The best way to prevent hydrate formation (and corrosion) is to keep the pipelines, tubing and equipment dry of liquid water. There are occasions, rightly or wrongly, when the decision is made to operate a line or process containing liquid water. If this decision is made, and the process temperature is below the hydrate point, inhibition of this water is necessary. Many materials may be added to water to depress both the hydrate and freezing temperatures. For many practical reasons, an alcohol or one of the glycols is injected as an inhibitor, usually methanol, diethylene glycol (DEG) or ethylene glycol (EG). All may be recovered and recirculated, but the economics of methanol recovery may not be favorable in many cases. Total injection rate is that needed to provide the necessary inhibitor concentration in the liquid water plus that inhibitor which enters the vapor and hydrocarbon liquid phases. Any inhibitor in the vapor phase or liquid hydrocarbon phase has little effect on hydrate formation conditions. Determination of the amount and concentration of inhibitors and their distribution in different phases are very important for practical purposes and industrial applications. Therefore, in order to determine the required amount and concentration of these inhibitors, several thermodynamic models for hand and rigorous calculations have been developed and incorporated into the computer software.

In this Tip of the Month, we will demonstrate the effect of total inhibitor circulation rate on hydrate depression in a sales gas dew point correction plant. Let’s consider the process flow diagram shown in Figure 1. The feed composition and conditions are shown in Table 1. The wet gas analysis was used in this study.

It is desired to process this feed gas to produce a sales gas with a dew point of -5˚F (-20.6˚C) at 990 psia (6.826 MPa) The feed gas is mixed with recycle gas from the NGL/Glycol separator, compressed and cooled to 120˚F (48.9˚C) and 1000 psia (6.895 MPa), and EG lean solution and then cooled in the HTX-1, and finally in the HTX-2 to -5˚F (-20.6˚C) before entering the separator at 990 psia (6.826 MPa). For the sake of simplicity, the refrigeration cycle for HTX-2 system is not shown in the above process diagram. Figure 2 represents the feed phase envelope, its hydrate curve, sales gas envelope and the cooling path. As seen in this figure, the gas temperature drops below the hydrate formation temperature of about 65˚F (18.3˚C) in the HTX-1 and to -5˚F (-20.6˚C) in HTX-2. Therefore; to prevent the hydrate formation in the HTX-1 and HTX-2, it is decided to inject an 80 weight percent lean EG solution to the gas stream. The concentration of rich inhibitor solution can be calculated by the shortcut method of Hammerschmidt [1] as described in Chapter 6, Volume 1 of Gas Conditioning and Processing [2] or using rigorous thermodynamic models. The required circulation rate is then determined by material balance and phase equilibrium calculations.

In order to show the effect of EG solution circulation rate on the depression of the hydrate formation temperature, the whole process shown in Figure 1. Figure 3 shows how the hydrate formation temperature in HTX-2 drops with increasing total inhibitor circulation rate. The corresponding calculated weight percent of EG in rich solution is also shown as a function of circulation rate. The concentration of EG in the rich inhibitor solution increases with the increase in the inhibitor solution circulation rate. Figure 3 indicates that for a 10˚F (5.6˚C) depression, or the hydrate formation temperature of -15˚F (-26.1˚C), a circulation rate of 975 lb_m/hr (442 kg/h) is required. At this circulation rate the corresponding weight percent of EG in rich solution drops to 74.

To learn more about similar cases and how to find the optimum inhibitor circulation rate and concentration for prevention of hydrate formation, we suggest attending our G4 and G5 courses.

Dr. Mahmood Moshfeghian

References:

1. Hammerschmidt, E.G., “Formation of gas hydrates in natural gas transmission lines,” Ind. & Eng. Chem., Vol. 26, p. 851, 1934
2. Campbell, J. M., “Gas Conditioning and Processing, Vol. 1, the Equipment Modules, 8^th Ed., J. M. Campbell and Company, Norman, Oklahoma, 2001

In this Tip of the Month, we will explore the regeneration of molecular sieve dehydrators. Can you save energy by ending the heating cycle when the regeneration outlet temperature reaches approximately 90% of the regeneration inlet heating temperature? Frequent readers of the Tip of the Month surely know the answer: it depends!

But on what does it depend? There are many factors to consider, but to simplify this discussion we shall assume a molecular sieve unit designed to do the following:

• Dehydrate natural gas to less than 1 ppmv
• Heating and cooling are done countercurrent to adsorption and the regeneration medium is bone dry
• The regeneration inlet temperature to the molecular sieves is 288˚C (550˚F)

First, let’s take a brief review of the entire cycle.  Molecular sieve dehydrators are typically two or three tower units. While one vessel is being regenerated, the remaining vessels are adsorbing water from the flowing natural gas.

