In the post-World war II period, the steels used in the oil and gas industry were quite different from what we use today. This tip of the month (TOTM) presents a brief overview of improvements in the steels used in oil and gas processing equipment for safer and more reliable operations.
Plate was SA-285C a 55,000 psi (379 MPa) tensile steel that was relatively soft and easy to fabricate. It was not killed steel and therefore, not fine grain steel. The low tensile strength meant thicker vessels and because of poor welding techniques, spot or no radiography at all was common, making the items even thicker. Figure 1 shows a vacuum tower made of SA-285C from the 1950’s. This tower was constructed in 1961 by Chicago Bridge and Iron for the Shell Martinez Refinery in California.
Figure 1. A vacuum tower made of SA-285C from the 1950’s. Shell Martinez serial #C-4201
A plate designated SA-212B Firebox was in use for higher tensile applications. It had a 70,000 psi (482 MPa) tensile, but was coarse grained and had the undesirable characteristic of fracturing in the parent metal after thermal expansion and contraction over a period of time. Due to repeated failures in service, this material was removed from the ASME Boiler and Pressure Vessel Code Section II in 1968 as being unfit for thermal cycling. Figure 2 presents a high pressure molecular sieve tower which was fractured by thermal cycling.
Figure 2. An example of fracturing in a vessel made from the SA-212B Firebox steel.
Pipe used in the 1950’s was SA-53B, which could be Electric Resistance Welded or seamless. It was not killed steel. It had a 60,000 psi (413 MPa) tensile and was the pipe of choice for vessel, tank, and piping fabrication at the time.
The forging of the 1950’s was SA-181, a 60,000 psi (413 MPa) tensile steel used for flanges, forged steel fittings, and heavy nozzles. It was not killed steel.
Since none of these steels were killed, fine grained steels, their use declined rapidly as the industry moved into harsh environments such as the North Slope of Alaska and the processing of acid gases and sour crudes.
Killed steel came into wide use during the 1960’s. Killed steel is produced in the ladle by adding silicon or aluminum to prevent further deoxidation of the heat. Molten steel contains dissolved oxygen which can cause bubbles in the cooling and solidification process. The addition of silicon or aluminum stops the reaction of the oxygen with carbon, producing a fine grain steel free from dissolved gases, highly homogenous with excellent fabrication properties.
During the 1960’s, the SA-516 family of plate steels was introduced. These steels were silicon killed, fine grained, and produced excellent properties. The fine grain gave the steel impact resistance at temperatures down to -50 °F (-45.5 °C). The SA-516 suffix defines the tensile strength, 55,000, 60,000, 65,000, and 70,000 psi (379, 413, 448, and 482 MPa).
SA-516-55 was designed to replace SA-285C
SA-516-60 was designed for use in very cold service.
SA-516-65 was for intermediate tensile requirements
SA-516-70 was to replace SA-212B Firebox plate
The chemical and mechanical properties of these four grades of steel overlap to the extent that one plate can actually meet all four specifications.
Approximately 90% of all custom carbon steel pressure vessels manufactured for the oil and gas industry in the world today are made from SA-516-70 or its UNS (Unified Numbering System) equivalent. Figure 3 presents an example of a vertical drum made of SA-516-70.
Figure 3. An example of a vertical drum made of SA-516-70
During the 1960’s SA-106 pipe replaced SA-53 as the pipe of choice. Unlike SA-53B, SA-106B is seamless, killed, fine grain steel. It has a 60,000 psi (413 MPa) psi tensile.
In 1978, SA-105 forgings replaced the SA-181 as the forging material of choice. SA-105 has a tensile of 70,000 psi (482 MPa), so the pressure ratings of B16.5 carbon steel flanges increased.
Around the year 2000, the pipe manufactures improved their processes making SA-106 pipe to the point that they are able to meet the chemical and mechanical properties of SA-106B and SA-106C in the same heat.
Since 2003, basically all SA-106 pipe is dual certified to SA-106B and SA-106C. This means that all three major components of a pressure vessel or shell and tube heat exchanger now have the same tensile strength, 70,000 psi (482 MPa). Figure 4 presents pipes made of SA-106.
Figure 4. Pipes made of SA-106
Austenitic Stainless steels (300 series) fifty years ago were made to straight grade (0.08 carbon) or “L” grade (0.03 carbon). Steel service centers had to maintain stocks of both grades. About 45 years ago, the stainless mills improved their manufacturing techniques to produce dual certified stainless steel, meaning that virtually all stainless in the steel service centers meets the criteria of 0.03 carbon for “L” grade but also meets the mechanical properties of straight grade. Straight grades have a higher tensile allowing for the use of a thinner plate than “L” grade plate.
Figure 5 presents an example of a separator made of stainless vessel. This 316 stainless separator is the first to be used offshore in place of a clad vessel. Since the temperature was low, the higher tensile allowed this item to be thinner, saving weight, and not require PWHT (Post Weld Heat Treatment), impact testing or special paint.
Figure 5. This 316 stainless separator is the first to be used offshore in place of a clad vessel
Summary:
In the last half century, the adoption of new technology in the manufacturing of fine grain steel plates, pipes and forgings has vastly improved the quality of the steels used in Oil and Gas Processing Equipment. Along with improvements in the welding processes used to construct Oil and Gas Processing Equipment, vessels, exchangers, piping and storage tanks are safer than ever before.
