LPG Process Equipments

Gas gas exchanger
Within the contact section of an LPG recovery plant heat is generated from the exothermic reaction of LPG. The multi-stage nature of the process requires that this heat is removed from the process gas in order to obtain the highest possible conversion efficiency.  In a gas condensate plant much of this excess heat is used to generate high pressure steam.  Some of the heat is also used to reheat the process gases, such the gas returning from the intermediate absorption system.  In a metallurgical acid plant most of the heat is use to reheat either the cold gas coming from the gas cleaning system or reheating the gas from the intermediate absorption system.

The transfer of heat from one process gas to another gas is done in a gas-to-gas heat exchanger, so called because gas is flowing on both the shell and tube side of the heat exchanger.  A gas-to-gas heat exchanger usually takes the form of a vertical shell and tube heat exchanger.  Variations of the standard design are horizontal units, plate type heat exchangers and heat exchangers that are located inside the converter vessel.
Conventional Gas-to-Gas Heat Exchangers
The conventional gas-to-gas heat exchanger is vertical shell and tube heat exchanger with single segmental baffles.  These exchangers are typically over designed for the intended duty because of the following:

  • Single segmental baffles create a dead zones in the shell side flow path
  • Nozzle orientations are restricted due to the single segmental baffles
  • Gas entering the shell side impinges only one side of the tube bundle
  • High thermal stresses are induced by the cross flow arrangement
  • Large tube bundles makes it more difficult to obtain uniform flow on the shell and tube side
  • High shell side pressure drops required to obtain acceptable heat transfer coefficients
  • Non-uniform and non-symmetrical temperature profiles occur on the shell side

In a heat exchanger with single segmental baffles, gas will flow across the tube bundle first and then parallel to the tubes before encountering the next baffle
which forces the gas back across the tubes.  The different flow patterns on the shell side makes it difficult to design and predict heat transfer coefficients.
The non-uniform temperature profile across the tube bundle leads to thermal stresses which can lead to tube sheet warp ages and tube-to-tube sheet joint failures.
An improvement on the single segmental baffle is the double segmental baffle.  The gas is introduced to the two bundle at two points 180° apart.   The gas flows across the tube towards the centre of the tube bundle by two baffles.   At the centre of the tube bundle the space between the baffles allows the gas to flow along tubes into the next baffle section.  A single baffle located in the centre of the tube bundle forces the gas back out across the tube bundle towards the shell.   This pattern of flow is repeated until the desired thermal characteristics are achieved.  The advantage of the double segmental baffle is that the flow is split in two inside the tube bundle such that only half the gas flow is crossing the bundle.   This means that the pressure drop will be less than the single segmental baffle arrangement, everything else being equal.  However, there are still a mixture of cross and longitudinal flow within the tube bundle.  Double segmental heat exchangers require large bustles to distribute the gas to points 180° apart on the shell.  The flexibility of shell side nozzle placement is improved but the inlet and outlets must still be in-line with each other.
A further improvement is the so-called no-tubes-in-window design which omits all the tubes within the baffle window.  The flow of gas is now essentially always across the tube bundle which makes predicting shell side heat transfer coefficients easier.  The disadvantage is that the shell of the exchanger is generally larger to accommodate all the tubes within the overlap of the baffles and to minimize overall pressure drop.
Radial Flow Heat Exchangers
The principal behind the radial flow heat exchangers is the use of disc and dount baffles.  The dount baffle forces the gas to flow from the outside to the inside and through the hole in the middle of the baffle while the disc baffle does the opposite.  Placing the tubes within the overlapping area

of the disc and donut baffles means the all the gas flow will be essentially across the tubes only.   The disc and donut baffles are symmetrical around their centre so it does not matter where the gas enters or exits the shell side of the heat exchanger.  This allows the maximum freedom to place shell side nozzles.
The symmetrical layout of the tubes means that the temperature profile across the tube sheet will also be symmetrical.  The stresses occurring in the tube sheet will be more uniform and hence there will be less chance of failures of the tube sheet or tube-to-tube sheet joints.
Many different tube layouts can be used with a disc and donut baffle arrangement.  The simplest is to utilize a standard 30° or 60° tube layout and remove all the tubes that fall in the baffle windows.  The tube layout will not be perfectly symmetrical but with enough tube ‘rows’ across the flow path the unsymmetrical layout becomes less of a factor.
Plate Type Heat Exchangers
Plate type heat exchangers are relatively new to the industry.  The principal of plate heat exchangers is similar to the plate heat exchangers used in acid cooling service.  The cooling and heating streams flow on either side of a thin plate through which heat transfer occurs.