Below is a simplified cycle schedule showing a three bed dehydrator where A1=First Half of Adsorption Cycle Time, A2=Second Half of Adsorption Cycle Time, C= Cooling Cycle Time and H=Heating Cycle Time:

Not shown in this cycle schedule are the times required to depressurize and re-pressurize the system. The time required for these steps are included in the regeneration cycle and must be done in such a manner to avoid lifting the molecular sieves and/or causing a mechanical bed support failure.

The astute reader will notice the total regeneration time per vessel (heating plus cooling time) must be equal to or less than the total adsorption time per vessel (A1+A2). Failure to accomplish this will result in early water breakthrough of the vessel on adsorption and will ruin someone’s day.

Let’s now explore the mechanisms that lie behind the regeneration cycle of a molecular sieve bed.  During regeneration sufficient sensible heat must be provided to heat everything in the vessel up to 260˚C (500˚F) or so. This includes the molecular sieves, the inert support media, the metal, etc.  The heat of desorption of the water must also be supplied. The sum of the sensible heat plus heat of desorption is the required regeneration heat and helps set the regeneration heating cycle. The heating time is larger than the cooling time because the heating cycle must provide the heat of desorption of water while the cooling cycle is concerned only with sensible heat.

There is more to regeneration than simply waiting until the temperature of the spent regeneration gas reaches a plateau before switching to cooling.  The heating cycle is not complete until the water is swept out of the system and the molecular sieves have reached their design residual water loading. During low pressure regeneration, the limiting step is simply getting the total required heat into the bed of molecular sieves. In this instance, once the outlet temperature reaches approximately 90% of the inlet temperature [262˚C (500˚F)], the heating cycle is completed. In some situations, the heating cycle is actually stopped before the outlet temperature reaches 262˚C (500˚F). Such a cycle is called a thermal pulse.

During high pressure regeneration, the limiting step is getting the water away from the molecular sieves. Despite the fact that the outlet temperature has reached a plateau, the regeneration may not be completed because the high pressure gas cannot carry away the water molecules. In this situation, you have to continue heating the molecular sieves for a period of time until the residual moisture left in the bed is equal to or less than what it was designed for.

This should lead the reader to ask “How can I tell what is limiting my regeneration?” The answer to this is not so simple. One could run field trials; however, this is time consuming and provisions must be taken to avoid premature water breakthrough. It would be much simpler to contact your molecular sieve vendor and ask them. They have design techniques available to determine what the limiting mechanism is.

A conservative rule of thumb is if the regeneration pressure is greater than 4 MPa [600 psia], be careful. Sweeping the water from the molecular sieves may be your limiting step. This discussion assumes you are operating between a minimum pressure drop during the regeneration heating cycle of 0.23 kPa/m [0.01 psi/ft] to ensure good flow distribution and a maximum pressure drop of 5.4 kPa/m [0.24 psi/ft] to avoid bed lifting.

Harvey M. Malino

In this Tip of the Month, we will show how to determine the optimum interstage pressure for a two-stage compression process. We will also study other operating condition such as feed temperature, heavy end in the feed, and water moisture. For this purpose, we used a commercial simulation package and the SRK EoS for the prediction of phase behavior and thermodynamic properties.

The gas mixture with the composition shown in Table1 at 37 °C (98.6 °F) and 31 bar(g) (450 psig) on a dry basis is compressed in two stages for injection into an oil field as a means of enhancing production. The process flow diagram is shown in Figure 1. The injection pressure is 131 bar (g) (1900 psig) and temperature is 65 °C (149 °F). The gas rate for stream 2 is 6.792×10⁶ std m³/d (240 MMSCFD). The suggested isentropic efficiency is 72 percent and a mechanical efficiency for each stage of compressor is 80 percent. The inlet temperature of each compressor stage should not exceed 56 °C (132.8 °F). The feed gas is saturated with water and 5 psi (34 kPa) pressure drop is allowed between each compressor discharge and exit of the flash separator.

Phase Envelope – The first step is to determine the state of the inlet to the 1^st stage suction scrubber. The phase envelope for the feed gas after being saturated with water (stream 2) is shown in Figure 2. This figure also presents the phase envelope for stream 6 which is the vapor stream at the suction to the first stage of compression. The red circle displays the condition at the first stage suction.

Optimization Scenarios-Figure 3 presents the variation of compression ratios as a function of 1^st stage discharge pressure. From this figure, it can be seen that equal compression ratios of 2.04 is obtained at a pressure of 65.2 bar. The ideal optimum interstage pressure for equal compression ratio is also found to be  bar. Figure 4 presents the variation of each heat exchanger cooling load and each stage compression power requirement as a function of 1^st stage discharge pressure. These variations are almost linear.

