Thermodynamic Inhibitors – Part 3 – Pipeline Water Wet Condensate Study

Currently, there are several different thermodynamic hydrate inhibitors that one has available to choose from. One could choose to use an oxygenate, such as methanol, or a glycol, such as mono-ethylene glycol (MEG) or diethylene glycol (DEG). But what determines the inhibitors hydrate suppression performance, and which one should I choose for my inhibition needs? The primary considerations for an application include:

1. Hydrate Suppression Effectiveness

2. Inhibitor Regeneration Requirements

3. Inhibitor Losses / Product Contamination

4. Inhibitor Cost

Hydrate Suppression Effectiveness

The lower the molecular weight of a thermodynamic inhibitor, the better the hydrate suppression performance. For example, in glycols MEG has better performance characteristics than DEG, and methanol MeOH will outperform ethanol EtOH. In Part 1 of this tip of the month (November 2020 TOTM [1]) provided insight into the relative performance of hydrate inhibitors, and discussed the limitations, pros and cons of the various options. Part 1 demonstrated that for a hydrate depression temperature of 10^oC (18^oF) to prevent hydrates formation in our gathering system the required concentration of inhibitor in the liquid water phase using the Nielsen and Bucklin [2] equation is 13 mole %. Figure 1 provides the corresponding weight % of the various inhibitors that would be required for this depression temperature (d). As shown in Figure 1, the inhibitors weight % is a function of molecular weight shown on the horizontal axes.

Figure 1. Comparison of Inhibitor Requirements in Liquid Water Phase for Depression Temperature of 10^oC (18^oF)

In this TOTM we will revisit the Case Study of December 2020 TOTM [3] in the application and selection of inhibitor for a gathering system inhibition application. For the same Case Study, we will include the entrained water in the hydrocarbon condensate entering the pipeline.

Revised Case Study

To prepare a hydrate prevention program for the subsea pipeline, let us consider a multiphase subsea pipeline transporting gas and condensate from an offshore production facility to shore. The flow rates and pipeline conditions are provided in Table 1 and Figure 2. The feed condensate to pipeline water-cut is 0.2 volume percent. The solubility of methanol to the hydrocarbon liquid phase is a strong function of both the water phase concentration of inhibitor, but also the hydrocarbon phase composition. The previous case study assumed that the condensate that was separated on the platform was bone-dry when it was pumped into the pipeline with the gas phase. This tip of the month is going to focus on the impact of water-cut in the condensate on the inhibitor injection rates. As with the previous case study, for the methanol injection estimates, we will include both the vapor and liquid hydrocarbon losses.

Table 1. Flow rates and pipeline conditions

Figure 2. Schematic of subsea pipeline

To determine the impact of the water-cut of the condensate on the inhibitor injection requirement, one needs to calculate the mass of additional water this adds to the pipeline. The additional water is estimated as follows:

► Liquid water rate = (water-cut fraction, W-C_x) (condensate flow rate, q_c) (fresh-water density, ρ_water)

► Liquid water rate = (0.2 /100) (90 m³/d) (1000 kg/m³) =180 kg/d

► Liquid water rate = (0.2 /100) (566 bbl/d) (5.615 ft³/bbl) (62.4 lbm/ft³) =397 lbm/d

The water content of lean sweet natural gas is a function of temperature and pressure and can be estimated from the charts in chapter 6 of volume 1 of Gas Conditioning and Processing [4]. The water content of the gas at the inlet and outlet of pipeline are used to the determine the amount of condensed water in the transmission line. The liquid water associated with watercut should be added to the condensed water. The total liquid water is a first key factor that establishes the amount of hydrate inhibitor required. Likewise, the hydrate formation temperature (HFT) is a function of pressure and relative density of gas and can be estimated from the chart in reference [4]. The second key factor is the required depression temperature (d=HFT-T_cold). The HFT is an important parameter for determination of the depression temperature. The summary of preliminary calculation results is presented in Table 2.

Table 2. Summary of preliminary calculation results

The hydrate inhibitor performance plots that are presented in this tip were developed using the Nielsen and Bucklin [2] correlation presented as equation 1.

(1)

Where: d = depression of hydrate point. °C (°F in FPS)

x_m = mole fraction inhibitor in the liquid water phase

A = constant, -72 (129.6 in FPS)

The third key factor is the required inhibitor weight % in the liquid water phase, X_R. From Figure 3 for depression temperature. d = 12.1 °C (22 °F), X_R for MeOH, MEG and DEG are determined and shown in Table 3. The inhibitor concentration in the same phase for all three inhibitors is X_R= 15.6 mole %.

Figure 3. Depression temperature vs mole % and weight % hydrate inhibitor in the liquid water phase for Typical Gathering Systems

From Figure 4 for depression temperature, d = 12.1 °C (22 °F), and the required ratio of mass of inhibitor in the liquid water phase to mass of condensed water, (m_I/ m_W) for three inhibitors are determined and shown in Table 3. The required inhibitor rate in liquid water phase, m_I, the fourth key factor, can be calculated by equation 2.

(2)

Where: m_I = required inhibitor mass rate in liquid water, kg/d (lbm/d in FPS)

m_W = condensed water mass rate, kg/d (lbm/d in FPS)

(m_I /m_w ) = Inhibitor effectiveness factor, function of d (From Figure 4)

From Figure 4 at the same depression temperature, the required mass ratio, (m_I/ m_W), for lean MeOH 100 %, MEG 100 wt% and DEG 100 wt%, are 0.33%, 0.63% and 1.08, respectively; and for MEG 80 wt% and DEG 80 wt%, are 0.96% and 1.86, respectively.

Figure 4. Depression temperature vs ratio of mass of inhibitor to mass of condensed water, m_I/m_W

The summary of the required inhibitors rate calculations for liquid water phase are presented in Table 3.

Table 3. Summary of required inhibitors injection rate for liquid water phase

Figure 5 and Table 3 present comparison of inhibitors to liquid-water mass ratio, (m_I/ m_W), requirements for MeOH, MEG and DEG at various lean inhibitor weight %. Figure 5 indicates that methanol with MW=32 is the most effective because of its lowest mass ratio requiring the lowest inhibitor mass rate while DEG with MW = 106 is the least effective because its largest mass ratio requiring the most inhibitor mass rate.

Figure 5. Comparison of inhibitors to liquid water mass ratio, (m_I/ m_W)

MeOH Losses to Liquid and Vapor Hydrocarbon Phase

Due to low vapor pressure and low solubility in liquid hydrocarbons, glycol losses to vapor and liquid hydrocarbon phases are negligible, and we will assume these to be zero. However, MeOH losses to vapor and liquid hydrocarbon phases are significant and estimated in this tip.

The MeOH losses to liquid hydrocarbon phase can be estimated by our proposed Figure 6 and equation 3.

(3)

Where: mLF = MeOH loss factor from Figure 6,

kg MeOH/m³ (lbm MeOH/bbl condensate in FPS)

q_L = condensate rate, m³/d (bbl/d in FPS)

ɣ = condensate relative density

MW = condensate molecular weight

From Figure 6 for depression temperature, d = 12.1 °C (22 °F), the required MeOH loss factor is 0.44 kg/m³ (0.154 lbm/bbl). For a condensate rate of 90 m³/d (566 bbl/d), relative density of ɣ = 0.59 and molecular weight of MW = 55, calculate rate of MeOH loss to liquid hydrocarbon phase by equation 3.

Figure 6. MeOH loss factor to liquid hydrocarbon phase

The MeOH losses to vapor hydrocarbon phase can be estimated by our proposed equation 4 and Figure 7 [4].

Where: K = MeOH vapor-liquid equilibrium ratio

q_S = gas rate, std m³/d (scfd in FPS)

x = MeOH mole fraction in liquid water phase

A = conversion factor, 1.35×10⁶ (84 300 in FPs)

From Figure 7 for cold temperature of 4.4 °C (40 °F) and pressure of 6.2 MPa (900 psia) the K value is 0.0019. For a gas rate of 1.4125 x 10⁶ std m³/d (50 MMscfd) and MeOH mole fraction of x=0.156, calculate rate of MeOH losses to vapor hydrocarbon phase by equation 4.

Figure 7. MeOH K-value in natural gas

The total required inhibitors injection rate to prevent hydrate formation is calculated by equation 5.

(5)

Where: m_T = total required inhibitor rate

m_I = required inhibitor rate in liquid water, this is the inhibitor amount that prevents hydrate formation

m_LHC = inhibitor rate loss to liquid hydrocarbon phase

m_VHC = inhibitor rate loss to vapor hydrocarbon phase

Table 4 presents summary of total required inhibitors injection rate to prevent hydrate formation. Figure 8 also presents comparison of water phase inhibitor rate compared (blue bars) to total required inhibitors injection rate (red bars) to prevent hydrate formation.

Table 4. Summary of total required inhibitors injection rate to prevent hydrate formation.

The inhibitors injection rate in parentheses are for the case of excluding condensate entrained liquid water.

Table 4 and Figure 9 indicate that including the (0.2 vol% watercut) entrained water increases the inhibitor injection rates by 6.7 % for MeOH and 20% for MEG and DEG.

Figure 8. Comparison of total required inhibitors injection rate to prevent hydrate formation for wet (water-cut 0.2 vol%) and bone-dry (water-cut 0.0 vol%) condensate.

Figure 9. Comparison of total required inhibitors injection ratio to prevent hydrate formation for wet (water-cut 0.2 vol%) and bone-dry (water-cut 0.0 vol%) condensate.

In fact, the losses are significantly greater than the rate required in the water phase which results in higher injection rates required as compared to MEG and DEG. For our Case Study the sum of MeOH losses to the vapor and liquid hydrocarbon phases is about 1.6 to two times greater than the required rate of inhibitor in the liquid water phase. The amount of MeOH losses to the vapor phase is a function of pressure, temperature, and mole fraction of MeOH in the liquid water phase. But the amount of MeOH losses to the condensate phase is a function of temperature, mole fraction of MeOH in the liquid water phase, condensate relative density, molecular weight, and concentration of aromatic hydrocarbons. Note that mole fraction of MeOH in the liquid water phase is a function of the required depression temperature.

Once you inject methanol into the pipeline, it distributes to all phases: it is in the gas, the total amount of water (condensed and entrained), and the condensate. This is becoming a significant issue recently. Too much MeOH in the produced water can kill the microbes in the produced water treatment system (and it is toxic to fish so produced water offshore could not go overboard). In addition, MeOH is a petrochemical catalyst poison. Many North American natural gas, y-grade (NGL), and liquid products pipeline tariffs are specifying a maximum oxygenate concentration – or specifying ZERO! In addition, the methanol in the natural gas will slowly kill your very expensive 4A molecular sieve dehydration bed over time if the natural gas is being processed in an expander plant.

Glycols on the other hand have very low vapor pressure, with minimal losses in the vapor phase. In addition, there is very little mutual solubility with liquid hydrocarbons, so the injected glycol remains in the phase that you want it to be in, the water phase. The difficulty of using glycols is that the producer / owner operator needs to own the equipment on both ends of the pipeline so that the glycol can be separated in the plant inlet slug-catcher and inlet separators, so that it can be regenerated and either shipped, trucked or pumped back to the injection points in the gathering system.

Inhibitor Cost

Methanol is cheap in comparison to glycols. As a rough rule of thumb, glycols are roughly three times the cost of methanol. That is why glycols are nearly always regenerated. You only need to invest in the initial chemical costs once. However, methanol is cheap, but you will need to continually buy the amount of chemical you are injecting as it goes down stream and is not recovered.

Summary

This TOTM reviewed a Case Study in the application and selection of inhibitor for a feed gas with water wet condensate in a gathering system inhibition application. The contribution of entrained water in condensate to the inhibitor were considered. Table 4 and Figure 9 indicate that including the (0.2 vol% water-cut) entrained water increases the inhibitor injection rates by 6.7 % for MeOH and 20% for MEG and DEG. If the entrained water in the condensate is not taken into account for the inhibitor injection rates, then there is a strong possibility that hydrates may form in the pipeline because there is insufficient inhibitor present in the water phase. In addition, the pro’s and con’s of oxygenates versus glycols was discussed in more detail.

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing) and G5 (Practical Computer Simulation, and Applications in Gas Processing) courses.

By: Kindra Snow-McGregor, P.E., and Mahmood Moshfeghian, Ph.D.

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References:

1. Snow–McGregor, K., and Moshfeghian, M. “Thermodynamic Hydrate Inhibitors – How Do They Compare? Part 1: General Considerations,” PetroSkills -John M. Campbell Tip of the Month, November 2020.

2. Nielsen, R.B., and R.W. Bucklin, “Why not Use Methanol for Hydrate Control?”, Hyd. Proc., Vol. 62, No. 4 (Apr. 1983), p. 71.

3. Snow–McGregor, K., and Moshfeghian, M. “Thermodynamic Hydrate Inhibitors – How Do They Compare? Part 2: Case Study,” PetroSkills -John M. Campbell Tip of the Month, December 2020.

4. Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9^th Edition, 3^rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018.

Written on February 1, 2021 at 2:10 pm, by admin

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Solubility of Gases in Water: Part 1

Part 1: Light Hydrocarbon Gases

Why should we care about the solubility of gases in water?

The solubility of hydrocarbons and non-hydrocarbons like carbon dioxide and hydrogen sulfide in water is of interest for oil and gas production and processing facilities dealing with water treatment. It is also important for disposal facilities from an environmental aspect. In addition, determination of the solubility of these components in aqueous phase is critical in the study of the gas hydrates kinetics [1].

In Part 1 of this series, we will focus on the solubility of light hydrocarbon gases such as CH₄, C₂H₆, C₃H₈, iC₄H₁₀, and nC₄H₁₀. We will review and present example available experimental solubility data and a thermodynamic model. For the case study, we will present diagrams to show the effect of temperature and pressure on the solubility of these light hydrocarbon gases in water.

Solubility of Light Hydrocarbon Gases

In general, the solubility of a gas (solute) in water (solvent) is a function of temperature, pressure, and composition. The solubility of gases in water can be measured experimentally or estimated by thermodynamic models using equation of state or a correlation based the Henry’s law.

In the thermodynamic model [2] developed by Zirrahi et al., the Peng-Robinson equation of state coupled with a non-random mixing rule is used to model gas phase (applicable to sweet, sour, or acid gases). The aqueous phase is modeled by Henry’s law approach (applicable to the fresh and brine water). They reported good agreement between their model and experimental data available in the literature. The aqueous phase solubility and water content of pure CH₄, CO₂, and H₂S were represented with absolute average relative deviations of less than 6.3% and 5.6%, respectively.