In this Tip of the Month (TOTM) we will discuss how to determine CO2 solubility and flashing issues in water at pressures and temperatures commonly associated with gathering systems and production facilities. This is mainly important for CO2 Enhanced Oil Recovery (EOR) floods as the CO2 concentration is high and the initial separation is at higher pressures than is common in non-CO2 EOR oilfields. These two conditions result in significant dissolving of CO2 into the produced water with resultant flash gas from downstream Free Water Knockouts (FWKO), treaters, and tanks. In mature CO2 EOR floods with Water-Alternating-Gas (WAG) injection schemes, it is likely that most of the flash gas in the downstream separations will be from the produced water.
While this TOTM is significant mainly for CO2 EOR floods, the following analysis is general in nature; it would apply to other situations involving CO2 solubility in water issues. This analysis assumes that there is no H2S. H2S would have somewhat higher solubility than CO2 which would force more gas to flash from the FWKO and tanks. Higher H2S than about 5% would begin to appreciably increase the solubility of H2S into water.
A hypothetical field production example will demonstrate how to calculate the CO2 solubility and flash gas volumes. The field configuration is shown in Figure 1.
A common configuration for a CO2 EOR flood is to have the field gathering system consist of three flashes. The first flash is at the Field Separators, the second flash is at the FWKO, and the third flash is at the tanks. A number of wells are gathered into Field Separators for the first flash of gas from the liquids. For the purposes of this case, all separators are assumed to operate at a single pressure and temperature. The amount of gas from the Field Separators is not able to be calculated from the analysis presented in this TOTM. Instead, the gas from this separation is the result of the complex flows within the reservoir through the downhole equipment to the separator. This analysis is applicable to the next two flashes that result from the liquid flowing to the FWKO and then to the Water Tank. Also note that the gas from the oil (which includes both hydrocarbon gas and CO2 and also will be a significant volume) is beyond the scope of this analysis.
Figure 1. A simplified production flow diagram
With the above figure and discussion in mind, the problem statement is as follows:
A CO2 EOR flood is in operation. The Field Separators operate 15.6 °C (60 °F) and 1483 kPag (215 psig) and the gathering system gas is 87.5% CO2. The liquids flow from the Field Separators to a central tank battery where the unheated FWKO is operating 15.6 °C ( 60 °F) and 207 kPag (30 psig) and the gas from it is 89.7% CO2. The tanks operate at 3.45 kPag (0.5 psig.) The atmospheric pressure is 93.1 kPaa (13.5 psia). The water from the field is 3180 STm3/d (20,000 bbl/day) and has 4 wt% associated salts.
Determine the amount of gas flashed from the produced water at the FWKO and water tanks.
The analysis begins by determining the CO2 in solution within the Field Separators. The solution gas is determined from Figure 2 below.
Figure 2 (SI). Solubility of CO2 in fresh water as a function of its partial pressureFigure 2 (FPS). Solubility of CO2 in fresh water as a function of its partial pressure
Step 1 – Find the amount of CO2 in solution in the water in the Field Separator from Figure 1. Determine the CO2 partial pressure in the gas phase of the Field Separator which is operating at 1380 kPaa (200 psia). Then read the gas saturation at the separation temperature which is 15.6 °C (60 °F). The solution gas is determined to be 11.23 Sm3/STm3 (63 scf/bbl). The key is to measure the gas composition at the outlet of the Field Separator.
Note that the curves for 0 °C (32 °F) stops at 2069 kPaa (300 psia), 4.4 °C (40 °F) stops at 2759 kPaa (400 psia), and 15.6 °C (60 °F) stops at 4828 kPaa (700 psia) CO2 partial pressure. That is because the CO2 becomes liquid as it increases much past these pressures. No design should contemplate separation where the CO2 is still liquid. Its density would be too close to oil so separation would not be feasible. The curves themselves will not extrapolate successfully from these endpoints; the character of the curve changes as the CO2 moves into the liquid phase.
It is not easy to read the saturation at the FWKO and even more difficult to read the tanks so a close-up of the graph has been prepared as Figure 3. The “X” axis is now zero to 690 kPaa (100 psia) rather than zero to 5517 kPaa (800 psig). The “Y” axis is now zero to 8.92 Sm3/STm3 (50 scf/bbl) rather than zero to 35.66 Sm3/STm3 (200 scf/bbl). The various curves now are reasonably approximated as straight lines which intercept the “Y” axis at zero. The equations are included.
Figure 3 (SI). Close-up of solubility of CO2 in fresh water as a function of its partial pressureFigure 3 (FPS). Close-up of solubility of CO2 in fresh water as a function of pressure
Step 2 – The CO2 solubility of the water in the FWKO is determined in the same fashion as for the Field Separator. The composition of the gas is the gas exiting the FWKO is the composition to use to determine the CO2 partial pressure. In the FWKO 2.23 Sm3/STm3 (12.5 scf/bbl) remains in solution.