Table 2 presents the simulation results for two cases of optimizations. In case A , the total power requirements was minimized by finding the 1^st stage discharge pressure with the constraint of equal stage compression ratio. This results in approximately equal compression power requirements for the two stages. It should be noted that the slight difference in compression ratio and stage compression power is due to the 34 kPa (5 psi) pressure-drop between discharge of 1^st stage and suction of 2^nd stage. However, in Case B, the total energy requirement was minimized by finding the 1^st stage discharge pressure without the constraint of equal stage compression ratios. The results summarized in Table 2 indicate that there is a big difference between the case A and case B 1^st stage discharge pressures. It can also be seen that the case A total power requirement (W₁+W₂) is clearly larger than case A (about 40 % higher).

The variation of total compression power and total cooling load requirement as a function of 1^st stage discharge pressure are shown on the left hand side y-axis of Figure 5. This figure indicates clearly that the minimum power requirement occurs when the 1^st stage discharge pressure is 77.8 bars. Figure 5 also provides an indication that the total power requirement changes very little for 1^st Stage discharge pressures between about 76 bars and 80 bars.

Effect of Water Vapor in the Feed-The detail of simulation results based on unequal compression ratio for the two options of wet feed (saturated with water vapor) and dry feed is shown in Table 3. As can be seen from this table, water vapor has little effect on the performance of the process. The 0.23 % increased compression power requirement for the wet feed is due to 0.28 % increase in feed flow rate for the presence of water. It should be noted that the dry feed flow rate is 11953 kmol/hr and the wet feed flow rate is 11987 kmol/hr (34 kmol water/hr + 11953 kmol of dry gas/hr).

Effect of Heavy Ends in the Feed-In order to study the effect of heavy ends on the performance of the process, normal octane (nC₈H₁₈) was replaced with normal decane (nC₁₀H₂₂) and the simulation was repeated. The detail of simulation results based on unequal compression ratio for these two options of heavy ends is shown in Table 4. As can be seen from this table, the total compression power requirements decreases slightly for the case using nC₁₀H₂₂ due to the fact that more of the heavy component is removed in the first separator. The compressor power for the stages and heat exchanger duties are not affected by the presence of heavier components in the feed stream. In other words, the feed flow rate to the compressor decreases when nC₈H₁₈ is replaced by nC₁₀H₂₂.

Effect of Feed Temperature – In order to study the effect of feed temperature on the performance of the process, the feed temperature was increased from 37 °C to 56 °C and the simulation results are shown in Table 5.

Table 5. The effect of feed temperature on the performance of the process

The feed at 56 °C represents the actual condition during the summer season. As can be seen from this table, the warmer feed requires an increase of 5.34 % in total compression power consumption. So the feed temperature is an important parameter and its variation, especially due to seasonal change, should be taken into considerations.

For the unequal compression ratio case having compression ratios of 2.397 and 1.733 for stages 1 and 2, respectively, the variation of energy requirement with feed temperature is shown in Figure 6. Stage 2 was not affected with the variation of feed temperature; therefore, the compression power and cooling load for stage 2 remained constant at 7.254 MW and 8.407 MW, respectively. Since, the compression ratio was constant, the compression power requirement for stage 1 varied from 12.2 to 13.28 MW; however, the cooling load varied drastically from 10.85 to 14.83 MW.

In the light of preceding discussion, the following tips are suggested:

• Be sure to check the phase of the compressor suction stream. This also includes the interstage condition to ensure that liquid does not enter the compressor.
• If economically possible, lower the interstage suction temperature since this will reduce the overall compression power requirement.
• Be sure to check the water content at the interstage conditions since there may be water drop out which would impact equipment performance.
• The choice of equal pressure ratios for minimizing the compression power requirement is close to an optimum choice when the suction temperatures are equal.
• Characterization of the heavy ends (C₇₊) does not greatly impact the compressor power requirement since heavy components are mostly removed in the inlet scrubber.
• Characterization of the C₇₊ will impact the condensation that takes place in the inlet suction scrubber and thus the molecular weight of the compressed gas will be affected.

Dr. Mahmood Moshfeghian

Problems in meeting sales-gas dew point specifications are not unusual.  A facility engineer often suspects separator carryover when trouble-shooting such a plant.  Proper sizing of equipment for vapor-liquid separation is essential to almost all processes.  The fundamentals of a simple separator design may be extended to several other processes such as fractionation towers, two-phase flow, slug catcher design etc.  Many facility operating problems are related to improperly designed or under-sized gas-liquid separators.

