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  1. Re: Fundamentals of Oil and Gas Processing Book "Full text"

    Crude Oil Stabilization and Sweetening - Chapter 6

    Fundamentals of Oil and Gas Processing Book
    Basics of Gas Field Processing Book
    Prediction and Inhibition of Gas Hydrates Book
    Basics of Corrosion in Oil and Gas Industry Book

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    Chapter 6 193
    Crude Oil Stabilization and Sweetening 193
    6.1: Introduction 193
    6-1-1: Crude oil treatment steps 193
    6.2: Process Schemes 194
    6.2.1: Multi-Stage Separation 194
    6.2.2: Oil Heater-Treaters 194
    6.2.3: Liquid Hydrocarbon Stabilizer 195
    6.2.4: Cold-Feed Stabilizer 197
    6.2.5: Stabilizer with Reflux 197
    6.3: Stabilization Equipment 199
    6.3.1: Stabilizer Tower 199
    6.4: Stabilizer Design 205
    6.5: Crude Oil Sweetening 206
    6.6.1: Stage vaporization with stripping gas. 206
    6.6.2: Trayed stabilization with stripping gas. 207
    6.6.3: Reboiled trayed stabilization. 208

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    Chapter 6

    Crude Oil Stabilization and Sweetening

    6.1: Introduction
    Once degassed and dehydrated–desalted, crude oil is pumped to gathering facilities to be stored in storage tanks.
    These liquids contain a large percentage of methane and ethane, which will flash to gas in the tank. This lowers the partial pressure of all other components in the tank and increases their tendency to flash to vapors. Stabilization is the process of increasing the amount of intermediate (C3 to C5) and heavy (C6+) components in the liquid phase. In an oil field this process is called crude stabilization and in a gas field it is called condensate stabilization.
    In stabilization, adjusting the pentanes and lighter fractions retained in the stock tank liquid can change the crude oil gravity. The economic value of the crude oil is accordingly influenced by stabilization because of the following reasons:
    1- Liquids can be stored and transported to the market more profitably than gas.
    2- It is advantageous to minimize gas losses from light crude oil when stored.
    This chapter deals with methods for stabilizing the crude oil to maximize the volume of production as well as its API gravity, against two important constraints imposed by its vapor pressure and the allowable hydrogen sulfide content.
    In addition, the requirement to treat the oil at high temperature is more important than stabilizing the liquid and may require the flashing of both intermediate and heavy components to the gas stream.
    Gas condensate, on the other hand, may contain a relatively high percentage of intermediate components. Thus, some sort of condensate stabilization should be considered for each gas well production facility.
    The most common method used to remove the light components from hydrocarbon liquids before the liquid enters a stock tank or a pipeline is stage separation.
    Crude oil is considered ‘‘sweet’’ if the dangerous acidic gases are removed from it. On the other hand, it is classified as ‘‘sour’’ if it contains as much as 0.05 ft3 of dissolved H2S in 100 gal of oil. Hydrogen sulfide gas is a poison hazard because 0.1% in air is toxically fatal in 30 min.
    Additional processing is mandatory—via this dual operation—in order to release any residual associated gases along with H2S present in the crude.
    A stabilizer can achieve a stable specification product with a higher liquid recovery, but usually results in higher capital expenditures’ (CAPEX) and operating expenses (OPEX). The addition of a stabilizer requires additional space which is normally not a factor for onshore applications, but may be a major consideration for an offshore installation.

    6-1-1: Crude oil treatment steps
    Produced hydrocarbons from wells normally flow to a separator for removal of the hydrocarbon gas. The hydrocarbon crude or condensate oil outflow from the separator usually goes through additional stages of separation or treatment before reaching the sales point. In each of these stages the liquid reaches near equilibrium at a different condition of pressure and temperature thus to some extent “stabilizing” the crude or condensate.
    The following methods of crude stabilization are normally used:
    • Multi-stage separation
    • Weathering in a stock tank
    • Heater-treater after separation
    • Stabilizer.
    The method one selects for stabilization depends primarily on contract specifications and economics. Factors that favor the installation of a stabilization unit include:
    • An oil contract specification that requires a low crude vapor pressure that cannot easily be obtained by stage separation.
    • A sour crude with a contract specification that limits the H2S content to less than 60 ppm.
    • Condensate production with 500API or higher and flow rates in excess of 5,000 bpd.
    6.2: Process Schemes
    6.2.1: Multi-Stage Separation
    Figure 6-1 shows a multi-stage separation system. This is the most common method of separating oil and gas. This system typically requires from two to four separation stages, each occurring in a separator vessel.

    Figure 6-1. Schematic of a three-stage separation system.

    6.2.2: Oil Heater-Treaters
    Three-phase separators, which utilize gravity separation, often are not adequate to separate the water from the oil. Heating the emulsion is commonly used to break the emulsion. Heater treaters not only improve the oil-water separation process, but also stabilize the crude by vaporizing the light hydrocarbons prior to the crude flowing to an atmosphere pressure storage tank. Utilizing heater-treaters alone often results in higher than desired losses of intermediate components to the vapor phase when the hot crude is flashed entering the storage tank.
    The crude departing the treater can be cooled before going to the storage tank by exchanging heat with the colder emulsion upstream of the treater. This will lead to fewer vapor losses and will help stabilize the intermediate components when the crude is flashed at storage tank conditions. For small flow rates, the oil-treating temperature is kept as low as possible to prevent stock tank losses, since the treated oil will normally go directly to the stock tank without cooling.

    6.2.3: Liquid Hydrocarbon Stabilizer
    It is possible to stabilize a hydrocarbon liquid at constant pressure by successively flashing the hydrocarbon liquid at increasing temperatures as shown in Figure 6-2. At each successive stage the partial pressure of the intermediate components is higher than it could have been at that temperature if some of the lighter components had not been removed by the previous stage. It would be very costly to arrange a process as shown in Figure 6-2 and thus never done. Instead, the same effect can be obtained in a tall, vertical pressure vessel with a cold temperature at the top and a hot temperature at the bottom. This unit is called a “stabilizer.”

    Figure 6-2. Multiple flashes at constant pressure and increasing temperature.

    A stabilizer applies the same principles as multi-stage separation except that the flashes take place in a stabilizer tower operating at a constant pressure, but with varying temperatures. The stabilizer tower is normally a trayed vertical pressure vessel; however, structured packing may also be used. As heat is added to the bottom of the stabilizer tower, vapors are generated on the bottom tray. The hot vapors rise to the tray above, where they bubble through the liquid. The liquid is heated by the hot vapors, which vaporize some of the hydrocarbon liquid. The vapors, in turn, are cooled by the liquid, and a portion of the vapor is condensed.
    This process of vaporization and condensation is repeated on each tray in the stabilizer tower. As the liquids fall down the stabilizer tower, the heavier hydrocarbons are condensed so that the hydrocarbon liquids leaving the stabilizer tower contain almost none of the light hydrocarbon components, and the vapor leaving the top of the stabilizer tower contains almost none of the heavier components.
    The vapor pressure of the liquid hydrocarbon leaving the bottom of the tower is controlled by controlling the stabilizer tower pressure and bottom temperature. At a constant pressure, the liquid hydrocarbon product’s vapor pressure can be increased by lowering the bottom temperature, or decreased by increasing the bottom temperature.

    Figure 6-3 illustrates a liquid hydrocarbon stabilizer system. The well stream flows to a high pressure, three-phase separator. Liquids containing a high fraction of light ends are cooled and enter the stabilizer tower at a pressure between 100 to 200 psi.

    Figure 6-3. Cold-feed stabilization system.

    As the hydrocarbon liquid falls from tray to tray in the stabilizer tower, it is heated by the hot gases bubbling through the liquid. On each tray some of the liquids are vaporized and some of the hot gases are condensed. The liquids falling down the stabilizer tower become richer and richer in heavy hydrocarbon components and leaner and leaner in light hydrocarbons. At the bottom of the stabilizer tower, some of the liquids are cycled to a reboiler where they receive heat to provide the necessary bottom temperature which is normally in the range of 2000 to 4000F. The reboiler could be a direct-fired bath, an indirect-fired bath, or a heating media exchanger. For a specific bottom product’s vapor pressure, a lower stabilizer tower operating pressure requires a lower bottom temperature, but more compression is required for the overhead vapors.
    The hydrocarbon liquid leaving the stabilizer tower at the bottom tray temperature is in equilibrium with the vapors and is at its bubble point.
    The liquid leaving the stabilizer tower is cooled before going to storage or pipeline. The hydrocarbon vapors leaving the top of the stabilizer tower are in equilibrium with the liquids on the top tray and are at their dew point.
    One design consideration that needs to be addressed in the design of a stabilizer system is whether to use a cold feed or reflux. A cold-feed stabilizer without reflux such as that shown in Figure 6-3 does not achieve as good a split between the light and heavy components as a column with reflux (see Figure 6-4 and the following discussion); thus, recoveries are not as high. However, a stabilizer with reflux requires additional equipment, higher CAPEX, and higher OPEX, but achieves a higher recovery. Descriptions of both a cold-feed stabilizer and a stabilizer with reflux follow.

    6.2.4: Cold-Feed Stabilizer
    A conventional stabilizer tower is a distillation column with a reboiler, but no overhead condenser (refer to Figure 6-3). The lack of an overhead condenser means that there is no liquid reflux from the overhead stream.
    Thus, the feed is introduced on the top tray and must provide all the cold liquid for the stabilization tower. Since the feed is introduced on the top tray, it is important to minimize the flashing of the feed so that intermediate components are not lost overhead. To lower the feed stream temperature and reduce flashing, a cooler is sometimes added on the inlet feed stream.
    Adding a cooler on the inlet feed stream lowers the temperature of the inlet hydrocarbon liquid, lowers the fraction of intermediate components that flash to vapor on the top tray and increases the recovery of these components in the liquid bottoms. However, the colder the feed, the more heat is required from the reboiler to remove light components from the liquid bottoms. If too many light components remain in the liquid, the vapor pressure limitations for the liquid may be exceeded. Light components may also encourage flashing of intermediate components (by lowering their partial pressure) in the storage tank. There is a balance between the amount of inlet cooling and the amount of reboiling required.
    The hydrocarbon liquid out the bottom of the stabilizer tower must meet a specified vapor pressure. The stabilizer tower is designed to maximize the molecules of intermediate components in the liquid without exceeding the vapor pressure specification. This is accomplished by driving the maximum number of molecules of methane and ethane out of the liquid and keeping as much of the heavier ends as possible from going out with the gas. The hot liquid from the stabilizer is at its bubble point at the pressure and temperature in the stabilizer. It must be cooled sufficiently to avoid flashing when it enters the atmospheric storage tank.

    The overhead gas can be used as fuel, or compressed and included with the sales gas. Any water that enters the column in the feed stream will collect in the middle of the column due to the range of temperatures involved. This water cannot leave with the bottom product or with the overhead stream; therefore, provisions should be made to remove this water from a tray near the middle of the column. The heating of the liquid hydrocarbon in the stabilizer tower acts as a demulsifier to remove water from hydrocarbon liquid. The excellent water-separating ability of the stabilizer usually eliminates the need for a hydrocarbon liquid dehydration system.

    6.2.5: Stabilizer with Reflux
    Figure 6-4 shows a typical stabilizer system with reflux and a feed/bottom heat exchanger. In this configuration, the well fluid is heated by the bottom product and injected into the stabilizer tower, below the top, where the temperature in the stabilizer tower is equal to the temperature of the feed. The stabilizer tower’s top temperature is controlled by cooling and condensing part of the hydrocarbon vapors leaving the stabilizer and pumping the resulting hydrocarbon liquids back to the tower. This replaces the cold feed configuration and allows better control of the overhead product and, consequently, slightly higher recovery of the heavier components. This configuration minimizes the amount of flashing.
    The principles of this configuration are the same as in a cold-feed stabilizer or any other stabilizer tower. As the liquid falls through the tower, it goes from tray to tray, and gets increasingly richer in the heavier components and increasingly leaner in the lighter components. The stabilized hydrocarbon liquid is cooled in the heat exchanger by the feed stream before flowing to the stock tank or pipeline.
    At the top of the stabilizer tower intermediate components going out with the gas are condensed, separated, pumped back to the stabilizer tower, and sprayed down on the top tray. This liquid is called “reflux,” and the two-phase separator that separates it from the hydrocarbon liquid from the gas is called a “reflux tank” or “reflux drum.” The reflux performs the same function as the cold feed in a cold feed stabilizer. Cold liquid hydrocarbons strip out the intermediate components from the gas as the gas rises.
    The heat required at the reboiler depends upon the amount of cooling done in the condenser. The colder the condenser, the purer the product, and the larger the percentage of the intermediate components that will be recovered in the separator and kept from going out with the gas.
    The hotter the bottom temperature, the greater the percentage of light components boiled out of the bottoms. The greater the percentage of light components boiled out of the bottoms liquid, the lower the vapor pressure of the bottoms liquid.
    A heat balance around the stabilizer tower is part of the design. The heat leaves the stabilizer tower in the form of vapors out the top, and the liquid bottom product has to be balanced by the heat entering in the feed and the reboiler. If the stabilizer tower has a reflux, this amount of heat has to be added to the column balance.
    A stabilizer tower with reflux will recover more intermediate components from the gas than a cold-feed stabilizer tower. However, it requires more equipment to purchase, install, and operate. This additional cost must be justified by the net benefit of the incremental hydrocarbon liquid recovery, less the cost of natural gas shrinkage and loss of heating value, over that obtained from a cold-feed stabilizer.


    Figure 6-4. Schematic of a typical crude stabilization with reflux and feed/bottom heat exchanger.

    6.3: Stabilization Equipment
    6.3.1: Stabilizer Tower
    The stabilizer tower is a fractionation tower using trays or packing.
    Figure 6-5 shows a stabilizer tower with bubble cap trays.
    Trays, structured packing, or random packing are used in the tower to promote intimate contact between the vapor and liquid phases, thereby permitting the transfer of mass and heat from one phase to the other. The feed to the stabilizer tower normally enters near the top of a cold-feed stabilizer, and at or near the tray where the stabilizer tower conditions and feed composition most nearly match the inlet feed conditions, in stabilizer towers with reflux. The liquids in the stabilizer tower fall down through the downcomer, across the tray, over the weir and into the down-comer to the next tray. The temperature on each tray increases as the liquids drop from tray to tray. Hot gases come up the stabilizer tower and bubble through the liquid on the tray above, where some of the heavier components in the gas are condensed and some of the lighter components in the liquid are vaporized. The gas gets leaner and leaner in heavy hydrocarbons as it travels up the stabilizer tower; the falling liquids get richer and richer in the heavier hydrocarbon components. The vapors leaving the top of the stabilizer tower contain a minimum amount of heavy hydrocarbons, and the liquid leaving the bottom of the tower contains a minimum of light hydrocarbons. Stabilizer columns commonly operate at pressures between 100 to 200 psig.

    6.3.1.1: Trays and Packing
    The more stages, the more complete the split, but the taller and more costly the tower. Most stabilizers will normally contain approximately five theoretical stages. In a refluxed tower, the section above the feed is known as the rectification section, while the section below the feed is known as the stripping section. The rectification section normally contains about two equilibrium stages above the feed, and the stripping section normally contains three equilibrium stages.

