CO2 mixtures transportation pipelines with emphasis on H2 as the key impurity, as it has the most impact on the physical properties of the mixture. Then, we present our key operating scenarios where there is potential for H2-rich gas breakout in certain parts of the gathering and injection system. We propose design methodologies/ guidelines to define and evaluate the risk of localized H2-rich gas pockets. It is worth to mention that material fitness assessment is not part of the scope of this paper. Finally, we provide potential mitigations; however, specific solutions to manage high concentration of H2 in gas breakout should be based on the particulars of the CCS project.

dispersion) of the

pipeline

In the event of a medium-term or long-term shutdown, it is possible that the onshore CO2 transportation pipeline will need to be depressurized. This minimizes the inventory of CO2 and therefore reduces the severity of any leak, which may be particularly important if maintenance is required. There are also likely to be long-term shutdown scenarios that require depressurization of an offshore CO2 pipeline. This process will reduce the pressure of CO2 within the pipeline such that the entire inventory is in the gaseous phase. The CO2 may either be vented or injected into the storage reservoir, and the phase transition, from liquid to gas, is crossed during this process. The reverse takes place during pressurization. During depressurisation/dispersion, gas breakout is expected to occur (and possibly CO2 freezing) but the duration of the high H2 concentration gas is, in general, shorter in comparison to other operating scenarios where long-term exposure to H2 might pose a problem. Therefore, this Scenario A is not part of the scope of this paper and is not investigated further.

• Scenario B - Upset conditions leading to low pressure at an onshore high point

Low pressure operation, possibly outside of the bounds of the normal operating envelope, may lead to pipeline pressure fall below the bubble point pressure, in certain locations. In general, there are two possible locations for low pressure: upstream of a pumping station during normal operation or at a high elevation point along the pipeline route during normal operation or during shutdown. While evolution of gas in these locations is operationally undesirable, it is probably beneficial for the pipeline system to be operating close to the pressure where this may take place. The minimum onshore pipeline normal operating pressure for a CCS network (operating in the liquid phase), is normally the bubble point pressure plus some safety margin. The maximum normal operating pressure is defined to be as low as possible, with allowances for frictional pressure drop, high and low points, operational flexibility, and other factors. If the minimum operating pressure is close to the bubble point, only a relatively small upset may be

2. Thermodynamic and flow modelling of CO2 pipelines with impurities One of the challenging aspects in CO2 pipeline transportation is an accurate understanding of the phase behaviour of carbon dioxide and relevant impurities as these greatly increase the possibility of two-phase flow (increasing the cricondenbar of the mixture). Knowledge about the 2