The heat exchanger is modular in construction allowing it to be fully shop fabricated and expanded easily.   Different materials of construction can be used for the modules taking advantage of better corrosion resistance or high temperature strength.
The MonPlex design uses a floating plate pack contained within a rigid housing.  The design allows for thermal expansion and contraction with minimal stress on the components.  The gas passages are easily accessible on both sides of the exchanger unlike a shell and tube unit where the tube side is easily inspected and cleaned while the shell side is near impossible to access.
Hot Exchangers
Hot Exchangers are generally cool the hot gases leaving the first catalyst pass before entering the second catalyst pass.  The cooling medium is generally the partially heated gas entering the plant going to the first catalyst pass.  Hot exchangers are exposed to the highest gas temperatures exit Bed 1 of the converter.  Maximum temperatures can reach as high as 625°C depending on gas strength..
Cold Exchangers
Cold exchangers are generally the first heating stage for the incoming gas and the final cooling stage for the gas entering the final absorbing tower.  For high gas strengths where excess heat is generated the cold exchanger maybe followed by an SO3 cooler or economizer. Old exchangers are generally constructed of carbon steel since the operating temperatures are not as in hot exchangers.  The problem that cold exchangers must contend with that hot exchanger do not, is the possibility of gas condensation.  From a corrosion point of view, carbon steel is not the best choice of materials.


Figure :Schematic diagram of a turboexpander driving a compressor.

A turboexpander, also referred to as a turbo-expander or an expansion turbine, is a centrifugal or axial flow turbine through which a high pressure gas is expanded to produce work that is often used to drive a compressor.
Because work is extracted from the expanding high pressure gas, the expansion is an isentropic process (i.e., a constant entropy process) and the low pressure exhaust gas from the turbine is at a very low temperature, sometimes as low as −90 °C or less.
Turboexpanders are very widely used as sources of refrigeration in industrial processes such as the extraction of ethane and natural gas liquids (NGLs) from natural gas, the liquefaction of gases (such as oxygen, nitrogen, helium, argon and krypton)[5][6] and other low-temperature processes.
Turboexpanders currently in operation range in size from about 750 W to about 7.5 MW (1 hp to about 10,000 hp).


Although turboexpanders are very commonly used in low-temperature processes, they are used in many other applications as well. This section discusses one of the low temperature processes as well as some of the other applications.

Extracting hydrocarbon liquids from natural gas

Figure :A schematic diagram of a demethanizer extracting hydrocarbon liquids from natural gas.
Raw natural gas consists primarily of methane (CH4), the shortest and lightest hydrocarbon molecule, as well as various amounts of heavier hydrocarbon gases such as ethane (C2H6), propane (C3H8), normal butane (n-C4H10), isobutane (i-C4H10), pentanes and even higher molecular weight hydrocarbons. The raw gas also contains various amounts of acid gases such as carbon dioxide (CO2), hydrogen sulfide (H2S) and mercaptans such as methanethiol (CH3SH) and ethanethiol (C2H5SH).
When processed into finished by-products (see Natural gas processing), these heavier hydrocarbons are collectively referred to as NGL (natural gas liquids). The extraction of the NGL often involves a turboexpander and a low-temperature distillation column (called a demethanizer) as shown in Figure 2. The inlet gas to the demethanizer is first cooled to about −51 °C in a heat exchanger (referred to as a cold box) which partially condenses the inlet gas. The resultant gas-liquid mixture is then separated into a gas stream and a liquid stream.
The liquid stream from the gas-liquid separator flows through a valve and undergoes a throttling expansion from an absolute pressure of 62 bar to 21 bar (6.2 to 2.1 MPa), which is an enthalpic process (i.e., a constant enthalpy process) that
results in lowering the temperature of the stream from about −51 °C to about −81 °C as the stream enters the demethanizer.
The gas stream from the gas-liquid separator enters the turboexpander where it undergoes an isentropic expansion from an absolute pressure of 62 bar to 21 bar (6.2 to 2.1 MPa) that lowers the gas stream temperature from about −51 °C to about −91 °C as it enters the demethanizer to serve as distillation reflux.
Liquid from the top tray of the demethanizer (at about −90 °C) is routed through the cold box where it is warmed to about 0 °C as it cools the inlet gas, and is then returned to the lower section of the demethanizer. Another liquid stream from the lower section of the demethanizer (at about 2 °C) is routed through the cold box and returned to the demethanizer at about 12 °C. In effect, the inlet gas provides the heat required to “reboil” the bottom of the demethanizer and the turboexpander removes the heat required to provide reflux in the top of the demethanizer.
The overhead gas product from the demethanizer at about −90 °C is processed natural gas that is of suitable quality for distribution to end-use consumers by pipeline. It is routed through the cold box where it is warmed as it cools the inlet gas. It is then compressed in the gas compressor which is driven by the turbo expander and further compressed in a second-stage gas compressor driven by an electrical motor before entering the distribution pipeline.
The bottom product from the demethanizer is also warmed in the cold box, as it cools the inlet gas, before it leaves the system as NGL.

Power generation

Figure :Schematic diagram of power generation system using a turboexpander.