Step 3 – This shows the gas in solution in the storage tanks. An assumption is made that the gas is liberated from the tankage that the tanks are 100% CO2. It will be slightly less than 100% but should be close enough for design purposes. This assumption will slightly overstate the amount of CO2 remaining in solution. Observe that in the storage tanks the water still has 0.8 Sm3/STm3 (4.5 scf/bbl) of CO2. This is enough CO2 to make the water substantially more corrosive than water from non-CO2 EOR floods.
The determination of CO2 solubility has thus far been made for fresh water but the produced water contains salts (Na, Mg, Ca). An adjustment is to be made for each of the three saturations based on the 15.6 °C (60 °F) temperature and the salts percentage by weight of the produced water. This is accomplished with Figure 4 which calculates the appropriate reduction in solubility.
Figure 4. Brine Correction for CO2 Solubility in Water
Step 4 – Make the adjustments for salt in the water by applying the Salinity Reduction Factor (SRF). The temperature impact is small enough that only three curves are shown. It would be too difficult to read the graph if all curves were shown. To read this point it may be easier to use the corresponding Table 1 which contains more temperature data.
Table 1. Salt Reduction Factor (SRF) [1]
The data obtained from the graphs and data from the problem statement are summarized in Table 2. With some calculations the appropriate gas volumes can be determined.
Table 2. Example of CO2 Flashing for 3180 STm3/d (20,000 bbl/day) of Water, 4% Salts
Step 5 – Column (D) – Input the data from Figure 2 and Figure 3 into the appropriate row
Step 6 – Column (C) – Input the SRF from Figure 4 into the last row
Step 7 – Column (E) – Subtract 2.23 Sm3/STm3 (12.54 scf/bbl) from 11.23 Sm3/STm3 (63 scf/bbl) = 9 Sm3/STm3 (50.46 scf/bbl) to calculate the amount of gas that evolves from solution into the gas phase from the Field Separators to the FWKO.
Step 8 – Column (E) – Subtract 0.8 Sm3/STm3 (4.5 scf/bbl) from 2.23 Sm3/STm3 (12.54 scf/bbl) = 1.43 Sm3/STm3 (8.04 scf/bbl) to calculate the amount of gas that evolves from solution into the gas phase from the FWKO to the tankage.
Step 9 – Column (F) – Multiply 9.00 Sm3/STm3 (50.46 scf/bbl) times 3180 STm3/d (20,000 bbl/d) = 28.58×103 Sm3/d (1009 MSCFD) to calculate the amount of gas to compress from the FWKO
Step 10 – Column (F) – Multiply 1.43 Sm3/STm3 (8.04 scf/bbl) times 3180 3180 STm3/d (20,000 bbl/d) = 4.56×103 Sm3/d (161 MSCFD) to calculate the amount of gas to compress from the tankage
Step 11 – Column (G) – Multiply the quantities of Column (F) times the SRF to determine the gas as adjusted for salinity. The total gas is 28.29×103 Sm3/d (999 MSCFD).
Supplemental Data:
A 2nd order polynomial for each of the curves in Figure/s 2 has been prepared. These equations will allow the user to determine the solution gas by calculation rather than by reading the graphs directly.
Figure 2 can be presented by Equation 1 [2].
SCO2 = A (PCO2)2 + B (PCO2) (1)
Where:
Conclusions:
The combination of relatively high Field Separator pressures and high CO2 content in gas from CO2 EOR floods results in substantial flashing of CO2 from water in the FWKO and tanks. This has the potential of overwhelming the piping, Vapor Recovery Units, and Flash Gas compressors.
To learn more about similar cases and how to determine the impact of CO2 on field operational problems, we suggest attending our PF81 (CO2 Surface Facilities) course.
John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at http://petroskills.com/consulting, or email us at consulting@PetroSkills.com.
By: Paul Carmody
Reference:
Chang, Y.-B., Coats, B. K., & Nolen, J. S. (1998, April 1). A Compositional Model for CO2 Floods Including CO2 Solubility in Water. Society of Petroleum Engineers. doi:10.2118/35164-PA.
VMG Sim v8.0 (Build 85) August 2014, Thermo = “APR for Natural Gas”.
One of the most common problems in Oil and Gas Processing facilities is underperforming vapor / liquid separators. The most common types of gas-liquid separators are:
Slug catchers
Vessel / Finger-type
“Conventional” separators
Vertical / Horizontal
Scrubbers (i.e. Compressor Suction Scrubbers)
Gas “polishers”
Coalescing Filters / Filter Separators
Underperforming separators generally result from either: 1. the wrong type of equipment was selected for the application, or 2. the correct type of equipment was selected, but the sizing methodology was inadequate. The type of separator required for an application depends largely on the gas-liquid ratio of the stream to be treated, and the flow variability of the process, as shown in Figure 1. As the flow variability increases with low to moderate gas-liquid ratios, the separator selection will move from a conventional separator to a slug catcher. For applications where there is a high gas-liquid ratio (i.e. very low liquid content), and the flow variability is moderately low, scrubbers and gas polishers would be the appropriate equipment selection depending upon the gas quality requirement for the treated stream.
Figure 1. Gas-Liquid Separation Equipment Selection Map [1]Figure 2. Troubleshooting Methodology [1]Unfortunately, once the equipment has been selected and installed, it is very costly to replace if the separator was not specified properly. Common separator performance issues are: too much liquid in the separated gas, inadequate slug/surge capacity, and too much gas in the separated liquid.