    Trays
    For most trays, liquid flows across an “active area” of the tray and then into a “down-comer” to the next tray below, etc. Inlet and/or outlet weirs control the liquid distribution across the tray. Vapor flows up the stabilizer tower and passes through the tray active area, bubbling up through (and thus contacting) the liquid flowing across the tray. The vapor distribution is controlled by:
    • Perforations in the tray deck (sieve trays),
    • Bubble caps (bubble cap trays), or
    • Valves (valve trays).

    Trays are generally divided into four categories:
    • Sieve trays,
    • Valve trays,
    • Bubble cap trays, and
    • High capacity/high efficiency trays.

    Sieve Trays
    Sieve trays are the least expensive tray option. In sieve trays, vapor flowing up through the tower contacts the liquid by passing through small perforations in the tray floor (Figure 6-6). Sieve trays rely on vapor velocity to exclude liquid from falling through the perforations in the tray floor. If the vapor velocity is much lower than design, liquid will begin to flow through the perforations rather than into the downcomer.
    This condition is known as weeping. Where weeping is severe, the equilibrium efficiency will be very low. For this reason, sieve trays have a very small turndown ratio.


    Figure 6-5. Schematic of a stabilizer tower.


    Figure 6-6. Vapor flow through a sieve tray.
    Valve Trays
    Valve trays are essentially modified sieve trays. Like sieve trays, holes are punched in the tray floor. However, these holes are much larger than those in sieve trays. Each of these holes is fitted with a device called a “valve.” Vapor flowing up through the tower contacts the liquid by passing through valves in the tray floor (Figure 6-7). Valves can be fixed or moving. Fixed valves are permanently open and operate as deflector plates for the vapor coming up through the tray floor. For moving valves, vapor passing through the tray floor lifts the valves and contacts the liquid. Moving valves come in a variety of designs, depending on the manufacturer and the application. At low vapor rates, valves will close, helping to keep liquid from falling through the holes in the deck.
    At sufficiently low vapor rates, a valve tray will begin to weep. That is, some liquid will leak through the valves rather than flowing to the tray down-comers. At very low vapor rates, it is possible that all the liquid will fall through the valves and no liquid will reach the down-comers.
    This severe weeping is known as “dumping.” At this point, the efficiency of the tray is nearly zero.

    Figure 6-7. Vapor flow through valve tray
    Bubble Cap Trays
    In bubble cap trays, vapor flowing up through the tower contacts the liquid by passing through bubble caps (Figure 6-8).
    Each bubble cap assembly consists of a riser and a cap. The vapor rising through the tower passes up through the riser in the tray floor and then is turned downward to bubble into the liquid surrounding the cap. Because of their design, bubble cap trays cannot weep. However, bubble cap trays are also more expensive and have a lower vapor capacity/higher pressure drop than valve trays or sieve trays.


    Figure 6-8. Vapor flow through bubble cap tray

    High Capacity/High Efficiency Trays
    High capacity/high efficiency trays have valves or sieve holes or both. They typically achieve higher efficiencies and capacities by taking advantage of the active area under the down-comer. At this time, each of the major vendors have their own version of these trays, and the designs are proprietary.

    Bubble Cap Trays vs. Valve Trays
    At low vapor rates, valve trays will weep. Bubble cap trays cannot weep (unless they are damaged). For this reason, it is generally assumed that bubble cap trays have nearly an infinite turndown ratio. This is true in absorption processes (e.g., glycol dehydration), in which it is more important to contact the vapor with liquid than the liquid with vapor. However, this is not true of distillation processes (e.g., stabilization), in which it is more important to contact the liquid with the vapor. As vapor rates decrease, the tray activity also decreases. There eventually comes a point at which some of the active devices (valves or bubble caps) become inactive. Liquid passing these inactive devices gets very little contact with vapor. At this point, it is possible that liquid may flow across the entire active area without ever contacting a significant amount of vapor. This will result in very low efficiencies for a distillation process.
    Nothing can be done with a bubble cap tray to compensate for this.
    However, a valve tray can be designed with heavy valves and light valves. At high vapor rates, all the valves will be open. As the vapor rate decreases, the valves will begin to close. With light and heavy valves on the tray, the heavy valves will close first, and some or all of the light valves will remain open. If the light valves are properly distributed over the active area, even though the tray activity is diminished at low vapor rates, what activity remains will be distributed across the tray. All liquid flowing across the tray will contact some vapor, and mass transfer will continue. Of course, even with weighted valves, if the vapor rate is reduced enough, the tray will weep and eventually become inoperable.
    However, with a properly designed valve tray this point may be reached after the loss in efficiency of a comparable bubble cap tray. So, in distillation applications, valve trays can have a greater vapor turndown ratio than bubble cap trays.

    Tray Efficiency and Stabilizer Height
    In general, stabilizer trays generally have a 70% equilibrium stage efficiency. That is, 1.4 actual trays are required to provide one theoretical stage. The spacing between trays is a function of the spray height and the down-comer backup (the height of clear liquid established in the down-comer). The tray spacing will typically range from 20 to 30 inches (with 24 inches being the most common), depending on the specific design and the internal vapor and liquid traffic. The tray spacing may increase at higher operating pressures (greater than 165 psia) because of the difficulty in disengaging vapor from liquid in the active areas of the tray.

    Packing
    Packing typically comes in two types: random and structured. Liquid distribution in a packed bed is a function of the internal vapor/liquid traffic, the type of packing employed, and the quality of the liquid distributors mounted above the packed bed. Packing material can be plastic, metal, or ceramic. Packing efficiencies can be expressed as height equivalent to a theoretical plate (HETP).

    Random Packing
    A bed of random packing typically consists of a bed support (typically a gas injection support plate) upon which pieces of packing material are randomly arranged (they are usually poured or dumped onto this support plate). Bed limiters, or hold-downs, are sometimes set above random beds to prevent the pieces of packing from migrating or entraining upward. Random packing comes in a variety of shapes and sizes. For a given shape (design) of packing, small sizes have higher efficiencies and lower capacities than large sizes.
    Figure 6-9 shows a variety of random packing designs. An early design is known as a Rasching ring. Rasching rings are short sections of tubing and are low-capacity, low-efficiency, high-pressure drop devices. Today’s industry standard is the slotted metal (Pall) ring. A packed bed made of 1-inch slotted metal rings will have a higher mass transfer efficiency and a higher capacity than will a bed of 1-inch Rasching rings. The HETP for a 2-inch slotted metal ring in a stabilizer is about 36 inches. This is slightly more than a typical tray design, which would require 34 inches (1.4 trays × 24-inch tray spacing) for one theoretical plate or stage.

    Structured Packing
    A bed of structured packing consists of a bed support upon which elements of structured packing are placed. Beds of structured packing typically have lower pressure drops than beds of random packing of comparable mass transfer efficiency. Structured packing elements are composed of grids (metal or plastic) or woven (metal or plastic) or of thin vertical crimped sheets (metal, plastic, or ceramic) stacked parallel to each other. Figure 6-10 shows examples of the vertical crimped sheet style of structured packing.
    The grid types of structured packing have very high capacities and very low efficiencies, and are typically used for heat transfer or for vapor scrubbing. The wire mesh and the crimped sheet types of structured packing typically have lower capacities and higher efficiencies than the grid type.

    Trays or Packing ?
    There is no umbrella answer. The choice is dictated by project scope (new tower or retrofit), current economics, operating pressures, anticipated operating flexibility, and physical properties.

    Distillation Service
    For distillation services, as in hydrocarbon stabilization, tray design is well understood, and many engineers are more comfortable with trays than with packing. In the past, bubble cap trays were the standard. However, they are not commonly used in this service anymore. Sieve trays are inexpensive but offer a very narrow operating range when compared with valve trays. Although valve trays offer wider operating range than sieve trays, they have moving parts and so may require more maintenance. High capacity/high efficiency trays can be more expensive than standard valve trays. However, high capacity/high efficiency trays require smaller diameter stabilization towers, so they can offer significant savings in the overall cost of the distillation tower. Random packing has traditionally been used in small diameter (<20 inches) towers. This is because it is easier and less expensive to pack these small diameter towers. However, random packed beds are prone to channeling and have poor turndown characteristics when compared with trays. For these reasons, trays were preferred for tower diameters greater than 20 inches.

    Stripping Service
    For stripping service, as in a glycol or amine contactor, bubble cap trays are the most common. In recent years, there has been a growing movement toward crimped sheet structured packing. Improved vapor and liquid distributor design in conjunction with structured packing can lead to smaller-diameter and shorter stripping towers than can be obtained with trays.

    6.3.1.2: Stabilizer Reboiler
    The stabilizer reboiler boils the bottom product from the stabilizer tower.
    The source of all heat used to generate vapor in a stabilizer is the reboiler. The boiling point of the bottom product is controlled by controlling the heat input of the reboiler together with the stabilizer operating pressure, this actions control the vapor pressure of the bottom product.
    Reboiler temperatures typically range from 2000 to 4000F (900 to 2000C) depending on operating pressure, bottom product composition, and vapor pressure requirements. It’s important to note that reboiler temperatures should be kept to a minimum to decrease the heat requirements, limit salt buildup, and prevent corrosion problems.
    Maintaining stabilizer operating pressures below 200 psig will result in reboiler temperatures below 3000F. A water-glycol heating medium can then be used to provide heat. Higher stabilizer pressures require the use of steam or hydrocarbon-based heating mediums.
    However, operating at high pressures decreases the flashing of the feed when entering the stabilizer tower and decreases the amount of feed cooling required. In general, a liquid hydrocarbon stabilizer should be designed to operate between 100 to 200 psig.
    Selection of a stabilizer heat source depends on the medium and tower operating pressure. The source of reboiler heat should be considered when a crude stabilizer is being evaluated. If turbine generators or compressors are installed nearby, then waste heat recovery should be considered.


    Figure 6-9. Various types of random packing.


    Figure 6-10. Structured packing can offer better mass transfer than trays.

    6.3.1.3: Stabilizer Cooler
    The stabilizer cooler is used to cool the bottom product leaving the tower before it goes to a tank or pipeline. The temperature of the bottom product may be dictated by contract specification or by efforts to prevent loss of vapors from an atmospheric storage tank.
    For a stabilizer with a reflux system, the bottom product may be cooled by exchanging heat with the feed to the stabilizer.

    6.3.1.4: Stabilizer Reflux System
    The stabilizer reflux system consists of a reflux condenser, reflux accumulator, and reflux pumps. The system is designed to operate at a temperature required to condense a portion of the vapors leaving the top of the stabilizer.
    The temperature range is determined by calculating the overhead vapor’s dew point temperature. The heat duty required is determined by the amount of reflux required.
    The type of exchanger selected for the reflux depends on the design temperature required to condense the reflux. The lower the operating pressure of the stabilizer, the lower the temperature required for condensing the reflux. In most installations, air-cooled exchangers may be used. Some installations may require refrigeration and a shell-and-tube exchanger configuration.
    The reflux accumulator consists of a two-phase separator with several minutes of retention time to allow separation of the vapors and liquids.
    The reflux accumulator is normally located below the reflux condenser, with the line sloped from the condenser to the accumulator. The size of the reflux accumulator depends on the amount of reflux required and the total amount of vapors leaving the stabilization tower.
    Reflux pumps are sized to pump the required reflux from the reflux accumulator back to the top of the stabilizer tower. Depending upon the reflux circulation rate, two 100 percent pumps or three 50 percent pumps may be installed. This allows either a 100 percent spare or a 50 percent spare pump.

    6.3.1.5: Stabilizer Feed Cooler
    An inlet feed cooler may be required if a cold feed stabilizer tower is used. Calculations are required to determine the design feed temperature and the heat duty exchanger. This exchanger is usually a shell-and-tube type with some type of refrigerant required to cool the feed sufficiently.

    6.3.1.6: Stabilizer-Heater
    A feed heater may be required for stabilizers with a reflux system. If a feed heater is used, it is normally a shell-and-tube type exchanger that exchanges heat between the cold feed and the hot bottom product, which is then cooled before going to storage or pipeline.
    The selection of equipment and the decision whether to use cold-feed or a reflux system depends on a number of factors. The availability of heat sources for reboiler and streams for cooling the system influence the final decision. Economics of product recovery, CAPEX, and OPEX are major considerations.
    6.4: Stabilizer Design
    It can be seen from the previous description that the design of both a cold-feed stabilizer and a stabilizer with a reflux is a rather complex and involved procedure. Distillation computer simulations are available that can be used to optimize the design of any stabilizer if the properties of the feed stream and desired vapor pressure of the bottom product are known. Cases should be run of both a cold-feed stabilizer and one with reflux before a selection is made. Because of the large number of calculations required, it is not advisable to use hand calculation techniques to design a distillation process. There is too much opportunity for computational error.
    Normally, the contract specification will specify a maximum Reid Vapor Pressure (RVP). This pressure is measured according to a specific American Society of Testing Materials (ASTM) testing procedure.
    A sample is placed in an evacuated container such that the ratio of the vapor volume to the liquid volume is 4 to 1. The sample is then immersed in a 1000F liquid bath. The absolute pressure then measured is the RVP of the mixture.
    The vapor pressures of various hydrocarbon components at 1000F are given in Table 6-1.
    The vapor pressure of a mixture is given by:
    VP = ∑ [ VPn x MFn ] Eq. 6-1
    Where
    Where VP = vapor pressure of mixture, psia
    VPn = vapor pressure of component n, psia
    MFn = mole fraction of component n in liquid


    Table 6-1, Vapor pressure of relative light components.
    6.5: Crude Oil Sweetening
    Apart from stabilization problems of ‘‘sweet’’ crude oil, ‘‘sour’’ crude oils containing hydrogen sulfide, mercaptans, and other sulfur compounds present unusual processing problems in oil field production facilities. The presence of hydrogen sulfide and other sulfur compounds in the well stream impose many constraints. Most important are the following:
    * Personnel safety and corrosion considerations require that H2S concentration be lowered to a safe level.
    * Brass and copper materials are particularly reactive with sulfur compounds; their use should be prohibited.
    * Sulfide stress *****ing problems occur in steel structures.
    * Mercaptans compounds have an objectionable odor.
    Along with stabilization, crude oil sweetening brings in what is called a ‘‘dual operation,’’ which permits easier and safe downstream handling and improves and upgrades the crude marketability.
    Three general schemes are used to sweeten crude oil at the production facilities:

    6.6.1: Stage vaporization with stripping gas.
    This process—as its name implies—utilizes stage separation along with a stripping agent.
    Hydrogen sulfide is normally the major sour component having a vapor pressure greater than propane but less than ethane.
    Normal stage separation will, therefore, liberate ethane and propane from the stock tank liquid along with hydrogen sulfide. Stripping efficiency of the system can be improved by mixing a lean (sweet) stripping gas along with the separator liquid between each separation stage. Figure 6-11, represents typical stage vaporization with stripping gas for crude oil sweetening/stabilization. The effectiveness of this process depends on the pressure available at the first-stage separator (as a driving force), well stream composition, and the final specifications set for the sweet oil.