In this Tip of the Month, we will demonstrate the impact of liquid carry over on the sales gas “spec dew point”. Let’s consider the process flow diagram shown in Figure 1 for a simple gas plant. The feed composition and condition are shown in Table 1.

It is desired to process this feed gas to produce a sales gas with a dew point of 20 ˚F (-6.7˚C) at 540 psig (3.723 MPa) The feed gas is mixed with recycle gas from stabilizer, compressed and cooled to 110˚F (43.3˚C) and 555 psig (3.827 MPa), then cooled in the gas-gas exchanger, gas-liquid exchanger and finally in the chiller to 20˚F (-6.7˚C) before entering the separator at 540 psig (3.723 MPa). The phase envelopes for the feed, vapor and liquid streams of the separator are shown in Figure 2. Figure 2 presents the phase behavior of separator if it was operating perfectly, without any liquid carry over or entrainment. In real life, the situation is different and most probably there would be some liquid carry over. In order to show the impact of liquid carry over, we withdraw a small portion of liquid stream from separator and remixed it with the vapor stream and recalculated the dew point temperature. Figure 3 shows how the sales gas dew point curves shift to the right as the percentage of liquid carry over increases. Figure 4 shows the calculated sales gas dew point temperature, at 540 psig (3.723 MPa), as a function of percentage of liquid carry over. It is interesting to note that even a 5% liquid carry over shifts the dew point temperature by more than 20˚F (11˚C) which may results in unexpected large amount of condensate as the gas transported in the pipeline or cause severe damage to the downstream compressor. To learn more about similar cases and how to prevent operational problems such as liquid carry over, we suggest attending our G4 and G5 courses.

By: Dr. Mahmood Moshfeghian

In a previous “Tip of the Month” we discussed several methods of heavy ends characterization and as an example, for a rich natural gas, we tuned the heavy end parameters to match the experimentally measured saturation pressure. After tuning the heavy end parameters, we obtained a phase envelope for each method that passed through the experimental dew point; however, their shapes and specifically, their cricondentherm points were different. At the end, we were faced with the question of “which one is the right phase envelope?”

In this tip, we will explain a procedure for selecting an appropriate C₆₊ characterization method which results in a broader match with the experimental information. For further detail, please refer to Gas Conditioning and Processing, Volume 3, Advanced Techniques and Applications.

Let’s consider a lean natural gas with the composition shown in the first two columns of Table 1. This lean gas contains only 0.067 mol% C₆₊ and even though the amount of C₆₊ is very small, we will see that it has a large impact on the condensate. The detail laboratory analyses [1] of C₆₊ are shown in the 1^st and 2^nd columns of Table 1. We used the SRK EoS in GCAP Software for Vol. 3 of Gas Conditioning and Processing and determined the C₆₊ properties of MW=94.12, SG=0.738, andNBP=195 °F. All other calculations were performed using theSRK EoS.

Table 1. Characterization/Distribution of C₆₊ by different methods

We have plotted the experimentally measured Potential Hydrocarbon Liquid Condensed (PHLC) at 594.7 psia as a function temperature in Figure 1, and by extrapolation, a dew point temperature of 52.1 °F was obtained. We used the methods discussed in the previous tip of the month to characterize the C₆₊ by matching the experimentally determined dew point. For each method, we have presented the tuned MW and distribution of components in Table 1 and the predicted PHLC as a function of temperature and the phase envelope were plotted in Figures 1 and 2, respectively. It should be pointed out that the cricondentherm for the lumped C₆₊ method was below 15 °F; therefore, zero values for PHLC were predicted.

Figure 2 indicates that three of the characterization methods practically generate the same phase envelope. Are all of these phase envelopes presenting the true phase behavior of this lean gas? The answer is “Yes” and “NO”. Figure 1 indicates that, for temperatures close to the dew point, all three methods predict the PHLCs very close to the experimental values; however, at lower temperatures (about 20 °F) the deviation from experimental values increases. Only the normal alkane distribution method gives accurate values even at lower temperatures.

In summary, for sound process design and/or operation we suggest using at least one experimental saturation measurement such dew point or bubble point near the potential operating conditions to characterize the heavy ends. Once the characterization is established, additional experimental measurements should be utilized to verify the accuracy and validity of the tuning technique.

By: Dr. Mahmood Moshfeghian

Reference:

1. Derks, P. A. H., van der Meulen-Kuijk, L., and Smit, A. L. C., “Detailed Analysis of Natural Gas for an Improved Prediction of Condensation Behavior,” Proceedings, 72^nd Annual Convention, Gas Processors Association, 1993.
2. Riazi, M.R. and T.E. Daubert, Hydr. Proc. P. 115, (March) 1980

In a previous “Tip of the Month” we explained how a phase envelope is generated and what factors affect the shape and accuracy of a phase envelope.