Figure 3 depicts a electric power generation system that uses a heat source, a cooling medium (air, water or other), a circulating working fluid and a turboexpander. The system can accommodate a wide variety of heat sources such

  • Geothermal hot water
  • Exhaust gas from internal combustion engines burning a variety of fuels (natural gas, landfill gas, diesel oil, or fuel oil)
  • A variety of waste heat sources (in the form of either gas or liquid)

Referring to Figure 3, the circulating working fluid  is pumped to a high pressure and then vaporized in the evaporator by heat exchange with the available heat source. The resulting high-pressure vapor flows to the turboexpander where it undergoes an isentropic expansion and exits as a vapor-liquid mixture which is then condensed into a liquid by heat exchange with the available cooling medium. The condensed liquid is pumped back to the evaporator to complete the cycle.
The system in Figure 3 is a Rankine cycle as is used in fossil fuel power plants where water is the working fluid and the heat source is derived from the combustion of natural gas, fuel oil or coal used to generate high-pressure steam. The high-pressure steam then undergoes an isentropic expansion in a conventional steam turbine. The steam turbine exhaust steam is next condensed into liquid water which is then pumped back to steam generator to complete the cycle.
When an organic working fluid such as R-134a is used in the Rankine cycle, the cycle is sometimes referred to as an Organic Rankine Cycle (ORC).

Refrigeration system

Figure : Schematic diagram of a refrigeration system using a turboexpander, compressor and a motor.

Figure  depicts a refrigeration system with a capacity of about 100 to 1000 tons of refrigeration (i.e., 352 to 3,520 kW). The system utilizes a compressor, a turboexpander and an electric motor.
Depending on the operating conditions, the turboexpander reduces the load on the electric motor by some 6 to 15% as compared to a conventional vapor-compression refrigeration system that uses a throttling expansion valve rather than a turboexpander.
The system employs a high-pressure refrigerant (i.e., one with a low normal boiling point) such as:

  • Chlorodifluoromethane (CHClF2) known as R-22, with a normal boiling point of −47 °C
  • 1,1,1,2 Tetrafluoroethane (C2H2F4) known as R-134a, with a normal boiling point of −26 °C.

As shown in Figure 4, refrigerant vapor is compressed to a higher pressure resulting in a higher temperature as well. The hot, compressed vapor is then condensed into a liquid. The condenser is where heat is expelled from the circulating refrigerant and is carried away by whatever cooling medium is used in the condenser (air, water, etc.).
The refrigerant liquid flows through the turboexpander where it is vaporized and the vapor undergoes an isentropic expansion which results in a low-temperature mixture of vapor and liquid. The vapor-liquid mixture is then routed through the evaporator where it is vaporized by heat absorbed from the space being cooled. The vaporized refrigerant flows to the compressor inlet to complete the cycle.

                                                                                                                                                                                                         Slug catcher

The scope of the slug catcher is the separation of a two-phase flow.
The two-phase stream flow from the pipeline enters the slug catcher section at about 11.5 m elevation from ground.
The last part of the pipeline is also flanged pieces of pipe slightly sloped downward to permit the liquid to accumulate and be separated.
When the flow comes to the separator inlet, the first separation occurs between the gas that flows in the separator and the liquids that goes straight to the end section of the pipeline where it is discharged by a level control valve.
The gas phase flows through the separator and the demisters where the remaining entrained liquids are separated to the bottom of the vessel and returned into the pipeline section (liquid section).
Any excess gas retained in the liquid section is returned to the separator with the equalizer line.
The demister space can be divided into multi equivalent sections with separate gas outlets and washing devices.
If salt or pipeline solids plug the demisters, it is possible to wash one section at a time, while the other sections continue to work. If the amount of liquid phase increases, it is possible to add a flanged piece of pipe at the end of the pipeline to accommodate the unexpected flow increase.

To sum up the main features of our proprietary slug catcher are the following:

  • If the liquid flow increases it is possible to handle the unexpected liquid stream simply by adding flanged pieces of pipe as needed;
  • The plugging of the demister will not stop the operation.

Liquid slugs can exist in a region between the superficial liquid velocities of 1 to 4 m/s and a superficial gas velocity of 4 to 20 m/s. These liquid slugs travel at the same velocity as the gas through a pipeline and can cause significant damage to, or operational problems for, gas processing equipment. Gas pipelines have typically used “Slug Catchers” to dissipate the energy of the liquid slugs that intermittently propagate through a gas pipeline.
Current “Slug Catcher” designs are based on reducing fluid velocities to promote a “stratified” flow regime and subsequent gravity separation. To attain this the slug catcher must control and dissipate the energy of the incoming gas stream as it enters the slug catcher to minimize turbulence and ensure that the gas and liquid flow rates are low enough so that gravity segregation can occur. Velocity reduction is achieved by enlarging the pipe n. diameter. A rule of thumb is that the gas velocity cannot exceed 1.5 m/s (5 f/s) for liquid removal to occur. The current slug catcher design must not only promote stratification, but must also be capable of handling the largest slug volume without permitting slug formation in the slug catcher. Thus, selecting slug catcher length is an important part of the slug catcher design. To conserve on land area, it is common to use “finger storage” which is essentially a mesh of interconnected slug catchers
Basic Operation
In its simplest form, the Slug Stabilizer System consists of a branch connection, a slug stabilizing loop, a bypass loop, and another branch connection. In normal operation the gas flow enters the Slug Stabilizer System and divides at the first branch connection between the slug catching loop and the bypass loop before recombining at the second branch connection and continuing along the pipeline. The branch connections will normally be a tee and can be oriented from horizontal trough to vertical position. The branch positions and sizes can be optimized to meet individual gas/liquid pipeline flow characteristics and minimize the volume of liquid entering the branch connection while still allowing sufficient gas flow to bypass ahead of the liquid slug.