    Figure 6-11. Crude sweetening by stage vaporization with stripping gas.

    6.6.2: Trayed stabilization with stripping gas.
    In this process, a tray stabilizer (nonreflux) with sweet gas as a stripping agent is used as shown in Figure 6-12. Oil leaving a primary separator is fed to the top tray of the column countercurrent to the stripping sweet gas. The tower bottom is flashed in a low-pressure stripper. Sweetened crude is sent to stock tanks, whereas vapors collected from the top of the gas separator and the tank are normally incinerated. These vapors cannot be vented to the atmosphere because of safety considerations. Hydrogen sulfide is hazardous and slightly heavier than air; it can collect in sumps or terrain depressions.
    This process is more efficient than the previous one.

    Figure 6-12. Crude sweetening by trayed stabilization with stripping gas.
    6.6.3: Reboiled trayed stabilization.
    The reboiled trayed stabilizer is the most effective means to sweeten sour crude oils. A typical reboiled trayed stabilizer is shown in Figure 6-13. Its operation is similar to a stabilizer with stripping gas, except that a reboiler generates the stripping vapors flowing up the column rather than using a stripping gas. These vapors are more effective because they possess energy momentum due to elevated temperature.
    Because hydrogen sulfide has a vapor pressure higher than propane, it is relatively easy to drive hydrogen sulfide from the oil.

    Figure 6-13. Crude sweetening by reboiled trayed stabilization.


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    Oil and Gas Measurements - Chapter 7

    Fundamentals of Oil and Gas Processing Book
    Basics of Gas Field Processing Book
    Prediction and Inhibition of Gas Hydrates Book
    Basics of Corrosion in Oil and Gas Industry Book

    [link Point to another website Only the registered members can access]

    -------
    Chapter 7 209
    Fluid Measurements 209
    7.1: Gas Measurement 209
    7.1.1: Orifice-Meter Measurement 209
    7.1.1.5: Meter Tubes 213
    7.1.2: Ultrasonic Measurement 220
    7.2: Liquid Measurements 221
    7.2.1: Volumetric Measurement Meters (Orifice Meters) 221
    7.2.2: Turbine Meters 223
    7.2.3: Positive Displacement Meters 224
    7.2.4: Turbine and Positive Displacement Meter Selection 224
    7.2.5: Mass Measurement Meters 225
    ---
    Chapter 7

    Fluid Measurements
    7.1: Gas Measurement
    7.1.1: Orifice-Meter Measurement
    The most commonly used differential measurement device, the orifice meter, is widely accepted for use in measuring volumes of liquids or vapors.
    The orifice meter consists of static pressure and differential pressure recording gauges connected to an orifice flange or orifice fitting. The orifice meter tube (meter run) consists of upstream and downstream sections of pipe for which size and tolerance have been determined through calculation and which conform to specifications set forth in API Chapter 14.3.
    The orifice plate is held perpendicular to flow by flanges or a fitting. Bore, circumference, edge sharpness, and other tolerances must meet specifications as set forth in API Chapter 14.3.
    For additional calculation procedures, refer to in Chapter 14.3 of the API Manual of Petroleum Measurement Standards.

    7.1.1.1: Orifice Flanges (Fig. 7-1a)
    Orifice flanges require that the line be shut down and depressurized in order to inspect or change the orifice plate. The flange bolts are loosened and removed.
    The flanges are spread by use of "jack" bolts, and the plate is removed.
    When orifice flanges are used, the pressure tap hole placement may be determined by measuring from the face of the flange to the pressure tap hole center.

    7.1.1.2: Single Chamber Orifice Fitting (Fig. 7-1b)
    This fitting also requires that the line be shut down and depressurized in order to inspect or change the orifice plate. However, this fitting does not require breaking apart the flanges. Instead, the bolts are loosened on the cover plate and the cover plate removed. The orifice plate holder and orifice plate are then removed from the fitting. These fittings provide precise alignment of the orifice plate.

    7.1.1.3: Dual-Chamber Orifice Fitting (senior fitting) (Fig. 7-1c).
    This fitting allows the removal and inspection of an orifice plate while the line remains under pressure. It allows the orifice plate holder and orifice plate to be raised into the upper cavity of the fitting by the use of a crank handle. A valve is then closed to separate the upper cavity from the lower cavity of the fitting. The upper cavity is then depressurized, the top cover plate removed, and the orifice plate cranked out.

    7.1.1.4: Orifice Plates
    The minimum, maximum, and recommended thicknesses of orifice plates for various pipe sizes are given in Table 7-1. Also shown in this figure are maximum allowable differential pressures for stainless steel plates of the recommended thickness at a maximum temperature of 150 degrees F.
    The thickness of the orifice plate at the orifice edge (e) shall meet the following:
    The minimum thickness is defined by e ≥ 0.01d or e > 0.005 in. whichever is greater.
    The maximum thickness is defined by e ≥ 0.02d or ≤ 0.125d whichever is smaller
    If the thickness of the orifice plate must be greater than permitted by these limitations, the downstream edge shall be cut away (beveled or recessed) at an angle of 45°± 15° or less to the face of the plate, leaving the thickness of the orifice edge within these requirements. All orifice plates which are beveled should have the square-edge side (i.e., the side opposite the beveling) stamped "inlet" or the beveled side stamped "outlet".
    The upstream face of the orifice plate shall be flat and perpendicular to the axis of the meter tube, when in position between the orifice flanges or in the orifice fitting. Any plate that does not depart from flatness along any diameter by more than 0.010 inch per inch of the dam height, (D-d)/2, shall be considered flat. See Figure 7-2, table. 7-1.
    The upstream edge of the orifice shall be square and sharp so that it will not show a beam of light when checked with an orifice edge gauge, or alternately will not reflect a beam of light when viewed without magnification. The orifice shall not have a burred or feathered edge. It shall be maintained in this condition at all times. Moreover, the orifice plate shall be kept clean at all times and free from accumulation of dirt, ice, and other extraneous material. Orifice plates with small nicks in the edge can be expected to increase the flow measurement uncertainty.


    Fig. 7-1. Orfice flange and orifice fitting.

    Fig. 7-1. Continued.


    Fig. 7-2. Orifice Plate Dimensions



    Table. 7-1. Orifice Plate Dimensions
    Notes:
    1. The maximum allowable ΔP (In. of H2O) orifice flange is 1000 for all orifice sizes.

    7.1.1.5: Meter Tubes
    The term "meter tube" shall mean the straight upstream pipe ahead of the orifice fitting of the same internal diameter as the orifice fitting (length UL on Fig. 7-3) and the similar downstream pipe (length DL on Fig. 7-3) following the orifice.
    The sections of pipe to which the orifice flanges are attached or the sections adjacent to the orifice flange or fitting shall comply with API Chapter 14.3 (AGA Report No. 3).
    See Tables 7-2, 7-3, and 7-4 for proper meter tube lengths.


    FIG. 7-3. Orifice Meter Tube Layout for Flanged or Welded Limit


    Table. 7-2.Orifice Meter Installation Requirements without a flow conditioner. Minimum straight unobstructed meter tube length from the upstream and downstream side of the orifice plate (in internal pipe diameter, Di)



    Table. 7-2. (continued) Orifice Meter Installation Requirements without a flow conditioner.
    UL – Minimum meter tube length upstream of the orifice plate in internal pipe diameter, Di (See Fig. 7-3). Straight length shall be measured from the downstream end of the curved portion of the nearest (or only) elbow or of the tee or the downstream end of the conical portion of reducer or expander
    DL – Minimum downstream meter tube length in internal pipe diameters, Di (See Fig. 7-3).
    Sep – Separation distance between piping elements in internal pipe diameter, Di, measured from the downstream end of the curved portion of the upstream elbow to the upstream end of the curved portion of the downstream elbow.
    Note: The tolerance on specific length for UL and DL is +/- 0.25Di.


    Table. 7-3. Orifice Meter Installation Requirements with 1998 uniform Concentric 19-Tube Bundle Flow Straightener for Meter Tube Upstream Length of 17Di =< UL<29Di.

    Notes:
    Length shown under the UL2 column are the dimension shown in figure 7-3 expressed as the number of published internal pipe diameters (Di) between the downstream end of the 1998 Uniform Concentric 19-Tube Bundle Flow Straightener and the upstream surface of the orifice plate.
    ( * ) – 13 Di allowed for up to β = 0.54
    ( ** ) – 9.5 Di allowed for up to β = 0.47
    ( *** ) – 9.5 Di allowed for up to β = 0.46
    Sep – separation distance defined in previous table.
    UL1 = UL – UL2
    Note: The tolerance on specific length for UL, UL2, and DL is +/- 0.25Di.
    Not allowed means that it is not possible to find an acceptable location for the 1998 Uniform Concentric 19-Tube Bundle Flow Straightener downstream of the particular fitting for all values of UL.


    Table. 7-4. Orifice Meter Installation Requirements with 1998 Uniform Concentric 19-Tube Bundle Flow Straightener for Meter Tube Upstream Length of UL=>29Di.
    Notes:
    Length shown under the UL2 column are the dimension shown in figure 7-3 expressed as the number of published internal pipe diameters (Di) between the downstream end of the 1998 Uniform Concentric 19-Tube Bundle Flow Straightener and the upstream surface of the orifice plate.
    Sep – separation distance defined in previous table.
    UL1 = UL – UL2
    Note: The tolerance on specific length for UL, UL2, and DL is +/- 0.25Di.
    Not allowed means that it is not possible to find an acceptable location for the 1998 Uniform Concentric 19-Tube Bundle Flow Straightener downstream of the particular fitting for all values of UL.
    7.1.1.6: Flow Conditioners (Fig. 7-4)
    The purpose of flow conditioners is to eliminate swirls and cross currents set up by the pipe, fittings and valves upstream of the meter tube.
    Please refer to API Chapter 14.3 (AGA Report No. 3) for detailed specifications for flow conditioners.


    FIG. 7-4. 1998 Uniform Concentric 19-Tube Bundle Flow Straightener


    7.1.1.7: Gas Orifice Calculations
    Orifice Sizing
    A simple calculation is often needed to properly size an orifice plate for new or changing flow rates through existing meter tubes. The procedure uses an existing or assumed flow quantity, a desired differential pressure at a specific static pressure, an estimated flowing temperature, and a determined or assumed specific gravity. The Key orifice coefficient is calculated from the gas flow equation. This calculated value is then compared to Table. 7-5, and the next larger size is usually selected.

    To determine the approximate orifice size required, the corresponding Keyg (natural gas) is calculated using appropriate terms of Eq. 7-1;
    Qh = Keyg x Ftf x Fg (hw x Pf)0.5 Eq. 7-1

    Where
    Qh = rate of flow, std. cu ft/hr
    Keyg = Fn (Fc + Fsl) = orifice factor
    Ftf = flowing temperature factor to change from the assumed flowing temperature of 60°F to the actual flowing temperature = [520/(460+Tf)]0.5
    Fg = specific gravity factor applied to change from a specific gravity of 1.0 (air) to the specific gravity of the flowing gas = (1/ G gas Sp. Gr.)0.5
    hw = differential pressure measured across the orifice plate in inches of water at 60°F (1 psi = 2.77 inches of water)
    Pf = flowing pressure psia.

    Table. 7-5. Plate Sizing and Approximate Flow rate, Natural Gas, Natural Gas Liquids and Steam.
    (For Flange Taps).

    Example 7-1: Size an orifice plate in gas service.
    Given Data:
    Line Size, D = 4.026 in.
    Flange Taps
    Specific Gravity = 0.700
    Flowing Temperature = 100°F
    Flowing Pressure = 75 psia
    Flow Rate = 14,200 cu ft/hr
    (14.73 psia @ 600F)
    Desired Differential = 50 in. of water

    Solution (using equation 7-1)
    Ftf = (520/560)0.5 = 0.9636
    Fg = (1/0.7)0.5 = 1.1952

    14,200 = Keyg x 0.9636 (1.1952) (50 x 75)0.5
    Keyg = 201.342

    Referring to Keyg (Table. 7-5) for a 4.026 inch line with flange taps, access the Keyg value which approximates the calculated number. A 1.000 in. orifice size would be selected which has a
    Key value of 201. More precise calculations would include other corrections. For more precise custody transfer calculations, please refer to API Chapter 14.3 (AGA Report No. 3).

    Flow Rate Calculation
    The following example illustrates a calculation of flow rate through an orifice.

    Example 3-2: Calculate an approximate flow rate for the orifice using appropriate terms from Eq 7-1.
    Given Data:
    Line Size, D = 6.065 in.
    Orifice Size, d = 3.500 in.
    Flange Taps
    Flowing Temperature = 70°F
    Flowing Pressure = 90 psia
    Differential = 60 in. of water
    Specific Gravity = 0.750

    Qh = Keyg x Ftf x Fg (hw x Pf)0.5 Eq. 7-1
    From table 7-5, Keyg = 2646
    Ftf = (520/530)0.5 = 0.9905
    Fg = (1/0.75)0.5 = 1.1547

    Qh = 2646 (0.9905) (1.1547) (60 x 90)0.5

    Qh = 222,387 cu ft/hr @ 14.73 psia and 60°F

    More precise calculations would include other corrections.
    For more precise custody transfer calculations, please refer to API Chapter 14.3 (AGA Report No. 3).
    Well Test Calculation
    Often it is necessary to determine an approximate flow quantity from a well head or field separator vent to the atmosphere for test purposes. The use of a "well head tester" has been a common practice since the early days of the oil and gas industry. See Figure 7-5. An orifice is installed between a pair of flanges, at the outlet of a pipe nipple which is at least eight pipe diameters long. The square edge of the orifice faces the flow. The diameter of the pipe nipple should not be greater than the preceding fittings. The pressure connection may be made in the upstream flange or at any point in the pipe nipple within three diameters from the orifice. The pressure differential across the orifice is the difference between the upstream pressure and atmospheric pressure.

    FIG. 7-5. Typical Test Set-Up for Measuring Gas from a Separator Vent

    An approximate flow rate may be calculated from:

    Q = 16,330 (1 + β4) (d2) Ftf x Cg [H(29.32+0.3H)]0.5 Eq. 7-2

    For conditions other than 60°F (flowing) and G of 0.6, correction factors must be applied.
    Ftf = [520/(460+Tr)]0.5 Eq. 7-3
    Cg = (0.6/ G gas Sp. Gr.)0.5 Eq. 7-4

    Where
    Q = gas flow rate, scfd
    β = ratio of the orifice or throat diameter to the internal diameter of the meter run, dimensionless
    d = orifice diameter, in.
    Ftf = flowing temperature factor to change from the assumed flowing temperature of 60°F to the actual flowing temperature
    Cg= gravity correction factor for orifice well tester to change from a gas specific gravity of 0.6
    H = pressure, inches of mercury (1 psi = 2.04 inches of mercury)
    Tf = flowing temperature, °F
    G = specific gravity at 60°F

    Example 3-3 — Calculate the daily gas flow through a 1-inch orifice in a nominal 3-inch pipe. The gas gravity is 0.70, the flowing temperature is 60°F, and the pressure upstream of the orifice is 5 inches Hg. The published ID of a 3-inch pipe is 3.068 in.
    Solution:
    β = 1/3.068 = 0.3259
    β4 = 0.01128
    1 + β4 = 1.01128
    Ftf = 1
    Cg= (0.6/0.7)0.5 = 0.9258

    Q = 16,330 (1 + β4) (d2) Ftf x Cg [H(29.32+0.3H)]0.5 Eq. 7-2

    Q = 16,330 (1.01128) (1) x 1 x 0.9258 [5(29.32+1.5)]0.5

    = 190,000 scfd.

    7.1.2: Ultrasonic Measurement
    This section gives a short overview of ultrasonic meters. If meter design and custody quality calculations are required, please refer to American Gas Association Report #9. These meters are designed for measurement of single-phase fluid only.
    An ultrasonic meter (UM) is a fluid velocity-sensing device.

    Fig. 7- 6. Ultrasonic Flow Meter

    (See Figure 7-6) The flowing gas velocity is determined by the transit times of high frequency pulses between two matched transducers. One is designated as upstream and one as downstream to the position in the meter and the direction of flow. These transducers attach into the pipe wall but do not protrude into the gas stream, thus creating a zero pressure drop. There are simple, single path meters that consist of one pair of transducers and multi-path meters with three or more pairs of transducers. Each pair of transducers measures the transit time of each sound pulse transmitted from the up-stream transducer to the downstream transducer with the flow (t1), and from the downstream to the upstream transducers against the flow (t2). The transit time for a signal traveling with the gas flow is less than travel time against the gas flow. The difference in these transit times relates to the gas velocity along that specific path. Various calculations and methodologies are then used to calculate the average gas velocity and flow rate at line conditions.
    A single path meter monitors only one path’s mean velocity at one elevation in the gas flow. Since most gas flow is not fully symmetrical, the use of a single path UM would have inaccuracies dependent on the flow velocity profile. Single path meters are generally used for operational balancing and flare measurement and are generally not accepted for custody transfer measurement.
    A multi-path UM continuously monitors three or more mean velocities at different elevations in the gas stream of the metered area. The averages of these mean velocities are used to calculate the gas flow rate. Meter designs of various meter manufactures are able to minimize the effect of non-symmetrical flow profiles on the overall meter accuracy. It is recommended that UM’s with three or more paths be used for custody measurement (based on available data).

    7.2: Liquid Measurements
    7.2.1: Volumetric Measurement Meters (Orifice Meters)
    Liquid volume measurement by an orifice meter can be determined by following the guidelines established in API Chapter 14.8. As with gas measurement, the primary element should consist of an orifice plate, the orifice holder with it’s associated tap holes to sense the differential and static pressure, and the upstream and downstream piping “meter tube”.
    The differential and static pressure readings are sensed at the flange taps by a secondary element sensor or transducer. The temperature of the fluid should also be recorded by the temperature sensor or transducer. Note that the meter is the tertiary device that records the output of the sensors/transducers.
    The Reader-Harris, Gallagher equation used with orifice meters produces discharge coefficients accurate within +/- 0.5%. Measurement using orifice meters must include this uncertainty, as well as the uncertainty in the metering equipment, unless the metering system is proven against a traceable standard (see API Chapter 4), similar to the way turbine meters and PD meters are typically proven. Then the overall system uncertainty may be reduced to +/- 0.25%.
    Some fluid physical properties also need to be known. Examples may include density, viscosity, and compressibility to accurately determine volume using the AGA Report #3 method. For systems performing custody transfer mass measurement for light hydrocarbons such as ethane, ethylene, E/P Mix, high ethane raw mix NGLs, etc., the flowing density of the stream should be measured with a density meter. Then the mass of the delivery may be determined by multiplying the volume at flowing conditions from the meter/ELM, times the density of the flowing stream from the density meter. Details of this method can be found in API Chapters 14.4, 14.6, 14.7, 14.8, and 21.2.

    The following equation is to determine orifice size required, or liquid flow rate,.
    Qh = Keyl x Fgt (hw)0.5 Eq. 7-5
    Where
    Qh = rate of flow, gal./hr
    Keyl = Fn (Fc + Fsl) = orifice factor
    Fgt = gravity-temperature factor for liquids

    Fgt = [1.0057/ (Gl)0.5]x (Gf/Gl)0.5 Eq. 7-6

    hw = differential pressure measured across the orifice plate in inches of water at 60°F
    Gf = specific gravity at flowing temperature (Extracted from tables of specific gravity correction factors if flowing temperature differ than 600F or from fig 7-7, Extract API value at flowing temperature, and convert it to specific gravity)
    Gl = specific gravity at 60°F



    Example 3-4: Calculate an approximate orifice size for the given flow rate and line size.
    Line Size, D = 3.068 in.
    Flange Taps
    Specific Gravity at 60°F = 0.690
    Flowing Temperature = 60°F
    Flow Rate = 3400 gal/hr.
    Desired Differential = 50 in. of water

    Solution:
    Fgt = [1.0057/ (Gl)0.5]x (Gf/Gl)0.5
    Fgt = [1.0057/ (0.69)0.5]x 1
    = 1.2107
    Qh = Keyl x Fgt (hw)0.5
    3400 = Keyl x 1.2107(50)0.5
    Keyl = 397

    Referring to the Key values (Table. 7-5) for a 3.068 inch line with flange taps, access the value listed which approximates the calculated Keyl. A 1.5 inch orifice diameter would be selected, which has a 471 Keyl value.

    Flow Rate calculation
    The liquid flow rate through an orifice is calculated using eq. 7-5.
    The initial calculation can be completed using only the Keyl and the Fgt correction factors to solve for Qh since those factors are most significant.

    Example 3-5 — Calculate a liquid flow rate for the given orifice setting.
    Line Size, D = 8.071 in.
    Orifice Size, d = 4.000 in.
    Flange Taps
    Specific Gravity at 60°F = 0.630
    Flowing Temperature = 80°F
    Differential = 36 in. of water

    Solution:
    Qh = Keyl x Fgt (hw)0.5 Eq. 7-5

    The value of Keyl from table. 7-5 is 3345 for an 8.071 in. line with a 4.0 in. orifice. The value of Fgt is calculated using eq. 7-6.
    Fgt = [1.0057/ (Gl)0.5]x ( Gf/Gl)0.5 Eq. 7-6

    API =89 at 80 0F from fig. 7-7
    Sp.gr at 80 0F (Gf) = 141.5/220.5 = 0.6417
    Fgt = [1.0057/ (0.63)0.5]x ( 0.6417/0.63)0.5
    Fgt =1.267 x 1.009 = 1.278
    Therefore,
    Qh = 3345 x 1.278 (36)0.5 = 25,649 gal/hr.
    For more precise calculations, refer to Chapter 14.8 of the API Manual of Petroleum Measurement Standards.

    Figure 7-7. Specific gravity of petroleum fractions. kW is Watson characterization factor, (use approximate value 11, if data is not available), for more precious calculations use API correction tables for temperature.

    7.2.2: Turbine Meters
    Turbine meters are velocity-sensing devices. The direction of flow through the meter is parallel to a turbine’s rotary axis and the speed of rotation of the rotor is proportional to the rate of flow.
    The turbine meter normally consists of one moving part; an impeller held in place by high pressure, low drag bearings. A magnetic transducer mounted in the meter body is used to count revolutions as the flow passes. The pulses from the transducer are determined for a known volume passing through the meter to develop a factor in pulses per gallon, or other desired unit volume. Turbine meter components are shown in Fig. 7-8. Expected accuracies of plus or minus 0.25% can be attained by certain turbine meters where proper stream conditions are maintained and the meter is properly installed and proven.
    Doing mass measurement with turbine meters is often preferred where conditions in temperature, pressure, intermolecular adhesion and solution mixing present difficulty in converting volumes from flowing conditions to standard conditions, such as with ethane, natural gas liquids (NGL), or ethane-propane mixes. To do this properly an online densitometer needs to be used. Refer to GPA 8182 or API Chapter 14, Section 7 (14.7) for further details on mass measurement for NGLs.

    7.2.3: Positive Displacement Meters
    Displacement meters take a physically enclosed volume of fluid and move it from upstream to downstream of the metering point. The sum of these operations is an indication of the amount of liquid, which is moved over a period of time.
    An expected accuracy of 0.25% for a positive displacement (PD) meter can be attained when it is properly installed and proven. Application is normally limited to those fluids that exhibit some lubricating properties because of the multiple moving parts of a positive displacement meter. Typical applications are butane and heavier products since ethane and propane have minimal lubricating properties. Fig. 7-9 shows some internal details of a positive displacement meter. PD meters may perform mass or volumetric measurement, depending on their configuration and companion equipment.


    Fig. 7-8. Turbine meter.

    7.2.4: Turbine and Positive Displacement Meter Selection
    Turbine and positive displacement meter installations should include the following considerations:
    • Application to proper flow ranges
    • Upstream strainers to protect meter internals from foreign material
    • Pulsation and vibration
    • Proper upstream flow conditioning
    • Significant rate changes
    • Changes in flow temperature, pressure, and density
    • Back pressure (2 times “ DP” across meter plus 1.25 times equilibrium vapor pressure is minimum recommended). [“DP” is the difference between the flowing pressure and the equilibrium vapor pressure of the liquid.]
    • Connections to prove the meter
    • Verification that Liquid temperature correction factor (Ctl) and liquid pressure correction factor (Cpl) will not be applied when the meters are performing mass measurement, except during provings.
    The normally acceptable performance of a turbine or positive displacement meter will result in a change in the pulse count of less than 0.05% between meter prover runs, and less than 0.25% between provings. If the factor changes more than 0.25% between provings:
    • meter maintenance may be required
    • a total flow adjustment must be made
    If the factor changes more than 0.5% between provings:
    • the turbine must be pulled and inspected for damage or wear
    • a total flow adjustment must be made
    • the turbine must be proven again following inspection
    More details about turbine and positive displacement meter installations, operation, and proving are available in Chapters 4, 5, 6, and 12 of the API Manual of Petroleum Measurement Standards.

    7.2.5: Mass Measurement Meters
    Mass measurement of a flowing fluid is advantageous where the physical properties of the fluid are not well defined or available. Mass measurement is especially important in measuring streams containing ethane and methane because of substantial solution mixing effects. Mass measurement is accomplished by multiplying the volume of the fluid at flowing conditions, over a defined period of time, by the density of the fluid at flowing conditions during that same time. This procedure eliminates the need for the correction factors (Ctl and Cpl) for the metered volume. The total stream mass can be converted into pure components by using a weight analysis of the fluid. Refer to GPA 8182 or API Chapter 14, Section 7 (14.7) for further details on mass measurement for NGLs.
    Several more different techniques and processes have been developed to directly measure the mass of a flowing fluid. The devices utilize the principle that angular momentum of a mass is directly proportional to the mass velocity. The resistance of a mass to change direction is measured by different types of devices using combinations of sensitive mechanical and electrical sensors and transmitters that can result in a variety of electronic signals. Mass flow meter installations may not require upstream and downstream piping usually associated with other types of measurement. Proving mass flow meters may involve a complicated arrangement of flow and density measuring equipment, or access to an alternate proving station, or use of a master mass meter comparison.

    7.2.5.1: Coriolis Meters
    there is no flow, the two sine waves produced are in phase. When there is flow, the induced Coriolis force causes the tubes to twist, resulting in two out-of-phase sine waves.
    The time difference in the sine waves is directly proportional to the mass flow rate through the tubes (this The Coriolis meter is a mass-measuring device. It consists of a sensor, a transmitter and peripheral devices to provide monitoring, alarm, and/or control functions.
    The sensor consists of two flow tubes, the drive coil and magnet, two pick-off coils and magnets and the RTD “Resistance Temperature Devices.”. During operation, process fluid entering the sensor is split, half passing through each flow tube. The drive coil is energized causing the tubes to oscillate up and down in opposition to one another.
    The pick-off coils are mounted on one tube while the magnets are mounted on the other. Each coil moves through the uniform magnetic field of the adjacent magnet as the two tubes move. The voltage generated from each pickoff coil creates a sine wave representing the motion of one tube relative to the other. When may only be true at a fixed pressure).
    The density of the fluid is calculated from the frequency of oscillation of the tubes.
    The transmitter provides three actions. First, it sends a pulsed current to the sensor drive coil causing the flow tubes to vibrate. Second, it processes the sensor input signals, performs calculations, and produces various outputs to peripheral devices. Most commonly, the output of the meter is a pulsed output. Third, it allows communication with an operator or control system. Figure 7-10 shows the components of a Coriolis meter.

    Fig. 7-9. Positive displacement meter.



    FIG. 7-10. Components of a Coriolis Meter
    Detail description of how a Coriolis meter operates can be found in appendix A of the API Coriolis Liquid Measurement Draft Standards.
    When configuring the meter, users should provide some means to block in the flow so the zero flow condition can be verified. Zero verification of the meter is required from time to time as part of the normal operating procedures. Zeroing is necessary when the zero offset has shifted outside the defined limits. Since the meter should be proven after each zero, unnecessary zeroing should be avoided to minimize potential errors associated with meter factor reproducibility.
    The Coriolis meter should be proven under conditions as close to normal operating conditions as practical. The result of a meter proving will be a new or reaffirmed meter factor (MF). This meter factor may be entered in accessory equipment, the Coriolis transmitter, or applied manually to the quantity indicated. The preferred method is to input the meter factor into the accessory equipment due to its audit trail capabilities. A Coriolis meter is normally set up with calibration factors from the manufacturer. These factors, although adjustable, should not be changed. Figure 7-11 shows a typical schematic of a Coriolis meter installation. For more information on Coriolis meter, please refer to the API draft standard, Measurement of Single-phase Intermediate and Finished Hydrocarbon Fluids by Coriolis Meters.


    FIG. 7-11. Typical Installation of a Liquid Coriolis Meter

    The Coriolis meter has gained popularity in recent years as it presents a number of advantages over other types of meters. A Coriolis meter has an accuracy range of (+/-0.1%) and acceptable repeatability. It provides multi-variable measurement in one device: mass flow rate, volumetric flow rate, density and temperature. It is very tolerant of the changes in the fluid quality and flow rate. It may also be used as a bi-directional meter. Ease of installation and low maintenance are other bonuses as there are no special mounting, no flow conditioning, no straight pipe run requirements and no moving parts.
    Like all other types of meters, the Coriolis meter has its own down side. There is a significant pressure drop across the meter making it unsuitable for an existing operation where additional pressure drop cannot be tolerated.

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    Instrumentation and Control - Chapter 8
    Fundamentals of Oil and Gas Processing Book
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    Chapter 8 228
    Instrumentation and Control 228
    8.1: Introduction 228
    8.2: Type Selection and Identification 228
    8.2.1: Pneumatic Power Supplies 228
    8.2.2: Electronic Power Supplies 229
    8.3: Sensing Devices 230
    8.3.1: Pressure Sensors 230
    8.3.1.3: Bellows (Fig. 8-3) 230
    8.3.2: Level Sensors 232
    8.3.3: Temperature Sensors 237
    8.3.4: Flow Sensors 239
    8.4: Signal Transmitters 241
    8.4.1: Pneumatic Transmitters 241
    8.4.2: Electronic Transmitters 241
    8.5: Signal Converters 241
    8.5.1: Pneumatic-to-electronic (P/I) 242
    8.5.2: Electronic-to-pneumatic (I/P) 242
    8.5.3: Isolators 242
    8.5.4: Electric signal converters 242
    8.5.5: Frequency converters 242
    8.6: Recorders and Indicators 242
    8.6.1: Recorders 242
    8.6.2: Indicators 242
    8.7: Control Concepts 243
    8.7.1: Control Loops 243
    8.8: Control Modes and Controllers 245
    8.8.1: Two-Position (on-off) Controllers 245
    8.8.2: Proportional Control Mode 245
    8.9: Control Valves 246
    8.9.1: Control-Valve Bodies 247
    8.9.2: Control-Valve Actuators 248
    8.9.3: Flow Characteristics and Valve Selection 249
    8.9.4: Fundamentals of Control Valve Sizing 250

    ------------------

    Chapter 8


    Instrumentation and Control


    8.1: Introduction
    Instrumentation in oil and gas processing plant is usually comprised of a system of pneumatic, hydraulic, and electronic devices for measurement and control of all the process variables (pressure, flow, temperature, etc.) which are pertinent to the operation of the plant. In addition, computers are normally included in the instrumentation system to handle functions such as data gathering and transmission, bulk data storage, display, alarms, logging, computations, and control. Since the advent of integrated circuit electronics, specifically the microprocessor, many types of instruments are becoming more intelligent or “computerized.” An instrument may perform a single function such as a temperature indicator (TI), or a combination of functions such as a flow recording controller (FRC).
    8.2: Type Selection and Identification
    Often the type selection of an instrument is pre-determined by whatever is available, or what will be compatible with the rest of a system. There are cases, however, where the choice to install pneumatic or electronic instrumentation must be made by comparing the features of each type.

    8.2.1: Pneumatic Power Supplies
    The pneumatic power supply is more commonly known as the instrument air system. The main considerations of an instrument air system are:
    1. Adequate Capacity: The minimum capacity of the system should be the sum of the individual requirements of each air-consuming instrument in the system, plus a supplemental volume for purges, leaks, additions, etc. If accurate consumption figures are not available, an estimated consumption volume of 0.5 cubic foot per minute for each air-consuming device is usually adequate.
    The air storage tank should have sufficient capacity to maintain this rate for about five minutes or such time as is considered adequate to perform an emergency shut-down of the plant or to switch over to a backup air system. Also the air storage tank capacity should be large enough to prevent excessive cycling of the compressor.
    2. Filtering and Regulation: Instrument air systems are normally designed for pressures up to 125 psig and should be protected by relief valves. Instrument air should be free from all contamination such as oil, water, and any hazardous or corrosive gases. Non-lubricated compressors should be used if possible. Where lubricated compressors are used, an oil removal separator is required. The presence of oil may cause instrument contamination and possibly create a combustible mixture. After being compressed, instrument air must be cooled to remove the major portion of the contained water. A final drying system must be used to reduce the water dew point to at least 10°F below the ambient temperature at line pressure. An after filter may be required to remove particulate carryover from the dehydrators.
    3. Proper Distribution: The air distribution system should be free of any “pockets” where liquid could accumulate. If this is not possible, drain valves should be installed.
    All supply lines should connect to the top of the air manifold or “header.” Instrument air filter-regulators should be provided at each air-consuming device to reduce the line pressure to the supply pressure recommended by the instrument manufacturer. Instrument Society of America Standards ISA-S7.3 and ISA-S7.4 are references for additional information.
    4. Non-Air Systems: Natural gas has been used instead of instrument air in some remote installations where compressed air was not available. This practice should be avoided if at all possible due to safety and pollution problems and the additional filtering and clean-up of the gas which must be done to protect the instruments. The user must be cognizant of all applicable regulations when considering the use of any combustible gas in instrumentation service. Some small-scale systems have used bottled nitrogen for instrument gas. This is quite acceptable, but non-bleed type instruments should be used to keep the consumption to a minimum.
    5. Hydraulic Powered Devices: Hydraulic actuators are sometimes used on valves or rams where very high thrusts (up to 50,000 pounds force) are required for operation.
    Due to the problems of transmitting very high pressure signals, a local pump powered by an electric motor is often used to form what is commonly known as an “electro-hydraulic actuator.”

    8.2.2: Electronic Power Supplies
    Electronic systems send and transfer signals in cables distributed all over the process plant, working in a voltage range 0-1 Volt and current density 4-20 mA. Specifications and codes related to electronic system is more specific for electrical and instrumentations engineers and cannot be presented here.
    In the following, table a comparison between pneumatic and electrical process control.
    Pneumatic Electronic
    Advantages
    1. Intrinsically safe, no electrical circuits. 1. Greater accuracy.
    2. Compatible with valves. 2. More compatible with computers.
    3. Reliable during power outage for short period of time, dependent on size of air surge vessel. 3. Fast signal transit time.

    4. No signal integrity loss if current loop is used and signal is segregated from A.C. current.
    Disadvantages
    1. Subject to air system contaminants. 1. Contacts subject to corrosion.
    2. Subject to air leaks.
    2. Must be air purged, explosion proof, or intrinsically safe to be used in hazardous areas.
    3. Mechanical parts may fail due
    to dirt, sand, water, etc. 3. Subject to electrical interference

    4. Signal boosters often needed on transmission lines of over 300 feet. 4. More difficult to provide for positive fail-safe operation.
    5. Subject to freezing with moisture present. 5. Requires consideration of installation details to minimize points 1, 2, 3, and 4.
    6. Control speed is limited to velocity of sound.

    Table. 8-1. Instrument Type Features.

    8.3: Sensing Devices
    Some of the more common types of sensing devices for the measurement of process variables are described as follows:
    8.3.1: Pressure Sensors
    8.3.1.1: Manometer (Fig. 8-1)
    Two different pressures are applied to two separate openings in a transparent vessel containing a liquid. The difference in the heights of the liquid is used as a measure of the differential pressure. This difference should be corrected for temperature and gravity of the liquid in the manometer (usually either water or mercury). Pressures are often expressed in units such as “inches of water” or “millimeters of mercury.”

    FIG. 8-1. Types of Manometers

    8.3.1.2: Bourdon tubes (Fig. 8-2)
    A Bourdon tube is a metallic coil constructed from a metal tube having the desired elastic quality and corrosion resistance. The tendency of the tube to straighten under pressure causes a mechanical linkage to move a pointer or initiate pneumatic or electronic transmission of the measured pressure. Dampeners should be used where pulsation is a problem. Condensate traps should be used upstream of the device in steam service. The pressure indicated is “gauge” pressure which is relative to that of the surroundings. Bourdon gauges are also available as “compound” types which indicate vacuum as well as positive pressure.

    8.3.1.3: Bellows (Fig. 8-3)
    A tubular device with pleated sections somewhat like an accordion. It is flexible along its axis and lengthens or shortens according to the applied pressure.
    The bellows is usually used in low pressure or vacuum service but types are available for use with high pressures (up to several thousand psi). Typical diameters range from 1/2" to 12".
    They are often used in force-balance type transmitters and other applications where small displacements are required. Like the Bourdon tube, it indicates pressures as “gauge” or relative to its surroundings.

    FIG. 8-2. Types of Bourdon Tubes

    FIG. 8-3. Types of Bellows

    8.3.1.4: Diaphragm (Fig. 8-4)
    A flat or curved seal with a link attached to an indicator or transmission device. A diaphragm may have its own deflection properties such as with a metallic type or it may be attached to a spring or other elastic member such as with non-metallic diaphragms.

    8.3.1.5: Electrical Pressure Transducers
    The primary sensing element of many electrical pressure transducers usually takes the form of a Bourdon tube, bellows, or diaphragm to generate a movement which is transmitted to a strain gauge. A strain gauge is a device using resistance wire connected in a Wheatstone bridge configuration to generate an electrical signal proportional to the movement and hence proportional to the process variable being measured. Other types of electrical pressure transducers use properties of inductance, capacitance, or magnetic coupling to convert a pressure measurement to an electrical signal.


    FIG. 8-4. Pressure diaphragm elements.

    8.3.2: Level Sensors
    8.3.2.1: Gauge glass (Fig. 8-5)
    This is the most commonly used visual process-level device. Gauge glasses are generally classified as either transparent or reflex types. A transparent gauge glass consists of either a glass tube or an arrangement of flat glass plates in some type of holder. Since the process fluid level is viewed directly, the transparent gauge glass is normally used with opaque fluids. The reflex type has reflecting prisms to aid in viewing transparent fluids.

    FIG. 8-5. Flat glass gauge glasses


    8.3.2.2: Chain and tape float gauges (Fig. 8-6)
    Used in large, unpressurized storage tanks where the entire full-to empty range must be measured.

    FIG. 8-6. Chain and Tape Float Gauge

    8.3.2.3: Lever and shaft float gauges (Fig. 8-7)
    Used on either unpressurized or pressurized vessels where only a small range of level must be measured. The range of measurement is determined by the length of the float arm, but usually is between a few inches and a few feet.

    FIG. 8-7. Lever and Shaft Float Gauge
    8.3.2.4: Displacer level measuring device (Fig. 8-8)
    One of the most frequently used level measuring devices is the torque tube displacer. It is attached to the free end of a torque tube which has elastic properties that permit it to twist as the displacer tries to float. This slight turning of the free end of the torque tube is connected to an indicator or transmitter.
    Torque tube displacement gauges are normally limited to level spans of ten feet.

    Fig. 8-8. Displacer Level Measuring Device

    8.3.2.5: Head-pressure level gauges (Fig. 4-13)
    The true level of a liquid can be determined by dividing the measured hydrostatic head by the density of the liquid. This method requires a knowledge of the densities of all phases of the liquid. Some of these methods are: pressure gauge, bubble tube, and differential pressure measurement. The bubbler (Fig. 4-13a) is used at vacuum and low pressures and is especially good for services such as molten sulfur and dirty liquids. In "boiling-liquid" service (Fig. 4-13b), a condensate trap must be used on the vapor leg. The level of trapped condensate in the vapor leg will usually be different than the vessel liquid level, requiring compensation of the transmitter.

    8.3.2.6: Electrical type level gauges and switches (Fig.8-10)
    Two common types of level gauges are the float-magnetic gauge configuration and the conductive type shown in Fig. 8-10. Slight tension on the tape reel permits the follower magnet to track the float at the liquid level in the device in Fig. 8-10a. The position of the reel represents the level and is either connected to an indicating device or a transmitter. The device shown in Fig. 8-10b illustrates the use of a conductive fluid for high and low level alarm indication.

    Fig. 8-9. Head Pressure Level Gauges

    Fig. 8-10. Electrical Level Gauges/Switches

    8.3.2.7: Capacitance probes
    A continuous method of level measurement based on electrical properties.
    This method uses an electrode placed inside a vessel (or in a protective shell inside the vessel). The capacitance between the electrode and the wall of the vessel or shell varies as the dielectric constant varies. The dielectric in this case is the fluid, hence the capacitance varies in proportion to the liquid level. This capacitance is then measured, and converted to a level measurement to be indicated or transmitted.(figure 8-11)
    These devices operate on the fact that process fluids generally have dielectric Oils have dielectric constants from 1.8 to 5. constants, , significantly different from that of air, which is very close to 1.0.
    Pure glycol is 37; aqueous solutions are between 50 and 80. This technology requires a change in capacitance that varies with the liquid level, created by either an insulated rod attached to the transmitter and the process fluid, or an uninsulated rod attached to the transmitter and either the vessel wall or a reference probe. As the fluid level rises and fills more of the space between the plates, the overall capacitance rises proportionately. An electronic circuit called a capacitance bridge measures the overall capacitance and provides a continuous level measurement.

    Figure 8-11. Capacitive level sensors measure the change in capacitance between two plates produced by changes in level. Two versions are available, one for fluids with high dielectric constants (A) and another for those with low dielectric constants (B).

    8.3.2.8: Ultrasonic Level Transmitters.
    Ultrasonic level sensors (see Figure 8-12) measure the distance between the transducer and the surface using the time required for an ultrasound pulse to travel from a transducer to the fluid surface and back. These sensors use frequencies in the tens of kilohertz range; transit times are ~6 ms/m. The speed of sound (340 m/s in air at 15°C (1115 fps at 60°F) depends on the mixture of gases in the headspace and their temperature. While the sensor temperature is compensated for (assuming that the sensor is at the same temperature as the air in the headspace), this technology is limited to atmospheric pressure measurements in air or nitrogen.
    8.3.2.9: Radar Level Transmitters.
    Through-air radar systems beam microwaves downward from either a horn or a rod antenna at the top of a vessel. The signal reflects off the fluid surface back to the antenna, and a timing circuit calculates the distance to the fluid level by measuring the round-trip time. these systems can be installed either vertically, or in some cases horizontally with the guide being bent up to 90° or angled, and provide a clear measurement signal.

    Figure 8-12. Ultrasonic level measurement system, and Guided wave radar (GWR) systems
    8.3.3: Temperature Sensors
    8.3.3.1: Thermocouples
    An ordinary thermocouple consists of two different kinds of wires (dissimilar metals) joined together at one end to form the measuring or “hot” junction. Where the free ends are connected to the measuring instrument, a reference or “cold” junction is formed. The millivolt readings measured by the instrument represent the difference in the temperatures of the two junctions and can be converted to temperature by various methods using conversion data from thermocouple tables. The reference temperatures normally used to generate thermocouple tables are 32°F and 70°F. Figs.8-13, 8-14.

    Fig.8-13. Schematic drawing of a thermocouple

    Fig.8-14. Schematic drawing of a thermocouple

    Table. 8-2 shows some of the common thermocouple types, their usable temperature ranges, and the materials of construction.
    Thermocouples used for process measurements are usually protected by a thermowell. The mass of the thermowell should be kept to a minimum in the interest of faster response. The thermocouple must be in thermal contact with the thermowell.
    This is accomplished by the use of a thermally conductive lubricant or physical contact between the thermocouple and the well. In many measurement and control applications, electrical grounding of the thermocouple at the measurement point must be avoided.
    Various series arrangements of thermocouples may be made to obtain differential temperatures or temperature averages.
    Qualified personnel may check indicating or recording temperature devices measuring thermocouple potentials using portable equipment compatible with the thermocouple and with compensating circuitry identical to the primary device.
    The use of incompatible equipment could result in erroneous results, especially in low temperature applications. At low temperatures, extreme care must be taken to eliminate sources of moisture in thermocouple installations. Common properties for different types of thermocouples are given in Table. 8-2. Conversion tables for converting millivolts to temperatures can be found in NBS Circular #561, or obtained from thermocouple suppliers for common types.


    Table. 8-2. Thermocouple types.

    8.3.3.2: Resistance thermometers
    These are often called RTD’s for “Resistance Temperature Devices.” Since the resistance of metals changes as the temperature changes, a resistance thermometer can be constructed using this principle.
    The metals that fit a near linear resistance temperature relationship requirement best are platinum, copper, and nickel. An accurate resistance measuring device utilizing a Wheatstone bridge is calibrated in units of temperature rather than resistance.
    RTD’s are used in applications where faster responses and greater accuracies are required than may be obtained with thermocouples. Also RTD’s have a fairly high electrical output which is suitable for direct connection to indicators, controllers, recorders, etc. The use of RTD’s may also be more economical in some installations since the extension wires may be of copper rather than the more expensive thermocouple extension wire. A reference temperature source is not required for calibration. A special class of resistance thermometer is the thermistor device. It is low in cost, has fast response, and is very stable, but is limited to use at temperatures below 600°F.

    8.3.3.3: Filled-system thermometers
    These are simple, reliable, low cost devices. A bulb is attached to a capillary tube which is connected to a measuring element (bellows, Bourdon tube, etc.) in an indicating or transmitting device. The system is filled with a liquid or gas which changes in volume or pressure as the temperature of the bulb changes.

    8.3.3.4: Glass stem thermometers
    These devices are normally used in the office, laboratory, or other non-process areas.
    Breakage is a problem; accuracy is from 0.1 to 2.0 degrees depending upon the range.

    8.3.3.5: Bimetallic thermometers
    The sensing element consists of two metals with different coefficients of expansion bonded together and attached to an indicator. These are inexpensive, but not very accurate and are normally used in on-off temperature thermostats where precise control is not required, or in process applications where relative changes are to be monitored. They should be calibrated at or near the normal operating point of the temperature being monitored.
    8.3.4: Flow Sensors
    8.3.4.1: Variable head flow meters
    Flow meters in this class detect a pressure difference across a flow element specially designed to create that pressure difference. The most common flow element is the orifice plate, but other elements also in use are flow nozzles, venturi tubes, pitot tubes, averaging pitot tubes, target plates, and pipe elbows.

    8.3.4.2: Variable area flow meters (Fig. 8-15)
    This type includes the familiar rotameter. The differential pressure across the device is held constant, and the area through which the fluid passes changes due to the movement of the float up and down the tapered tube. These are usually limited to use with relatively small flows where visual indication is sufficient.

    Fig.8-15. Rotameter

    8.3.4.3: Turbine meters
    These use a small permanent magnet mounted on the meter tube to create a magnetic field. A small turbine is mounted inside the tube and turns with a speed proportional to the flow rate. As each vane of the turbine passes through the magnetic field the magnetic flux is disturbed which induces a pulse in a pickup coil mounted on the outside of the meter. The pulse rate is proportional to the flow rate. Pulses are then counted and converted to standard flow units.

    8.3.4.4: Positive displacement meters
    Positive displacement meters and metering pumps measure discrete quantities of the flowing fluid. The rotating element is mechanically coupled to a transmitter or counter which integrates or totals the counts to provide an indication in units of gallons, liters, cubic feet, etc. Some common types are: rotating vane, bi-rotor, rotating paddle, oscillating piston, and oval gear meters. They are used for custody transfer devices such as gas meters or gasoline pumps.

    8.3.4.5: Electromagnetic Flowmeter
    If an electrical conductor is moved in a magnetic field, an electrical voltage is introduced in the conductor which is perpendicular to both the direction of motion and the magnet field and whose magnitude is proportional to the magnetic field strength and the velocity of the movement. The characterization of the laws of induction also applies to the movement of a conductive fluid in a pipe through a magnetic field and is the basis for the electrostatic flowmeter.
    8.3.4.6: Ultrasonic Flow Meters
    Ultrasonic flow measurement is based on sending and receiving acoustic signals through the flow. The difference in transit time between transducers, built in at opposite sides of the pipe gives signals that can be transferred to flow. A sound wave travels faster with the flow than one propagated against the flow. The difference in transit times is proportional to the medium’s mean flow velocity.
    By installing more than one pair of transducers, a larger range of the flow profiles across the metering section can be covered and thereby increase the accuracy of the meter.

    8.3.4.7: Other flowmeters
    Some other flowmeter types occasionally encountered are:
    • Doppler effect or ultrasonic flowmeters
    • Vortex shedding flowmeters
    • Laser velocimeters
    • Thermal meters
    • Nuclear Magnetic Resonance meters
    • Gas ionization meters
    • Cross-correlation devices
    • Mass flowmeters
    All flow meters should be calibrated using the fluid being measured, or, if a different fluid is used for calibration, the properties of the calibrating fluid must be related to the fluid of measurement.

    8.4: Signal Transmitters
    8.4.1: Pneumatic Transmitters
    A pneumatic transmitter is a device that senses some process variable and translates the measured value into an air pressure which is transmitted to various receiver devices for indication, recording, alarm, and control. The signal range of 3-15 psig is the accepted industry standard; however, other ranges may be encountered. This signal is proportional to the range of measurement of the process variable. For example, 3-15 psig can represent 0-100 psi, 500-1000 gpm, –50 to +50°F, etc.
    The prime function of a transmitter is to reproduce the low energy measurement signal with sufficient energy that it may be transmitted over an appreciable distance or used as a power source to a control device. The low-energy measurement signal is that position or movement associated with the action of the process variable on the sensing element (bellows, diaphragm, Bourdon tube, etc.). Pneumatic transmitters operate in a manner similar to proportional controllers.
    8.4.2: Electronic Transmitters
    Electronic transmitters perform the same function as pneumatic transmitters: a low energy process-related signal is converted into a higher energy signal suitable to connect to other instruments in the system. The output signal of most electronic transmitters is a 4-20 mA, 10-50 mA, or 1-5 Vdc signal.
    Other ranges often encountered are: 0-10 Vdc, 2-10 Vdc, and 0.25-1.25 Vdc.
    8.5: Signal Converters
    Signal converters are used either to achieve compatibility between different types of instruments or for isolation purposes.
    Some common forms of signal converters are:
    8.5.1: Pneumatic-to-electronic (P/I)
    These are electronic pressure transmitters designed for 3-15 psig input range and the desired output range (4-20 ma, etc.).
    8.5.2: Electronic-to-pneumatic (I/P)
    I/P converters are pneumatic transmitters with an electro-magnetic device connected to a nozzle-baffle arrangement which generates a pneumatic output signal which is proportional to the input signal.
    8.5.3: Isolators
    These are usually electronic current-to-current or voltage-to-voltage converters which provide electrical isolation to eliminate unwanted ground loop currents or common mode voltages.
    8.5.4: Electric signal converters
    These fit the same category as I/Ps and P/Is in that they change the signal from one range to another. Examples are 4-20 mA to 0-10 vdc, 1-5 Vdc to 10-50 mA, etc.
    8.5.5: Frequency converters
    Frequency to DC converters typically receive pulse inputs from turbines or positive displacement flowmeters and provide a proportional 4-20 ma, 10-50 mA or voltage output. Voltage output converters are often referred to as F/V (frequency-to-voltage) converters or transmitters.
    V/F (voltage-to-frequency) converters are often used to interface standard “current-loop” type instrumentation to control devices requiring frequency or pulse-train set point inputs. These are commonly used in speed indicators for high speed centrifugal equipment.

    8.6: Recorders and Indicators
    8.6.1: Recorders
    A recorder is a device used to plot the value of one or more measured variables, generally against time, but in some cases against another associated variable or variables. Recorders are often classified in the following ways:
    1. According to use, i.e., whether the recorder is an integral part of the measuring/controlling system or is a general purpose type such as would be used in a laboratory or with a chromatograph.
    2. According to method used to drive the pen(s). This refers to whether the pen is directly connected to the sensing element or to some type of pen positioning mechanism activated by the measuring signal.
    3. According to chart type. This primarily refers to whether the recorder is of the circular or strip chart type and whether the time-axis drive is powered by a mechanical spring, electrical motor, or pneumatic drive.
    4. Analog or Digital. Analog recorders are the more familiar strip chart and circular types. Digital recorders include such things as strip printers, data loggers, electronic totalizers, and computer-related devices such as data terminals and printers.
    8.6.2: Indicators
    An indicator is any device which presents a visual display of a measured quantity such as temperature, pressure, humidity, voltage, etc. Indicators are included in an instrumentation system either as independent devices (denoted as TI, PI, FI, etc.), or as a part of a controlling device (TIC, PIC, etc.). Indicators may be classified in the following groups:

    8.6.2.1: Mechanical type
    In these indicators the measured quantity causes the movement of a pointer along a graduated scale. This movement is due to the action of the measured quantity on a diaphragm, bellows, electromagnetic coil, or other sensing device which is mechanically linked to the pointer. This includes pressure gauges, filled tube dial thermometers, voltage and current meters, level gauges, etc.

    8.6.2.2: Electronic analog type
    These are analog indicators with no moving parts. A signal from the sensing device activates an optical display attached to graduated scale. A common type uses a bank of 200 tiny gas filled tubes which are illuminated additively in proportion to the magnitude of the process signal.

    8.6.2.3: Digital type
    Digital indicators include an analog-to digital converter which changes the electrical process signal to binary format which is then displayed in numerical form.
    Typical displays consist of light emitting diodes (LED’s), liquid crystal displays (LCD’s), gas filled tubes, etc.

    8.7: Control Concepts
    8.7.1: Control Loops
    A control circuit is commonly referred to as a “loop.” A control loop may be classified as either “open” or “closed” depending upon whether the control adjustments are manual settings (open loop) or automatically determined by some type of feedback controller (closed-loop).
    8.7.1.1: Open loop (Fig. 8-16a)
    In an open-loop control system, an operator makes a manual adjustment to a device (valve) which controls the flow of a manipulated variable (steam) to attempt to achieve some set-point (desired temperature) value of a controlled variable (hot water). However, this adjustment is only valid for the conditions under which the operator made the adjustment. Any disturbance such as a change in inlet water temperature, steam temperature, heat loss to the surroundings, or throughput will cause the outlet temperature to change.

    Fig.8-16.a. Open loop.
    8.7.1.2: Closed loop (Fig. 8-16b)
    If appropriate measuring and controlling elements are added to the system, the loop is closed by the inclusion of an automatic feedback controller. The controller detects any difference between the set-point and measurement signals (error signal) and produces an output signal to drive the valve in the proper direction to adjust the heat input to cause the measurement to reach the set-point value.

    Fig.8-16.b. closed loop.

    8.7.1.3: Feedback control (Fig. 8-16c)
    The basic components of a feedback control loop are shown in block diagram form in the figure. The “comparator” actually represents the entire controller and any associated signal converters. The “control element” is the valve, the “feedback element” is the transmitter, and the “process” is the mixing of the steam and cold water inside the water heater.

    Fig.8-16.c. feedback control loop

    8.7.1.4: Feedforward control (Fig. 8-16d)
    Feed forward control (often called “Predictive Control”) is actually a form of open-loop control. An input variable (cold water temperature) is monitored and the manipulated variable (steam flow) is adjusted accordingly to compensate for changes in the input variable.
    Feedforward control is almost always used in conjunction with feedback control to overcome the effects of some expected disturbance.

    Fig.8-16.d. feedforward control loop

    8.8: Control Modes and Controllers
    Basic forms of control action or “modes” used in most process control are: two-position or “on-off” control, proportional control, integral or “reset” control, and derivative or “rate” control.
    The latter three modes are often used in various combinations with each other.

    8.8.1: Two-Position (on-off) Controllers
    The simplest form of control action is “on-off” control, in which the controller output either energizes or de-energizes some two-state device such as a relay or an open-shut type valve. The two-position controller is used extensively in home heating and cooling systems, refrigerators, hot water tanks, air compressors, and other applications where the cost of more precise control is not justified. Most two-position controllers are reverse-acting, i.e., when the measured variable is above the set-point, the controller turns the manipulated variable OFF, and when the measured variable is below the set point, the controller turns the manipulated variable ON. A “deadband” or differential gap exists around the zero error condition to minimize cycling. This is often implemented as a pair of control points: one where the controller will “kick-on” and the other where the controller will “kick-off” as opposed to a single set point. Fig.8-17.
    8.8.2: Proportional Control Mode
    In the proportional control mode, the final control element is throttled to various positions that are dependent on the process system conditions. For example, a proportional controller provides a linear stepless output that can position a valve at intermediate positions, as well as "full open" or "full shut."


    Fig.8-17. two positions control system.

    Fig.8-18. Proportional system control system.

    8.9: Control Valves
    Selecting the proper control valve for each application involves many factors. The valve body design, actuator style, and plug characteristic are critical items for selection. Proper valve sizing is necessary for accurate, efficient, economical process control. In areas where personnel will be affected, noise prediction and control becomes a significant factor.
    Engineering application guidelines, nomographs, and equations presented in the following pages may be used to determine the correct control valve configuration, size and flow characteristics, and to predict noise levels for most applications.
    The material presented here may also be used to evaluate the performance of valves installed in existing plants. The equations given in this section are used to calculate the flow coefficient (Cv or Cg) required for a valve to pass the required flow. Most valve manufacturers publish flow coefficients for each valve style and size.
    A brief description of the two major components of a control valve, the valve body and the actuator, is presented in Fig. 8-19.

    FIG. 8-19. Relationship of Major Components

    8.9.1: Control-Valve Bodies
    The control-valve body (see Fig. 8-20) regulates the rate of fluid flow as the position of the valve plug is changed by force from the actuator. Therefore, the valve body must permit actuator thrust transmission, resist chemical and physical effects of the process, and provide the appropriate end connections to mate with the adjacent piping. It must do all of this without external leakage. Most valve body designs are of the globe style, but other configurations such as ball and butterfly styles are available. Final selection depends upon detailed review of the engineering application.

    FIG. 8-20. Control valve body.

    8.9.2: Control-Valve Actuators
    Pneumatically operated control-valve actuators are the most popular type in use, but electric, hydraulic, and manual actuators are also widely used. The spring-and-diaphragm pneumatic actuator (see Figs. 8-21a and b.) is commonly specified, due to its dependability and its simplicity of design. Pneumatically operated piston actuators provide integral positioner capability and high stem-force output for demanding service conditions, such as high differential pressure or long valve stem travel distance.



    FIG. 8-21A . Direct Acting Spring-and-Diaphragm Actuator Assemblies (Air to close, fail open)


    FIG. 8-21b. Typical Reverse Acting Spring-and-Diaphragm Actuator Assemblies. (Air to open, fail close)

    8.9.3: Flow Characteristics and Valve Selection
    The flow characteristic of a control valve is the relationship between the flow rate through the valve and the valve travel as the travel is varied from 0 to 100%.
    Fig. 8-20 illustrates typical flow-characteristic curves.
    • The quick-opening flow characteristic provides for maximum change in flow rate at low valve travel with a fairly linear relationship. Additional increases in valve travel give sharply reduced changes in flow rate. When the valve plug nears the wide open position, the change in flow rate approaches zero.
    In a control valve, the quick-opening valve plug is used primarily for on-off service; however, it is also suitable for many applications where a linear valve plug would normally be specified.
    • The linear flow-characteristic curve shows that the flow rate is directly proportional to the valve travel. This proportional relationship produces a characteristic with a constant slope so that with constant pressure drop (ΔP), the valve gain will be the same at all flows. (Valve gain is the ratio of an incremental change in flow rate to an incremental change in valve plug position. Gain is a function of valve size and configuration, system operating conditions, and valve plug characteristic.)
    The linear-valve plug is commonly specified for liquid level control and for certain flow control applications requiring constant gain.
    • In the equal-percentage flow characteristic, equal increments of valve travel produce equal percentage changes in the existing flow. The change in flow rate is always proportional to the flow rate just before the change in position is made for a valve plug, disc, or ball position. When the valve plug, disc, or ball is near its seat and the flow is small, the change in flow rate will be small; with a large flow, the change in flow rate will be large. Valves with an equal-percentage flow characteristic are generally used for pressure control applications.
    They are also used for other applications where a large percentage of the total system pressure drop is normally absorbed by the system itself, with only a relatively small percentage by the control valve. Valves with an equal-percentage characteristic should also be considered where highly varying pressure drop conditions can be expected.
    • The modified parabolic-flow characteristic curve falls between the linear and the equal-percentage curve.

    Note: Where detailed process knowledge is lacking, as a rule of thumb, use equal-percentage characteristics at 70% opening for the valve sizing.

    FIG. 8-22. Example Flow Characteristic Curves.

    8.9.4: Fundamentals of Control Valve Sizing
    8.9.4.1: Gas Service
    Critical Pressure Drop — Critical flow limitation is a significant problem with sizing valves for gaseous service.
    Critical flow is a choked flow condition caused by increasing gas velocity at the vena contracta. The vena contracta is the point of minimum cross-sectional area of the flow stream which occurs just downstream of the actual physical restriction.
    When the velocity at the vena contracta reaches sonic velocity, additional increases in pressure drop, ΔP, (by reducing downstream pressure) produce no increase in flow.
    In the ISA sizing procedure critical flow limitations i addressed by calculating (Î¥), the expansion factor, for utilization within the actual sizing equation.

    Î¥ = 1 – X / (3FkXc) Eqn. 8-1
    Where
    Fk = k /1.4 Eqn. 8-2
    Where
    Υ = expansion factor, ratio of flow coefficient for a gas to that for a liquid at the same Reynolds number, dimensionless
    X = ratio of pressure drop to absolute inlet pressure (ΔP/P1), dimensionless
    Fk = ratio of specific heats factor, dimensionless, for natural gas use k = 1.27, so Fk = 0.9071
    Xc = pressure drop ratio for the subject valve at critical flow, with Fk = 1.0, dimensionless (from table 8-3)
    k = ratio of specific heats, dimensionless. For natural gas use k = 1.27

    Critical pressure drop, and thus critical flow, is realized when X ≥ FkXc. Therefore, since the flow can’t exceed that produced at the critical pressure drop the value of Î¥ in the following sizing equations should never be less than 0.67.

    Î¥ = 1 – X / (3FkXc) = 1- (1/3) = 0.67 Eq 8-3

    Likewise the value of X in the equations should never exceed FkXc.

    Sizing Calculation Procedure – The compressible fluid sizing equation can be used to determine the flow of gas or vapor through any style of valve. Absolute units of temperature and pressure must be used in the equation.
    Equations used to calculate the required Cv (valve flow coefficient) and thus valve size for a given set of service conditions is as follows:.

    Q = 32640 FpCvP1Y [X/(GgTZ)]0.5 Eq.8-4

    Where
    Q = volumetric flow rate scfd
    Fp=piping geometry factor, dimensionless (If the valve has no such fittings attached, e.g., the nominal valve size and nominal pipe size are the same, then Fp = 1.0), for other configuration Fp value can be calculated per the ANSI/ISA S75.01 standard.
    Cv = valve flow coefficient. Given from valve data sheets or to be calculated using known flow rate for valve design.
    P1 = upstream absolute static pressure, measured two nominal pipe diameters upstream of valve fitting assembly, psia
    Y = expansion factor, calculated from eq. 8-1 and 8-2 and should never be less than 0.67
    X = (ΔP/P1)
    Gg = Gas specific gravity.
    T = Temperature 0R.
    Z = Compressibility factor.

    The equations can likewise be rearranged to calculate the flow or pressure drop for a given valve and set of service conditions.
    For a new valve selection, a valve size is typically chosen such that the maximum, calculated Cv is close to 75% to 85% of valve travel. This allows for process variability while maintaining flow capability. The minimum, calculated Cv should typically occur at or about 10% of valve travel.

    8.9.4.2: Liquid Service
    The procedure used to size control valves for liquid service should consider the possibility of cavitation and flashing since they can limit the capacity and produce physical damage to the valve. In order to understand the problems more thoroughly, a brief discussion of the cavitation and flashing process is presented below.

    Flow Capacity --- The valve sizing coefficient most commonly used as a measure of the capacity of the body and trim of a control valve is Cv. One Cv is defined as one U.S. gallon per minute of 60 0F water that flows through a valve with a one psi pressure drop. The general equation for Cv is as follows:
    Cv = flow x (sp.gr at flowing temperature/ΔP)0.5 Eq.8-5
    When selecting a control valve for an application, the calculated Cv is used to determine the valve size and the trim size that will allow the valve to pass the desired flow rate and provide stable control of the process fluid.


    Table 8-3. Typical Cv, Xc and FL Values for Valves

    Cavitation — In liquids, when the pressure anywhere in the liquid drops below the vapor pressure of the fluid, vapor bubbles begin to form in the fluid stream. As the fluid decelerates there is a resultant increase in pressure. If this pressure is higher than the vapor pressure, the bubbles collapse (or implode) as the vapor returns to the liquid phase.
    Cavitation occurs in pumps, control valves, and any flow chocking devices.
    Determining when a problem-causing level of cavitation is present represents a considerable challenge. The reader is referred to ISA RP75.23, “Considerations for Evaluating Control Valve Cavitation.” This recommended practice provides more information on the cavitation process as well as suggesting a common terminology and methodology for making safe valve selections in cavitating applications.
    As discussed in the recommended practice, the selection of the appropriate operating limit for a given situation is dependent on the service conditions but should also consider other influences such as duty cycle, location, desired life, and past experience.
    All of these point to the need to consult the valve manufacturer when selecting a valve for cavitation control.

    Fig. 8-23. Bubble formations.

    Fig. 8-24. Bubble compression and collision with metal surface.

    Flashing — If the downstream pressure is equal to or less than the vapor pressure, the vapor bubbles created at the vena contracta do not collapse, resulting in a liquid-gas mixture downstream of the valve. This is commonly called flashing. When flashing of a liquid occurs, the inlet fluid is 100 percent liquid which experiences pressures in and downstream of the control valve which are at or below vapor pressure. The result is a two phase mixture (vapor and liquid) at the valve outlet and in the downstream piping. Velocity of this two phase flow is usually very high and results in the possibility for erosion of the valve and piping components.

    Sizing Information — The following section is based on ISA-S75.01, “Flow Equations for Sizing Control Valves.” The reader is referred to that standard for more complete discussion of these equations and methods. As that standard points out, these equations are not intended for situations involving mixed-phase fluids, dense slurries, dry solids, or non-Newtonian liquids. In these cases the valve manufacturer should be consulted for sizing assistance.
    The ISA methodology recognizes the impact of service conditions that will cause the liquid to vaporize at some point between the inlet and outlet of the valve. This vaporization results in either cavitation or flashing, causing a breakdown in the normal relationship between Cv and √ΔP and ultimately a limit to the flow through the valve regardless of an increasing pressure drop caused by decreasing P2. The recognition of this comes in the form of a separate sizing equation for each regime, nonvaporizing and vaporizing. Each must be solved and then the larger calculated Cv chosen as the required value.
    This discussion of liquid sizing will be further restricted to:
    1. Turbulent flow streams: There are usually flow streams that are not either high viscosity or low velocity. The majority of process plant control valves do operate in the turbulent regime, however if the Reynolds number for a process is less than 4000 the reader is referred to the ISA standard where a non-turbulent flow correction method can be found.
    2. Valve installed without fittings attached to the valve ends: When fittings are present there are, as with the previous gas sizing discussion, necessary modifications to the sizing equations to accommodate the additional disturbance to flow. This discussion will be limited to the case where there are no fittings attached, therefore the valve size and pipe size are the same, Fp = 1.0. Refer to the full ISA standard for the proper methods if fittings are present.

    Sizing Calculation Procedure —

    Nonvaporizing Volumetric flow (gpm) q = FpCv (ΔP/Gf)0.5 Eq. 8-6

    Vaporizing Volumetric flow (gpm) q = FLCv [(P1-FFPv)/Gf]0.5 Eq. 8-7

    Where
    q = Volumetric flow (gpm)
    Fp = piping geometry factor, dimensionless (If the valve has no such fittings attached, e.g., the nominal valve size and nominal pipe size are the same, then Fp = 1.0), for other configuration Fp value can be calculated per the ANSI/ISA S75.01 standard.
    Cv = valve flow coefficient. Given from valve data sheets or to be calculated using known flow rate for valve design.
    P1 = upstream absolute static pressure, measured two nominal pipe diameters upstream of valve fitting assembly, psia
    FF = liquid critical pressure ratio factor, dimensionless. From fig. 8-25 based on the critical pressure and inlet vapor pressure for subject liquid using Table. 8-4, which lists critical pressures for some common fluids.
    FL = liquid pressure recovery factor of a valve without attached fittings, dimensionless. Table 8-3.
    Pv = vapor pressure of liquid at valve inlet temperature, psia
    Gf = liquid specific gravity at upstream conditions, dimensionless


    Fig. 8-25. Liquid critical pressure ratio factor, FF


    Table. 8-4. Critical pressure for selected liquids, psia.

  6. Re: Fundamentals of Oil and Gas Processing Book "Full text"

    Pressure Relief System - Chapter 9

    Fundamentals of Oil and Gas Processing Book
    Basics of Gas Field Processing Book
    Prediction and Inhibition of Gas Hydrates Book
    Basics of Corrosion in Oil and Gas Industry Book

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    ------------

    Chapter 9 256
    Process Relief Systems 256
    9.1: Introduction 256
    9.2: Relief Device Design and Requirements: 256
    9.2.1: Blocked Discharge 257
    9.2.2: Fire Exposure 257
    9.2.3: Tube Rupture 257
    9.2.4: Control Valve Failure 257
    9.2.5: Thermal Expansion 257
    9.2.6: Utility Failure 257
    9.3: General discussion 258
    9.4: Special Relief System Considerations 260
    9.4.1: Pumps and storage equipment 260
    9.4.2: Low Temperature Flaring 260
    9.5: Relieving Devices 260
    9.5.1: Conventional Relief Valves 260
    9.5.2: Balanced Relief Valves 262
    9.5.3: Pilot Operated Relief Valves 262
    9.5.4: Resilient Seat Relief Valves 264
    9.5.5: Rupture Disk 265

    -----------------
    Chapter 9

    Process Relief Systems

    9.1: Introduction
    The most important safety devices in a production facility are the pressure relief valves, which ensure that pipes, valves, fittings, and pressure vessels can never be subjected to pressures higher than their design pressures.
    Relief valves must be designed to open rapidly and fully, and be adequately sized to handle the total flow of gas and liquids that could potentially cause an overpressure situation. They relieve the pressure by routing this stream to a safe location where it can be vented to atmosphere or burned.
    As long as pressure, level, and temperature control devices are operating correctly, the safety system is not needed “In case of steady conditions and working within design range”. If the control system malfunctions, then pressure, level, and temperature safety switches sense the problem so the inflow can be shut off. If the control system fails and the safety switches don't work, then relief valves are needed to protect against overpressure. Relief valves are essential because safety switches do fail or can be bypassed for operational reasons. Also, even when safety switches operate correctly, shutdown valves take time to operate, and there may be pressure stored in upstream vessels that can overpressure downstream equipment while the system is shutting down. Relief valves are an essential element in the facility safety system.

    9.2: Relief Device Design and Requirements:
    The ASME code requires every pressure vessel that can be blocked in to have a relief valve to alleviate pressure build up due to thermal expansion of trapped gases or liquids. In addition, the American Petroleum Institute Recommended Practice (API RP) 14C, "Analysis, Design, Installation and Testing of Basic Surface Safety Systems on Offshore Production Platforms," recommends that relief valves be installed at various locations in the production system; and API RP 520, "Design and Installation of Pressure Relieving Systems in Refineries," recommends various conditions for sizing relief valves.
    Proper selection, use, location, and maintenance of relief devices are essential to protect personnel and equipment as well as to comply with codes and laws.
    Determination of the maximum relief required may be difficult.
    Loads for complex systems are determined by conservative assumptions and detailed analysis. By general assumption, two unrelated emergency conditions caused by unrelated equipment failures or operator error will not occur simultaneously (no double jeopardy). The sequence of events must be considered. The development of relief loads requires the engineer to be familiar with overall process design, including the type of pump drives used, cooling water source, spares provided, plant layout, instrumentation, and emergency shutdown philosophy.
    In production facility design, the most common relieving conditions are as follows:

    9.2.1: Blocked Discharge
    The outlet of almost any vessel, pump, compressor, fired heater, or other equipment item can be blocked by mechanical failure or human error. In this case, the relief load is usually the maximum flow which the pump, compressor, or other flow source produces at relief conditions.

    9.2.2: Fire Exposure
    Fire is one of the least predictable events which may occur in a gas processing facility, but is a condition that may create the greatest relieving requirements. If fire can occur on a plant-wide basis, this condition may dictate the sizing of the entire relief system; however, since equipment may be dispersed geographically, the effect of fire exposure on the relief system may be limited to a specific plot area. Vapor generation will be higher in any area which contains a large number of uninsulated vessels. Various empirical equations have been developed to determine relief loads from vessels exposed to fire. Formula selection varies with the system and fluid considered. Fire conditions may overpressure vapor-filled, liquid-filled, or mixed-phase systems.

    9.2.3: Tube Rupture
    When a large difference exists between the design pressure of the shell and tube sides of an exchanger (usually a ratio of 1.5 to 1 or greater), provisions are required for relieving the low pressure side. Normally, for design, only one tube is considered to rupture. Relief volume for one tube rupture can be calculated using appropriate sizing equations in this section.
    When a cool media contacts a hot stream, the effects of flashing should be considered. Also the possibility of a transient overpressure caused by the sudden release of vapor into an all-liquid system should be considered.
    9.2.4: Control Valve Failure
    The failure positions of instruments and control valves must be carefully evaluated. In practice, the control valve may not fail in the desired position. A valve may stick in the wrong position, or a control loop may fail. Relief protection for these factors must be provided. Relief valve sizing requirements for these conditions should be based on flow coefficients (manufacturer data) and pressure differentials for the specific control valves and the facility involved.
    9.2.5: Thermal Expansion
    If isolation of a process line on the cold side of an exchanger can result in excess pressure due to heat input from the warm side, then the line or cold side of the exchanger should be protected by a relief valve.
    If any equipment item or line can be isolated while full of liquid, a relief valve should be provided for thermal expansion of the contained liquid. Low process temperatures, solar radiation, or changes in atmospheric temperature can necessitate thermal protection. Flashing across the relief valve needs to be considered.
    9.2.6: Utility Failure
    Loss of cooling water may occur on an area-wide or plant-wide basis. Affected are fractionating columns and other equipment utilizing water cooling. Cooling water failure is often the governing case in sizing flare systems. Electric power failure, similar to cooling water failure, may occur on an area-wide or plant-wide basis and may have a variety of effects. Since electric pump and air cooler fan drives are often employed in process units, a power failure may cause the immediate loss of reflux to fractionators. Motor driven compressors will also shut down. Power failures may result in major relief loads.
    Instrument air system failure, whether related to electric power failure or not, must be considered in sizing of the flare system since pneumatic control loops will be interrupted. Also control valves will assume the position as specified on "loss of air" and the resulting effect on the flare system must be considered.
    A vessel may be subject to more than one condition under different failure scenarios.
    For example, a low pressure separator may be subject to blocked discharge, gas blowby from the high pressure separator, and fire. Only one of these failures is assumed to happen at any time. The relief valve size needs to be calculated for each pertinent relieving rate and the largest size used. The usual controlling cases for common vessels and piping are shown in Table 9-1.