This paper is focused on troubleshooting inadequate liquid removal from the gas for conventional separators (moderate to high liquid loads) and scrubbers (very low liquid loads).
A troubleshooting methodology is provided in Figure 2 [Reference No]. The problem, in this case, is too much liquid in the separated gas stream. In order to effectively troubleshoot separator performance, it is required to understand the metrics of good performance, and the functions and analysis of the various components of the separation equipment.
Typical performance metrics for separators are provided in Table 1. The specific performance requirements for a given separator are set by the sensitivity of the downstream process / equipment to the presence of liquids. For example, the product gas (sales gas) off of the cold separator in an NGL Extraction facility is sensitive to the presence of entrained liquids. The product gas can go off specification if there is too much carryover of liquids from the cold separator. On the other hand, the sensitivity of the downstream equipment from the facility inlet separator is much less, and the amount of liquids entrained in the gas is more tolerable.Table 1. Example separator performance metrics [1]
Separator Components
The main components of a separator, shown in Figure 3, are the feed pipe, inlet device, gas gravity separation section, mist extractor and the liquid gravity separation section. The gas/liquid separator components will be briefly discussed in regards to their effects on gas/liquid separation performance. These effects need to be understood and quantified in order to troubleshoot separator operations, and to identify modifications that can be made to improve performance. The liquid gravity separation section will not be discussed.
Figure 3. Parts of a Conventional Separator [2]Inlet Feed Pipe
The inlet feed pipe sizing and geometry is important as it is desired to keep the multiphase flow pattern “stabilized” in the piping to minimize the production of small liquid droplets, and liquid entrainment into the gas phase. Figure 4 [2] shows the effect of feed pipe velocity on liquid entrainment. Figure 5 [2] demonstrates how quickly the liquid entrainment increases once the entrainment inception point is reached.
Figure 4. Effect of feed pipe velocity on liquid entrainment [2]Figure 5. Example of liquid entrainment behavior in a gas-liquid system [2]Some general guidelines for inlet piping to minimize liquid entrainment are:
Provide 10 diameters of straight pipe upstream of the inlet nozzle without valves, expansions/contractions or elbows.
If a valve is required, only use full port gate or ball valves.
Inlet Device
The main purpose of an inlet device is to improve separation performance. This is achieved by maximizing the amount of gas-liquid separation occurring in the feed pipe, minimizing droplet shearing, and optimizing the downstream velocity distributions of the separated phases into the separator. Schematics for inlet devices are shown in Figure 6. In large capacity, more critical separation applications, the vane-type and cyclonic inlet devices are commonly used. The simpler, and less expensive, impact (or diverter plates) are often used where the separation performance is less critical.
Figure 6. Various separation equipment inlet devices [2]Table 2 provides a comparison of the performance of various inlet devices.
Table 2. Comparison of inlet devices [2]
The inlet momentum (ρmV2m – density*velocity2 of the mixture) of the feed stream is typically used to select and size inlet devices. Table 3 provides the suggested upper limits of inlet momentum values. For conditions where it is not practical to avoid higher feed pipe velocities, it must be recognized that failure to do so will result in higher entrainment loads, smaller droplet sizes, and more difficult separation conditions.
Table 3. Inlet device ρV2 upper limits [3]
The quality of the flow distribution downstream of the inlet device is critical to mist extractor performance. Historically, tracer surveys have been used to provide an approximate indication of the continuous phase velocity within separators. In more recent years, the use of CFD (Computational Fluid Dynamics) has provided insight into the flow behavior of fluids, and has resulted in significant improvement in separator internals design. Separator performance is to a large degree dependent on the removal of droplets/ bubbles from the continuous phase. The efficiency of this removal is a function of the continuous phase velocity, thus the importance of understanding the factors that affect velocity profiles. Figure 7 provides an example of ideal versus actual gas velocity profiles within a separator.
Figure 7. Ideal and actual gas velocity profiles [3]Gas Gravity Separation Section
The gas gravity separation section of a separator has two main functions: 1) reduction of entrained liquid load not removed by the inlet device, 2) improvement / straightening of the gas velocity profile.
Most mist extractors have limitations on the amount of entrained liquid droplets that can be efficiently removed from the gas, thus the importance of the gas gravity section to remove the liquids to an acceptable level upstream of the mist extractor. This is particularly important for separators handling higher liquid loads. For scrubber applications with low liquid loadings, the Ks value will be primarily dependent on the mist extractor type, and the gas gravity separation section becomes less important.
For the higher liquid load applications, there are two approaches for sizing the gravity separation section to remove liquid droplets from the gas: 1) Ks method, 2) Droplet settling theory.
Historically the Ks method has been employed as it can provide reasonable results and is easy to use, but has shortcomings in terms of quantifying separator performance. References 3-5 provide the details on the droplet settling theory methods which can be used to more accurately quantify separator performance. The Ks approach is limited in that it only informs of the average droplet size, but cannot quantify the amount of liquid droplets exiting the gas gravity section.
The Ks method (Equation 1) is an empirical approach to estimate the maximum allowable gas velocity to achieve a desired droplet separation.