In this tip, we will show several methods of C₇₊ (heavy ends) characterization and check the accuracy of each method and present tips to improve the accuracy of each method. For more detail, please refer to Gas Conditioning and Processing, Volume 3, Advanced Techniques and Applications.

Let’s consider a rich gas with the composition shown in the first two column of Table 1. This rich gas contains 3.1 mol% C₇₊ with reported MW=132 and SG=0.774. The experimentally measured dew point pressure is 4075.6 psia at 180.4 °F. For this case study, we will use PR EoS, needless to say that similar results are obtained by SRK EoS.

Method A: In this method we will treat C₇₊ as a single cut and based on its MW and SG, we will predict its normal boiling point (NBP=317.2 °F), critical temperature (Tc=645.2 °F), critical pressure (Pc=371.6 psia), and acentric factor (ω=0.425574) using correlations similar to the ones by Riazi and Duabert [1] which are on page 107, Chapter 4, Gas Conditioning and Processing, Volume 1, Basic Principles. The predicted dew point pressure is 3433 psia far away from the measured value of 4076 psia. The predicted phase envelope for this method is shown in Figure 1. As can be seen, this phase envelope does not pass through the measured dew point.

Method B: We break the C₇₊ into 11 Single Carbon Numbers (SCN) ranging from SCN 7 to SCN 17+ using the exponential decay procedure presented by Katz [2] and applied by others [3-5]. The resulting distributions are shown in the 3^rd and 4^th column of Table 1. The predicted dew point pressure is 4030 psia which is relatively accurate. The corresponding phase envelope is shown in Figure 1 which almost passes through the measured dew point.

Method C: We lumped the 11 SCN components of Method B into 4 cuts. These cuts and compositions are shown in the 5^th and 6^th column of Table 1. The predicted dew point pressure is 3885 psia and the corresponding phase envelope is plotted in Figure 1.

Method D: This method is similar to Method B, except that we used 12 normal parafins (alkanes) instead of SCN components to represent the C₆₊. The 7^th and 8^th column of Table 1 present the distribution of nC₆ to nC₁₇. The predicted dew point pressure is 3730 psia and the corresponding phase envelope is also shown on Figure 1. The advantage of this method is that n-alkane components are readily available in many commercial software where as the SCNs may not.

Figure 1 indicates that none of the methods, except method B which come close, matches the experimentally measured dew point. In the next section, we show how the prediction of each method can be improved to match the experimentally measured dew point.

Adjusting MW (or Tc) in Method A: As can be seen in Figure 2, by changing MW of C₇₊ to 152.5, we can match the measured dew point, perfectly. Please note that based on MW=152.5 and SG=0.774, the new values for the estimated properties will be NBP= 381 °F, Pc=308.4 psia, Tc=706.9 °F, and ω= 0.468456. An alternative option is to adjust only Tc to 703.5 °F. Again, the tuned Tc makes the phase envelope pass through the measured dew point as shown in Figure 2. The Tc adjustment is preferred because less work is involved.

Tuning binary interaction parameters, k_ij, in Methods B and C: A common correlation to estimate the binary interaction parameter is:

In the above equation, Vc_i and Vc_j represent the critical volumes of components i and j, respectively. The default value of exponent n is normally set to 1.2 but it can be used as a tuning parameter to match the experimentally measured dew point. For our case study, we obtained n=1.3011 for Method B and n=1.6573 for Method C. The resulting phase envelopes for these two methods are shown in Figure 2.

Tuning MW in Method D: The distribution (i.e. mole %) of the alkane part of C₆₊ depends on the assumed value of C₆₊ MW. As shown in Figure 2, by changing MW of C₆₊ to 154, we can match the measured dew point, perfectly.

As can be seen from Figure 2, all of the generated (tuned) phase envelopes are passing through the measured dew point; but they have different cricondentherm point. Are all of these phase envelopes correct? In the next tip of the month we will demonstrate how to choose the most accurate phase envelope!

By: Dr. Mahmood Moshfeghian

In a previous “Tip of the Month” we briefly discussed the need for understanding a phase diagram in a gas processing system. We also defined the areas of a light mixture phase envelope and the terms necessary to “talk intelligently” about the shape of a mixture phase diagram. This allowed us to look at the methods of calculation and their limitations in another tip and eventually defined our areas of risk in the operation or design of a facility based on the phase diagram.

In this Tip we will explain how a phase envelope is generated and what factors affect the shape and accuracy of a phase envelope. There are two methods of generating a phase envelope: a) by conducting a series of bubble point and dew point measurements in a PVT laboratory b) Using a cubic EoS such as SRK or PR and performing a series of bubble point and dew point calculations. The triangle symbols in Figure 1 present a phase envelope measured in a laboratory for a synthetic natural gas.

For the same mixture, we used GCAP for Volume 2 of Gas Conditioning and Processing Software to generate the phase envelope using SRK and PR EoSs. The dashed line represents the SRK EoS and the solid line represents the PR EoS. As can be seen, for this case the SRK EoS gives a perfect match with the experimentally measured dew point curve but PR predicts a lower cricondentherm point. The built in and in–house binary interaction parameters for SRK and PR were used to generate these two diagrams; however, an experienced engineer is able to produce a close match for either of these two EoSs by tuning the binary interaction parameters and/or heavy end properties.

The best practical method for generating an accurate phase envelope by any commercial software is to utilize a limited number of VLE measurements and tune one or more properties of the heavy ends (C₇₊). In other words, we suggest using combination of methods “a” and “b”.

For real natural gas mixtures, the components and composition of the heavy ends are not well defined and laboratory measurements are not accurate enough. Therefore, different techniques as described in the literature and Volume 3 of Gas Conditioning and Processing are used to properly characterize the heavy ends resulting in an overall acceptable match.

Since the laboratory reported values of molecular weight (MW) and specific gravity (SG) of C₇₊ are questionable, in the proceeding section we will demonstrate the impact of these two properties on the shape of overall phase envelope.

Figure 2 presents the impact of molecular weight of C₇₊ on the phase envelope of a rich natural gas mixture. As the C₇₊ molecular weight increases, the cricondentherm temperature and cricondenbar pressure will increase and the two phase region expands.

Figure 3 shows that as the C₇₊ specific gravity increases, the cricondentherm temperature and cricondenbar pressure will also increase and the two phase region expands.

From these two diagrams, one can see that improper characterization of heavy ends may result in a bad design or troublesome operation (e.g. designing a dry gas pipeline instead of two phase gas-liquid flow).

In the next tip of the month we will demonstrate how to tune computer software to generate proper and relatively accurate phase envelopes.

By: Dr. Mahmood Moshfeghian

Reference:

1. C. Jarne, S. Avila, S. T. Blanco, E. Rauzy, S. Otin, I. Velaso, Ind. Eng. Chem. Res. 43 (2004) 209-217.

In a previous “Tip of the Month” we briefly discussed the problems of how to “name” the various regions of the phase diagram. To decide if this “naming game” can be a problem, look at Figure 1.

This is a typical two-stage compressor system with intercoolers. This particular system however is being used to compress gas to a pressure above the cricondenbar of the gas for “dense phase” transportation. Note that the streams are only labeled with letters because I used numbers on the phase diagram. As a “process engineer” I find it difficult to discuss streams with out seven digit numbers identifying them.

Point A is easily “defined” as a saturated vapor. Point B will be superheated vapor or possibly dense phase. Point C might be 2 phase, super heated vapor, dense phase, or liquid phase. From an equipment point of few in all cases but the two-phase case, separator V-2 is not needed, in theory. In practice, we would always install V-2 as very minor changes in temperature or pressure can sometimes dramatically change the quantity of liquid in stream E.

For computer simulation of this process, the designation of the “phases” of stream C becomes quite important. Assuming the simulator has the ability to calculate Vapor-Liquid equilibrium properly, the two-phase case is the only case where the simulator can be trusted to give the proper duty requirements for compressor C-2.

Figure 2 is the “generic” phase diagram for a light mixture that we previously discussed. The general areas are: to the left of the phase diagram – liquid; area to the right of the phase diagram – vapor; the area inside the phase diagram – two phase; and the area above the phase diagram – dense phase.

My personal preferences for the sub areas are listed below:
Area 1 – Vapor
Area 2 – Vapor
Area 3 – Vapor
Area 4 – Vapor
Area 5 – Dense phase
Area 6 – Liquid
Area 7 – Liquid
Area 8 – Liquid
Area 9 – Vapor

The stream labels from figure 1 are plotted on the diagram. Note this assumes that the phase diagram represents the composition of the gas leaving separator V-1.

If stream C is “named” vapor by the simulator, then most computer simulations will properly calculate the power requirements of C-2. If stream C is “named” liquid by the simulation program then C-2 will have no flow through it. The program will either “fail” or worse yet give meaningless answers. Worse yet some simulators identify point C as two-phase with an arbitrary split between the flow rates of streams D and E.

The current trend in commercial simulators is to always “give an answer” no mater how “stupid the question is.” Most simulators will generally give warnings when calculating a “meaningless” answer, but with the current shortage of engineers, the chances of these being identified are fairly small.

This is a simple example of how “naming” can get you in trouble. There are many more that occur subtly inside simulations. Many of these errors go undetected because the program is quite “happy” with whatever phase it calculates.

For examples of how the choice of the name might be important in your process simulation calculations, contact JMC@jmcampbell.com for test simulations that you can use with your process simulator.

At a future date we will discuss the meaning of “condensate” and why this is so politically important even today.

In the meantime Chapter 4 and 5 of Volume I and chapter 15 of Volume II of Gas Conditioning and Processing and Chapter discusses these topics in more detail.

The stream labels from figure 1 are plotted on the diagram. Note this assumes that the phase diagram represents the composition of the gas leaving separator V-1.

WHAT If stream C is “named” vapor by the simulator, then most computer simulations will properly calculate the power requirements of C-2. If stream C is “named” liquid by the simulation program then C-2 will have no flow through it. The program will either “fail” or worse yet give meaningless answers. Worse yet some simulators identify point C as two-phase with an arbitrary split between the flow rates of streams D and E.

The current trend in commercial simulators is to always “give an answer” no mater how “stupid the question is.” Most simulators will generally give warnings when calculating a “meaningless” answer, but with the current shortage of engineers, the chances of these being identified are fairly small.

This is a simple example of how “naming” can get you in trouble. There are many more that occur subtly inside simulations. Many of these errors go undetected because the program is quite “happy” with whatever phase it calculates.

For examples of how the choice of the name might be important in your process simulation calculations, contact JMC@jmcampbell.com for test simulations that you can use with your process simulator.

At a future date we will discuss the meaning of “condensate” and why this is so politically important even today.

In the mean time Chapter 4 and 5 of Volume I and chapter 15 of Volume II of Gas Conditioning and Processing and Chapter discusses these topics in more detail.

By: Dr. Larry L. Lilly

In a previous “Tip of the Month” we briefly discussed the need for understanding a phase diagram in a gas processing system.  In this Tip we will try to clearly define the areas of a light mixture phase envelope and the terms necessary to “talk intelligently” about the shape of a mixture phase diagram. This will allow us to look at the methods of calculation and their limitations in another tip and eventually define our areas of risk in the operation or design of a facility based on the phase diagram.

The figure is a “generic” phase diagram. The general areas are 1) to the left of the phase diagram – liquid; 2) to the right of the phase diagram – vapor; 3) inside the phase diagram – two phase; and 4) above the phase diagram – dense phase.

However, note that there are several sub-areas that might be questioned as to their exact phase name.  For instance the area marked “3” is “above the phase diagram but below the highest pressure that two phases can exist.  Should this region be called “vapor” or “dense phase?”  Or perhaps the more important question is:  “Do I care?”  The answer to the second question is “yes, sometimes.”  The answer to the first question is “maybe, sometimes.” With these two very clear answers in mind, maybe the general criteria needs to be further defined.

For a quick review let’s start with pure component and mixture phase diagram differences.

For a pure component an acceptable definition of the CRITICAL POINT is “the highest temperature and pressure that two phases can exist for a given component.”  The critical point for our generic mixture is point C. This is an experimentally determinable point and is the point where the mixture properties in the vapor phase and the liquid phase are the same.  As shown on the diagram there is a considerable region at temperatures higher than the critical point that is two phase and a smaller region of pressures higher than the critical pressure that is also two phase.

Two new terms are introduced to describe these regions and their limits.  The first term is the name given to the highest temperature at which two phases can coexist.  This is point T in the figure.  This point is called the cricondentherm. The second term is the highest pressure at which two phases can coexist – point B. This point is called the cricondenbar. Any facility that operates at a temperature always higher than the cricondentherm will never condense liquids.  Any process that is operated at a pressure higher than the cricondenbar will not be two-phase. A process in this region can be liquid, vapor or “dense phase”, but never two at the same time. Dense phase pipelines are designed to always be at pressures above the cricondenbar. Dewpoint control is often a matter of controlling the temperature of the cricondentherm for sales gas so that the gas pipeline is not two-phase at the coldest temperature in the system.

Now back to defining each of the areas in the figure.  The easy areas first:  Areas 1 and 2 are generally called vapor. Areas 7 and 8 are generally called liquid.  Areas 5a and 5b are generally called dense phase.

Area 3 (vapor or dense phase) is generally called a vapor because it is below the cricondenbar (B) and condenses liquid when the pressure is decreased at constant temperature.