Incoming liquid slugs enter the Slug Stabilizer System and the majority of the slug is propelled past the first branch connection. Most of the liquid will continue past the branch connection and into the slug-stabilizing loop due to the liquid’s momentum energy and its resistance to change flow direction. The liquid slug will continue passing the branch connection (with some liquids entering the branch) until the gas behind the slug reaches the same branch connection. The gas pressure behind the liquid slug is normally higher (to overcome the slug’s friction losses and provide the liquid momentum energy) than the pressure in the slug stabilizing loop and the bypass piping. When the higher gas pressure pushing the slug reaches the branch connection the gas will bypass into the branch, as the gas pressure is lower in the branch connection. As the gas flows into the bypass branch the pressure will increase in the branch connection and decrease in the piping behind the liquid slug.
The branch by-pass piping directs the high-pressure gas from behind the liquid slug to the other end of the slug-stabilizing loop where the second branch connection is located. The gas will divide with most of the gas continuing down the pipeline with some gas entering the slug stabilizing loop and increasing the pressure in front of the liquid slug. The second branch connection can also be designed such that most of the liquids that entered the bypass branch (while the slug was passing the first branch connection) will be carried past the second branch connection into the slug stabilizing loop. With the gas pressure increasing in front of the slug and the gas pressure reducing behind the slug (due to gas bypassing around the slug) the momentum of the slug will be reduced until the pressure stabilizes and effectively reduces the slug velocity to nil. The preferred gas flow is now through the bypass piping and the trapped liquid slug along with
the bypass liquids are essentially suspended in the slug-stabilizing loop. The size, length and orientation of the slug stabilizing loop can be optimized for the maximum slug expected and the applicable gas flow rates and pressures.
The slug-stabilizing loop can be designed with a grade to collect the trapped liquids and discharge through a drain. An alternative would be to grade the slug-stabilizing loop to allow a stratified flow reintroduction of liquids into the primary gas stream. The liquids can then be removed with downstream equipment that only has to be designed for stratified flow and sized for a liquids withdrawal rate equivalent to the re-introduction liquid flow rate. In effect, the slug stabilizer system allows downstream separation equipment to have liquid capacity design in line with the re-introduction rate which can occur over an hour or longer instead of having to separate the liquid slug in a minute or less. This is a major improvement over the design criterion of existing slug catchers which rely on gas velocity reduction to enable liquid separation in the short time available that the slug travels through the conventionally designed slug catcher.
Slugcatcher is the name of a unit in the gas refinery or petroleum industry in which slugs at the outlet of pipelines are collected or ‘caught’. A slug is a large quantity of gas or liquid that exits the pipeline.


Pipelines that transport both gas and liquids together, known as Two-phase flow, can operate in a flow regime known as slugging flow or slug flow. Under the influence of gravity liquids will tend to settle on the bottom of the pipeline, while the gasses occupy the top section of the pipeline. Under certain operating conditions gas and liquid are not evenly distributed throughout the pipeline, but travel as large plugs with mostly liquids or mostly gasses through the pipeline. These large plugs are called slugs.
Slugs exiting the pipeline can overload the gas/liquid handling capacity of the plant at the pipeline outlet, as they are often produced at a much larger rate than the equipment is designed for.
Slugs can be generated by different mechanisms in a pipeline:

  • Terrain slugging is caused by the elevations in the pipeline, which follows the ground elevation or the sea bed. Liquid can accumulate at a low point of the pipeline until sufficient pressure builds up behind it. Once the liquid is pushed out of the low point, it can form a slug.
  • Hydrodynamic slugging is caused by gas flowing at a fast rate over a slower flowing liquid phase. The gas will form waves on the liquid surface, which
  • may grow to bridge the whole cross-section of the line. This creates a blockage on the gas flow, which travels as a slug through the line.
  • Riser-based slugging, also known as severe slugging, is associated with the pipeline risers often found in offshore oil production facilties. Liquids accumulate at the bottom of the riser until sufficient pressure is generated behind it to push the liquids over the top of the riser, overcoming the static head. Behind this slug of liquid follows a slug of gas, until sufficient liquids have accumulated at the bottom to form a new liquid slug.
  • Pigging slugs are caused by Pigging operations in the pipeline. The pig is designed to push all or most of the liquids contents of the pipeline to the outlet. This intentionally creates a liquid slug.

Slugs formed by terrain slugging, hydrodynamic slugging or riser-based slugging are periodical in nature. Whether a slug is able to reach the outlet of the pipeline depends on the rate at which liquids are added to the slug at the front (i.e. in the direction of flow) and the rate at which liquids leave the slug at the back. Some slugs will grow as they travel the pipeline, while others are dampened and disappear before reaching the outlet of the pipeline.

Purpose of the slug catcher

A slug catcher is a vessel with sufficient buffer volume to store the largest slugs expected from the upstream system. The slug catcher is located between the outlet of the pipeline and the processing equipment. The buffered liquids can be drained to the processing equipment at a much slower rate to prevent overloading the system. As slugs are a periodical phenomenon, the slug catcher should be emptied before the next slug arrives.
Slug catchers can be used continuously or on-demand. A slug catcher permanently connected to the pipeline will buffer all production, including the slugs, before it is sent to the gas and liquid handling facilities. This is used for difficult to predict slugging behaviour found in terrain slugging, hydrodynamic slugging or riser-based slugging. Alternatively, the slug catcher can bypassed in normal operation and be brought online when a slug is expected, usually during pigging operations. An advantage of this set-up is that inspection and maintenance on the slug catcher can be done without interrupting the normal operation.

Slug catcher design

Slug catchers are designed in different forms:

  • A vessel type slug catcher is essentially a conventional vessel. This type is simple in design and maintenance.
  • A finger type slug catcher consists of several long pieces of pipe (‘fingers’), which together form the buffer volume. The advantage of this type of slug catcher is that pipe segments are simpler to design for high pressures, which are often encountered in pipeline systems, than a large vessel. A disadvantage is that its footprint can become excessively large. An example of a large finger-type slug catcher can be seen in Den Helder, The Netherlands, using Google Maps.

A basic slug catcher design contains the buffer volume for gas and liquid. A control system is used for controlled outflow of gas and liquid to the downstream processing facilities. The inlet section is designed to promote the separation of gas and liquid.

Molecular Sieves
Molecular sieve adsorbents were developed to remove water from liquids and gases more efficiently than the older silica gel and alumina desiccants. Molecular sieves, because of their crystalline composition, will yield virtually water-free products. This feature makes molecular sieves particularly useful in cryogenic operations where liquefaction of gases is required and water must be eliminated to avoid freezing.
Molecular sieves offer adsorption selectivity based on molecular size, molecular affinity for the sieve crystal surface, and shape of the molecule. Molecular sieve systems are extremely simple to use for reducing contaminants to nearly undetectable levels compared to other technologies such as liquid separation systems. This, along with the higher efficiency compared to other adsorption systems, has made molecular sieve systems a popular engineering choice.
Less than 30-35%RH, these are effective.
Using a dry purge gas at an inlet temperature of 200-350°C and a pressure of 0.3 ~
0.5 kg/cm², heat the molecular sieve bed. As the temperature in outlet reaches 150 ~ 180°C, let the bed cool off.
Handling and Storage: Product must be stored in a dry location (relative humidity < 90%) to prevent re-adsorption of water. Opened packages should be resealed to prevent contamination of the product. Whenever possible, rotate stock to use oldest material first.
Types of the Moleculer sieves beds

13X Molecular Sieve 4×8

Applications: Multi-purpose. Used for the dehydration and purification of various hydrocarbon and non-hydrocarbon gas and liquid streams.

Size Qualification ≥ 95.0%
Water Adsorption ≥ 25.5 mg/g
CO2 Adsorption ≥ 18.0 mg/g
Compressive Strength ≥ 25.0 n
Bulk Density ≥ 0.64 g/ml
Wear Rate ≤ 0.30% weight
Package: Residual Moisture ≤ 1.5% weight
Country of Origin China
Packaging 140 kg steel drum
(Supersacks Available)


13X Molecular Sieve 8×12

Applications: Multi-purpose. Used for the dehydration and purification of various hydrocarbon and non-hydrocarbon gas and liquid streams.

Size Qualification ≥ 95.0%
Water Adsorption ≥ 25.5 mg/g
CO2 Adsorption ≥ 18.0 mg/g
Compressive Strength ≥ 25.0 n
Bulk Density ≥ 0.64 g/ml
Wear Rate ≤ 0.30% weight
Package: Residual Moisture ≤ 1.5% weight
Country of Origin China
Packaging 140 kg steel drum
(Supersacks Available)


13X Molecular Sieve APG

Application: 13X Molecular Sieve APG is used commercially for air plant feed purification (simultaneous removal of H2O and CO2 ) and general gas drying.

Properties Measure Unit Spherical Cylindrical
Diameter mm 1.6-2.5 3-5 1.6 3.2
Static Water Adsorption ≥ mg/g 270 270 270 270
CO2 Adsorption ≥ mg/g 180 180 180 180
Bulk Density g/ml 0.67±0.02 0.66±0.02 0.62±0.02 0.62±0.02
Crush Strength ≥ N 30 80 30 65
Wear Rate ≤ %wt 0.1 0.1 0.2 0.1
Package Moisture ≤ %wt 1.0 1.0 1.0 1.5


13X Molecular Sieve HP

Application: 13X Molecular Sieve HP is used commercially for air plant feed purification and PSA oxygen generation process.