    Vessel Relieving Scenario
    Production Separators Blocked Discharge
    Test Separators Blocked Discharge
    Low Pressure Separators Blocked Discharge or Blowby
    Glycol Contact Tower Fire
    Oil Treater Gas Blowby or Fire
    Utility or Fuel Gas Scrubber Regulator Failure
    Heat Exchanger Tube Rupture
    Compressor Scrubber Fire
    Compressor Discharge Blocked Discharge

    Table 9-1. Maximum Rate Relieving Scenarios
    9.3: General discussion
    A vessel can only be overpressurized if the upstream vessel has a higher pressure than the vessel in question. A compressor scrubber with a MAWP of 285 that gets flow from a 285 MAWP separator does not need to have a relief valve sized for blocked discharge. The upstream relief valve will keep the upstream separator pressure from going higher than 285, so there is no way it can overpressure the downstream scrubber. The scrubber PSV only needs to be sized for fire.

    Good engineering judgment should be used to determine the relief rate when the separator MAWP is higher than the well SITP (shut in tubing pressure). Unexpected things can happen with a well. Production reservoirs at different pressures within the well bore can communicate in unexpected ways (for example, as the result of a poor cement job). Where flow is coming from a well, it is a good idea to provide an extra margin of safety. In the same time, the relief valve should be sized for blocked discharge of the full production rate.

    Figure 9-1 shows the various relationships between MAWP and the relief valve set pressure. The primary relief valve should be set to open at no more than 100% of MAWP and to relieve the worst case flow rates, not counting fire (i.e., blocked discharge or gas blowby), at a pressure of 1.10 MAWP. If two relief valves are used to handle the worst case flow rates, the first must be set no higher than 100% MAWP and the second at 1.05 MAWP. They must relieve the worst case flow rates, not counting fire, at 1.16 MAWP. The maximum pressure for relieving fire relief rates is 1.21 MAWP. Thus, under relief conditions, the pressure in the vessel may actually exceed MAWP. This buildup of pressure in the vessel above the MAWP as the relief valve opens is called "overpressure." This is taken into account by the various safety factors in the ASME Code and is one of the reasons the vessel is originally tested to 1.5 MAWP.

    The relief valve must be installed so that gases are routed to a safe location. In small facilities and remote locations this is accomplished with a simple "tail pipe," which points the discharge vertically upward and creates a jet in excess of 500 feet per second. The jet action dilutes the discharge gases to below the lower flammable limit in approximately 120 pipe diameters. Liquids may fall back on the equipment.
    In large facilities and offshore platforms where the escaping gases and liquids could present a source of pollution or ignition, it is common to route the relief valve discharges into a common "header" that discharges at a remote safe location. Often a vent scrubber is installed in this header to separate the bulk of the liquids and to minimize the possibility of liquid discharges to atmosphere.


    Pressure relief valve is a generic term applied to relief valves, safety valves, or safety relief valves. Definition by type of relief valve is covered in the relief device description. Relief valve characteristics related to pressure vessel requirements are illustrated in Fig. 9-1.


    Fig. 9-1. Characteristics of Safety Relief Valves for Vessel Protection

    Notes:
    1. The operating pressure may be any lower pressure required.
    2. The set pressure and all other values related to it may be moved downward if the operating pressure permits.
    3. This figure conforms with the requirements of the ASME Boiler and Pressure Vessel Code, Section VIII.
    4. The pressure conditions shown are for safety relief valves installed on a pressure vessel (vapor phase).
    9.4: Special Relief System Considerations
    9.4.1: Pumps and storage equipment
    The following considerations should be followed for relief system design in the following equipment.
    Pumps — Relief valves are required on the discharge of each positive displacement pump. Normally, these reliefs are piped back to the process upstream of the pump rather than to the flare system. Isolation valves around the relief valves may not be required if the pump itself can be isolated for maintenance.
    Vessels and Tanks — Vessels or tanks which are subject to atmospheric "breathing" due to cooling of gas or liquid contents are normally protected by "breather" valves or vacuum relief valves.
    Compressors — Each positive displacement compressor must have a relief valve on the discharge of each stage upstream of the block and check valves in order to protect the compressor.

    9.4.2: Low Temperature Flaring
    When low temperature streams are relieved, the flare system piping and equipment exposed to cryogenic temperature may require stainless steel or other acceptable alloys.
    The system should be designed for the coldest process stream to be relieved plus the lower temperature effect of the expanding fluid (Joule-Thomson effect). Materials selection should be made according to applicable code recommendations.
    9.5: Relieving Devices
    Valves that activate automatically to relieve pressure are called "safety valves," "relief valves," or "safety relief valves."
    Safety valves are spring loaded and characterized by a rapid full opening or "pop" action. They are used primarily for steam or air service. Sometimes they are referred to as "pop valves."
    Relief valves are spring loaded and open more slowly. They reach full opening at 25% over set pressure and are used primarily for liquid services.
    Safety relief valves can be either spring loaded or pilot operated and are designed to provide full opening with little overpressure.
    Most automatically-actuating relief devices used in production facilities are actually safety relief valves; however, they are commonly referred to as relief valves or safety valves. In this book, the term "relief valve" is used in the generic sense of any automatically-actuating pressure relieving device.
    There are three types of relief valves: conventional, balanced-bellows, and spring loaded.

    9.5.1: Conventional Relief Valves
    In a conventional relief valve, the inlet pressure to the valve is directly opposed by a spring. Tension on the spring is set to keep the valve shut at normal operating pressure but allow the valve to open when the pressure reaches relieving conditions.
    This is a differential pressure valve. Most conventional safety-relief valves available to the petroleum industry have disks which have a greater area, AD, than the nozzle seat area, AN. The effect of back pressure on such valves is illustrated in Fig. 9-2. If the bonnet is vented to atmospheric pressure, the back pressure acts with the vessel pressure so as to overcome the spring force, FS, thus making the relieving pressure less than when set with atmospheric pressure on the outlet. However, if the spring bonnet is vented to the valve discharge rather than to the atmosphere, the back pressure acts with the spring pressure so as to increase the opening pressure. If the back pressure were constant, it could be taken into account in adjusting the set pressure. In operation the back pressure is not constant when a number of valves discharge into a manifold. A cut-away of a conventional relief valve is shown in Fig. 9-2.



    FIG. 9-2.Conventional Safety-Relief Valve, and effect of back pressure on valve setting.
    Conventional relief valves should only be used where the discharge is routed independently to atmosphere, or if installed in a header system, the back-pressure build-up when the device is relieving must be kept below 10% of the set pressure so the set point is not significantly affected. (The set point increases directly with back-pressure).
    Conventional relief valves may be equipped with lifting levers or screwed caps. The lifting lever permits mechanical operations of the valve for testing or clean-out of foreign material from under the seat.

    9.5.2: Balanced Relief Valves
    Balanced safety-relief valves incorporate means for minimizing the effect of back pressure on the performance characteristics — opening pressure, closing pressure, lift, and relieving capacity.
    These valves are of two types, the piston type and the bellows type, as shown diagrammatically in Fig. 9-3. In the piston type, of which several variations are manufactured, the guide is vented so that the back pressure on opposing faces of the valve disk cancels itself; the top face of the piston, which has the same area, AP, as the nozzle seat area, AN, is subjected to atmospheric pressure by venting the bonnet. The bonnet vented gases from balanced piston-type valves should be disposed of with a minimum restriction and in a safe manner.
    In the bellows type of balanced valve, the effective bellows area, AB, is the same as the nozzle seat area, AN, and, by attachment to the valve body, excludes the back pressure from acting on the top side of that area of the disk. The disk area extending beyond the bellows and seat area cancel, so that there are no unbalanced forces under any downstream pressure.
    The bellows covers the disk guide so as to exclude the working fluid from the bonnet. To provide for a possible bellows failure or leak, the bonnet must be vented separately from the discharge. The balanced safety-relief valve makes higher pressures in the relief discharge manifolds possible.
    Both balanced-type valves shown in Fig. 9-3 should have bonnet vents large enough to assure no appreciable back pressure during design flow conditions. If the valve is in a location in which atmospheric venting (usually not a large amount) presents a hazard, the vent should be piped to a safe location. The user should obtain performance data on the specific type of valve being considered.
    A cross section drawing of a balanced (bellows) relief valve is shown in Fig. 9-3.

    9.5.3: Pilot Operated Relief Valves
    A pilot operated relief valve consists of two principal parts, a main valve and a pilot. The valve utilizes a piston instead of a shaft. Inlet pressure is directed to the top of the main valve piston. More area is exposed to pressure on the top of the piston than on the bottom so pressure, instead of a spring, holds the main valve closed. At the set pressure, the pilot opens, reducing the pressure on top of the piston thereby allowing the main valve to open fully. For some applications, pilot-operated relief valves are available in minimum friction, light-weight diaphragm construction (in place of heavy pistons). Pilot operated valves can allow backflow if downstream pressure exceeds set points. Backflow prevention is required on valves connected to common relief headers. A check valve, split piston type valve, or backflow preventer in the pilot line can be used.

    A typical pilot operated relief valve is shown in Fig 9-4. This style valve should be considered for applications involving high back pressure, high operating pressure, or where premium seat tightness is desired.



    FIG. 9-3. Balanced Safety-Relief Valve and the effect of back pressure on set pressure.


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    FIG. 9-4. Pilot Operated Relief Valve

    9.5.4: Resilient Seat Relief Valves
    With the use of metal-to-metal seat conventional or balanced type relief valves where the operating pressure is close to the set pressure, some leakage can be expected through the seats of the valve (refer to API Standard 527). Resilient seat relief valves with either an O-ring seat seal or plastic seats can provide seat integrities which exceed API Standard 527 (Fig. 9-5); however, there are limitations of temperature and material compatibility when using these valves, and manufacturer guidelines should be consulted. Although such valves provide near zero leakage until seat damage occurs, the resilient seats may erode rapidly once leakage begins.


    [link Point to another website Only the registered members can access] FIG. 9-5. O-Ring Seals - Conventional and Bellow Valves

    9.5.5: Rupture Disk
    A rupture disk consists of a thin diaphragm held between flanges. The disk is designed to rupture and relieve pressure within tolerances established by ASME Code. Rupture disks can be used in gas processing plants, upstream of relief valves, to reduce minor leakage and valve deterioration. In these installations, the pressure in the cavity between the rupture disk and the relief valve should be monitored to detect a ruptured disk. In some applications a rupture disk with a higher pressure rating is installed parallel to a relief valve. A rupture disk is subject to fatigue failure due to operating pressure cycles.
    To establish recommended replacement intervals, consult rupture disk suppliers.
    Rupture disks should be used as the primary relieving device only if using a pressure relief valve is not practical. Some examples of such situations are:
    (a) Rapid rates of pressure rise. A pressure relief valve system does not react fast enough or cannot be made large enough to prevent overpressure, e.g., an exchanger ruptured tube case or a runaway reaction in a vessel.
    (b) Large relieving area required. Because of extremely high flow rates and/or low relieving pressure, providing the required relieving area with a pressure relief valve system is not practical.
    (c) A pressure relief valve system is susceptible to being plugged, and thus inoperable, during service.




    Figure.9-6. Rupture disk.


    Figure. 9-7. Relief valve and safety relief valve  
    Fundamentals of Oil and Gas Processing
    References.
    1- Gas Processors Suppliers Association GPSA Engineering Data Book 11th, 12th, & 15th Editions. Tusla, OK.
    2- Arnold, K. and Stewart, M., Surface prod operations V1_ 2E, Surface prod operations V2_ 2E, & Surface prod operations V1_ 3E, Gulf Publishing Co., Richardson, TX.
    3- Abdel-Aal, H. K., Surface Petroleum Operations, Saudi Publishing & Distributing House, Jeddah, 1998.
    4- H.K. Abdel-Aal and Mohamed Aggour, Petroleum and gas field processing, 2003 by Marcel Dekker, Inc.
    5- Crude-Oil-Treating-Systems-Design-Manual-Sivalls-Inc.
    6- API Spec. 12J (Specification of Oil and Gas Separator) 7th.ed. Oct. 1998.
    7- API Spec. 12K (Indrict Fired Heater) 7th.ed.Jun.1989.
    8- API Spec. 12L (Vertical and Horizental Emulsion Treater) 4th ed. Nov. 1994.
    9- API Std. 560 (Fired Heaters for General Refinery Service) 3rd. ed. May 2001
    10- Standard Handbook of Petroleum and Natural Gas 2nd ed. William C. Lyons, Ph.D., P.E. Gary J. Plisga, B.S. 2005, Elsevier Inc.
    11- Gas Pipeline Hydraulics, E. Shashi Menon, P.E. PDH Engineering course material.
    12- Chilingarian, G. V., Robertson, J. O., Jr., and Kumar, S., Surface Operations in Petroleum Production, I, 1987, Elsevier Science, Amsterdam.
    13- The Chemistry and Technology of Petroleum, James G. Speight
    14- Flow Management for Engineers and Scientists, Nicholas P. Cheremisinoff and Paul N. Cheremisinoff.
    15- Al-Tahini, A., Crude oil Emulsions, Co-op Report, Department of Chemical Engineering, KFUPM, Dhahran, Saudi Arabia, 1996.
    16- Thro, M. E. and Arnold, K. E., Water droplet size determination for improved oil treater sizing, SPE 69th Annual Technical Conference and Exhibition, 1994.
    17- Basseler, O. U., De-emulsification of Enhanced Oil Recovery Produced Fluids, Tretolite Div., Petrolite Corp., St. Louis, MO, 1983.
    18- Manning, F. S. and Thomson, R., Oil-Field Processing of Petroleum, Penn-well Publishing, Tulsa, OK, 1991.
    19- Campbell, John M., ‘‘Gas Conditioning and Processing,’’ Vol. 2, published by Campbell Petroleum Series, Norman, Oklahoma, 1976.
    --------------

    Fundamentals of Oil and Gas Processing Book
    Basics of Gas Field Processing Book
    Prediction and Inhibition of Gas Hydrates Book
    Basics of Corrosion in Oil and Gas Industry Book

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  7. Re: Fundamentals of Oil and Gas Processing Book "Full text"

    Hey, can you attach this material or provide a link. Its useful but difficult to use this way.
    Thanks

    My threads; efficiency_247 :


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    • Re: Fundamentals of Oil and Gas Processing Book "Full text"

      Sorry
      the book is not available in pdf format
      You can use the book here as a text
      or
      you can purchase electronic or paper copy from Amazon
      Best regards

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