Where:
ρL = liquid density kg/m3 (lbm/ft3)
ρg = gas density kg/m3 (lbm/ft3)
Vgmax = maximum allowable gas velocity m/s (ft/sec)
KS = an empirical constant m/s (ft/sec)
Figure 8 provides the relationship of Ks values for various droplet sizes and separator operating pressures for the gas gravity section. Typically, a Ks value is selected that will achieve removal of all entrained droplets larger than a specified target droplet diameter in the original design of the separator. For conventional separators the target droplet diameter is typically 150 microns, and for scrubbers the target droplet size should not exceed ~500 microns. This correlation can also be used to determine the performance of the gas gravity section based upon current operating conditions. The separator Ks value can be estimated from the actual velocity and fluid conditions, and the droplet size removed in the gas gravity section can be estimated from Figure 8.
Figure 8. Ks vs. pressure and droplet size for empty vessels [2]Mist Extraction Section
The mist extractor is the final gas cleaning device in a conventional separator. The selection, and design to a large degree, determine the amount of liquid carryover remaining in the gas phase. The most common types include wire mesh pads (“mesh pads”), vane-type (vane “packs”) and axial flow demisting cyclones. Figure 9 shows the location and function of a typical mist extractor in a vertical separator.
Figure 9. Typical mist extractor in a vertical separator [2]Mist extractor capacity is defined by the gas velocity at which re-entrainment of the liquid collected in the device becomes appreciable. This is typically characterized by a Ks value, as shown in Equation 1. Mesh pads are the most common type of mist extractors used in vertical separator applications. The primary separation mechanism is liquid impingement onto the wires, followed by coalescence into droplets large enough to disengage from the mesh pad. Figure 10 provides some mesh pad examples. Table 4 provides a summary of mesh pad characteristics and performance parameters.
Figure 10. Mesh pad examples [1]
Table 4. Mesh pads summary of characteristics and performance parameters [1,4]
Notes:
Flow direction is vertical (upflow).
Assume mesh pad Ks values decline with pressure as shown in Table 5. Table 5 was originally developed for mesh pads, but is used as an approximation for other mist extractor types. [6].
If liquid loads reaching the mesh pad exceed the values given in Table 4, assume capacity (Ks) decreases by 10% per 42 L/min/m2 (10% per gal/min/ft2). [3-5].
These parameters are approximate.
Table 5. Mesh pad Ks deration factors as a function of pressure [2]
Vane packs, like mesh pads, capture droplets primarily by inertial impaction. The vane bend angles force the gas to change direction while the higher density liquid droplets tend to travel in a straight-line path, and impact the surface of the vane where they are collected and removed from the gas flow. Figure 11 shows a schematic of a single-pocket vane mist extractor. Table 6 provides vane pack performance characteristics.
Figure 11. Single-pocket vane schematic [2]
Table 6. Typical vane-pack characteristics [1,4]
Notes:
Assume vane-pack Ks values decline with pressure as shown in Table 5.
If liquid loads reaching the vane pack exceed the values given in Table 4, assume capacity Ks decreases by 10% per 42 L/min/m2 (10% per gal/min/ft2). [3-5].
These parameters are approximate only. The vane-pack manufacturer should be contacted for specific information.
In the case of demisting cyclones, the vendor should be consulted in regards to performance for the current operations of interest.
Troubleshooting
When troubleshooting a separator, one needs to quantify the acceptable performance of the separator in terms of the amount of liquids in the separated gas. The separator physical condition and design is then assessed to determine the liquid removal capability of the separation equipment installed. Each separator component should be analyzed in terms of the current operating conditions versus the original design specifications.
Table 7 provides a few common causes of inadequate liquid removal performance of a separator. The separator components that need to be reviewed are identified to determine the specific limitation. This table can serve as a road map for the calculations to work through when doing this type of analysis.
Table 7. Common conditions resulting in inadequate separated gas quality [1]
There are numerous options available to improve the performance of a separator depending upon what the cause for the poor performance is. Depending upon the size and construction of the separator, it may be possible to retro-fit the separator internals. Another option may be modification of the inlet feed piping geometry and number of fittings upstream of the vessel if this is found to be less than ideal. The inlet device may be damaged, or in the bottom of the vessel. Higher efficiency inlet devices may be an option for consideration. Frequently, different mist extraction equipment can be selected to improve performance. For example, if the mist extractor Ks value is greater than the original design, a different mist extraction device could improve performance. The separator internals modifications may or may not be possible without welding on the vessel (which adds additional complications and cost to the project).
The operating liquid levels should also be reviewed in terms of the distance of the normal operating liquid level in relation to the inlet feed device. If the liquid level is too high, the gas velocity from the inlet could be re-entraining liquids that were previously separated in the feed piping / inlet device. Unfortunately, in some cases the only way to improve performance is to cut rate (i.e. gas velocity) through the separator.
Campbell, J.M., Gas Conditioning and Processing, Volume 2: The Equipment Modules, 9th Edition, 2nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
Bothamley, M. 2013. Gas-Liquid Separators – Quantifying Separation Performance Part 1. SPE Oil and Gas Facilities, Aug. (22 – 29).
Bothamley, M. 2013. Gas-Liquid Separators – Quantifying Separation Performance Part 2. SPE Oil and Gas Facilities, Oct. (35 – 47)
Bothamley, M. 2013. Gas-Liquid Separators – Quantifying Separation Performance Part 2. SPE Oil and Gas Facilities, Dec. (34 – 47)
Fabian, P., Cusack, R., Hennessey, P., Neuman, M. 1993. Demystifying the Selection of Mist Eliminators, Part 1: The Basics.Chem Eng 11 (11): 148 – 156.