Figure 1 presents a comparison between the experimentally measured CH₄ solubility in water by different investigators and the results of the proposed model by Zirrahi et al. [2]. Solid lines are the results obtained from the proposed model at (a) 298.15 K (25°C), and (b) 338.2 K (65.1°C), 324.65 K (51.5°C), 310.93 (37.8°C) K, 303.2 K (30.1°C), and 285.67 K (12.5°C).

Figure 1. Comparison between solubility of CH₄ experimental data and results of a proposed model by Zirrahi et al. [2]. Solid lines are the results obtained from the model at (a) 298.15 K (25°C), and (b) 338.2 K (65.1°C), 324.65 K (51.5°C), 310.93 (37.8°C) K, 303.2 K (30.1°C), and 285.67 K (12.5°C).

CASE STUDY

To investigate the effect of temperature, pressure, and composition on the solubility of light hydrocarbon gases in water, we considered a lean gas mixture with the composition shown in Table 1.

Table 1. Composition of gas mixture

We obtained the solubility of each component in water from ProMax [3] using the Peng-Robinson equation of state. Four isotherms of 5°C, 10°C, 20°C, and 30°C (41°F, 50°F, 68°F, and 86°F) were selected. For each isotherm, the pressure was varied from 200 kPa to 8000 kPa (29 psia to 1160 psia). The calculated gas solubility results, std m³ of gas per m³ of water (SCF of gas per bbl of water) are presented in Figures 2 and 3.

Figures 2 and 3 indicate that:

1. For a given pressure and temperature, methane has the highest and n-butane has the lowest solubility in the aqueous phase.

2. For an isotherm, methane and ethane solubility increase as pressure increases.

3. For an isotherm, propane solubility increases as pressure increase up to 3000 kPa (435 psia) beyond which remains constant.

4. For an isotherm, i-butane and n-butane solubility increase as pressure increases up to 3000 kPa (435 psia) beyond which decreases with pressure increase.

Figure 2. Variation of solubility of CH₄,C₂H₆,C₃H₈, iC₄H₁₀, and nC₄H₁₀ with pressure and at low temperatures

Figure 3. Variation of solubility of CH₄,C₂H₆,C₃H₈, iC₄H₁₀, and nC₄H₁₀ with pressure and at moderate temperatures

Summary

This TOTM presented the solubility of light hydrocarbon gases, i.e., methane through normal butane, in water as a function of pressure and temperature. Simple charts (Figures 1a and b) to estimate the solubility of pure methane in water were presented. Using ProMax [3], the solubility charts for a mixture of methane through normal butane (Table 1) in water for four isotherms are also presented (Figures 2 and 3).

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Practical Computer Simulation, and Applications in Gas Processing) courses.

By: Mahmood Moshfeghian, Ph.D.

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References

1. Sloan, E.D., Koh, C.A., Clathrate Hydrates of Natural Gases, 3^rd ed., CRC Press, 2008.

2. Zirrahi, M., Azin, R., Hassanzadeh, H., Moshfeghian, M. “Mutual solubility of CH₄, CO₂, H₂S, and their mixtures in brine under subsurface disposal conditions,” Fluid Phase Equilibria, Fluid Phase Equilibria 324, 80–93. 2012

3. ProMax 5.0, Build 5.0.20034.0, Bryan Research and Engineering, Inc., Bryan, Texas, 2020.

Written on January 1, 2021 at 10:36 am, by admin

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Thermodynamic Hydrate Inhibitors – Part 2 – Pipeline Case Study

Part 2 – Thermodynamic Hydrate Inhibitors Series

1. Hydrate Suppression Effectiveness

2. Inhibitor Regeneration Requirements

3. Inhibitor Losses / Product Contamination

4. Inhibitor Cost

Hydrate Suppression Effectiveness

The lower the molecular weight of a thermodynamic inhibitor, the better the hydrate suppression performance. For example, in glycols MEG has better performance characteristics than DEG, and methanol MeOH will outperform ethanol EtOH. In Part 1 of this tip of the month (November 2020 TOTM [1]) provided insight into the relative performance of hydrate inhibitors, and discussed the limitations, pros and cons of the various options. Part 1 demonstrated that for a hydrate depression temperature of 10 ^oC (18^oF) to prevent hydrates formation in our gathering system the required concentration of inhibitor in the liquid water phase using the Nielsen and Bucklin [2] equation is 13 mole %. Figure 1 provides the corresponding weight % of the various inhibitors that would be required for this depression temperature (d). As shown in Figure 1, the inhibitors weight % is a function of molecular weight shown on the horizontal axes.

Figure 1. Comparison of Inhibitor Requirements in Liquid Water Phase for Depression Temperature of 10^oC (18^oF)

In this TOTM we will review a Case Study in the application and selection of inhibitor for a gathering system inhibition application. Read Part One here.

Case Study

Table 1. Flow rates and pipeline conditions

Figure 2. Schematic of subsea pipeline

The water content of lean sweet natural gas is a function of temperature and pressure and can be estimated from the charts in chapter 6 of volume 1 of Gas Conditioning and Processing [3]. The water content of the gas at the inlet and outlet of pipeline are used to the determine the amount of condensed water in the pipeline. The condensed water is a first key factor that establishes the amount of hydrate inhibitor required. Likewise, the hydrate formation temperature (HFT) is a function of pressure and relative density of gas and can be estimated from the chart in reference [3]. The second key factor is the required depression temperature (d=HFT-T_cold). The HFT is an important parameter for determination of the depression temperature. The summary of preliminary calculation results is presented in Table 2.

Table 2. Summary of preliminary calculation results

The hydrate inhibitor performance plots that are presented in this tip were developed using the Nielsen and Bucklin [2] correlation presented as equation 1.

Figure 3. Depression temperature vs mole % and weight % hydrate inhibitor in the liquid water phase for Typical Gathering Systems

From Figure 4 for depression temperature, d = 12.1 °C (22 °F), the required ratio of mass of inhibitor in the liquid water phase to mass of condensed water, (m_I/ m_W) for three inhibitor are determined and shown in Table 3. The required inhibitor rate in liquid water, m_I, the fourth key factor, can be calculated by equation 2.

From Figure 4 at the same depression temperature, the required mass ratio, (m_I/ m_W), for lean MEG 100 wt% and DEG 100 wt%, are 0.63% and 1.08, respectively.

Figure 4. Depression temperature vs ratio of mass of inhibitor to mass of condensed water, m_I/m_W

The summary of the required inhibitors rate calculations for liquid water phase are presented in Table 3.

Table 3. Summary of required inhibitors injection rate for liquid water phase

Figure 5 and Table 3 present comparison of inhibitors-to-condensed water mass ratio, (m_I/ m_W), requirements for MeOH, MEG and DEG at various lean inhibitor weight %. Figure 5 indicates that methanol with MW=32 is the most effective because of its lowest mass ratio requiring the lowest inhibitor mass rate while DEG with MW = 106 is the least effective because its largest mass ratio requiring the most inhibitor mass rate.

Figure 5. Comparison of inhibitors to liquid water mass ratio, (m_I/ m_W)

MeOH Losses to Liquid and Vapor Hydrocarbon Phase

The MeOH losses to liquid hydrocarbon phase can be estimated by our proposed Figure 6 and equation 3.

Figure 6. MeOH loss factor to liquid hydrocarbon phase

The MeOH losses to vapor hydrocarbon phase can be estimated by our proposed equation 4 and Figure 7 [3].

From Figure 7 for cold temperature of 4.4 °C (40 °F) and pressure of 6.2 MPa (900 psia) the K value is 0.0019. For agas rate of 1.4125 x 10⁶ std m³/d (50 MMscfd) and MeOH mole fraction of x=0.156, calculate rate of MeOH losses to vapor hydrocarbon phase by equation 4.

Figure 7. MeOH K-value in natural gas

The total required inhibitors injection rate to prevent hydrate formation is calculated by equation 5.

Table 4. Summary of total required inhibitors injection rate to prevent hydrate formation.

Figure 8. Comparison of total required inhibitors injection rate to prevent hydrate formation.

Despite MeOH excellent performance in terms of hydrate suppression, its biggest downfall is its distribution behavior. As shown in Table 4 and Figure 8, methanol was distributed in all three phases of liquid water, liquid and vapor hydrocarbon phases. In fact, the losses are significantly greater than the rate required in the water phase which results in higher injection rates required as compared to MEG and DEG. For our Case Study the sum of MeOH losses to the vapor and liquid hydrocarbon phases is two times greater than the required rate of inhibitor in the liquid water phase. The amount of MeOH loss to the vapor phase is a function of pressure, temperature, and mole fraction of MeOH in the liquid water phase. But the amount of MeOH loss to the condensate phase is a function of temperature, mole fraction of MeOH in the liquid water phase condensate relative density and molecular weight. Note that mole fraction of MeOH in the liquid water phase is a function of depression temperature.

Once you inject methanol into the pipeline, it is in the gas, the condensed water, and the liquid hydrocarbons. This is becoming a more significant issue these days. Too much MeOH in the produced water can kill the microbes in the produced water treatment system (and it is toxic to fish so produced water offshore could not go overboard). In addition, MeOH is a petrochemical catalyst poison. Many North American natural gas, y-grade (NGL), and liquid products pipeline tariffs are specifying a maximum oxygenate concentration – or specifying ZERO!. In addition, methanol will slowly kill your very expensive 4A molecular sieve dehydration bed over time.

Inhibitor Cost

Methanol is cheap in comparison to glycols. As a rough rule of thumb, glycols are roughly three times the cost of methanol. That is why glycols are nearly always regenerated. You only need to invest in the chemical costs once. However, methanol is cheap, but you will need to continually buy the amount of chemical you are injecting as it goes down stream and is not recovered.

Summary

This TOTM presented two new charts, Figures 4 and 6, and their corresponding equations. Both figures and equations are simple and easy to use. Figure 4 can be used to estimate the required inhibitor injection rate in liquid water phase for specified depression temperature and rate of water condensed. Figure 6 can be used to estimate the rate of MeOH loss to the liquid hydrocarbon phase if depression temperature, cold temperature and condensate relative density are known. This TOTM also reviewed a Case Study in the application and selection of inhibitor for a gathering system inhibition application. In addition, the pro’s and con’s of oxygenates versus glycols was discussed in more detail.

By: Kindra Snow-McGregor, P.E., and Mahmood Moshfeghian, Ph.D.

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References

2. Nielsen, R.B., and R.W. Bucklin, “Why not Use Methanol for Hydrate Control?”, Hyd. Proc., Vol. 62, No. 4 (Apr. 1983), p. 71.

3. Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9^th Edition, 3^rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018

Written on December 7, 2020 at 12:15 pm, by admin

1 Comment
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1 response to “Thermodynamic Hydrate Inhibitors – Part 2 – Pipeline Case Study”

Harab says:

January 3, 2021 at 3:14 am

would like to know about convergence pressure Pk, assume that we have reservoir fluid composition (Pres, Tres, Zi) and we would like to estimate Pb pressure (unknown), by using Rachford Rice equation(RR eq.) when we want to calculate the K value we need the value of Pk, Is there any equation to calculate Pk direct ,please explain how to estimate Pk ,

Thermodynamic Hydrate Inhibitors – How Do They Compare?

Part 1. General Considerations

Currently, there are several different thermodynamic hydrate inhibitors that one has available to choose from. One could choose to use an oxygenate, such as methanol, or a glycol, such as such as mono-ethylene glycol (MEG) or diethylene glycol (DEG). But what determines the inhibitors hydrate suppression performance, and which one should I choose for my inhibition needs? The primary considerations for an application include:

Hydrate Suppression Effectiveness
Inhibitor Regeneration Requirements
Inhibitor Losses / Product Contamination
Inhibitor Cost

Hydrate Suppression Effectiveness

The lower the molecular weight of a thermodynamic inhibitor, the better the hydrate suppression performance. For example, in glycols MEG has better performance characteristics than diethylene glycol DEG, and methanol MeOH will outperform ethanol EtOH. This tip of the month provides insight into the relative performance of hydrate inhibitors, and discusses the limitations, pros and cons of the various options. In this TOTM we will discuss the general considerations for thermodynamic inhibitor selection. In Part 2, we will review a case study on the difference in performance between MeOH, MEG and DEG for a gathering system inhibition application.

To give insight into the hydrate suppression performance of various inhibitors, we first want to highlight that thermodynamic hydrate inhibitor suppression is a function of the mol percent of the inhibitor in the free water in the system. The hydrate inhibitor performance plots that will be presented in this tip were developed using the Nielsen and Bucklin [1] correlation, and is plotted with the Hammerschmidt [2] correlation, as well as laboratory equilibrium hydrate inhibition data in Figure 1.

Figure 1. Hydrate Suppression vs. Inhibitor Concentration in Mol %, Comparison and Correlation of Data [3]

Reviewing Figure 1, we see that the Nielsen-Bucklin correlation does an excellent-to reasonable job for all inhibitor concentrations, as compared to the Hammerschmidt equation, in matching the laboratory equilibrium data (data points plotted and referenced in Figure 1).

The Nielsen-Bucklin equation is provided in Equation 1.

(1)

Table 1. provides the molecular weights of the inhibitors that will be used for comparison.

Table 1. Hydrate Inhibitor Molecular Weight (MW)

Inhibitor	MeOH	EtOH	NaCl	MEG	DEG	TEG
MW	32	46	58	62	106	150

Using Equation 1, and plotting x_m as a function of mol %, and then converting the mol% to weight percent (wt%) for each inhibitor results in Figure 2a (SI) and 2b (FPS). Recall, the wt% value is the wt% of the inhibitor required in the free water in the system to prevent hydrate formation.

Figure 2a. Hydrate Inhibitor Suppression Effectiveness

Figure 2b. Hydrate Inhibitor Suppression Effectiveness (FPS)

Given these plots are bit too small to read effectively, it is useful to create two separate, more focused plots. In most gathering systems, the depression temperatures required are typically no greater than 20 ^oC, or 36 ^oF. Figure 3 provides the useful correlation data that can be used to estimate the required wt% of hydrate inhibitor in the rich water phase for typical pipeline / gathering system conditions.

An example of how the various inhibitors compare in terms of suppression performance, let’s assume we needed to achieve a hydrate depression temperature of 10 ^oC [18^oF] to prevent hydrates from forming in our gathering system. Figure 4 provides the wt% of the various inhibitors that would be required for this depression temperature. The results are tabulated in Table 2 below. As shown in Figure 4, the concentration for all inhibitors is the same value of 13 mole %.