Properties Measurement Unit Spherical
Diameter mm 0.5-1.0 1.6-2.5
Static Water Adsorption ≥ mg/g 285 285
CO2 Adsorption ≥ mg/g 190 195
Bulk Density g/ml 0.64±0.02 0.62±0.02
Crush Strength ≥ N 20
Wear Rate ≤ %wt 1.0 1.0
Package Moisture ≤ %wt 1.0 1.0


13X Molecular Sieve PG

Application: 13X Molecular Sieve PG is used commercially for desulfurization of liquefied petroleum gas.It is also used for dehydration of gases and liquids.

Properties Measurement Unit Spherical Cylindrical
Diameter mm 1.6-2.5 3-5 5-8 1.6 3.2
Static Water Adsorption ≥ mg/g 260 260 260 260 260
CO2 Adsorption ≥ mg/g 175 175 175 175 175
Bulk Density g/ml 0.67±0.02 0.66±0.02 0.64±0.02 0.62±0.02 0.60±0.02
Crush Strength ≥ N 25 80 120 25 64
Wear Rate ≤ %wt 0.1 0.1 0.1 0.2 0.2
Package Moisture ≤ %wt 1.0 1.0 1.0 1.5 1.5


13X Activated Powder

Application: 13X Activated Powder is available for dehydration in coating industry and paint industry.

Properties Measurement Unit Specification
Form Powder
Pore Diameter X 13
Whiteness ≥ % 90
Particle Diameter μ m 1-4
Static Water Adsorption ≥ %wt 260
PH 11.5
Gravity ≥g/ml .33
Package Moisture ≤% Wt 3.0

Molecular sieves are Alumino Silicates, which have a uniform, crystalline
structure.  It is this structure, which is used to “filter” out molecules of contaminants.  The sieves are produced in pellets of different shapes and sizes to suit the application and bed weight.
 Regeneration (activation)
Regeneration in typical cyclic systems constitutes removal of the adsorbate from the molecular sieve bed by heating and purging with a carrier gas. Sufficient heat must be applied to raise the temperature of the adsorbate, the adsorbent and the vessel to vaporize the liquid and offset the heat of wetting the molecular-sieve surface. The bed temperature is critical in regeneration. Bed temperatures in the 175-260° range are usually employed for type 3A. This lower range minimizes polymerization of olefins on the molecular sieve surfaces when such materials are present in the gas. Slow heat up is recommended since most olefinic materials will be removed at minimum temperatures; 4A, 5A and 13X sieves require temperatures in the 200-315 °C range.
After regeneration, a cooling period is necessary to reduce the molecular sieve temperature to within 15° of the temperature of the stream to be processed. This is most conveniently done by using the same gas stream as for heating, but with no heat input. For optimum regeneration, gas flow should be countercurrent to adsorption during the heat up cycle, and concurrent (relative to the process stream) during cooling. Alternatively, small quantities of molecular sieves may be dried in the absence of a purge gas by oven heating followed by slow cooling in a closed system, such as a desiccator.
Molecular Sieves are crystalline zeolites that have been activated for adsorption by removing their water of hydration.  During this dehydration, highly porous adsorbents are formed that have strong affinity for water and certain other gases and liquids.
The pores of any particular type of Molecular Sieve are precisely uniform in size and of molecular dimensions.  Depending on the size of these pores, molecules may be readily adsorbed , slowly adsorbed or completely excluded.
Molecular Sieves have been used in many commercial and industrial systems for
drying and purifying liquids and gases.  Also, these adsorbents have made possible the development of large scale separation processes used to recover normal paraffins from branched chain and cyclic hydrocarbons.
Molecular Sieves can be used in the following types of adsorption systems:
Multiple-Bed Adsorption.  This method can be used for the majority of commercial large scale fluid purification operations.  Conventional fixed bed, heat regenerated adsorption systems are generally utilized.  A typical dual-bed installation places one bed on-stream purifying the fluid while the other is being heated, purged and cooled.
Single-Bed Adsorption.  When interrupted product flow can be tolerated, it is sometimes convenient to use a single adsorption bed.  When the adsorption capacity of the bed is reached, it can be regenerated for further use or discarded, depending on process economics.
Static Adsorption.  When formed into wafers or other shapes, Molecular Sieves can be used as static desiccants in closed gas or liquid systems.
The basic types of Molecular Sieves are the 3A, 4A, 5A and 13X sieves.  Each of these are available in the form of beads of 4 x 8 (2 mm) and 8 x 12 (4 mm) mesh sizes.

 An integrated deethanizer and ethylene fractionation column and process for separating a feed stream comprising ethylene, ethane and C3+ is disclosed. A single shell houses a refluxed upper portion and a lower portion of the column. A generally vertical wall partitions the lower portion of the column into a deethanizer section and an ethylene stripper section. The upper column portion is used as the absorption section of the ethylene fractionator. The feed is supplied to an intermediate stage in the deethanizer, and the deethanizer is operated at a lower pressure (and correspondingly lower temperature) matching that of the ethylene fractionation. The design allows the use of one slightly larger column in place of the two large columns previously used for separate deethanization and ethylene fractionation.

The present invention relates to an apparatus and method for the deethanization and ethylene fractionation in an olefin plant processing propane and heavier feedstocks, and particularly to the use of an integrated column which combines both the deethanizer and ethylene fractionator into a single column.
A typical process for the separation and recovery of olefins from pyrolysis furnaces operated with feedstocks heavier than ethane, is known as the front end depropanizer and front end acetylene hydrogenation scheme. A brief review of the typical front end depropanizer process is in order.
there are three stages of conventional compression to raise the pressure of the process gas from just above atmospheric to a pressure of about 15 bars (210 psia). Condensed liquids, i.e. hydrocarbons and water, are separated.
The gas is then treated in a conventional two or three stage caustic wash tower 10  for the removal of carbon dioxide and hydrogen sulfide. The gas is cooled and mildly chilled before entering the dryers. Water is removed completely.
The gas is then further chilled in propylene refrigerant exchanger, and enters the high pressure depropanizer  which does not really operate at high pressure but is only called that because there is also a low pressure depropanizer. The high pressure depropanizer typically operates at a pressure of 12 bars (170 psia), and the low pressure depropanizer 20 at a pressure of 8.5 bars (120 psia).
The overhead of the high pressure depropanizer  is usually compressed in compressor to a pressure of 38 bars (550 psia) and is then sent to the acetylene hydrogenation system which typically consists of two or three adiabatic reactors in series with inter-cooler for the removal of the heat of reaction. The reactor effluent is cooled in cooling water exchanger 26 and partially condensed in propylene refrigerant exchanger. A portion of the condensate is used as reflux via line  for the high pressure depropanizer. The rest is sent to the demethanizer stripper  via line.
In the stripping section  of the high pressure depropanizer  only ethane and lighter components are removed, resulting in a fairly low bottoms temperature of 56° C. (133° F.). The bottoms product is sent via line  to the low pressure depropanizer  where it is separated into C 3 ‘s and C 4 +. The C 3 is used as reflux in the high pressure depropanizer  via line, while the C 4 + is sent to the debutanizer (not shown) via line. Due to the low operating pressure, the bottoms temperatures in the depropanizers are quite low, namely 56° C. (133° F.) and 71° C. (160° F.). Therefore, there is no fouling in either tower or their respective reboilers.
The acetylene hydrogenation unit  is highly efficient and selective. The acetylene removal easily results in acetylene concentrations of less than 1 ppm in the final ethylene product while the ethylene gain amounts to 50% or more of the acetylene. Due to the high hydrogen content of the feed gas, no carbonaceous material is deposited on the catalyst. The catalyst needs no regeneration and thus the reactors  need no spares. Green oil formation is miniscule.
In the acetylene hydrogenation reactor  about 80% of the methyl-acetylene and 20% of the propadiene are converted to propylene. If the olefins plant produces polymer grade propylene the remaining C 3 H 4 can be easily fractionated into the propane product; the high conversion of methyl-acetylene and propadiene in the acetylene hydrogenation reactors obviates the need for an additional separate C 3 H 4 hydrogenation system.
The operational stability of the acetylene hydrogenation reactor  is enhanced by its location in the gross overhead loop of the depropanizer and in the minimum flow recycle circuit of the fourth stage of compression. These factors reduce the acetylene concentration in the inlet to the reactor  and stabilize the flow rate irrespective of the furnace throughput.
The vapor and liquid from the reflux accumulator of the high pressure depropanizer flow to the chilling and demethanization section). The liquid plus the condensate formed at -37° C. (-35° F.) is sent via respective lines  and  to the demethanizer stripper . The overhead vapor from the demethanizer stripper  plus the liquids formed at lower temperatures are sent to the main demethanizer  via respective lines  and . The tower  is reboiled by reboiler  with condensing propylene refrigerant, and reflux is condensed in heat exchanger  with low temperature ethylene refrigerant.
The respective bottoms products of the two demethanizers , after some heat exchange which is not shown, enter the prior art deethanizer . The tower  recovers approximately 40 percent of the ethylene contained in the two feeds as high purity product. Sixty percent of the ethylene and all the ethane leave the tower  as a side stream  and proceed to the low pressure ethylene fractionator . The deethanizer  is reboiled by reboiler  with quench water and reflux is condensed in exchanger  with -40° propylene refrigerant. The bottoms product  of the deethanizer  is a stream containing propylene, propane and the remaining C 3 H 4 . It flows to a conventional propylene fractionator Because of the ethylene fractionation in its top section , the deethanizer  has fifty more trays than a conventional deethanizer which produces a mixed ethylene and ethane overhead product in line .