Unlike the natural gas compressors which rely upon recycling the after-cooled compressed gas to compressor suction for surge prevention, the surge control in refrigeration systems is quite different because the refrigerant vapors are compressed only to their saturation pressure for dew point temperature, which when cooled for recycling immediately condense into a liquid. Consequently, the hot compressed refrigerant vapors, also known as the hot gas, before any cooling are recycled to compressor suction to make up for any loss of the process chilling load that reduces the amount of available refrigerant vapors to satisfy the minimum flow to the compressor suction, thus avoiding the phenomena known as “surge”.
Three-Stage Refrigeration Cycle
A three-stage propane refrigeration system [1, 2] is adapted into Figure 1 to illustrate a typical refrigeration cycle.
”]As shown in Figure 1, all three process chillers are located on the 1st stage and there is no process side chilling load at the two economizer levels. The shown schematic is a simplest illustration of a three-stage propane refrigeration cycle [1, 2].
The compressed refrigerant vapors are condensed with ambient air. The high pressure (D-0116) and low-pressure (D-0115) economizers separate the flashed vapors that enter the third and second stages of the refrigerant compressor, respectively. To protect the compressor from any liquid carry over, each stage has a suction drum, and all vapors entering the compressor must pass through these drums (D-0112, D-0113 and D-0114).
Surge Protection with Hot Gas and Quench Liquid
So how do you provide a cool vapor stream to prevent the refrigerant compressor from surging in order to keep the discharge temperature from getting too high? To address this challenge, hot compressed vapors from refrigerant compressor discharge are used to vaporize liquid refrigerant, in a similar manner to how the process stream is used to vaporize refrigerant in the chillers.
In the refrigeration cycle of Figure 1, the hot gas from compressor discharge flows into each suction drum as determined by the anti-surge controller to prevent each compression stage from surging. To cool the hot recycled vapors, a temperature controller introduces quench liquid directly from the refrigerant accumulator D‑0111 into the inlet of each compressor suction drum. Ideally, if the compressor anti-surge vapor flow and temperature control valve on quench liquid are perfectly matched and there is no time lag within the control system, the configuration of Figure 1 will work perfectly.
Unfortunately, if there is any mismatch, which in a real world situation always exists, either too much quench liquid is added to the inlet stream whereby there is liquid level in the suction drums or not enough quench liquid is added whereby the discharge temperature keeps rising. In either mismatch case, the refrigerant compressor will shut down either from high liquid level in the suction drum or from high discharge temperature caused by the high inlet temperature due to recycle of hot gas.
Surge Protection with Sparged Hot Gas using No Quench Liquid
A well-proven alternative system is shown in Figure 2 [3, 4] in which hot propane vapors from the compressor discharge, as controlled by the anti-surge system, flow through the sparged line submerged in the liquid filled portions of the HP (D-0116) and LP (D-0115) economizers, and under the tube bundle of the first stage chillers (E-0107, E-0108, and E-0111). There is no need for quench liquid as the hot vapors provide the heat to the available liquid propane in these services to generate the required propane vapors to prevent the compressor surge conditions.
”]The configuration of Figure 2 builds upon the recognition that:
The shortage of propane refrigerant entering the compressor stages is caused by reduction in heat exchanger duty from changes on the process side of the chillers (E-0107, E-0108 or E-0111). This reduction of heat transfer duty is offset by routing the required amount of hot propane vapor from compressor discharge directly to the kettle side of the chiller to supplement process thermal requirements.
The compressor suction drums are provided to prevent any carried-over liquid from entering the compressor and should essentially have no liquid in them.
By diverting the anti-surge hot vapor away from the compressor suction drums assures that the compressor suction drums stay dry.
By using the already present liquid refrigerant in the chillers and economizers it avoids any potential mismatch between the hot propane vapors and the quench liquid that may inadvertently buildup high liquid level and end up tripping the refrigerant compressor.
Conclusion
While both refrigerant compressor surge control systems are proven and used by the industry, the hot sparged gas with no quench liquid configuration of Figure 2 is more forgiving and reliable.
John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.
By: Yuv R. Mehra
References:
Al-Shahrani, Saleh M. and Mehra, Yuv R., “Start-up Experience for Value Recovery from CCR Net Gas at Yanbu’ Refinery,” Advances in Hydroprocessing II Session, Paper 26d, 2007 Spring National Meeting of AIChE, Houston, Texas, USA, April 22-26, 2007.
Al-Shahrani, Saleh M. and Mehra, Yuv R., “Saudi Aramco installs new LPG recovery unit at Yanbu’ Refinery,”Oil & Gas Journal, Vol. 105.21, Jun. 4, 2007, p. 60.
Mehra, Yuv R., “Focus on Value-Chain Contributions during Process Technology Selection,” Worldwide Developments Forum,86th Annual Meeting of the Gas Processors Association, San Antonio, Texas, USA, March 11-14, 2007.
Mehra, Yuv R., “Focus on Value-Chain Contributions during Process Technology Selection,” Saudi Aramco Journal of Technology, Summer 2007, p. 2.