Table 2. Comparison of Rich Inhibitor Requirements for 10^oC [18^oF] Depression T

Inhibitor	MeOH	EtOH	NaCl	MEG	DEG
Wt% Required	21%	27.5%	32.7%	34%	46.7%

As can be observed, as the inhibitors MW increases, the amount of inhibitor in the water phase increases for the same depression temperature. Notice TEG is off the chart and would require more than 50 wt% concentration. TEG is not an effective hydrate inhibitor, which is why it is not used in this application. However, it is exceptionally efficient at gas dehydration at very high wt % concentrations, which is why it is the workhorse in the industry in gas dehydration.

Reviewing the remaining inhibitors that can meet this depression temperature requirement, it appears that the oxygenates (MeOH, EtOH) would be the inhibitor of choice as they would require the smallest injection requirements. However, MeOH and EtOH have high vapor pressure, which results in the inhibitor being transported primarily in the vapor phase (which is not where we want the inhibitor to be). As a result of the high vapor pressure, it distributes preferentially into the vapor phase.

Figure 3. Hydrate Inhibitor Suppression Performance for Typical Gathering Systems

Figure 4. Hydrate Inhibitor Suppression Performance for Typical Gathering Systems

In addition, MeOH and EtOH have some mutual solubility with liquid hydrocarbons which is hard to characterize and model. These issues will be discussed further in the Inhibitor Losses and Product Contamination sections.

For deep depression temperature requirements, such as that occur in NGL extraction facilities (propane refrigeration or JT valve plants), the range of depression temperatures can be up to 80 ^oC [144 ^oF]. Figure 5 provides the hydrate suppression performance of the various thermodynamic inhibitors in this depression range. Again, you see the same characteristics in terms of performance of the inhibitors, however, at these very cold processing temperatures – minimum -40^oC [-40^oF], the hydrate inhibitor freezing points must be considered. The heavier the glycol becomes, the narrower the window of operation in terms of freezing points becomes. As a result, MEG is the inhibitor of choice for NGL extraction facilities that are operating at cold temperatures.

Figure 5. Hydrate Inhibitor Suppression Performance for Typical NGL Extraction Requirements

Inhibitor Regeneration Requirements

Both MEG and DEG are easily regenerated, whereas MeOH and EtOH present a more difficult regeneration problem as they form an azeotrope with water and aromatics require a complex distillation design. The relative volatility between water and methanol is only 2.5, which indicates a very difficult separation (i.e. very tall and expensive distillation tower). There are only a limited number of installations where this has been done.

When MeOH is used in gathering system hydrate inhibition applications, typically it will not be recovered. The methanol will travel with the various streams, gas, liquid, produced water which ultimately can result in operating problems downstream.

Glycols are nearly always regenerated because of its high purchase cost. Essentially the glycol is heated up in cross exchangers with the regenerated hot lean glycol, and then flows to a reboiler which boils the water concentration out of the glycol to the desired lean concentration. Typical lean glycol concentrations for hydrate inhibition range from 60 – 80 wt %. When MEG used in production systems vs. NGL extraction (refrigeration or JT valve) plants, they will have different regeneration issues.

For NGL extraction facilities, the general issues are emulsions with the liquid hydrocarbons, hydrocarbon carryover to the regenerator, and aromatic emissions (potentially), from the regenerator.

For production systems, the MEG is often times inhibiting produced water which will have some amount of salinity from the reservoir. The rich MEG from the inlet slug catcher of the facility will contain those water with the dissolved salts which will then flow to the regeneration system. In conventional MEG regeneration units, the salt will accumulate and precipitate in the regeneration system causing significant fouling, corrosion and likely result in MEG thermal decomposition [3]. Additional equipment will be required to remove the salts from the MEG. These processes are licensed and are an additional expense that needs to be considered and planned for in the hydrate inhibitor selection process.

Inhibitor Losses / Product Contamination

Despite oxygenates excellent performance in terms of hydrate suppression, their biggest downfall is their distribution behavior. Methanol will distribute in EVERY phase in the system. In fact, the losses are significantly greater than the concentration required in the water phase which results in higher injection rates required as compared to MEG for example

Once you inject methanol into the pipeline, it is in the gas, the condensed water, and the liquid hydrocarbons. This is becoming a more significant issue as of late. Too much MeOH in the produced water can kill the bugs in the produced water treatment system (and it is toxic to fish so produced water offshore could not go overboard). In addition, oxygenates are a petrochemical catalyst poison. Many North American natural gas, y-grade (NGL), and liquid products pipeline tariffs are specifying a maximum oxygenate concentration – or specifying ZERO! In addition, methanol will slowly kill your very expensive 4A molecular sieve dehydration bed over time.

Inhibitor Cost

Methanol is cheap in comparison to glycols. As a rough rule of thumb, glycols are roughly three times the cost of methanol. That is why glycols are nearly always regenerated. You only need to invest in the chemical costs once. However, methanol is cheap, but you will need to continually buy the amount of chemical you are injecting as it goes downstream and is not recovered.

Summary

This TOTM provided some insight into the various thermodynamic inhibitors’ suppression performance and other considerations. In Part 2 we will review a Case Study in the application and selection of inhibitor for a gathering system inhibition application. In addition, the pro’s and con’s of oxygenates versus glycols will be discussed in more detail.

By: Kindra Snow-McGregor, P.E. & Mahmood Moshfeghian, Ph.D.

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References:

Nielsen, R.B., and R.W. Bucklin, “Why not Use Methanol for Hydrate Control?”, Hyd. Proc., Vol. 62, No. 4 (Apr. 1983), p. 71.
Hammerschmidt, E,G., preprint of a paper given at 1939 Tulsa meeting of the Natural Gas Section of the American Gas Association.
Campbell, J.M., “Gas Conditioning and Processing, Volume 2: The Equipment Modules,” 9^thEdition, 3^rdPrinting, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018.
Moshfeghian, M.,http://www.jmcampbell.com/tip-of-the-month/2020/03/impact-of-process-gas-pressure-on-the-performance-of-a-mechanical-refrigeration-plant/, PetroSkills -John M. Campbell Tip of the Month, March 2020.
Moshfeghian, M.,http://www.jmcampbell.com/tip-of-the-month/2019/06/impact-of-temperature-approach-of-the-heat-exchangers-on-the-capex-and-opex-of-a-mechanical-refrigeration-plant-with-meg-injection/, PetroSkills -John M. Campbell Tip of the Month, April 2019
GCAP 9.3.2, Gas Conditioning and Processing, PetroSkills/Campbell, Tulsa, Oklahoma, 2020.
Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9^thEdition, 3^rdPrinting, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018

Written on November 1, 2020 at 10:58 am, by admin

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How to Reduce Compressor Power in a Mechanical Refrigeration Plant

Compression refrigeration is the most common mechanical refrigeration process. It has a wide range of applications in the gas processing industry, including [1]:

►Chilling natural gas for hydrocarbon dewpoint control

►Chilling natural gas for NGL extraction

►Condensation of reflux in deethanizers

►NGL product storage and transportation

►Natural gas liquefaction (LNG)

Continuing the March 2020 [2] and June 2019 [3] Tips of The Month (TOTM), this tip reviews several options for reducing the compression power requirement for a mechanical refrigeration plant with mono ethylene glycol (MEG) injection for hydrocarbon dewpoint (HCDP) control. Specifically, for each option the calculated required compression power is compared with the base case and percentage reduction reported. All of calculations are performed using GCAP software [4].

Figure 1 presents the process flow diagram for a simple propane refrigeration unit. The mixture of refrigerant vapor and liquid enters the chiller/evaporator typically 3-6°C (5-10°F) lower than the temperature to which the process stream is to be cooled. The refrigerant leaves the condenser as a saturated liquid or slightly subcooled. For air cooling, the condenser temperature will usually be 14-16°C (25-30°F) above the air dry bulb temperature. For water cooling the condensing temperature will be 5-10°C (9-18°F) above the water temperature. Refer to reference [1] for more detail.

The listed process conditions in Figure 1 are for the base case study. The specified values are shown in red color and the calculated values are shown in black color.

Figure 1. Example of process flow diagram for a simple propane refrigeration system

Figure 2 presents the process flow diagrams for a typical HCDP control plant using mechanical refrigeration with MEG injection system. The details of a mechanical refrigeration are given in Chapters 6 and 15 of the Gas Conditioning and Processing, Volumes 1 and 2 [1, 5].

Figure 2. Typical mechanical refrigeration plant with MEG Injection system for rich wet inlet gas [5]

Case Study:

Let’s consider a rich wet inlet gas (Stream 1 in Fig. 2) at 7500 kPa (1088 psia) and 12 °C (54 °F) with known compositions and a flow rate of 8.08×10⁶ std m³/d (284.9 MMscfd). The objective is to meet a hydrocarbon dewpoint specification of (Streams 3 and 4 in Fig. 2) -15 °C (5 °F) at 7500 kPa (1088 psia) for the sales gas by removing heat in the “Gas/Gas” heat exchanger (HX) with a hot end temperature approach (TA) of 5°C (9°F) and in a propane chiller with TA of 5°C (9°F). The heat will be rejected to the environment by a propane condenser (Air Cooler) with TA of 16.7°C (30°F) at 40°C (104°F). Pure propane is used as the working fluid in the simulation. The pressure drops in the “Gas/Gas” HX and on the process side of the propane chiller are assumed to be negligible. The tip will investigate the impact of a few options to reduce the compression power for the mechanical refrigeration presented in Figure 1.

Option 1: Minimizing Chiller Duty:

Chiller duty can be lowered by installation of a Gas/Gas HX (Fig. 2) and/or gas-liquid HX upstream of chiller. Table 1 presents the Gas/Gas, chiller, and total heat duties. Note that by lowering the hot end TA between streams 1 and 5 in Gas/Gas HX (Fig. 2) from 5°C (9°F) to 3°C (5.4°F), the chiller duty reduces from 2.4 to 1.85 MW (8.1 to 6.3 MMBtu/hr). Lower Gas/Gas TA results in bigger Gas/Gas HX and more CAPEX but smaller chiller.

Table 1. Impact of Gas/Gas HX TA on chiller duty

Table 2 indicates that by lowering Gas/Gas TA from 5°C (9°F) to 3°C (5.4°F), the compression power is reduced to 0.78 of the base case of TA = 5°C (9°F). For greenfield projects, even though Gas/Gas CAPEX goes up, the chiller and condenser CAPEX go down. For compressor both CAPEX and OPEX go down. For the existing units the saving will be only on compressor OPEX.

An optimization study can be completed costing out the various size options of the gas/gas heat exchanger temperature approach, and the propane chiller / refrigeration equipment capital cost and operating costs to determine the most cost-effective installation. For an example see June 2019 [3] Tips of The Month.

In fact, for all of the options discussed in this tip, completing the optimization studies along with equipment pricing and operating cost estimates can be used to determine the best process configuration for the various options.

Table 2. Impact of Gas/Gas HX TA on compression power and condenser duty

Option 2: Chiller TA Between Process Gas and Propane:

Propane vaporizing temperature and pressure are a function of process gas temperature and chiller TA. Lower TA results in higher chiller temperature and pressure. Therefore, suction pressure goes up and reduces the compressor power. Table 3 indicates that by lowering process gas-propane TA from 5°C (9°F) to 3°C (5.4°F), the compression power reduced to 0.71 of the base case of TA = 5°C (9°F). For greenfield projects, chiller and condenser CAPEX goes down and compressor CAPEX and OPEX go down. For existing units, the saving will be only on compressor OPEX.

Table 3. Impact of process gas-propane (chiller) TA on compression power and condenser duty

Option 3: Condenser TA Between Air and Propane:

Propane condensing temperature and pressure are a function cooling media (air) temperature and condenser TA. Lower TA results in lower propane condensing temperature and pressure. Therefore, discharge pressure goes down and reduces the compressor power. Table 4 indicates that by lowering process gas-propane TA from 16.7 to 13.9°C (30 to 25°F), the compression power reduced to 0.72 of the base case of TA = 5°C (9°F). For greenfield projects, the condenser CAPEX goes down and compressor CAPEX and OPEX go down. For existing units, the saving will be only on compressor OPEX.

Table 4. Impact of air-propane (condenser) TA on compression power and condenser duty

Option 4: A Flash Tank with Two Stages of Compression Economizer [1]:

Figure 3 presents a refrigeration system employing one flash tank economizer. In this system, the saturated liquid refrigerant leaving the accumulator is expanded across a valve to an intermediate pressure where vapor and liquid are separated. This separator liquid is expanded across the second valve to chiller pressure while the separator vapor goes to the compressor inter-stage suction. The refrigerant entering the chiller now has a higher liquid content. This reduces the refrigerant circulation rate through the chiller as well as the first stage compressor power requirement. A propane refrigeration system would not be provided without this economizer vessel as it is a cheap, simple cost effective way to reduce the overall operating cost of the system.

The pressure at point E is typically set at a value that minimizes the total compression power and is compatible with the compressor design. Total power is minimized when the first stage and second stage power are equal, but this leads to a first stage compression ratio that is higher than the second stage which may not be compatible with the compressor design, particularly when a centrifugal compressor is used. For purposes of preliminary calculations, assuming an equal compression ratio between compression stages is adequate.

Figure 3 also presents the simulation results for an interstage pressure of about 595 kPa [86 psia] which yield equal compression ratio.

Figure 3. An example of a flash tank with two stages of compression economizer

Option 5: Energy Integration – Utilizing an Available Cold Process Stream

The residue gas (Stream 5) of Figure 2 at T = 7 °C (45 °F), P = 7500 kPa (1088 psia), and a condenser TA= 5°C (9 °F) can be utilized as cooling medium for condensing propane at 12 °C (54 °F) prior compression to pipeline inlet pressure. Figure 3 presents the process flow diagram utilizing Stream 5 cold temperature of Figure 2 for condensing propane in a simple refrigeration system.

Tables 5a (SI units) and 5b (FPS units) present the simulation results for three cases; (1) a simple refrigeration system as the base case, (2) a flash tank economizer of option 4, and (3) the simple refrigeration plants shown in Figure 4. Table 5 indicates that the simple refrigeration utilizing the cold process stream for condensing propane requires the lowest power followed by the flash tank economizer. Compared to the base case, the power reduction ratio is 0.47 and 0.85 for the simple system with cold stream for condensing propane and the flash tank economizer, respectively.

Figure 4. Utilizing Stream 5 cold temperature for condensing propane in a simple refrigeration system

Table 5a. Comparison of flash tank economizer and cold gas condenser with the base case

Table 5b. Comparison of flash tank economizer and cold gas condenser with the base case

Summary:

This tip demonstrated the impact of several key parameters on the compression power of a mechanical refrigeration system (for a HCDP control plant) and the following observations were made.