The ethylene fractionator  is a relatively low pressure tower typically operating at 4 bars (60 psia) with approximately 100 trays. It uses an open heat pump. Ethylene refrigerant is condensed in the reboiler  and is then used as reflux via line . Effectively, the reboiler  also serves as the reflux condenser. There are no reflux pumps and there is no reflux drum.
The present invention combines the deethanizer and ethylene fractionator of the prior art into a single fractionation column, reduces the pressure of the deethanizer to that of the ethylene fractionator and locates the deethanizer and the stripping section of the ethylene fractionator in the bottom portion of a single distillation column divided by a vertical wall. Locating the deethanizer and the stripping sections of the ethylene fractionator in the bottom section of a single distillation column divided by a vertical wall has the capital cost savings of replacing two large columns with a slightly larger column; eliminates the deethanizer reflux condenser, drum and pumps; and employs a much smaller deethanizer reboiler.
In one aspect, the present invention provides an integrated deethanizer and ethylene fractionation column for separating a feed stream comprising ethylene, ethane and C 3 + into an ethylene stream, an ethane stream and a C 3 + stream. The integrated column is made of a single shell housing a refluxed upper portion and a lower portion. Each of the integrated column portions comprise multiple vapor-liquid contacting elements. A generally vertical wall partitions the lower portion into a deethanizer section and an ethylene stripper section. A feed line supplies at
least one feed stream to at least one feed stage of the deethanizer section of the lower portion of the column, between a plurality of absorption stages above the feed stage and a plurality of stripping stages below the feed stage, for producing an overhead vapor stream from the deethanizer section consisting essentially of ethylene and ethane and a bottoms stream consisting essentially of C 3 and heavier components. A distribution pan with vapor chimney(s) at the lower end of the upper portion of the column facilitates passage of vapors from the deethanizer and ethylene stripper sections into the upper column portion, and collects liquid for passage from the upper portion of the column into the upper stage of the deethanizer section and into the ethylene stripping section.
The integrated column can have a deethanizer section comprising from 20 to 60 trays. The upper and lower portions of the integrated column preferably have the same cross-sectional diameter. The integrated column can also include a reboiler for the deethanizer section heated with high pressure depropanizer gross overhead or some other suitable heating medium. The integrated column can also include a reboiler for the ethylene stripper section, heated by ethylene condensed at a relatively higher pressure than the integrated column. The integrated column preferably includes a line for refluxing the upper portion of the column with the ethylene condensed in the ethylene stripper section reboiler. The integrated column can include a compressor for compressing overhead vapor from the upper portion of the column to the pressure of the ethylene stripper section reboiler. The integrated column preferably comprises respective liquid lines from the distribution pan to the tops of the dethanizer section and the ethylene stripper section. The line from the distribution pan to the top of the deethanizer section can include a valve for controlling the amount of liquid supplied to the deethanizer section. The integrated column preferably has an operating pressure of from 2 to 20 bars (30 to 290 psia).
Process Description
The debutanizer column considered in this study .the LPG can be produced with high WLPG values (up to 15oC) and sent to the FCC unit to be reprocessed. Depending on the adopted operating strategy, completely different process conditions are obtained and consequently the system follows different models. Then, a robustness problem is created and the control system must deal with it. The schematic representation of the debutanizer is shown in Fig.3. The feed stream comes from the top of the crude pre-flash column at a temperature of 40oC and pressure of 8 kgf/cm2. It is pre-heated in the heat exchanger E-01 where the hot fluid is the stabilized gasoline that leaves the bottom of the debutanizer at approximately 163 oC. At the exit of E-01 the temperature of the feedstream is approximately 136oC and it is 6% vaporized. The feed is introduced in the debutanizer column (T-01) at stage 17. The column has 30 stages with a Murphree efficiency of approximately 75%. Measured points of the temperature profile are the top (about 54oC), bottom (about 163oC) and stage 5 (about 67oC). The top condenser has a heat duty of approximately 1.0 MMcal/h that corresponds to a distillate flow of about 108 m3/d and a reflux flow of about 400 m3/d. A small flow of fuel gas is also produced.

In the control scheme adopted in this study the manipulated variables are the reboiler heat duty (QREB) and the reflux flow rate (VRT) which operating window is represented in Fig.4. This operating window results from several process constraints that are defined in Table 1.

LPG Process Equipments

  • Gas-gas exchanger
    1. Introduction
    2. Conventional gas-gas heat exchanger
  • Radial floe heat exchanger
  1. Plate type heat exchanger
  2. Hot exchanger
  3. Cold exchanger
  • Turbo expander
    1. Applications
      1. Extracting hydrocarbon liquids from natural gas
      2. Power generation
      3. Refrigeration system
  • Slug catcher
    1. Basic operation
    2. Slugs
  • Purpose of slug catcher
  1. Slug catcher design
  • Molecular sieves
    1. Types of molecular sieve beds
    2. Regeneration
  • Applications
  • De-ethanizer


LPG Process Equipments