Area 4 (vapor or dense phase) could be called vapor or dense phase depending on the exact definition used. For instance, if all areas above the cricondenbar (B) are defined as dense phase then area 4 is dense phase. If all areas to the right of the cricondentherm (T) are defined as vapor then area 4 is vapor.

Area 6 (liquid or dense phase) could be called liquid or dense phase again depending on our exact definition.  If all areas above the cricondenbar (B) are defined as dense phase then it is dense phase. If all areas to the left of the critical point (C) are defined as liquid then it is liquid.

Area 9 (liquid, vapor, or dense phase) could be interpreted as any of the phases depending on your definitions.  Some people would define any area to the left of the cricondenbar (B) as a liquid.  This would imply that areas 5b, 6 and 9 are liquid.  Most people would define any single-phase fluid outside of the dewpoint line and below the cricondenbar (B) as a vapor.  This would imply that area 9 is a vapor.  Some people would define any single-phase fluid above the critical point (C) as dense phase.

My personal preferences are:

Area 1 – Vapor
Area 2 – Vapor
Area 3 – Vapor
Area 4 – Vapor
Area 5 – Dense phase
Area 6 – Liquid
Area 7 – Liquid
Area 8 – Liquid
Area 9 – Vapor

These choices are made based on the answer to the “Do I care?” question.  In a future tip we will discuss how these names (doesn’t a rose by any other name smell just as sweet?) do make a difference from a simulation point of view and from an interpretation of real world problems.

In the meantime Chapter 4 of Volume I Gas Conditioning and Processing discusses these topics in more detail.

By: Dr. Larry L. Lilly

In facilities operations the understanding of where the process is on a phase diagram can often help the engineer and operator avoid extremely embarrassing design and operating mistakes. The oil and gas industry is full of many “war stories” about “phase diagram disasters.” Most instances are never related back to the phase diagram misunderstanding.  In one well-documented but poorly published case a “dry gas” pipeline that was pigged flooded miles of sandy beach.  In another case thousands of kilowatts of compression power were installed to maintain the pressure of a reservoir above the dew point when in fact the reservoir was at a temperature above the cricondentherm.  In many cases equipment manufacturers and purchasers of gas have specifications of “superheat” or dew point that have not been met and led to upset customers and/or millions of dollars of lawsuits.

One of the first issues to be resolved by a facilities engineer working in a gas plant or gas production facility is where is the process operating with respect to the phase diagram.  A general knowledge, if not a detailed knowledge, will allow the design engineer and the facilities operator to make intelligent decisions that have significant impact on the profitability of a gas production facility.

The following figure is a “generic hydrocarbon mixture” phase diagram for a lean gas.  The area to the left of the Bubble Point line is the sub-cooled liquid region.

The area to the right of the Dew Point line is the super-heated gas region.  Between these two lines the mixture is two-phase. Other areas of interest are the retrograde region and the supercritical region. Each of these regions provides advantages and disadvantages for operations.

This month we will start to define the points of interest so that we may choose proper operating points for various types of processes.  The first point to define is the cricondentherm.  The definition of this point is the highest temperature at which two-phases (liquid and vapor for most processes) can coexist.  In the drawing above this is point M.  Point M has considerable theoretical and practical importance.  For example, if the cricondentherm for a sales gas (point M) is 0 ºC (32 ºF) cooling the gas to 4 ºC (40 ºF) at any pressure will not result in condensation of liquids.  This type of operation is typically the type used for cross-country transportation of gas in pipelines.  Operation with this type of system will not require “slug catchers” at the end of the pipeline and will significantly decrease pressure drop in the pipeline.

If the gas were processed in a cold separator such that point B (a dew point) was 0 ºC (32 ºF) problems could occur in the same conditions as the pipeline mentioned above.  If the pressure of the pipeline was between the pressure of point B and E and the pipeline cooled to 4 ºC (40 ºF) there could be significant quantities of liquid in the pipeline.  If the operations people were not familiar with the phase diagram they might increase the operating pressure of the cold separator and still keep the temperature at 0 ºC (32 ºF).  This action would result in increased liquids in the pipeline, not decreased.  However, if the cold separator was operated at the pressure of point M, at a temperature of 0 ºC (32 ºF), in theory there would be no liquids in the pipeline again. (More about the difference between theory and practice in future tips).

If you want more information about how to use phase diagrams to improve profitability of operations or how to generate a phase diagram similar to the calculated drawing below, from a process simulator using equations of state, use our search engine above.  You may look for a course to attend, books or software to buy or other articles on this web site.

Suggested search words: phase diagram, dew point control, equation of state, process simulation, gas conditioning and processing.

By: Dr. Larry L. Lilly