1. Option 1: The cold separator temperature (chilling temperature) and the chiller duty are the key design variable and have a significant impact on the compression power of mechanical refrigeration.

• Chiller duty may be reduced by utilizing a Gas/Gas HX and/or gas-liquid HX with optimum temperature approach (TA) upstream of chiller (Figure 2 and Tables 1 & 2).

• Lower TA increases the Gas/Gas HX duty but lowers the chiller duty (Table 1).

2. Option 2: The lower chiller TA between gas-propane, the higher chiller temperature and pressure which reduces compression power. (Table 3).

3. Option 3: The lower condenser TA between cooling media-propane, the lower condenser temperature and pressure which reduces compression power. (Table 4).

4. Option 4: Use of a flash tank economizer reduces compression power (Figure 3, Table 5)

5. Option 5: Using an available cold residue gas as cooling medium for condensing propane in a simple refrigeration system reduces compression power considerably (Fig. 4, Table 5).

Written By: Mahmood Moshfeghian, Ph.D.

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References:

1. Campbell, J.M., “Gas Conditioning and Processing, Volume 2: The Equipment Modules,” 9^th Edition, 3^rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018.

2. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2020/03/impact-of-process-gas-pressure-on-the-performance-of-a-mechanical-refrigeration-plant/ , PetroSkills -John M. Campbell Tip of the Month, March 2020.

3. Moshfeghian, M.,http://www.jmcampbell.com/tip-of-the-month/2019/06/impact-of-temperature-approach-of-the-heat-exchangers-on-the-capex-and-opex-of-a-mechanical-refrigeration-plant-with-meg-injection/, PetroSkills -John M. Campbell Tip of the Month, June 2019

4. GCAP 9.3.2, Gas Conditioning and Processing, PetroSkills/Campbell, Tulsa, Oklahoma, 2020.

5. Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9^th Edition, 3^rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018

Written on October 1, 2020 at 10:26 am, by admin

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The APCI C3/MR Process – Part 2: An In-plant Review of the APCI C3/MR Propane Pre-Cooling Process

In Part 1 of this Tip of the Month Series, a basic review of the APCI C3/MR process was provided. In addition, a brief introduction into the global LNG trade was discussed. Part 1 cited many Liquefaction Technologies available for LNG processing, showing that the most commonly applied LNG liquefaction technology was the APCI C3/SMR(MR) process which is present in over 40 % of the worlds’ 33 existing LNG Plants during 2018 – 2019.

As stated in Part 1, the existing APCI C3/MR liquefaction process has been proven to be very energy efficient, thermodynamically stable, operator friendly, as well as operationally reliable, and applies well known refrigerant components for the process. The current Part 2 of the Tip seeks to provide an insight into the operation of the propane pre-cooling cycle of an APCI Single Mixed Refrigerant (SMR/MR) facility processing 680 MMscfd (19.21 x 10⁶ std m³/d) of inlet gas. This is roughly equivalent to 14,000 tpd, or 5.1 mtpa of LNG production.

For this case study, the Inlet Process Gas has a composition very similar to the previous compositions and shown in Part 1 of the Tip.

An Analysis of the APCI C3/MR Propane Pre-cooling cycle

Table 1 [1] provides the assumed inlet gas composition and operating conditions for the assumed feed gas to the facility. Free liquids are separated from the inlet gas, in the inlet separator, and the gas is filtered to remove any particulates. From there, the gas flows to the acid gas removal unit to remove the acid gas components (H₂S / CO₂) to an acceptable level. For simplicity, the gas composition provided does not show the ppm values of H₂S and CO₂ that were present that required treating.

Table 1. Inlet Gas conditions and compositions [1]

The feed gas temperature increases when it flows through the amine contactor in the Acid Gas Removal Unit. The reaction between the amine and the acid gas components is exothermic. The process gas conditions entering the first Chiller are shown in Table 2. At this point, roughly 0.28 x 10⁶ std m³/d (10 MMscfd) of fuel gas has been taken off from the inlet stream, in addition, the gas pressure has also been adjusted to be at the optimum liquefication train operating pressure. The locations of the stream conditions are summarized on the process flow diagram shown in Figure 1.

Table 2. Treated Gas to First Propane Chiller

Figure 1. Process Flow Diagram – Inlet Gas Treating Conditions [2]

This first propane Chiller operates at 660 kPa (96 psia) and 11°C (52°F). The gas is cooled to 18°C (64°F), with a cold end temperature approach on the chiller of roughly 7°C (12°F). There is a small pressure drop of the process gas through the gas chiller, resulting in an outlet process gas pressure of 7300 kPa (1059 psia).

The gas then flows through the molecular sieve dehydration unit, and then on to the mercury removal unit (MRU). The assumed mercury content of the gas is 110 μgm/Nm³, which corresponds to roughly 0.01 ppm, or virtually nil in terms of mole %. For the assumed inlet flow rates, and mercury concentration, this translates into 0.16 kg Hg/d (0.35 lbm Hg/d). Often times, a mercury removal unit bed life could be on the order of 14 – 15 years. It really depends on how much mercury is present in the feed gas, and how the original guard bed was designed.

The stream conditions are now defined in Figure 2.

Figure 2. Process Flow Diagram – Process Gas Conditions after the first Propane Chiller [2]

Now let’s take a deep dive into the Propane Refrigeration – Pre-Cooling portion of the plant.

Propane Refrigerant Operating Conditions:

It is assumed that propane is available locally for delivery to the plant as 99-99.5 pure C₃. Within the Process Gas cooling cycles, effective application of propane thermodynamic performance is essential. A selection five (5) operating pressure stages are established for the propane processing cycle for both the MR, and Process Gas. The final Propane compression stage is designated to be readily condensable via aerial cooling units, and not to be considered a “cooling level”. The four (4) compression suction stages are also designated for cooling of the process gas, (PG) as well as the Mixed Refrigerant (MR).

These cooling operation pressures and temperatures will be referred to as LEVELS so as to differentiate compression “stages”, vs. cooling “levels”. Optimization of Propane compressor horsepower often is the driver that sets the facilities operating pressures and temperature for the four (4) chilling stages. Total propane mass flow required can be estimated once both process gas, and MR mass flows and cooling heat loads have been established. With reference to Figure 3 [2], that depicts the typical APCI C3MR Plant under consideration, Table 3 [2] summarizes the existing Propane pressure, and temperature conditions applicable to the Plant Operations for compression, and chilling of the Process Gas, as well as the MR pre-cooling and partial condensation:

Table 3. Propane Refrigeration Conditions for a Typical APCI C3MR Plant [2]

Figure 3. Propane Pre-cooling Cycle highlighted – APCI C3 MR PFD [2]

Figure 4. Simplified Process Gas Propane Refrigeration Pre-Cooling Schematic [2]

Notice in Figure 4, the PG temperatures leaving each chiller are noted in light blue text. Understanding what the PG gas temperature and pressure is leaving the chillers allows one to estimate the duty of each of the PG propane chiller services. Also note, that at each chilling stage there will be a resulting “unused” liquid propane that is required to provide cooling duty to the subsequent lower pressure and temperature chilling levels. In effect, each of the propane chillers is also functioning as an economizer, which reduces the overall required propane refrigeration compressor horsepower. Thus, the 3^d stage propane chilling at 197 kPaa (28.5 psia), and -25 ºC (5 ºF) denoted by the point of reference “2” is metered and directed to the MR chiller, as well as to the PG Chiller. As seen the propane vapor of this pressure level is directed to the compressor suction scrubber denoted as “5”. Finally, the 4^th stage propane chilling at 103 kPaa (15 psia), and – 40 ºC (- 40 ºF) are indicated by point “3” and the resulting vapors are directed to the Compressor Suction scrubber indicated by “6”. All the remaining liquid propane at the 3^d stage pressure/temperature chilling level is converted to a vapor phase at the 4^th stage chilling with no further liquid propane required.

An important point to consider is that the required propane liquid mass flow rate required to provide the process gas, and mixed refrigerant heat loads for all four (4) chilling levels will not be the final TOTAL propane mass to be handled by the four (4) stages of compressor operation due to the amount of propane vapor that flashes off at each level of refrigeration.

With understanding the basic thermodynamic principles of the propane refrigeration, one can set up the APCI C3MR propane pre-cooling process on a propane P-H diagram as shown below in Figure 5.

Figure 5. PH Diagram for Propane (SI and FPS units shown); kJ/kg (0.433 BTU/lbm) [3]

The detailed hand calculations to estimate the required propane flow rate for the PG service is out of the scope of this TOTM. However, a summary of the total cooling demand required for the pre-cooling propane service is provided in Table 4.

Table 4. Propane Pre-Cooling Chiller Summary [4]

Now that the propane pre-cooling section is summarized, we can take a look at the cooling demand that is provided by the MR.

Main Cryogenic Heat Exchanger Energy Balance for the PG:

The following conditions are present for the PG entering the MCHE:

► Q_g (Gas Rate) = 670 MMscfd (~19 x 10⁶ std m³/d)

► Q_gm (Gas Mass Rate) = 586 000 kg/hr (1.29 x 10⁶ lbm/hr)

► Inlet Pressure = (7000 kPa; 1015 psia)

► Inlet Temperature -38 ºC (-36.4 ºF)

► Discharge Pressure = (5300 kPa; 769 psia)

Figure 6 [4] presents the results of the first calculations at the above stated conditions for the PG indicating an energy load of – 355.179 x 10⁶ kJ/h (98.6 MW). This value corresponds to the heat to be removed from the PG resulting in the existing LNG at tower top conditions before the Isenthalpic expansion to atmospheric pressure and approximately -161 ºC (259 ºF). This heat must be removed by the two-phase MR entering as a cold liquid phase from the inlet at the bottom of the MCHE and progressing as a cold liquid through the “Warm Bundle” in the lower part of the exchanger. The liquid MR undergoes at JT expansion to pre-cool the incoming PG and upward flowing vapor phase MR. The cold vapor phase MR proceeds vertically through the exchanger to the “Cold Bundle” where it also undergoes a JT expansion resulting in temperature well below -170 ºC (-274 ºF). The MR exits the exchanger at low pressures and at temperatures some 5 – 8 ºC (9 -14 º) cooler than the incoming Process Gas.

Figure 6 [4]: Determination of the Process Gas (PG) Heat Load in the MCHE

The second calculation to be carried out would be to estimate the required MR rate that will yield the required heat exchange capacity to absorb the required MW (MMBtu/hr) of energy from the PG. Completing the MR refrigeration flow rate calculations is complex and requires the use of a process simulator. These calculations are out of the scope of this tip.

However, it is interesting to note that the propane pre-cooling cycle provides roughly 34% of the total cooling duty to the PG (propane cycle + MCHE = 33.4 + 98.6 = 132 MW).

It should also be noted that the heat sink to the propane refrigeration is atmosphere, and the heat sink for the MR is the propane refrigeration unit. For a plant running at these conditions, the estimated total duty of the propane cycle to provide the MR cooling and condensation is roughly 117.6 MW (401.4 MMBtu/hr). The propane is essentially removing the heat that the MR is absorbed in the MCHE from the PG, thus it makes sense that the propane duty required by the MR is significantly greater than that required to pre-cool the incoming natural gas.

Summary of Part 2 of the Tip of The Month

Part 2 has attempted to focus on a portion of the principal operational conditions and their functions as related to the PG, and MR in a typical APCI C3MR LNG facility. In this review, the following crucial items have been addressed for this purpose:

A typical APCI C3MR plant was chosen with corresponding inlet conditions considered for a Gas at 680 MMscfd (19.28 x 10⁶ std m³/d) to be converted to 670 MMscfd (18.93 x 10⁶ std m³/d) after initial Plant processing and preparation for liquefaction in the MCHE. The typical operational array of pressures and temperatures were discussed and reviewed.

A simplified analysis related to the Plant’s propane chilling configuration for a typical four (4) compression stage, and four (4) chilling (stage) level operation with corresponding pressures and temperatures was shown and discussed.

The thermodynamic criteria was discussed for determination of the Process Gas heat loads at each Chilling Level was presented on the propane Pressure-Enthalpy diagram. The temperatures, and pressures for the PG, were established in preparation for entry to the MCHE.

An energy heat load was computed for the PG with given inlet and outlet temperature and pressure at entering and MCHE tower top conditions. The heat required to be removed from the Process Gas to produce LNG was established.

To learn more about LNG, we suggest attending our G2 (Overview of Gas Processing), G29 LNG (Short Course: Technology and the LNG Chain) and G4 LNG (Gas Conditioning and Processing-LNG Emphasis) courses.

Written By: Dr. Frank E. Ashford & Kindra Snow-McGregor, P.E.

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References

1. Typical PROCESS GAS, and SMR MEDIUM Compositions for APCI Process: Private Communication

2. Typical APCI C3MR LNG Facility LNG: Private Communication

3. Gas Conditioning and Processing: Vol 2 (Edition 9.3) “The Equipment Modules”

4. PetroSkills/John M. Campbell: GCAP 9.

Written on September 2, 2020 at 3:09 pm, by admin

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Overview of Gas Conditioning and Processing: Why Is This Required, and What Drives the Unit Op Selection of a Facility?

Raw natural gas that is produced from the reservoir always requires some type of conditioning and processing before we can transport that gas to the customers. Why is gas conditioning and processing necessary? Because our sales gas, and liquid hydrocarbon products that we produce must meet the sales / product specifications. These specifications are set contractually by: 1) the buyer, 2) the transporter, or 3) by subsequent processing requirements. These product specifications are largely driven around safety in the transportation, end use of the various commodity streams, or the process operating conditions of downstream operations. Figure 1 provides a simplified schematic of the Natural Gas Value Chain.

Figure 1. Gas Gathering and Processing in the Total Production System[1]

As one can see from Fig. 1, natural gas can be produced from many different types of reservoirs. Unfortunately, there is no “typical” natural gas composition, the composition of the gas that comes out of the ground is set by the reservoir compositions and types, which can vary widely. The first constraint in figuring out the unit operations that are required in the gas conditioning and processing module, is understanding the inlet gas composition. So, let’s take a quick review of the various reservoirs that are shown in Figure 1 and discuss the possible hydrocarbon compositions and some common terminology used in the industry.

Let’s start with the oil well, shown in the figure on the far left. Notice, when the produced oil flows to production separator 1 in the figure, the oil flashes off some amount of solution gas. The solution gas flowing from the production separator was in the oil phase at reservoir conditions. As a result, this gas is very rich in heavy hydrocarbons. One of the key parameters in determining the gas conditioning and processing requirements is how much heavy hydrocarbons are present in the gas. Often times, the solution gas is processed to recover the additional heavy hydrocarbons from this gas stream as condensate and natural gas liquids (NGLs). The crude oil from the production separators would be processed and transferred to the end buyers. Notice the solution gas is shown going through some field treating (typically TEG Dehydration) and Compression (optional depending upon reservoir pressures and gathering system layout) before it flows to the Gas Processing Module. Notice, from the gas processing module, there are two streams that are shown going back to the oil well, CO₂ for EOR (Enhanced Oil Recovery) and Gas Lift and / or Injection. CO₂ EOR, is outside of the scope of the TOTM.

Commonly, solution gas that is produced with crude oil is used for gas lift. Gas is injected into the several gas-lift valves positioned along the well depth which are used during start-up. Once the well is started-up all of the valves are closed, except for the valve at the bottom remains open. The flow of gas is used to decrease the density of the crude in the well to allow for higher crude production rates at the same differential pressure. Or the gas may be reinjected into the reservoir to prolong and maintain the gas cap pressure which provides the drive mechanism for crude oil to flow up the well bores. This will also increase recoveries and the life of the field.

Let’s move to the next gas well with Production Separator 2, which is shown producing natural gas from the gas cap of the oil reservoir (often referred to as associated gas), AND gas from a conventional non-associated gas reservoir. Notice, in both cases, the gas that is being produced is in the vapor phase at reservoir conditions. This implies that these gases have a lighter composition than that of the solution gas. In terms of the conventional non-associated gas reservoir, non-associated gas only refers to the fact that the gas is “not associated” with any oil at reservoir conditions. Conventional refers to traditional reservoirs (hydrocarbons and water) that has sufficient porosity to hold economic quantities of hydrocarbons, and has permeability to allow the production of these fluids. The reservoir is sealed by non-permeable rock, which essentially traps the fluids in the reservoir. For conventional non-associated gas reservoirs, the composition can range from lean to rich really depending upon the reservoir type, for example is it a wet-gas (rich, or wet with heavy hydrocarbons) or dry-gas (lean – very little if any heavy hydrocarbons). Sometimes wet / dry gas can be confused with water content. It should be noted that all production from the reservoirs will be saturated with water, which is the most common contaminant in natural gas that must be removed.

Notice, from Production Separator 2, that field condensate is being separated from the produced natural gas. When rich gas is produced from a reservoir, the heavy hydrocarbons (C₅+) will be in the liquid phase at production separators temperatures and pressures. This field condensate is often blended into the crude oil provided the crude oil vapor pressure specification is not exceeded (as shown in Figure 1). This is done because crude oil had a higher commodity value than field condensate, thus maximizing revenue from the field production.

Let’s move now to Production Separator 3. This production separator is producing from an “unconventional gas reservoir”. A few examples would be a shale gas reservoir, a tight gas reservoir, or even coal bed methane (CBM). Shale gas is often referred to as “resource plays” because the gas is trapped in the rock (like air in pockets trapped in bread). The reservoir has very low permeability, thus fracking is required to produce these wells. Shale gas can range from very lean and dry to very rich in heavy hydrocarbons, with the composition changing within a single shale play depending upon the depth and zone of the production. Table 1 provides some example compositions from North American Shale Plays that highlights the variability in composition.

Table 1. Example Shale Gas Play Compositions [2]

CBM is similar to shale gas, except the gas is absorbed in the matrix of coal. The gas composition is primarily methane, with possibly some CO₂ / N₂ and water. As a result of CBM being primarily methane, this produced gas requires much less processing than the others as there are no potential hydrocarbon liquids to recover to sell.

So, as was discussed here, the inlet gas composition to a gas processing facility can vary significantly. The processing steps required for a given inlet composition to meet the sales gas and possibly liquid product specifications is dependent upon that inlet gas composition.

Essentially, STEP 1 – Know your fluids

Now that the inlet composition is understood, the next part of determining what processing will be required is determining what products the facility is going to produce. This decision sets the ultimate design of the plant. There are a number of factors that play here:

1. Probable forecast of reservoir performance / or size of shale play and drilling / production plans

2. Local commodity markets and access to those markets

3. Local infrastructure available

4. Capital and Operating Costs of the various processing options

If we look back at Figure 1, let’s focus on the products that a gas processing facility makes and their uses.

Sales Gas (often referred to as Residue Gas) – natural gas that we use in our homes

LNG – Liquified Natural Gas – means of transporting large volumes of gas to distant markets

GTL – Gas to Liquids – process that converts natural gas into lube oils, diesel, etc.

NGLs – Natural Gas Liquids – these can be sold as pure commodity products

Ethane – Petrochemical feedstock for manufacture of ethylene

Propane* – Petrochemical feedstock for manufacture of propylene / residential, commercial and transportation fuel

Iso-Butane – Refinery feedstock to alkylation unit / fuel use as LPG

Normal Butane – Gasoline (petrol) blending, petrochemical feedstock for manufacturing light olefins, fuel use as LPG, and can be isomerized to i-C4

Natural Gasoline** – Refinery feedstock to reformer or isomerization unit, petrochemical feedstock for manufacture of light olefins

* Often sold as liquified petroleum gas (LPG) which can be C3, C3-C4 mixture or predominantly C4

** Natural gasoline is a North American term, also referred to as raw gasoline, light naphtha or condensate in other regions

It should be noted, that heavier hydrocarbons have a higher market value as compared to natural gas. When a facility is designed to do NGL extraction, the processor is taking advantage of this price differential between the NGL products (ethane + heavier hydrocarbons) and its value as a natural gas constituent. The value of NGL product as a natural gas constituent is termed the “shrinkage value”. This represents the value foregone by extraction of the NGL from the gas. So, for example, if one were to take all of the propane out of a natural gas stream that would have been sold in the sales gas stream, the sales gas stream “shrinks” because those propane molecules are removed and being sold as a liquid product stream, that typically, has higher value than the sales gas stream. Shrinkage is the highest operating cost in terms of an NGL extraction facility and must be taken into account. I will write more about this in an forthcoming tip of the month.

STEP 2 – Define the products you want to make

Now that the products of a facility are defined, the appropriate gas conditioning and processing unit operations can be selected. Let’s take a quick look at what contaminants may be found in produced natural gas and their associated problems. Note, these contaminants will need to be removed to meet the appropriate sales gas and liquid product specifications.

Typical well-stream impurities along with their associated problems are listed in Table 2.

Table 2. Example Impurities and their Associated Problems [2]

STEP 3 – Identify the contaminants that need to be removed, and to what level

The determination of what contaminants need to be removed and at what level depends upon multiple factors. Each product commodity has its own product specifications, which as mentioned previously is either set by 1) the buyer, 2) the transporter, or 3) by subsequent processing requirements or regulations. For example, if we are in a situation where there is an ethane market is available, the gas composition fits, etc.., then we will need to go into a deep NGL extraction unit which requires cryogenic temperatures. This would be an example of subsequent processing requirements where the process operating temperature requires further contaminant removal (bone dry water content), than say a pipeline sales gas specification. For cryogenic processing, there are also tighter specifications on CO₂ concentrations as well, as CO₂ can freeze in the cold box (think dry ice).

With that said, let’s take a quick look at some typical pipeline sales gas specifications from around the world, as shown in Table 3.

Table 3. Example Sales-Gas / Transportation Specifications [2]

Liquid hydrocarbon products must be dry of water, and either have a vapor pressure or composition specification, along with a total sulfur specification, or a copper test.

In terms of our products, they must be safe to use, and that means that all free liquid water, or the possibility of free water condensing out in the system is eliminated. They must be safe to burn without excessive sulfur dioxide emissions (H₂S and total sulfur specifications). In addition, the hydrocarbon dewpoint is important to ensure that as the gas flows through the transmission line and cools to ambient conditions to the customers that no liquid hydrocarbons can condense. This is important as these pipelines and associated end users (your water heater for example) are not designed to deal with (or burn) liquid hydrocarbons. As a result, these specifications are driven by the local ambient conditions. The warmer the climate, the more water or heavy hydrocarbons are acceptable in the sales gas stream as a result.

The gas must also meet the appropriate heating value and Wobbe number so that the customers are assured the gas will burn safely in their gas turbines (for example in power generation plants), or in your natural gas cook top at home.

STEP 4 – Select the appropriate unit operations

Now that the facility inputs (feed composition and flow rates), and outputs (sales products we want to make) the process unit operations can be selected and optimized. Let’s take a look at the options, as shown in Figure 2.

Figure 2. Example Gas Processing Facility Block Flow Diagram [2]

Gas Conditioning Options

Notice the first block in the process is Gas Conditioning. This is where the contaminant removal takes place (CO₂, H₂S, H₂O, Hg).

Sweetening, in all intents and purposes refers to removing all acid gas contaminants (CO₂ and H₂S). The selection of the technology to meet the acid gas specifications depends largely on the concentration of the contaminants to be removed, and the inlet gas flow rate. Technically, carbon dioxide (CO₂), is not a sour component but is often removed in the sweetening unit (amine treating).

Amine treating is the work-horse in the gas processing industry for CO₂ and H₂S removal, and it is always done in the processing facility BEFORE dehydration because the amine solution is generally about 50% by weight water. The gas will be saturated with water after it leaves the acid gas treating unit (AGRU) as a result.

The second step is to dehydrate the gas, i.e., remove the water content to a “safe” level depending upon either downstream processing requirements or end users.

In pipeline applications, this would require TEG dehydration. The water only needs to be removed such that was are sure that it will not condense in the sales gas pipeline.

For cryogenic processing for deep NGL extraction (-101 C, -150 F), or LNG (-160 C, -260 F) this would require molecular sieve dehydration (bone dry gas is achieved but very expensive $$). And the lastly, if we are going into cryogenic processing where aluminum components are utilized, mercury removal is critical. In large base load LNG facilities, or locations with high levels of mercury in the natural gas, there is a mercury removal bed downstream of the mole sieve dehydration units.

Gas Processing – NGL Extraction Options

Once the gas is conditioned, it then flows on to the NGL extraction block. The technology selected for the NGL extraction depends upon the operating objective of the facility and the inlet gas composition. In modern gas processing facilities, the primary NGL extraction technologies utilized are identified under item 3 below the NGL Extraction block in Figure 2:

Condensation:

JT Valve (Valve Expansion)

Mechanical Refrigeration

Turboexpander

NGL Extraction, in a nutshell, is very simple. All we are doing is cooling the gas down to get the heavy hydrocarbons to condense out of the natural gas to meet one of two criteria:

1. A hydrocarbon dewpoint specification for a sales gas stream (minimal recoveries)

2. Deep NGL extraction to upgrade the value of the products the plant can sell

The different technologies listed under condensation simply cool the conditioned feed gas stream down to a desired temperature to meet the required hydrocarbon recovery that we want to achieve.

Typical process selections for hydrocarbon dewpoint facilities would be a JT plant or mechanical refrigeration, as best case cold temperature achieved would be -40 C [-40 F].

For deep NGL extraction facilities (C₂+ NGLs), a turboexpander plant is used. These plants can achieve temperatures to roughly -101 C [-150 F]. The Gas Sub-Cooled Process is the work-horse in North America in terms of our process operations, recommend that you get online and see what that process operation is about.

Stabilization / Fractionation / Product Treating

The natural gas liquids for the gas processing plant will need to be stabilized (all of the light hydrocarbons removed), so that they are safe to store and transport. These specifications can vary significantly depending upon end markets and transportation methods.

In some cases, this could be a stabilized condensate consisting of only C₅+ having a vapor pressure below atmospheric conditions (example of a hydrocarbon dewpoint plant). In deep NGL extraction plants, often time the composition (C₂/C₃ ratio) is specified. This is important as the fractionation facilities operations will only be equipped to process an NGL stream of a composition within a given range.

The fractionation facilities consist of distillation towers that can separate the hydrocarbons essentially based upon their boiling point differences, with the light hydrocarbons going out the top of the column as the distillate products, and the heavy hydrocarbons, going out the bottom as the bottoms product.

Product Treating – Why?

Depending upon the inlet gas composition, it is common for there to be CO₂ or H₂S in the inlet gas that, after the gas goes through processing will not exceed the sales specification. However, these contaminants will preferentially distribute to the hydrocarbon liquid phase and can cause the ethane, propane, etc.. to not meet sulfur, CO₂ specifications. It is more economical to allow the contaminants to build up in the product streams and treat the liquid product stream, rather than removing the contaminant from the inlet gas stream. The amount of liquids needing to be treated is significantly smaller (smaller equipment, lower capital cost, and lower operating cost).

There are a number of technologies than can be applied for liquid treating, and it really depends on a number of factors which is the best solution for the facility which is outside of the scope of this tip.

STEP 5 – Putting it all together – In Conclusion

I really hope that you found this tip useful. We reviewed at a very high level the issues associated with gas processing and conditioning, and the primary processing blocks that are required to go from the reservoir to market depending upon the gas composition.

So – why does my facility look the way it does? Well, it depended upon the initial forecasted feed gas composition, reservoir reserves, and production rates, combined with what the products were that they owner / operator decided was the most economic to sell based upon the local market access and infrastructure available.

To learn more about similar cases and how to minimize operational problems, we suggest attending our G2 (Overview of Gas Conditioning and Processing), G4 (Gas Conditioning and Processing), G5 (Practical Computer Simulation Applications in Gas Processing), and G6 (Gas Treating and Sulfur Recovery) courses.

Written By: Kindra Snow-McGregor, PE

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REFERENCES

1. Cannon, R., “The Gas Processing Industry: Origin and Evolution”, Gas Processors Association, 1993.

2. Gas Conditioning and Processing, Volume 1: The Basics, 9^th Edition, John M Campbell, Editors: Robert A. Hubbard, Kindra Snow-McGregor, 2017.

Written on August 1, 2020 at 1:53 pm, by admin

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An Introduction into the Air Products and Chemicals, Inc., Mixed Refrigerant LNG Liquefaction Process – What is it, and how does it work?

Part 1 – A Basic Review of the APCI C3/SMR World LNG Positioning and Impact

The birth of modern day LNG production and worldwide commerce commenced with the 1958 conversion of the NORGULF’S Shipping Line’s freighter, NORMANTI into the newly designed world’s first LNG carrier renamed the Methane Pioneer. The vessel was fabricated by the recently formed Constock International Methane (CIM) consortium. The Methane Pioneer was retrofitted for LNG transport, and designed to carry 5 000 m³ (32,000 Bbls) of LNG, equivalent to some 2,375 tonnes of LNG, corresponding to 3.23 x 10⁶ std m³/114 MMSCF of liquefied process gas. Her maiden voyage was from Constock’s LNG production facility on the Calcasieu River in Louisiana on 25 January 1959 to its destination at Canvey Island in England, arriving on 20 February.

Since that time, the worldwide LNG Export/Import evolution has risen from an initial 10-50 million tonnes per annum (mtpa) during the 1970 – 1990 era to a world Energy demand for LNG by 2018 of some 431 billion cubic meters – std. ( 15.2 trillion cubic feet – std.) equivalent to some 319 mtpa.

To date there are many Liquefaction Technologies available for LNG processing, that are part of the Exploration/Transportation/Liquefaction/Worldwide LNG Value Chain. This Tip of the Month will provide some current information regarding one of the most commonly applied LNG liquefaction technologies known as the APCI Propane Pre-cooled (C₃) / Single Mixed Refrigerant (SMR) process which is present in over 40% of the worlds’ 33 existing LNG Plants as of 2018.

The existing APCI C₃/SMR liquefaction process has been proven to be very energy efficient, thermodynamically stable, operator friendly, operationally reliable, and applies well known refrigerant components for the process. The Tip seeks to provide a current day insight into the intrinsic operation of a typical 500 MMSCFD (14.2 x 10⁶ std m³/d) process gas, (10,000 t/d; 3.65 mtpa) LNG APCI liquefaction facility applying a Single Mixed Refrigerant, sometimes denoted as SMR or frequently shortened to MR. In some specific installations worldwide, plant operators prefer the SMR reference.

A review of a typical inlet process gas, along with in plant LNG gas properties will be provided. The Propane Cycle operation in terms of mass, and energy flow for four (4) liquefaction stages of cooling of the Process Gas and MR is also discussed. Thermodynamic conditions for the APCI Main Cryogenic Heat Exchanger (MCHE) LNG liquefaction is presented. The intricate nature of the combined C₃/MR process will be reviewed in detail to yield an insight into the thermodynamic stability, and consequential reliability of the process that justifies its vast process application throughout the world.

Table 1 [1] reviews the typical compositional ranges for an LNG Process Gas expressed as: LIGHT, MEDIUM, and HEAVY. Notice that important parameters such as the GCV (Gross Calorific Value) given in units of kWhr/m³(n). Convenient conversions may be adapted recalling that a kWhr corresponds to 3600 kJ (3429 BTU), and 1 kcal = 4.184 kJ. Additionally, typical LNG densities are given.

Table 1. Typical LNG Compositions, GCV, and Cryogenic LNG Densities. [1]

As shown by Table 1 [1], there is a reasonable range for acceptable LNG Process gas compositions. The APCI C₃/MR cryogenic process is very sensitive to the inlet gas composition. In addition, the integral sections of “pre-processing” or “gas conditioning” processes prior to condensation such as: gas sweetening, pre-cooling, dehydration, and mercury removal are sensitive to the inlet gas composition and the required quantities of contaminants that need to be removed. Notice, the LNG densities that are provided for the various compositions reflect the overall “condensation ratios” in std m³/tonne (scf/tonne). These variations will have an important impact on the performance of the MCHE, selection of the MR composition, and mass flow ratio of the MR to LNG Process Gas. These issues will be addressed in detail in Part 2 of this TOTM, when an actual CASE STUDY will be analyzed.

To begin the discussion, it is convenient to observe the 2018 Natural Gas and LNG worldwide trade movements during 2018 [2]. Figure 1 [2] indicates that some 431 BCM (15.2 tcf), equivalent to 319 mta of LNG taken at an applicable conversion of 1360 std m³/tonne (48,000 scf/tonne) were commercialized. These imports emanated from 20 principal exporters indicated in Figure 2 [3].

This Figure indicates a total “Nominal Worldwide Capacity” of 383 mtpa, of which some 83% were actually exported.

Figure 1. Major Worldwide Natural Gas and LNG Commerce during 2018 [1]

Figure 2. Major Worldwide LNG Exporting Countries during 2018 [3]

Note: Nominal Liquefaction Capacity in mtpa and % Utilization – Total 383 mtpa

Figure 3 [3] shows the percentage application of the world’s major LNG liquefaction processes. Shown are the configurations of the APCI processes, as well as the most commonly applied other technologies. As can be observed, the APCI-C3MR process reflects some 41 % of all the worlds installed processes in 2018, and is clearly the most commonly applied liquefaction technology. Notice that APCI has four (4) processes that are active; however, not as highly utilized as the cited C3/MR. The AP-X process applies a Nitrogen Expander and HEX parallel refrigeration system that provides the final liquefaction, The AP-C3MR/Split MR applies two Turbines, with the refrigeration power equally split between the turbines. One compresses the propane refrigerant and the high pressure stage of the MR. The second is used for MR low, and medium stage compression. As the focus of the Tip is to track the basic APCI-C3MR process, the reader is referred to the LNG literature to find details on the other processes in operation.

Figure 3. Distribution of LNG Liquefaction Processes: 2018, 2019 and projected to 2024 [3]

In terms of Globalization of the LNG Market Figure 4 [3] indicates the major LNG Liquefaction facilities throughout the world. Some USA Units are not shown due to late 2018, or early 2019 incursion into the Export Market. Missing are Cheniere Corpus Christie, Texas, and Sempra Cameron, Hackberry, La. Of these Facilities, Dominion Cove Point, Cheniere Sabine Pass, and Trans-Foreland Kenai are Optimized COP installations, while the remaining two (2) are APCI C3MR.

Figure 4. Worldwide Major Liquefaction Facilities. 2018 [3]

Note: refer to text for missing USA Export Units not shown

An Introduction to the APCI C3/MR Process

As shown by Figure 5, the typical Process Flow Diagram (PFD) for an operating APCI C3/MR LNG Plant consists of various integral components:

1) Inlet Separation: Usually in range of 7-8 MPa; [1000-1160 psig] and to 20-30 °C [68 – 86 ºF].

2) Inlet Process Gas Sweetening: removal of CO₂/H₂S and mercaptans to nil levels.

3) Process Gas Pre-Cooling prior to Dehydration: Usually at pressure and cooled to 15-18 °C [60 to 64 ºF].

4) Process Gas Molecular Sieve Dehydration: water dewpoint temperature to -162 ºC [-260 ºF] at MCHE pressure levels

5) Mercury Removal: LNG Plant levels to 0.01 μg/Nm^3, or 0.01 ppt by volume.

6) Not shown in this APCI C3/MR Plant PFD are facilities that handle condensed C₄/C₅ + components that are directed to In-Plant Fractionation, where heavy liquids are separated and used for MR make-up, possible inlet stream reinjection and/or NGL product sales.

7) The MR is compressed from the MCHE discharge in a low temperature, low pressure superheated phase in range of – 46 to – 50 °C, [-51 to – 58 ºF], and typically 400 – 600 kPaa [ ~ 60 – 90 psia]. Subsequent compression is typically in three (3) stages takes the MR to its final pressure. The MR is cooled by four (4) stages of Propane chilling, and enters the MCHE as two (2) phases usually at 50 – 70 bar, 5 – 7 MPa; [725 – 1000 psig] and -32 to -38 °C [-26 to -36 ºF]. The separated vapor, LMR (light mixed refrigerant) and liquid, HMR, (Heavy Mixed Refrigerant) stream enters the MCHE at the bottom. The process gas is commonly cooled also in four (4) stages prior to entering MCHE inlet separator at temperatures also in range of -32 to – 38 °C [-26 to -36 ºF]. The MR is discharged as a superheated gas generally at low pressure, some 350 – 600 kPaa [51 – 90 psia], and temperature below the entering LNG gas temperature, in range of – 46 to – 50 °C, [-51 to – 58 ºF].

8) The MCHE operates at the cited inlet temperatures, and pressures usually in range of 50 – 70 bar, 5 – 7 MPa [725 – 1000 psig]. The outlet of the MCHE operates at pressures and temperatures in range of 50 – 60 bar, 5 – 6 MPa [725 – 870 psig], and -162.5 to -163 °C [-260.5 to -261.4 ºF].

9) The final ISENTHALPIC expansion of the Liquefied LNG to 101 kPaa (14.7 psia) may actually undergo a slight “HEATING” of some 0.5 °C [~ 1 ºF]. The Joule-Thomson (J-T) coefficient can either be positive or negative. For most gases, the J-T coefficient is positive, i.e. the gas cools upon expansion. However, for most liquids and high-pressure hydrocarbon gases at low temperatures, the J-T coefficient is negative, i.e., the fluids warm upon expansion. For the APCI C3/MR process, the final expansion of the LNG typically occurs high pressure, resulting in the final LNG product at -162 °C and 101 kPaa [-260 °F and 14.7 psia].

Figure 5: Process Flow Diagram – Example PFD APCI LNG Processing Plant

A Brief Insight into typical LNG and MR Compositions and Phase Envelopes

With the intention to provide some insight into the actual appearance of the phase envelopes of the mentioned LNG streams, as well as a typical MR stream [5], Table 2 [1/5], and Figure 6 [5] represent the composition and phase envelope of a LIGHT LNG (taken from Table 1). As can be seen the critical point for pressure/temperature are virtually coincident with the cricondenbar. This means that LNG MCHE liquefaction above 50 bar (5000 kPaa – 725 psia) will not result in an internal two phase LNG stream, but reflect a “dynamic” condensation criteria as the process stream is cryogenically cooled. As also may be observed, final “ISENTHALPIC” expansion to storage conditions theoretically yields a saturated LNG liquid at -162 °C and 101 kPaa [-260 °F and 14.7 psia].

Table 2 [1]. Composition for a Typical LIGHT LNG (from Table 1)

Figure 6. Phase Envelope for a Typical LIGHT LNG (from Table 1) [4]

It is also interesting to observe some selected typical information for an MR applied to an APCI Process. Table 3 [5], and Figure 7 [4] represent the composition and phase envelope of a given MR from an independent source. The selected MR has a small amount of ethylene to yield ample volatility. As can be seen, once again the critical region for pressure/temperature values are virtually coincident with the Cricondenbar. Also clearly shown is the capacity of the MR for required compression, and cooling at the pressure and temperature levels cited in Point 7) above. It is of interest to note that the MR displays a very distinct two – phase system in the pressure/temperature regions mentioned in Point 7) as the two streams are fed to the MCHE. The green line in the Phase Envelope represents 90% Vapor Quality. The MR discharged from the MCHE after absorbing the HEAT LOAD from the LNG process stream is seen to be a superheated vapor at low temperatures and pressures in range of – 46 to – 50 °C, [-51 to – 58 ºF], and 350 – 600 kPaa [51 – 90 psia].

Table 3. Composition for a Typical MEDIUM Molecular Weight MR [5]

Figure 7: Phase Envelopes for a Typical MEDIUM Molecular Weight MR [4]

It is obviously not the intent of the present TOTM to describe all LNG Plant Process Components. The purpose is to review the operation of the MCHE and its integral components and feed streams. These will include the review of the Propane Cycle mass flow, pressure and temperature levels, heat loads, and total compression power required to cool of the Process Gas plus the MR to requirements for entry into the MCHE and subsequent LNG liquefaction and production to the Storage Facility.

The Sequence of Numerical Calculations Required to Quantify the APCI C3/MR Process

Referring to the previous section, a specific Nine (9) items of analysis must be addressed in an interrelated fashion in order to quantify the performance of an APCI C3/MR LNG Process Plant.

Points 7), 8), and 9) are the essential analyses that must be addressed after the inlet feed gas has been pre-conditioned (sweetened, dehydrated, and liberated from mercury). This sequence will be detailed in this section before entering into an actual case study. It will be assumed that the processing of the inlet gas as per Points 1) through 6) has been performed; however, mention will be made so as to provide continuity of the Plant operations. In order to analyze any given APCI installation downstream of all previous plant processes, and before both the LNG Process Gas, and MR enter the MCHE, the following numerical analyses need to be addressed:

A) APCI C3/MR operations related to the LNG sweetened, dehydrated, and mercury free LNG Process Gas:

► It is assumed that all plant feed gas processing and treating upstream of the MCHE have been performed. These all require knowledge of inlet composition, volumetric and mass flows along with pressures, and temperatures pertaining to the upstream systems cited in Points 1) – 6). The inlet gas is sweetened, treated for removal of heavy components, dehydrated, and then treated for mercury removal. It should be noted in some facilities, the mercury removal unit maybe in the front of the plant prior to sweetening, but is a less common process configuration.

► The LNG process gas should be defined in units related to both MMscfd (10⁶ std m³/d) as well as in total mass flow in tonnes/d. The inlet temperature and pressure are defined as well. It should be noted with reference to Figure 5, the first Propane Chiller operating at the highest pressure and temperature cooling level is usually placed upstream of the molecular sieve dehydration unit to minimize the required size and operating costs of the gas dehydration unit. Pre-cooling the incoming wet feed gas condenses out a portion of the saturated water content where it can be removed upstream of the dehydration unit Point 4).

► With knowledge of the final composition of the LNG process gas, an Energy Balance must be performed for the mass flow entering and leaving the MCHE at the inlet/discharge ranges of pressures and temperatures cited. It should be remembered that the LNG process gas presents a superheated phase at inlet conditions, then is condensed into a two-phase region if the LNG phase diagram shows an operating pressure and temperature within this phase region. If the MCHE operating pressure is greater than the process gas Chricondenbar, the LNG will undergo dynamic condensation to a sub-cooled liquid within the MCHE discharge at the top of the Exchanger.

A) APCI C3/MR operations related to the LNG sweetened, dehydrated, and mercury free LNG Process Gas: Cont’d

At this point the LNG is at a high pressure, and cryogenic temperature as cited in Point 8). Recall that isenthalpic expansion of the sub-cooled LNG will probably yield a slight temperature increase, as per Point 8).

When prepared for storage, the cryogenic LNG product will be at near atmospheric pressure, and temperatures very close to -162 C [-260 F]. Figure (8) shows the results for this expansion for the LNG at a typical MCHE conditions. At this resting point the LNG is theoretically a saturated liquid at its atmospheric bubble point. The energy balance is generally expressed in MW (KWx1000). This final cryogenic condensation presents a good opportunity to validate the gas to LNG condensation ratio in scf/tonne (std m³/tonne). The normal ranges for the process is a direct function of the process gas composition (Molecular Weight, and LNG Relative Density); however values usually tend to be in the range of 1,300 – 1,360 std m³/tonne [46,000 – 48,000 approx. scf/tonne].

Figure 8: Isenthalpic Expansion of Light LNG, shown in Figures (6), indicating a slight temperature Increase when expanded for 43 bar to 1 bar at indications shown [4]

► An essential aspect relating to the APCI C3/MR process dynamics is to determine the MR to LNG process gas mass ratio. This procedure is performed after an Energy Balance is performed on the LNG process gas entering, and exiting the MCHE as cited in Point 8). There are various options that may be selected in choosing the components and their compositions in the MR, but a few essential conditions must be met: 1) As stipulated in Point 7) the MR is compressed typically in three (3) stages; 2) The MR enters the MCHE in the two-phase region as Liquid/Vapor and is discharged in the superheated vapor region as cited in Point 7); 3) Once the MR composition has been determined, the total mass flow required will be fixed by the Energy Balance of the LNG, yielding the energy to be extracted from the LNG process stream equal to the energy absorbed by the MR. It has been shown by many applicable studies that this Flowing Mass Ratio: Mass MR/ Mass LNG Process Gas is in the range of 2.2 – 2.6.

The units are usually expressed in tonnes/hr. As an arbitrary example, an LNG gas stream of 10,000 tonnes/d [417 tonnes/hr] will require an MR of some 22,000 – 26,000 tonnes/d [917 – 1084 tonne/hr]. The MR composition is essential to the thermal efficiency of the process, and will be discussed further in Part 2.

In Part 2 of this Tip of The Month, a particular CASE STUDY will be introduced and analyzed. An appropriate LNG composition will be presented, along with a representative gas flow rate in MMscfd [10⁶ std m³ /d]. Typical pressures, temperatures, for the LNG gas and MR will be applied. An applicable MR composition will be assumed, that will incorporate the necessary behavior to satisfy its essential phase behavior points discussed in this Part 1 of the TOTM. With the known LNG gas mass flow rate, and calculated MCHE cooling duty as per guidelines of this Part 1 of the TOTM, the known MR composition will yield its required mass flow rate. To determine the required Propane refrigerant cooling duties for the process LNG, and MR, five (5) propane pressure, and temperature levels will be employed between near atmospheric pressure to final refrigerant discharge pressure. This design will provide for four (4) stages of propane compression and cooling for both the process gas LNG, and MR. Compression levels for the MR Turbo-Compressors will be taken in accord with guidelines presented, allowing MR feed to the MCHE as a two phase stream., and discharge as a low pressure single phase gas. For the process gas LNG, guidelines of Part 1 of this TOTM will be followed for four (4) cooling levels in preparation for entry into the MCHE. The Total cooling duties for the MR and LNG Streams will be determined, along with total required propane refrigerant mass flow, and total stage duties.

To learn more about LNG, we suggest attending our G2 (Overview of Gas Processing), G29 LNG (Short Course : Technology and the LNG Chain), and G4 LNG (Gas Conditioning and Processing-LNG Emphasis) courses.

Written By: Dr. Frank E. Ashford & Kindra Snow-McGregor

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References

1. Accurate determination of LNG quality unloaded in Receiving Terminals: An Innovative Approach: Angel Benito, Enagás, S.A.

2. Statistical Review of World Energy, 2019 BP.

3. IGU Annual Report 2019 (IHS MARKIT)

4. GCAP 9.3.2, “Gas Conditioning and Processing Computer Program,” Editor Moshfeghian, M, PetroSkills, Katy, Texas, 2019.

5. Typical SMR MEDIUM Composition for APCI Process: Private Communication

Written on July 1, 2020 at 1:48 pm, by admin

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H2S and Hydrogen Charging

1 Jun, 2020

Have you ever wondered why hydrogen sulphide (H₂S) is so feared by pressure vessel and piping inspectors? Of course, anyone who works in the oil and gas industry understands the personal danger of H₂S. One whiff can kill you. Table 1 presents the health effects of H₂S concentrations on the human body according to the UK Health & Safety Executive.

Table 1. H₂S concentration health effect on the human body

Without going into the dangers to affected personnel and rescuers, we will now concentrate on the process safety dangers.

In 2006 I was running a Risk Based Inspection (RBI) workshop in Kuwait with 17 participants. In the early afternoon someone’s phone rang, then another, then another. Within ten minutes I was left alone with only two of the participants – because they had switched their phones off. We soon found out that there had been an explosion at an unmanned gathering station in the field. A contractor had a job at the station and the lead hand had been given the key. On reaching the station the lead hand exited the crew cab pickup leaving the other three contractors in the truck. He then unlocked the gate and stepped through, opening the gate. He waved the pickup through to enter the station and as it did, the truck ignited a gas cloud. The three contractors inside the truck were not badly hurt but the lead hand suffered 80% burns and died in hospital three days later. The investigation found a crack near a weld on the main gas line exporting from the station. The cause was hydrogen charging.

But what is hydrogen charging?

As the name suggests hydrogen charging is the process of hydrogen working its way through the steel wall of the pipe of pressure vessel. Let’s discuss where the hydrogen comes from.

Imagine a pipe transporting crude oil. Our crude has some hydrogen sulphide (H₂S) and crude oil always has some water content unless the water has been completely removed in a separation / drying process. Some of the H₂S and the water will naturally react.

(1) The sulphuric (and sulphurous) acid will react with the iron of the steel wall.

(2) At the same time the H₂S also reacts with the iron of the steel wall.

(3) In each case, we have free hydrogen released into the crude oil stream. This is not a huge amount of hydrogen but even if it were we could keep this in the stream as it contributes to the calorific value. However, the Iron Sulphide (FeS) now acts as a catalyst to break the molecular hydrogen (H₂) into atomic hydrogen. This is where things get interesting.

A molecule of hydrogen is made up of two protons each with its own electron and the complete molecule has a diameter of about (depends where you get the information from) 274 picometres (pm).

Figure 1. Body Centered Cubic (BCC) lattice crystal structure of steel [API 571 section 5.1.2.3]

The lattice parameter of 0.3% carbon steel is about 364 pm. As you can see in Figure 1 the molecular hydrogen is just too big to get through the gaps between the iron atoms. But remember the catalytic effect of the FeS on the molecular hydrogen – it causes the hydrogen to crack into two atoms. Notionally the atoms are also too big to get through the gaps but the BCC crystals of steel have many gaps (defects and vacancies) so it is possible for the hydrogen atoms to gradually work their way through the steel but it takes a long, long time for the hydrogen to work its way to the outside. The atomic hydrogen will occasionally lose its electron and becomes an ion of hydrogen – which is effectively a proton. The protons are much smaller than the atoms so are free to move through the steel very easily.

As shown in Figure 2, once the atomic or ionic hydrogen reaches a cavity in the steel lattice, such as a piece of dirt caught during manufacture, the ionic hydrogen can capture a free electron forming atomic hydrogen and the atomic hydrogen can combine to form molecular hydrogen. And the molecular hydrogen is now trapped.

Figure 2. The catalytic effect of the FeS on the molecular hydrogen – production of hydrogen atom and consequences [API 571 section 5.1.2.3]

Over months and years, more and more hydrogen becomes trapped in the space such that the hydrogen acquires a pressure. As the hydrogen pressure builds it results in stresses in the steel that are greater than the yield or even tensile strength of the steel causing a local failure. That failure can take several forms (Figure 3).

1. If the steel is ductile the high stresses will cause a tear or lamination that is usually parallel to the wall surface. The lamination now has a greater surface area so more hydrogen enters and the pressure continues to rise.

► If several laminations occur near to each other the stresses at the lamination tips cause cracks which tend to join lamination ends together – even at different steel depths. This is known as HIC (hydrogen induced cracking) or stepwise cracking.

► If the lamination occurs near the surface of the steel the smaller thickness yields more readily than the thicker portion and we end up with a blister rising from the steel surface.

2. If the steel is not ductile, such as in the Heat Affected Zone (HAZ) of a weld or perhaps a high strength steel or even an older steel that is simply not tough enough, then the high stresses cause cracks instead of plastic deformation.

► Stress Oriented Hydrogen Induced Cracking (SOHIC) is similar to HIC but appears as arrays of cracks stacked on top of each other resulting in a through-thickness crack that is perpendicular to the surface and is driven by high levels of stress (residual or applied). They usually appear in the base metal adjacent to the HAZ where they initiate from HIC damage or other cracks or defects including sulphide stress cracks

► Sulphide Stress Cracking (SSC) can initiate on the surface of steels in areas of high hardness in the weld metal and HAZ. Zones of high hardness can sometimes be found in weld cover passes and attachment welds which are not tempered (softened) by subsequent passes. PWHT (post weld heat treat) is beneficial in reducing the hardness and residual stresses that render steel susceptible to SSC.

Figure 3. Potential failure forms due to H₂S hydrogen charging [API 571 section 5.1.2.3]

So, what can we do to avoid these problems?

► NACE (National Association of Corrosion Engineers) has produced an international standard (MR 0175/ ISO 15156-1) that guides us in appropriate material selection. This standard was first introduced in the 1960’s but the defective pipe I described in Kuwait was installed in the 1950s so the steel was not clean enough and had many initiation locations (aka dirt in the steel) and was harder than we would accept now.

► If the partial pressure of the H₂S is more than 0.3kPa (0.05 psi) then MR 0175 shall apply. This also states that the weld procedure for all welds for sour service SHALL include a hardness test (for detail go to clause 7.3.3.2 of ANSI/NACE MR0175/ISO 15156-2:2009(E).

► Blisters may be carefully vented to halt the build-up of molecular hydrogen so the pressure is relieved. Check out API 579-1/ASME FFS-1 (2007) part 7 for a simple guide on how to qualify a blister for continued service.

► Take regular samples of free water to monitor any changes in process conditions that may cause hydrogen charging.

► During inspections pay special attention to weld seams and nozzles. NACE RP0296 has some recommendations that are very useful.

► Use Wet Fluorescent Magnetic Testing (WFMT), Eddy Current (EC), X-ray or gamma ray (RT) or Alternating Current Field Measurement (ACFM) to check for cracks and laminations. Ultrasonic techniques (UT) can also be used and short wave UT is especially useful for volumetric inspection and crack sizing. Acoustic emission can be used to monitor crack growth.

► All of these standards are extremely useful but always remember that your local jurisdiction takes priority in case of a conflict.

I hope you find this series of posts useful. To learn more about similar cases and how to minimize operational troubles, I suggest attending our ME41 (Piping Systems – Mechanical Design and Specification) and ME43 (Mechanical Specification of Pressure Vessels and Heat Exchangers) courses.

Stay Safe.

By: Ron Frend, Head of Facilities Training

Written on June 1, 2020 at 1:36 pm, by admin

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Piping Vibration: Causes, Limits & Remedies

Piping vibration is a major cause of concern in process plants, particularly in the oil and gas industry where the loss of containment could be catastrophic. This Tip of the Month explains the root causes of piping vibration, natural frequencies and how they may be changed using appropriate structural supports and layouts.

Imagine sitting in your office having a well-earned cup of coffee. The phone rings and on the other end is an irate production supervisor screaming that the plant is about to explode because the new section of pipe you installed last week is galloping up and down. Vast amounts of natural gas will be released into the plant when the pipe breaks. You calm him down and arrange to quickly take the section of pipe out of service. Now that everything is safe, you must figure out what went wrong. We are embarking on a voyage of discovery, that entails us calling at the following way stations:

PART 1 – Calculate the natural frequency of a pipe

PART 2 – Calculate VIV (Vortex Induced Vibration) affecting the pipe

PART 3 – Calculate the effect of flow induced vibration as flow rates change

PART 4 – Determine the severity of the vibration: Is it acceptable or does it need modifications?

PART 1: NATURAL FREQUENCY OF A PIPE

Why do pipes gallop?

Maybe galloping isn’t the right word; I would prefer to say the pipe is vibrating. But why is it vibrating so much? One word – resonance.

If you seat a child on a swing and you give a gentle nudge to the swing every time the child swings back towards you, you will quickly be able to get the swing higher and higher. You only need a very small force because you waited until the swing had reached its limit of movement (the limit of the swing) and only then did you give that gentle push. But you did it EVERY time the swing moved back towards you. A couple of interesting things are happening here.

1. You are pushing at exactly the same frequency as the frequency at which the child is swinging. Don’t forget that if you are troubleshooting a pipe vibration problem the natural frequency will depend not only on the steel of the pipe but also the mass of the fluid inside.

2. You are pushing at the same position in the swing cycle – in other words you have “locked phase” with the swing.

The occurrence of these two “interesting” points means you are in RESONANCE with the swing. When a pipe is vibrating heavily it is almost always because there is a resonance issue. The swing has a natural frequency because it has a period of oscillation that depends on the mass of the swing and the length of the pendulum. This is the reason why a pendulum is used on a clock. If you can work out the period of the cycle back and forth, then you know the frequency.

If the time period of oscillation is one second, then the frequency is one per second or 1Hz. If the time period is half a second, then the frequency is 2Hz and so on. By the same reasoning the pipe has a natural period of oscillation and so it has a natural frequency. The natural frequency of the pipe depends on its stiffness and its mass; the stiffer the pipe the higher the frequency, the more mass the pipe (including contents) has, the lower the natural frequency.

To calculate the natural frequency of a pipe with rigid supports use the following formula:

Where:

f_n = natural frequency of the pipe (Hz)
E = Young’s modulus of elasticity (200 GPa or 30E6 psi for steel – approximately but close enough)
I = 4^th polar moment of inertia for the pipe (0.049*[OD⁴-ID⁴]) in inches or metres
µ = mass per unit length of the pipe (remember to include the mass of the fluid) lbs/inch or kg/m
L = distance between pipe supports (inches or metres)

Let’s find the natural frequency of a 12” (300 mm) pipe made of A-106 GrB schedule 80 that is first empty and then filled with water. I’ll use SI units to make the math easier. Pipe supports are 5 m apart.

Given:

Pipe OD = 323.8 mm
Pipe ID = 288.84 mm
Mass/length empty = 132.05 kg/m
Mass/length full = 132.05 + π (0.28884)²/4 * 1000 = 197.57 kg/m
E = 200 x 10⁹ Pa

Calculate I:

Natural Frequency:

Now you know the natural frequency of the pipe you ask your vibration techs to take a vibration measurement on the pipe when it’s in operation. If the frequency is the same as your calculated frequency, then the pipe has a resonance problem and the next step is to identify what is the force that is exciting the natural frequency.

PART 2 – CALCULATE VIV (VORTEX INDUCED VIBRATION)

In part 1 we determined the natural frequency of the pipe so now we know the pipe natural frequency, but you are not standing there pushing it – so what is? It could be any of a number of things:

►Vibration at the same frequency coming from a pump or compressor (usually speed related). This could be caused by unbalance, misalignment, something may have come loose or just about any fault on a machine that causes a vibration at the same frequency as the natural frequency.