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cubic EoS fail near the critical region (see for example Abbott, 1973; Michels and Meijer, 1984) and have poor prediction of liquid densities and density-derived properties (see for example Martin, 1979; LopezEcheverry et al., 2017). Moreover, the thermodynamically behaviour of complex CO2 (strong dipole moments) mixtures with polar or self-associating components (such as water of alcohols) cannot be captured adequately with cubic EoS. Indeed, some of these limitations can be overcome by different modelling approaches. For example, using different mixing rules (i.e. classical van der Waals, Huron-Vidal (Huron and Vidal, 1979), Michelsen-Huron-Vidal (Michelsen, 1990)) could be used to improve phase equilibria calculations with polar components. Adjusting Binary Interaction Parameters (BIPs), to match experimental bubble and dew points, is another technique to improve the performance of cubic EoS. BIPs can be either constant or temperature-dependent but their tuned values are specific to the EoS/mixing rule of choice and the alpha-function contained in the attraction parameter of the cubic EoS (i.e. Soave, 1972; Graboski and Daubert, 1979). This is particularly important for systems containing H2 as shown by Jaubert et al. (2013). Volume shift factors (Péneloux et al., 1982) could be used to correct liquid density estimations. Even though the described approaches could improve the performance of cubic EoS, the inaccuracies near the critical point will not be alleviated. Here, it is important to note that cubic EoS will only be suitable for scenarios where two-phases (liquid-vapor) appear. Any scenario where the system is below the triple point (solid CO2) a more elaborate EoS will be required. Al-Siyabi (2013) showed that H2 as an impurity has the greatest impact on the thermodynamic and physical properties of a CO2 stream. Experimental data of various binary CO2 systems are used to tune three cubic EOS (PR, Soave-Redlich-Kwong (SRK) (Soave, 1972) and Valderrama-Patel-Teja (VPT) (Valderrama, 1990)) at different conditions. Although the models show a good agreement with the experimental data, they are weak near the critical point. The average deviation in estimation of saturation pressures, for all the EoS, is as high as 10% for CO2-H2 due to the low critical temperature of hydrogen. Nazeri et al. (2017) measured densities for various CO2-rich mixtures containing various impurities including hydrogen (up to 0.81% mol) for a relatively large pressure and temperature range. The experimental data was then used to evaluate several cubic EoS such as SRK and PR (both without and with Péneloux volume correction). It was concluded that performance of cubic EoS (for the mixtures investigated) was not as high as with GERG-2008. Tenorio et al. (2015) measured vapor-liquid equilibria data of binary and ternary mixtures containing H2 up to 5% mol for a limited temperature range. The experimental data is used to evaluate and assess the deviations from GERG-2004 and PR. The model that matches the bubble point data most accurately is PR EoS with temperature-dependent binary interaction parameters. However, it is noted that GERG2004 is more accurate than PR EoS for density calculations especially for low hydrogen concentration mixtures as PR EoS cannot reproduce the density of pure CO2 in liquid phase. Fandiño et al. (2015) presented equilibrium data for CO2-H2 binary systems for a limited temperature range. The fluid mixtures studied were modelled with the PR with two binary parameter interactions. It is shown that the model can match the experimental data in a satisfactory way except in the critical region. Xu et al. (2017) propose an enhanced predictive EoS based on PR78 (Robinson and Peng, 1978) for CCS fluids. The model calculates temperature-dependent binary interaction parameters with a group contribution correlation. The parameters can be accurately predicted even if experimental data is not available. For CO2-H2 mixtures, the model has reasonable accuracy at low pressures EoS considering association interactions effects such as Statistical Associating Fluid Theory (SAFT) models (e.g. Chapman et al., 1989; Gross and Sadowski, 2001; Papaioannou et al., 2014) have also been studied for CCS applications. Diamantonis et al. (2013a) showed SAFT-type equations do not

thermo-physical and transport properties of the mixture is fundamental for the design and operation of all stages of CCS systems (Tan et al., 2016). The presence of impurities in CO2 streams can lead to flow assurance, metallurgical, integrity, safety, processing, economic and geological storage issues (e.g. Verma et al., 2011; Blanco et al., 2012; Brown et al., 2013; Halseid et al., 2014; Huh et al., 2014; Wetenhall et al., 2014; Wang et al., 2015; Skaugen et al., 2016; Chen et al., 2018). To tackle these challenges, an accurate modelling and estimation of the fluid phase equilibria and thermo-physical properties with Equations of State (EoS) for a wide range of operating pressure and temperature conditions (including supercritical). Also, it is important that the model of choice does capture the interactions of the key contaminants with CO2.This section summarises the key available thermodynamic models relevant to CO2 transportation with H2 as the relevant impurity. We note that other impurities may be present in CO2 streams; however, for the operating cases described in this paper, H2 plays the crucial component due to the consequences of a pipeline being exposed to a H2-rich gas breakout. Similarly, interactions between impurities have not been revised in this paper. Several comprehensive reviews of thermodynamic models for CCS systems are available elsewhere and the reader is referred to them (e.g. Li et al., 2011; Aursand et al., 2013; Diamantonis et al., 2013b; Munkejord et al., 2016; Shin et al., 2018). While there is good availability of vapor-liquid equilibria experimental data for binary mixtures of CO2 with some impurities (e.g. H2O, CH4, N2), data for some other pairs is scarce (such as CO2-H2). In Munkejord et al., 2016 an extensive review on available thermo-physical and transport properties data for CCS systems is presented. Also, Coquelet et al. (2017) offer another recent survey of the available thermo-physical experimental data in the open literature. For pure CO2 streams, the Span-Wagner EoS (Span and Wagner, 1996) is considered the most accurate model. For CO2-rich systems, the Groupe Européen de Recherches Gazières GERG-2008 (Kunz and Wagner, 2012) is known to give high quality results for selected mixtures related to CCS processes (Diamantonis et al., 2013b). This noncubic EoS describes the thermodynamic properties as a mixture equation of various reference equations of pure components. It is formulated in terms of the reduced Helmholtz energy by combining pure component reduced properties. A similar formulation (in terms of reduced Helmholtz energy and reference equations) is proposed by Lemmon and Tillner-Roth (1999) which gives high quality results and it is implemented in the widely used property package REFPROP (Lemmon et al., 2018). Another accurate non-cubic EoS is proposed by Gernet and Span (2016), which can outperform GERG-2008, particularly for mixtures containing water. This EoS formulation does not consider hydrogen but an update to the model which includes hydrogen in the mixture model is implemented in the property package TREND (Span et al., 2019). TREND (Thermodynamic Reference and Engineering Data) has the additional capabilities of modelling phase equilibrium below the triple point which makes it suitable for modelling freezing of CO2 during dispersion/depressurization scenarios (e.g. Porter et al., 2016; Wareing et al., 2016), though these operating scenarios are out of scope of this paper. Demetriades and Graham (2016) developed a pressure-explicit EoS as function of volume, temperature and composition. For CO2-H2 mixtures, it is shown that the proposed model is more accurate than the GERG-2004 EoS (Kunz et al., 2007). However, the EoS is limited to a narrow temperature region (below the preferred operating region of CO2 transportation pipelines). The application and performance of cubic EoS for CCS processes has been extensively reported (e.g. Li and Yan, 2009; Li et al., 2011; Mazzoccoli et al., 2012; Wilhelmsen et al., 2012). DNV-RP-J202 (DNV, 2010) considers that, for a predominantly CO2 mixture, Peng-Robinson (PR) EoS (Peng and Robinson, 1976) provides sufficient accuracy. For significant levels of contaminants, the DNV practice recommends tuning the EoS to experimental data. However, it is well known that 3

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3. Methodology

significantly increases accuracy in comparison to a cubic EoS with binary parameters fitted to experimental data (if available). Wareing et al. (2016) show that a SAFT equation is not suitable for dispersion modelling (and possibly for pipeline simulation) due to a discontinuity in the dew line at low temperatures. Instead, it is suggested that the mixture model implemented in TREND is more adequate. On the other hand, Ke et al. (2017) compared a SAFT-type equation (Papaioannou et al., 2014), GERG-2008, the mixture model of TREND and PR against experimental data of a ternary mixture containing 3% mol H2. It was shown that the SAFT EoS outperformed all the models when predicting bubble point pressures. Although there are apparent advantages in using SAFT methods (they require limited or no experimental data to tune), their performance is still questionable, and they have not been widely adopted yet in conjunction with standard off-the-shelf multiphase flow simulators such as OLGA® or LedaFlow®. Multiphase and single-phase flow modelling of CO2 (pure or mixtures) transport in pipelines using commercial simulators has been studied in several published literature. Most of these flow assurance assessments (both steady-state and transient and for conditions above the triple point of CO2) has been focused on the determining the normal operating window, hydrate formation propensity, shutdown (planned and unplanned), start-up and injection into wellbore (e.g. Clausen et al., 2012; Peters et al., 2012; Böser and Belfroid, 2013; Oosterkamp et al., 2013; Yang et al., 2015; Munkejord et al., 2016). To our knowledge, there are not published studies addressing specific operating scenarios that could lead to high H2 content in gas breakout. Aursand et al. (2013) reviewed the most common available multiphase flow simulation software and their applicability to CO2 transportation. Commercial tools developed specifically for transient simulation of CO2 flows are scarce. Similarly, experimental or field validation of multiphase flow models is limited for transportation of CO2 with impurities (Clausen et al., 2012; de Koeijer et al., 2014; Drescher et al., 2017). The available leading modelling tools used in oil and gas industry, OLGA® and LedaFlow®, can be used for CO2 transportation in pipelines. However, flow assurance modelling with these tools is challenging. Yang et al. (2012) discussed the limitations associated with the use of standard multiphase flow simulators. Discontinuities in volume dependent physical properties near supercritical region or across a saturation boundary represent a numerical issue for most mathematical formulations implemented (for multi-component systems) in commercial flow simulators (see Ruden et al., 2014). It is highlighted that physical properties provided to the multiphase simulator (for example as a look-up table) need to have high resolution particularly near phase boundaries. Martynov et al. (2015) analyzed the performance of a number of integral thermo-hydraulic models in terms of pressure and temperature drops in a pipeline system transporting dense-phase and supercritical CO2. Although these models can be used to define pipeline sizing and pressure boosting requirements, these cannot be used to study full flow assurance/operability issues. Joint-industry and academic projects such as CO2PIPEHAZ, CO2QUEST and COOLTRANS have developed extensive theoretical methods capable of pipeline flow simulation (see for example Wareing et al., 2013; Porter et al., 2016; Wen et al., 2016). All these models have been used to study, mainly, gas dispersion scenarios (including CO2 solid formation) and, to our knowledge, have not been used to investigate flow assurance scenarios such as the ones addressed in this work.

3.1. Multiphase flow simulations and thermodynamic package Multiphase (two-phase) flow simulations were carried out using the commercial software LedaFlow® version 2.3. The thermodynamic properties required by the simulator, in form of a look-up table, have been determined with GERG-2008 EoS (Kunz and Wagner, 2012) using the thermodynamic property package Multiflash™ version 6.2. The look-up table consisted of 90000 pressure and temperature points with a fine grid near the phase boundaries. For the GERG-2008 EoS, the generalized mixture of Kunz et al (2007) has been used with all the BIPs and weighting factors of the departure functions set to 1.0. It is important to note that this work only focuses in presenting key operating cases where H2-rich gas breakout occurs, so no specific basis has been paid to the selection of EoS other than GERG-2008 outperforms other EoS as described earlier. Other EoS such as TREND can also be used with, potentially, different results but the design cases in which H2 breakout will occur remain the same. 3.2. CO2 mixtures compositions Table 1 shows the CO2 stream compositions considered in this work. 3.3. Phase behaviour of CO2 mixtures The presence of H2 in the CO2 mixtures drives the bubble point of the mixture to higher pressures. The phase envelopes for each of the three compositions in Table 1 are shown in Fig. 1. It is worth mentioning that the thermodynamic model is able to calculate both dew and bubble lines without convergence issues, but it fails around 0.5 °C from the critical point (we have smoothed both saturation lines around the critical point). The composition of the gas stream at the bubble point represents the worst case for H2 content. The first gas evolved at the bubble point pressure will have a higher concentration of H2 than the overall stream. This is because the H2 is the most volatile compound in the mixture. Table 2 and Table 3 show the equilibrium gas composition at the bubble point for the high and low H2 content compositions (see Table 1) at various temperatures. It is seen that, for the high H2 content composition, the H2 concentration of the incipient gas bubble is 17.7% mol at 5 °C. For most CCS systems this is, in general, unrealistically onerous and should only be considered as the absolute worst-case scenario. 3.4. Operating scenarios Very short-term exposure of H2 gas, such as a bubble that is produced but quickly recondensed into the main stream, is unlikely to present a design issue for the CO2 transportation pipeline. The equilibrium H2 concentration presented in Tables 2 and 3 will be the “first incipient bubble”. As the pressure in the system falls below the bubble point, a condition practically necessary for gas build-up, the H2 will be diluted and therefore the actual H2 concentration in a persistent gas bubble will be lower. Hence, the criteria to determine the governing design cases for H2 rich gas breakout are: Table 1 CO2 stream compositions.

4

Component

Pure CO2 % by volume/mole

Low H2 case % by volume/mole

High H2 case % by volume/mole

CO2 H2 N2 Ar CO

100.0 – – – –

98.0 0.10 1.30 0.40 0.20

96.0 2.0 1.4 0.4 0.2

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Fig. 2. Schematic of onshore CO2 transportation pipeline.

Fig. 1. Phase diagrams for the three compositions of Table 1 with various concentrations of H2. Solid line is the saturation curve of pure CO2. Broken line is the phase envelope of the low H2 content case. Dotted line is the phase diagram of the high H2 content case.

Table 4 Schematic of onshore CO2 transportation pipeline.

Table 2 Bubble point gas composition – High H2 case composition. Pressure (bar)

65.2

67.1

69.2

72

75.4

Temperature (°C) CO2 (% mol): H2 (% mol): N2 (% mol) Ar (% mol): CO (% mol)

−9.9 53.0 32.1 10.1 2.5 2.3

0 65.8 22.0 8.3 2.1 1.7

5 71.7 17.7 7.3 1.9 1.5

10 77.0 13.9 6.3 1.6 1.2

15 81.7 10.7 5.2 1.4 1.0

Parameter (Unit)

Value

Fluid inlet temperature (°C) Ambient temperature (°C) Pipeline diameter (m) Pipeline outlet pressure (bar) Low flowrate (kg/s) High flowrate (kg/s) Design Pressure (bar)

20 4 0.3 100 18.4 200 136

Table 3 Bubble point gas composition – Low H2 case composition. Pressure (bar)

24.9

46.4

50.8

55.6

60.8

Temperature (°C) CO2 (% mol): H2 (% mol): N2 (% mol) Ar (% mol): CO (% mol)

−53.2 29.0 14.3 36.8 8.7 11.2

0 83.0 1.5 10.5 2.8 2.3

5 85.8 1.2 8.8 2.3 1.8

10 88.4 0.9 7.3 1.9 1.5

15 90.8 0.7 5.9 1.6 1.1

Fig. 3. Steady state pressure profile along the CO2 (high H2 concentration) transportation pipeline. Solid line is the profile corresponds to the high flowrate case (200 kg/s). Broken-dotted line corresponds to the low flowrate case (18.4 kg/s). Broken line is the elevation profile of the pipeline.

• Gas can form under either normal or upset conditions; • The H concentration or partial pressure in the gas is higher than the bulk stream; and, Exposure times are prolonged. • 2

These criteria mean that the short-term presence of H2 rich gas during pipeline depressurization (Scenario A described in Section 1), or repressurization is not a design case. During pressurization, for example prior to start-up, there will likely be the requirement for the pipeline to be filled with injection gas, and then pressurised such that the contents condense into the liquid phase. During the condensation process, heavier components (such as CO2) will tend to condense first and lighter components (such as H2) will tend to condense last. Hence the gas in the pipeline during the phase change will tend to be H2-rich. (The reverse will happen during depressurization, again creating H2-rich gas during the pressure change transient.) However, this will only be the case for a short period of time. It follows that there are only two critical operating scenarios:

Fig. 4. Steady state temperature profile along the CO2 (high H2 concentration) transportation pipeline. Solid line is the profile corresponds to the high flowrate case (200 kg/s). Broken-dotted line corresponds to the low flowrate case (18.4 kg/s). Broken line is the elevation profile of the pipeline.

Table 1. Additional system information is shown in Table 4. Firstly, Figs. 3 and 4 show the steady state operating pressure and temperature profiles, respectively, for the low (18.4 kg/s) and high (200 kg/s) flowrate cases. The top of hill (position A) and downstream bottom of hill (position B) are also shown. It can be observed that the total pressure drop in the pipeline is very low (< 1 bar) for the low flowrate case. For the high flowrate case, the total pressure drop in the pipeline is 18 bar. The minimum pressure along the pipeline is at position A due to hydrostatic head. Table 5 shows the temperature and pressure at both positions A and B. The system is designed to operate at steady-state in single-phase (see Fig. 5) to prevent phase splitting (gas breakout), as this could cause cavitation in the pumps. Following a pipeline extended shutdown, the fluid temperature will cool down to ambient conditions. This will cause the pressure to drop due to contraction of the fluid during cooling. For both high and low

• High point bubble after shutdown (Scenario B); and, • Gas cap at the well head during well shut-in (Scenario C). 4. Results and discussion 4.1. Scenario B - high point gas pocket Let us consider the onshore CO2 transportation buried pipeline traversing a small hill shown in Fig. 2. This pipeline is a simplification of a National Grid CCS project (National Grid, 2015). For this scenario, we have considered the high H2 content composition (2% mol) from 5

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Table 5 Steady-state conditions at the top and bottom of the hill. Maximum fluid velocities corresponding to low and high flowrates are 0.32 and 3.5 m/s, respectively. Low flowrate Position

Top of hill (A) Bottom of hill (B)

Table 7 Required operating steady-state pressure to ensure no gas breakout after prolonged shutdown (settle-out pressure at top of the hill greater than 69 bar – bubble point at 4 °C).