Stainless steel is a family of corrosion resistant steels containing chromium in which the chromium forms a passive film of chromium oxide (Cr2O3) when exposed to oxygen [1]. This phenomenon is called passivation and is seen in other metals, such as aluminium and titanium.
The film layer is impervious to water and air and quickly reforms when the surface is scratched. This protects the metal beneath – preventing further surface corrosion. Since the layer only forms in the presence of oxygen, corrosion-resistance can be adversely affected if the component is used in a non-oxygenated environment e.g. underwater bolts on a platform support structure.
Such passivation only occurs if the proportion of chromium is high enough and is normally achieved with addition of at least 13% (by weight) chromium. Progressively higher levels of corrosion resistance and strength is achieved by the addition of other alloying elements – each offering specific attributes in respect of strength and corrosion resistance.
Classification issues
The need to classify stainless steel has led to a fundamental problem of which method to use. Probably the best known system derives from of the Society of Automobile Engineers (SAE) e.g. 316 Cr/Ni/Mo 17/12/2. This is interpreted as stainless steel containing the proportions of 17% chromium, 12% nickel, and 2% molybdenum.
However, the waters are somewhat muddied by a variety of international and country-based systems that include EN (European Norm); and UNS (Unified Numbering System). For example, SAE 304 Cr/Ni 18/10 stainless steel is EN 1.4301 which is UNS S30400.
Stainless steels may also be graded into five basic families or phases determined by their crystalline structure: the stable phases austenitic or ferritic; a duplex mix of the two; the martensitic phase created when some steels are quenched from a high temperature; and precipitation-hardenable.
Ferritic stainless steel
In ferritic stainless steel, the iron and chromium atoms are arranged in what is termed a body-centred cubic structure in which the atoms are arranged on the corners of the cube and one in the centre (Figure 1). As well as being ferro-magnetic, ferritic stainless steel exhibits very high stress corrosion cracking resistance.
Ferritic stainless steels are plain chromium (10.5 to 18%) grades characterized by moderate corrosion resistance and poor fabrication properties. These characteristics may be improved with the addition of molybdenum; some, aluminium or titanium.
Austenitic stainless steel
With the addition of nickel, the properties change dramatically. As shown (Figure 3) the atoms are re-arranged so that they occur on the corners of the cube and also in the centre of each of the faces. In this manner it becomes what is termed austenitic stainless steel.
It can thus be seen from Table 1 that unless you are specifically looking for a ferro-magnetic material, austenitic stainless steel would be the most obvious choice. Indeed this is borne out by the fact that austenitic stainless steels account for about 70% or more of all stainless steel used worldwide – with ferritic stainless steels making up about 25%. The other families each represent less than 1% of the total market.
Austenitic stainless steels are designated by numbers in the 200 and 300 series.
Series 300
The relationship between the 300 austenitic grades is shown in Figures 4.
The basic grade 304 contains about 18% chromium and 8% nickel (often referred to as 18/8) and range through to the high alloy or ‘super austenitics’ such as 904L and 6% molybdenum grades.
Additional elements can be added such as molybdenum, titanium or copper, to modify or improve their properties, making them suitable for many critical applications involving high temperature as well as corrosion resistance. This group of steels is also suitable for cryogenic applications because the effect of the nickel content in making the steel austenitic avoids the problems of brittleness at low temperatures, which is a characteristic of other types of steel.
Generally, the 300 grade alloys are subject to crevice and pitting corrosion.
Low-carbon versions, (indicated by the letter suffix L) include 304L, 316L and 317L, in which the carbon content of the alloy is below 0.03%. This reduces the effect of ‘sensitisation’ in which chromium carbides precipitate at the grain boundaries due to the high temperatures involved in welding. The relatively high nickel content also inhibits the brittleness exhibited by ferritic materials at low temperatures and thus makes austenitic steels suitable for cryogenic applications.
200 series
We have seen earlier how the addition of nickel is used in the creation of the classic chrome-nickel 300 series austenitic stainless steel.
The reduced nickel content of the 200 series chrome-manganese grades makes them significantly cheaper. However, depending on their chemistry, they also offer good formability (ductility) and/or strength. Indeed, certain grades (201, 202 and 205 series) even offer about 30% higher yield strength than the classic 304-series chrome-nickel grade – allowing designers to cut weight (Table C.2).
Reducing nickel, on the other hand, reduces the maximum chromium content possible in the alloy. Less chromium means less corrosion resistance and a consequent narrowing of the range of applications for which the material is suitable.
A word of warning comes from the International Stainless Steel Forum (ISSF). Continuous pressure to cut costs, especially from the Asian market, has resulted in the development of austenitic grades ever lower in nickel and chromium, often not covered by international codes or specifications. In fact, numerous chrome-manganese grades are company-specific and identified simply by a title given to them by the producer.
Duplex stainless steels
Duplex stainless steels [6] are a mixture of austenite and ferrite microstructures that combine some of the features of each class:
resistance to stress corrosion cracking – but inferior to ferritic steel;
superior toughness to ferritic steel – but inferior to austenitic steel;
roughly twice the strength of austenitic steel;
superior resistance to pitting, crevice corrosion and stress corrosion cracking;
high resistance to chloride ions attack; and
high weldability.