►Flow induced vibration. Now, this could be from the internal flow of the fluid through the pipe or even from wind flowing across the outside of the pipe. We have all seen the effect of wind induced vibration on street lamps or poles so when the wind hits a particular speed the pole starts to sway. That’s because the oscillation force associated with the wind is a function of the pipe outside diameter and the wind speed so the wind vortexes or swirls on the downwind side of the pipe and the vortexing induces KARMAN vibration.

Check this out for more information.

I usually start with the easiest option. Have a look around the pipe and see if there is any rotating equipment that has a run speed (or a harmonic of run speed) that is very close to the pipe natural frequency. If there is can either change the speed of the machine (if possible, because that is the easiest option) or change the natural frequency of the pipe. An easy way to change the natural frequency of the pipe is change the value of L – in other words change the location of the pipe supports or maybe just add another support. If you add another support be careful that you put it at a location of high amplitude – in other words an antinode of vibration.

You can see from the following image that the modifier we used (a=22.4) is only applicable to the first mode of vibration of a “clamped-clamped” beam or pipe. If you install another pipe support halfway between existing supports but the pipe is vibrating at the second mode (a=61.7) the amplitude of vibration will be unaffected. I usually hammer in a stout piece of wood as a (very) temporary measure to see if that indeed reduces the vibration.

If you find that there is no rotating equipment nearby that could affect your pipe we could see if there are any other possible causes and an easy one to check is vortex induced vibration.

For the same reason that a flag flutters, a pipe (or any object) will experience an oscillatory force when placed in a fluid flow. As the wind flows across the pipe there are tiny differences in air pressure from one external side of the pipe to the other, so the wind finds slightly less resistance on one side and more wind flows towards the lower pressure. As more wind flows towards the lower pressure side that side experiences an increase in air pressure so the flow of wind flips over to the other side. The other side experiences the increase in air pressure so the flow flops back again. The flow of wind is now flip flopping back and forth causing a transverse oscillating force on the pipe.

By Cesareo de La Rosa Siqueira – http://www.mcef.ep.usp.br/staff/jmeneg/cesareo/vort2.gif, Copyrighted free use

Going back to our flag analogy, the pipe flutters as the wind passes the flagpole and vortexes on the downwind side. The vortex is traveling along the flag and the flag “flutters.”

We, though, are interested in what is happening to our pipe. According to Strouhal and Karman there is a distinct relationship between the speed of the wind, the diameter of the pipe and the frequency of the oscillating force.

St = fD / V

Where:

f is the frequency (Hz)
D is the diameter of the pipe
V is the wind velocity.

St is the Strouhal number. This does vary somewhat with Reynolds number, but we can assume it to be 0.22.

So, if we have a wind velocity of 10 m/s and we know that our pipe has an OD of 0.3238 m (see Part 1 where we calculated natural frequency) the VIV frequency is 0.22*10/0.3238 = 6.79 Hz. Easy isn’t it?

However, the vortexing frequency effect is not limited to external wind. You will get the same effect from fluid flowing inside the pipe as it flows across an obstruction. So if we have a gate valve with a non-rising stem with a stem diameter of 3.5 cm and a fluid flow of 10 m/s we would have a vortexing frequency of 62.8 Hz. Remember that our pipe has a natural frequency of 63.77 Hz which is close enough to ensure resonance.

In our next article, we will examine the effect of changing the fluid velocity and see how that affects resonance in our pipe.

PART 3 – VIBRATION FROM FLOW VELOCITY

In Part 1 we figured out the natural frequency of a pipe and in Part 2 we looked to see if the resonance excitation was from a nearby rotating equipment or perhaps vortex induced vibration. If neither of these options came close to identifying the forcing frequency, we need to look at slightly more exotic causes.

Most people who work in process, power, oil & gas or refining will have come across a problem in which a perfectly normal section of piping with no significant vibration “suddenly” starts to vibrate for no apparent cause apart from a slight change in flow rate. But that doesn’t seem to make any sense. We have made no change to the mass or stiffness of the pipe, so the natural frequency hasn’t changed. We have not changed any run speed of nearby equipment and even if we check VIV vibration it doesn’t even come close to the problem frequency. What the heck could it be?

Let’s take a trip down to the train station. The express through train is coming down the track and as it passes us, we hear a definite change in pitch. A high pitch as the train travels towards us and a lower pitch as the train moves further away. We are talking about the DOPPLER effect.

To understand Doppler, we need a good understanding of noise and sound. When we speak the sound, we make travels through the air at about 330 m/s. That doesn’t mean we are expelling air from our mouth at that speed – that would be rather unpleasant. As our voicebox vibrates it creates an area of high air density as it pushes onto the air molecules. That high-density rams into the air next to it and bounces back transmitting the energy to the adjacent air molecules. The rate at which the energy is transmitted to adjacent air molecules is the speed of the sound. Remember from high school physics class that the speed of sound is function of frequency and wavelength?

Where:

C is the speed of sound in the fluid
f is the frequency of the sound
λ is the wavelength between the high-pressure pulsations of the sound

As a stationary observer of a stationary object that is making a noise the speed of sound is set by the air density and the air pressure. The pitch or frequency is determined by the distance between the high-pressure pulsations. So, let’s see what happens if we start moving the object making the noise.

When our train is stationary the sound moves away from the train at 330 m/s in all directions. The wavelength is the same as the same travels in all directions. When the train starts to move, we have the sound AND the train traveling in the same direction to the front of the train but moving in OPPOSITE directions when viewed from behind the train. The speed of the sound hasn’t changed but because the train and the sound are traveling in the same direction at the front of the train the wavelength is compressed. Using C = fλ as the wavelength is compressed and the speed of sound stays the same then frequency must increase. The opposite effect happens as the train moves away from us, so we hear a lower frequency or deeper tone.

Factor C is a physical value depending on the properties of the fluid. However, the wavelength is affected by the speed of the fluid flow through the pipe – the Doppler effect.

But we have a pipe with fluid traveling along the inside of the pipe. What does this have to do with moving trains? Only that in both cases we have to think about Doppler.

Let’s move to wind musical instruments. When air is blown into a trumpet or trombone you only get a distinct tone if you blow into the mouthpiece at a particular rate – that is why trumpeters “purse” their lips to get the right airspeed. When you get the correct airspeed the wavelength becomes the same as the length of tubing, so you get a standing wave and the instrument sounds the desired note. The length of the tubing can be adjusted on a trumpet using valves or on a trombone using the slide. In effect, you are changing the wavelength of the air and that is changing the frequency. All of us have tried blowing across a part empty bottle and we get a tone if the speed of the blow is “just right.”

The nearest our pipe is to the analogy of a trumpet is a section of pipe between two bends. The half wavelength is actually longer than the distance between the bends (add about 15%). The part open bottle analogy equates to a dead leg at a tee and in that case, we have a quarter wavelength.

So now let’s combine the standing wave and the Doppler effect in our pipe. Instead of a train moving we now have liquid or gas moving along the pipe. In effect, we are causing the wavelength to change by changing the flow velocity relative to the speed of sound. This can get us into trouble in one of two ways:

1. If you have a forcing vibration at the same frequency you get resonance of the fluid inside the pipe. This happened to me once on a pump running at 2970 rpm discharging into a line that had a length between the discharge flange and the next tee that equated to the standing wave that had a frequency of 50 Hz. Vibration at about 20 mm/s rms and a bearing life of 3 months. We changed the configuration of the discharge piping and vibration came down to less than 2 mm/s and no more bearing failures.

2. If the frequency of the standing wave is close to the natural frequency of the mechanical section of pipe you have resonance. This happened on the discharge of a large blower in China. The distance between the discharge flange of the blower and the NRV gave a standing wave with a frequency that was very close to the natural frequency of the piping. We moved the NRV and the vibration problem disappeared.

Remember that standing waves can occur not only at fundamental frequency but also at “overtones”. Don’t ignore the overtones.

We still use C = fλ to calculate the fundamental frequency so with a speed of sound of 330 m/s and an end-corrected wavelength of 10 m we would have a fundamental tone of 33 Hz. But as soon as fluid flows through the pipe we have to modify that wavelength. So, if we flow at 33 m/s that means a 10% change in apparent speed of sound so we would get a new frequency of (330+33)/10 = 36.3 Hz.

Our pipe is carrying 22 MW natural gas at 80 ֯C with a speed of sound of 365 m/s. Our section of piping has a length of 8.5 m between bends which equates to an open/open pipe of 8.5 m length. which gives a wavelength of 17 m. Add 15% to that length for “end effect” correction and we get a fundamental tone or frequency of 365/ (10*1.15) = 18.54 Hz.

But we now flow gas through the pipe at 20 m/s which, when considering the Doppler effect, changes that frequency to 19.72 Hz. But what makes it really worrying is that the first overtone is 39.45 Hz, the second overtone is 49.31 Hz but the 3^rd overtone is 78.9 Hz which is rather close to our mechanical natural frequency of 78 Hz.

On very large diameter piping there is a possibility of shell wall resonance but that is quite rare and tends to happen on trunking rather than pipework (large diameter and thin wall) so we won’t get into it here.

PART 4 – LIMITS FOR PIPING VIBRATION

Okay – let’s recap. We have a pipe with a natural frequency and a force with the same frequency. This means high amplitude vibration. So what?

If we leave a vibrating pipe in place long enough and the vibration is severe enough the pipe will develop a crack and we get a leak. We are talking about fatigue failure. To make things easy for us there are several versions of fatigue limits we can apply to piping and the one I will mention is API STD 618. Now before you start jumping up and down complaining that is a standard for reciprocating compressors let me say that yes, you are right. But this section of the standard works for all steel piping because it is VERY conservative.

Let’s look at some of the detail.

Section 7.9.4.2.5.2.4 Piping Design Vibration Criteria

The predicted piping vibration magnitude shall be limited to the following:

► A constant allowable vibration amplitude of 0.5 mm peak-to-peak (20 mils peak-to-peak) for frequencies below 10 Hz (the frequency of 10 Hz is also according to ISO 10816).

► A constant allowable vibration velocity of approximately 32 mm/s peak-to-peak (1.25 in./s peak-to-peak) for frequencies between 10 and 200 Hz.

• We need to be aware that 32 mm/s pk-pk is the same as 16 mm/s peak and 11.3 mm/s rms. To convert displacement to velocity:

• V = 2π. Displacement. Frequency

• So 0.5 mm at 10 Hz gives a velocity of 2π*0.5*10 = 31.42 mm/s

• Vibration measurements are almost always displayed in terms of peak-peak for displacement (total movement) but velocity readings are only ever shown as zero to peak or even rms so don’t be confused by the API 618 limits. To see the relationship between pk-pk, peak and rms look at the image below.

The rms. or root mean square velocity is a quantitative measure of the effective velocity and reflects the power or energy being used to vibrate the machine mass. Peak value is the maximum amplitude seen during the measurement referenced to zero velocity and peak to peak is a measure of the total movement so is usually only used for displacement.

This image illustrates the movement of a pendulum. The figure above shows that at position B and C, the velocity is zero, and at position A the velocity is maximum, first to the right, then to the left. The negative peak velocity differs only in direction, not magnitude. The rate of change of displacement is the velocity, therefore if D is expressed in terms of mm, instead of the usual micron, then the product 2pfD will be the velocity in mm/s which are the units used for velocity in vibration work.

This relationship between velocity and displacement is an important factor when considering severity of piping vibration – if you want to be accurate. I really don’t care about vibration amplitudes or movement in pipes, what I care about is the stresses that have been imparted to the pipe. If we consider a resonant pipe, then the actual forces are quite low but the physical movement (or displacement) could be quite high.

If we consider the stress-strain diagram for carbon steel, A106GrB has a tensile stress specified at 415 MPa and yield stress of 240 MPa (see below).

We need to make sure the stress value due to the bending effect of the piping vibration is well below the yield value and we can use our high school physics knowledge to do this.

Where:

E = Youngs modulus

D = pipe outside diameter

Δ = peak to peak displacement

L = length of pipe between supports

So, let’s say our 5 m pipe has a maximum vibration amplitude of 0.5 mm pk-pk. The bending stress is

S = 8 * 200E9 Pa * 0.3238 m * 0.5 / 5² = 10.36 MPa

That sounds quite reasonable but what happens if the vibration is 5 mm pk-pk? The stress is now 103 MPa. Compare that to our yield stress of 240 MPa and you would think our pipe should be fine, but API 618 warns us that if there is a cyclic stress level in the piping that stress level must not exceed the endurance limits of the piping materials. So, what is the endurance level? It is the value of the stress below which a material can presumably endure an infinite number of stress cycles, that is, the stress at which the S-N diagram becomes and appears to remain horizontal. The existence of a fatigue limit is typical for carbon and low alloy steels.

Looking at the image below we have a plot of applied stress against the number of cycles the metal endured before rupturing of a carbon steel of a known UTS. It is possible to predict (very approximately) the life to failure of the equipment if we have this plot. Our A-106 pipe steel follows the dashed line fairly closely so we can see that if we have a stress of 103 MPa the steel would be expected to fail due to cyclic loading at about 200,000 cycles.

If our pipe is vibrating at 67 Hz that means 67 cycles per second, so we have a life before failure of 200,000 / 67 = 2958 seconds – less than an hour.

To operate our pipe safely for an extended period then the vibration MUST induce a stress that is less than the endurance limit. And that is why the API recommendation is so conservative. But if you really need to operate the equipment you can carry out a fatigue analysis as we have done or you could use API 579-1/ASME FFS-1 which can guide you through a step by step procedure to determine if you can use the pipe for an extended time.

In conclusion, if your pipe is vibrating, you should do the following:

►Find out if the vibration is resonant – that means find the natural frequency

►If the vibration is resonant find out what is the forcing frequency – that could be a nearby piece of equipment, external effects such as wind or fluid flow induced vibration

►Determine if the vibration levels are acceptable – you could use API 618 for that or carry out a full fatigue analysis.

To remedy the problem, you simply have to separate the natural frequency from the forcing frequency.

To change the natural frequency, you must change either the stiffness or the mass. It is usually much easier to change the effective mass by changing the length between supports either by moving one of the existing supports or by adding a new one.

Otherwise, you have to attack the forcing frequency. If it’s a rotating machine you can change the speed, isolate the machine from the pipework or fix the vibration issue such as unbalance or misalignment. If the problem is flow induced or VIV you will have to do the analysis to positively identify the culprit and make changes in flow rate, piping arrangement or even production procedures.

I hope you find this series of posts useful. To learn more about similar cases and how to minimize operational troubles, we suggest attending our ME41 (Piping Systems – Mechanical Design and Specification) and PF49 (Troubleshooting Oil and Gas Processing Facilities) courses.

By: Ron Frend, Head of Facilities Training

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Written on May 5, 2020 at 10:32 am, by admin

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