High flowrate

Pressure (bar)

Temperature (°C)

Pressure (bar)

Temperature (°C)

88.4

15.9

97.7

17.3

98.5

16.7

105

18.2

CO2 H2 N2 Ar CO

78 14 6 1 1

150.1/150 170.1/150

135.3/147.3 145.1/152.7

PRes − PStat − PBub ≤ 0

(1)

where PRes is the pressure in the target storage reservoir; PStat is the static pressure of the injection fluid from injection point (bottom-hole) to wellhead choke, and PBub is the injection fluid bubble point pressure at ambient wellhead conditions. The size of the gas cap in the closed well is a product of the pressure and temperature gradient in the injection well. If storage reservoir pressures are very low, the entire well may be in the gas phase. However, based on the experience of the authors, a gas-filled tubing after a well shut-in is unlikely. During early life/first load, when low reservoir pressures are expected, operating at low flowrates could result in two-phase flow downstream of the injection choke valve. In general, phase splitting is undesirable as gas bubbles create cavitation leading to mechanical damage in the well tubing and tree. Normally, the system is designed to operate within the single-phase region at normal steady-state conditions. The worst-case for H2 concentration in a gas cap will likely occur during early life of the storage, assuming the initial reservoir pressure is low enough to allow for gas cap formation after shut-in. The pressure at the injection point will increase over time from the initial starting pressure to the ultimate “full” seal-and-abandon pressure. Therefore, the minimum well shut-in pressure is equal to the bubble point (at minimum ambient condition) will lead to the highest H2 concentration (and exposure time) in the gas cap.

Gas Composition (% mol)

71 18 8 2 2

Low flowrate High flowrate

A gas cap at the wellhead can occur where the storage pressure is low enough such that the top of the vertical injection column in the well is below the bubble point pressure of the injection fluid. This situation will occur at shut-in when Eq. (1) holds:

Table 6 Composition of flashed gas at ambient temperature (4 °C) and bubble point and settle-out condition.

57 bar (settle-out pressure)

Pressure (bar) at position A/ B

4.2. Scenario C - gas cap at wellhead

Fig. 6. Pressure-temperature phase envelope with settle-out condition at top of the hill (diamond) shown. For the high flowrate case (200 kg/s) the settle-out pressure is 57.8 bar. For the low flowrate case (18.4 kg/s) the settle-out pressure is 57.6 bar.

69 bar (bubble point)

Required inlet/outlet pressure (bar)

two-phase region is not entered during cooldown. Transient (cooldown) simulations were performed to determine the required initial steady state inlet and outlet pressures to ensure no gas breakout (settle-out pressure at top of the hill greater than 69 bar which is the bubble point at 4 °C) would occur following a prolonged shutdown for the low and high flowrate cases. The results are summarised in Table 7. It is noted the estimated operating pressures are significantly higher than the pipeline design pressure of 136 bar. For existing pipelines, increasing operating pressures to avoid gas breakout may not be feasible (or indeed for new pipelines, where increasing the design pressure may have a significant economic impact). In order to manage H2-rich gas breakout following a shutdown, the design should consider, as well as pipeline materials ability to be safely exposed to the worst case high H2 concentration gas (bubble-point), operational measures to reduce potential exposure time. Routine depressurization of onshore pipelines, such that the whole of the pipeline is in the single gas phase, should be considered in the event of a long shutdown as part of the pipeline operating philosophy. Additionally, the integrity management plan should be revised to ensure safe longterm operation. Fig. 7 shows a high-level design philosophy for when H2-rich gas breakout is expected during an extended shutdown. It is important to note that this is a generalized philosophy and each CCS system will have different project drivers and specific design/operating philosophy.

Fig. 5. Pressure-temperature phase envelope (CO2 stream with high H2 content) with steady-state conditions at top of the hill (A). Diamond is the condition for low flowrate case. Circle is the condition for high flowrate.

Component

Case

flowrate cases, the settle-out pressure at the top of the hill (position A) lies within the two-phase region (see Fig. 6). Gas breakout will occur and collect at the highest point in the pipeline. Table 6 shows the composition of gas evolved after the prolonged shutdown. The concentration of H2 of the persistent gas bubble at top of the hill is 14% mol. Although, this is lower than the H2 concentration in the incipient gas phase at the bubble point of 69 bar, it may represent a challenge for the metallurgy if the H2-rich gas bubble is persistent over time (depending on the maximum shutdown duration before restart or some other operational action that will remove or dilute the bubble). For no gas breakout to occur after a prolonged shutdown, the initial steady-state operating pressure must be high enough to ensure that the

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Fig. 7. High-level design philosophy for the operating scenario where H2-rich gas breakout during shutdown is expected (Scenario B).

and limited as it only considers shut-in (no flow) conditions. The minimum shut-in pressure equal to the bubble point pressure, implies that the flowing pressure downstream of the choke will be below the bubble pressure (this may be particularly true during early life when flowrates are expected to be low). This will lead to undesirable gas flashing across the injection choke (albeit of a lower H2 concentration in the gas). Fig. 10 shows the flowing pressure-temperature conditions along the injection well for the low flowrate case assuming the minimum reservoir pressure of 157 bar. It is seen that, at top the well, the system operates inside the two-phase region (flowing pressure and temperature is 59 bar and-3.5 °C). At these conditions, the concentration of H2 in the gas phase is 22.3%. Thus, it follows that, for the reservoir pressure of 165 bar (as per Table 8), the system will operate in the liquid phase region at both steady-state and shut-in conditions. Fig. 11 shows flowing pressuretemperature conditions along the injection well for the low flowrate case). It is shown that no gas breakout is expected for both normal steady-state and shut-in conditions. Table 9 provides a summary of these results. Fig. 12 shows a proposed design methodology for H2-rich gas cap during well shut-in. Note that this is a generalized philosophy and each

Let us consider the offshore CO2 injection well shown in Fig. 8. Table 8 outlines generic well tubing parameters (Yang et al., 2015) and process/subsurface conditions considered in a National Grid CCS project (National Grid, 2015). For this scenario, we have considered the high H2 content composition (2% mol) from Table 1. At the minimum ambient of -3.5 °C, the bubble point pressure is circa 65 bar with a H2 concentration in the gas phase of 25.2% mol. This provides a design basis for engineering analysis by defining the maximum concentration of H2. Fig. 9 shows the estimated pressure profile (hydrostatic head) using liquid density (obtained from flash calculations) and assuming a linear geothermal gradient (as per Fig. 8). Therefore, for the shut-in pressure of 65 bar at wellhead, the reservoir pressure required for a gas cap to occur during a prolonged shutdown is 157 bar. If the reservoir pressure is higher, the subsequent well tubing pressures would be high enough to keep the mixture within the liquid single-phase region. Complex subsurface model or rigorous flow assurance modelling is not required to set a (conservative) basis for design for maximum H2 concentration in a persistent gas cap. The design point is a product of ambient temperature (at the wellhead) and the maximum design H2 composition of the injection gas. However, it is noted that the above approach may be conservative 7

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Fig. 9. Pressure (solid line) and density (broken line) profile along the well at shut-in condition assuming geothermal profile (with reservoir pressure of 157 bar and shut-in pressure top of well of 65 bar).

Fig. 10. Pressure-temperature phase envelope (CO2 stream with high H2 content) with steady-state (solid line) and shut-in/settled-out (broken line) pressure-temperature conditions along injection well for low flowrate case and minimum reservoir pressure of 157 bar.

Fig. 11. Pressure-temperature phase envelope (CO2 stream with high H2 content) with steady-state (solid line) and shut-in/settled-out (broken line) pressure-temperature conditions along injection well for low flowrate case and reservoir pressure of 165 bar.

Fig. 8. Schematic of offshore CO2 well injection. Ambient and geothermal gradient temperatures are shown.

Table 8 Generic injection well parameters.

5. Conclusions and recommendations

Parameter (unit)

Value

Water depth (m) Well tubing depth (m) Well tubing size (in) Injectivity index (kg/s/bar) Reservoir temperature (°C) Initial reservoir pressure (bar) Air temperature (°C) Seawater temperature (°C) Pressure upstream injection choke(bara) Temperature downstream injection choke (°C) Reservoir type (-)

60 1300 5 0.8 60 165 −3.5 4 100 −3.5

Where pre-combustion capture or hydrogen manufacture is a source of CO2, H2 will be a possible contaminant in the main stream. The potential for long-term exposure of H2-rich gas to certain parts of a CCS gathering and injection network must be considered in the design. In this work, we have presented two operating scenarios in which the CCS transportation network H2-rich gas, higher in concentration than the bulk stream, will concentrate and persist where small amounts of gas are present in a mostly-liquid network. The first scenario considers H2rich gas breakout following shutdown at a high-point (such as at top of a hill) of an onshore CO2 transportation pipeline. The second case is a H2-rich gas cap formed following well shut-in. In both design scenarios, the maximum concentration of H2 that a system may be exposed to corresponds to the bubble-point pressure (at minimum ambient conditions). Our results are based on the maximum local concentration of H2 of 18%–25% mol which corresponds to the composition of the incipient gas phase at bubble-point conditions. For conservatism, the design (material fitness for service) should be tested against these high H2 concentrations. However, if these assumptions are too onerous, driving the design to expensive materials or coatings, conservatism of the H2-rich gas breakout can be reduced if actual settle-out/shut-in conditions are within the two-phase region (where H2 content of the gas phase will be lower than at bubble-point). To our knowledge, there are not specific guidelines/standards that

Depleted reservoir with partial CO2 storage

project will have different drivers and constraints. Should the design case prove unreasonably onerous (if a gas cap is formed), by driving the design towards expensive well materials or linings for example, there are mitigative options that can be considered. These might include dilution/displacement with nitrogen, or reduction of free H2 by adsorption onto an injected compound.

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Table 9 Summary of key results of Scenario C (Gas cap after well shut-in) for the low flowrate case. Reservoir pressure (bar)

157 165

Flowing conditions

Shut-in conditions

Bottom-hole pressure (bar)

Top of well pressure (bar)

H2 concentration in gas phase (% mol)

Top of well pressure (bar)

H2 concentration in gas phase (% mol)

180 211

59 87

22.3 No gas breakout

65 94

25.2 No gas breakout

1 (NACE, 2001) or for hydrogen service such as ASME B31.12-2011 (ASME, 2012). Furthermore, we note that the presence of hydrogen in natural gas pipelines has been studied and concentrations greater than 15% mol present significant challenges with respect to material compatibility and safety (Melania et al., 2013). Nevertheless, it is worth to

CCS transportation pipelines must comply regarding the presence of hydrogen. While CCS transportation pipelines should adhere to standards such ISO 27913 (ISO, 2016) and DNV-RP-J202 (DNV, 2010), we also recommend that CO2 pipelines (with H2 as one of the impurities) should be compliant for sour service such as NACE MR0175/ISO 15156-

Fig. 12. High-level design philosophy for H2 gas cap during well shut-in (Scenario C). 9

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state that natural gas distribution networks are not operated in the same way that CCS gathering and injection networks and are unlikely to have the same topographic/design features. In order to ensure safe long-term operation, an integrity management plan should consider the risks and assess the suitability of (existing) pipelines for CO2 transportation (Rabindran et al., 2011) with impurities like hydrogen. Also, we recommend that the design of CCS transportation should consider mitigative measures to reduce the exposure time of the system to H2 rich gas. Specifically:

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• The • •

onshore pipeline could be depressurized in the event of a medium- or long-term shutdown, which will reduce the maximum local H2 content of the gas to that of the bulk fluid. Options to displace, preferentially adsorb, or dilute H2 rich gas through chemical or inert gas injection. Options to bleed-off H2-rich gas at the wellhead as long this is a safe operation.

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