These features are achieved by adding less nickel than would be necessary for making a fully austenitic stainless steel. Thus, Grade 2304 comprises 23% chromium and 4% nickel whilst Grade 2205 comprises 22% chromium and 5% nickel – with both grades containing further minor alloying additions.
On the negative side, austenitic-ferritic duplex stainless steels are only usable between temperature limits of about -50°C and 300°C – outside which they suffer reduced toughness.
Martensitic stainless steel
Named after the German metallurgist, Adolf Martens, the martensitic Grade 400 series (Figure 5) are low carbon (0.1–1%), low nickel (less than 2%) steels containing chromium (12 to 14%) and molybdenum (0.2–1%).
Stainless steels hardened by transformation to martensite are tempered to give the desired engineering properties. At high temperatures they have an austenitic structure that is transformed into martensitic structure upon cooling to room temperature. Unfortunately, this tempering can influence corrosion susceptibility. For example, corrosion susceptibility of type 420 stainless steel is at its maximum when the alloy is tempered at temperatures in the range of 450° to 600°C. So, aalthough not as corrosion-resistant as the 200 and 300 classes, martensitic stainless steels are magnetic, extremely strong (if not a little brittle), highly machinable, and can be hardened by heat treatment.
Martensitic stainless steels are subject to both uniform and non-uniform attack in seawater. And the incubation time for non-uniform attack in even weak chlorides is often only a few hours or days.
These chromium- and nickel-containing steels can be precipitation hardened to develop very high tensile strengths. Precipitation-hardening stainless steels are usually designated by a trade name rather than by their AISI 600 series designations.
The most common grade in this group is ‘17-4 PH’, also known as Grade 630, with a composition of 17% chromium, 4% nickel, 4% copper and 0.3% niobium. The main advantage of these steels is that they can be supplied in the ‘solution treated’ condition – in which state the steel is just machinable. Following machining, forming, etc. the steel can be hardened by a single, fairly low temperature ‘ageing’ heat treatment that causes no distortion of the component.
Precipitation hardening generally results in a slight increase in corrosion susceptibility and an increased susceptibility to hydrogen embrittlement.
By: Mick Crabtree
References
1. T. Sourmail and H. K. D. H. Bhadeshia, ‘Stainless Steels’, University of Cambridge.
3. ‘Stainless Steel and Corrosion’, ArcelorMittal, Stainless Europe.
4. A.U. Malik, M. Kutty, Nadeem Ahmad Siddiqi, Ismaeel N. Andijani, and Shahreer Ahmad, ‘Corrosion Studies on SS 316 L in low pH high Chloride product water medium’, 1990.
5. ‘The Stainless Steel Family’, International Stainless Steel Forum.
6. API Technical Report 938-C: ‘Use of Duplex Stainless Steels in the Oil Refining Industry’.
![Figure 1. Gas-Liquid Separation Equipment Selection Map [1]](https://i0.wp.com/www.jmcampbell.com/tip-of-the-month/wp-content/uploads/2014/11/fig11.png?resize=554%2C339)
![Figure 2. Troubleshooting Methodology [1]](https://i0.wp.com/www.jmcampbell.com/tip-of-the-month/wp-content/uploads/2014/11/fig2.png?resize=321%2C526)
![Figure 3. Parts of a Conventional Separator [2]](https://i0.wp.com/www.jmcampbell.com/tip-of-the-month/wp-content/uploads/2014/11/fig31.png?resize=604%2C276)
![Figure 4. Effect of feed pipe velocity on liquid entrainment [2]](https://i0.wp.com/www.jmcampbell.com/tip-of-the-month/wp-content/uploads/2014/11/fig41.png?resize=455%2C346)
![Figure 5. Example of liquid entrainment behavior in a gas-liquid system [2]](https://i0.wp.com/www.jmcampbell.com/tip-of-the-month/wp-content/uploads/2014/11/fig51.png?resize=384%2C280)
![Figure 6. Various separation equipment inlet devices [2]](https://i0.wp.com/www.jmcampbell.com/tip-of-the-month/wp-content/uploads/2014/11/fig61.png?resize=593%2C371)
![Figure 8. Ks vs. pressure and droplet size for empty vessels [2]](https://i0.wp.com/www.jmcampbell.com/tip-of-the-month/wp-content/uploads/2014/12/fig8.png?resize=618%2C481)
![Figure 9. Typical mist extractor in a vertical separator [2]](https://i0.wp.com/www.jmcampbell.com/tip-of-the-month/wp-content/uploads/2014/12/fig9.png?resize=447%2C451)
![Figure 10. Mesh pad examples [1]](https://i0.wp.com/www.jmcampbell.com/tip-of-the-month/wp-content/uploads/2014/12/fig10.png?resize=599%2C228)
![Figure 11. Single-pocket vane schematic [2]](https://i0.wp.com/www.jmcampbell.com/tip-of-the-month/wp-content/uploads/2014/12/fig11.png?resize=488%2C201)
As shown in Figure 1, all three process chillers are located on the 1st stage and there is no process side chilling load at the two economizer levels. The shown schematic is a simplest illustration of a three-stage propane refrigeration cycle [1, 2].
The configuration of Figure 2 builds upon the recognition that:
