VENEZUELAN EXTRACTION PLANT EXPANDING TO 1 BCFD RATING

Dec. 3, 1990
Henry Jimenez-Gomez Corpoven S.A. San Joaquin, Venezuela Modification of the two process trains at Corpoven's San Joaquin, Venezuela, extraction plant will increase each train's sustained capacity to 500 MMscfd and 1 bcfd total plant capacity. The San Joaquin 1000 project, set for completion in 1991, will represent an increase in NGL production of 8,000 b/d, with a gross revenue of $25 million/year for Corpoven S.A., a subsidiary of Petroleos de Venezuela.
Henry Jimenez-Gomez
Corpoven S.A.
San Joaquin, Venezuela

Modification of the two process trains at Corpoven's San Joaquin, Venezuela, extraction plant will increase each train's sustained capacity to 500 MMscfd and 1 bcfd total plant capacity.

The San Joaquin 1000 project, set for completion in 1991, will represent an increase in NGL production of 8,000 b/d, with a gross revenue of $25 million/year for Corpoven S.A., a subsidiary of Petroleos de Venezuela.

The San Joaquin extraction plant currently consists of two identical process trains, each designed to handle 400 MMscfd of rich gas. Original design also provided for future conversion for 70% ethane recovery with the installation Of CO2-removal units for the inlet gas.

Because actual inlet gas has been much leaner than design, and excess capacity has existed in some vessels resulting from considerations for the 70% ethane-recovery case, true capacity of the plant is considerably higher than design.

In addition, stream tests have shown peak processing capacity approaching 500 MMscfd/train.

EASTERN CRYOGENIC COMPLEX

In November 1989, Venezuela's eastern cryogenic complex completed its fourth year of operation. The cryogenic complex consists of the San Joaquin extraction plant, the fractionation plant and marine terminal at Jose, and interplant pipelines.

Design of the San Joaquin extraction plant was as follows:

  • Capacity of 800 MMscfd in two identical process trains, each train handling 400 MMscfd via two turboexpanders operating in series.

  • Inlet gas composition of 74.88 mol % methane and 6.01 MOI % CO2 with 3.08 gpm C3 +.

  • Propane recovery of 90% in the inlet-gas stream and provision in the towers and other vessels to handle flows for the 70% ethane-recovery case.

This case requires additional CO2-removal units for the inlet gas, several heat exchangers, and pipeline modifications. This future expansion will depend primarily on the development of Venezuela's petrochemical industry.

Changes in gas from fields used as design criteria and development of new gas fields resulted in a substantially leaner inlet gas: 78.44 mol % methane and 7.10 mol % carbon dioxide with 2.30 gpm C3 +.

Today's configuration includes two turbocompressors per train capable of handling 250 MMscfd each with ethane rejection for actual 94% of propane recovery.

This additional available compression capacity is the main justification for increasing gas flow and processing capacity of the extraction plant.

In addition to the fact that leaner inlet gas implied a potential increase in gas-processing capacity, Corpoven's management initiated a program in 1988 to find this new gas processing limit.

During the development of this project, gas flow to the extraction plant was increased a little at a time, and the effects of each increment were evaluated with respect to product recovery, pressure drops, and other parameters. The test was conducted simultaneously in both process trains over a period of several months.

Results showed the actual sustained processing capacity of the extraction plant to be approximately 900 MMscfd. The test also revealed the first process bottleneck to be the glycol contactors at the inlet to each process train.

At this point, Corpoven initiated a plan to eliminate the remaining bottleneck that would permit the plant's sustained processing capacity to be increased to at least 1 bscfd.

GAS SEPARATION, DEHYDRATION

The process description is shown in Fig. 1.

Inlet gas from slugcatchers outside the battery limits is pressure controlled into four scrubbers at the inlet to the extraction plant. These inlet scrubbers are three-phase separators that operate at 100 F. and 990 psig.

Any free water separated from the gas streams is drained from the bottom of the scrubbers. Separated condensate is pressured to the stabilization section of the process trains.

Vapors from the scrubbers combine in a common 30-ft header and flow to the two process trains. Flow through each train is determined by the speed of the residue-gas compressors associated with that train.

Feed gas from the inlet-gas scrubbers combines with a gas stream from the stabilizer-overhead compressors, flows through a filter separator, and enters the bottom of the glycol contactor.

As gas flows upwards through the contactor against a counter current of 98% triethylene glycol (TEG) solution, it is dehydrated to a water dew point of approximately 20 F. The dehydrated gas leaves the TEG contactor at 110 F. and 980 psig.

Dehydrated gas leaving the contactor is cooled to 54 F. in four heat exchangers operating in parallel.

One stream is cooled to 45 F. by exchange with cool residue gas. Another stream is cooled to 45 F. by exchange with condensed liquid from the high-pressure expander outlet separator.

The third and fourth streams are cooled to 62 F., one by providing reboil heat to the de-ethanizer side reboiler, and the other by exchange with condensed liquid from the high-pressure expander inlet separator.

Streams then recombine and flow to the high-pressure expander inlet separator where the condensed liquid separates from the vapor. This separator operates at 54 F. and 970 psig.

The condensed liquid flows from the separator under level control and is heated to 90 F. by exchange with the warm gas stream from the TEG contactor. Vapor liquid mixture from the exchanger flows to the stabilizer-feed flash tank.

HIGH-PRESSURE EXPANSION

Gas from the expander inlet separator flows through the high-pressure expander, where its pressure is reduced to 710 psig and its temperature to 27 F. Energy extracted from the expander is used to drive the high-pressure expander-compressor.

An expander bypass (Joule-Thomson valve) is provided so that the plant can be operated when the high-pressure expander is down.

Liquid condensed in the expander is separated from the gas in the high-pressure expander outlet separator. This separator operates at 27 F. and 710 psig.

Liquid flows from the separator and is heated to 85 F. by exchange with the warm gas stream from the TEG contactor. Vapor-liquid mixture from this exchanger also flows to the stabilizer feed flash tank. Gas from the outlet separator flows to the molecular sieve dehydration unit.

Gas from the high-pressure expander flows through a filter separator, then through the mol-sieve dehydrators where the water content is reduced to less than 0.1 ppm (volume).

The system consists of four adsorption towers; three are in service at any time while the fourth is being regenerated. The dehydrated gas from the beds flows through a dust filter where fine mol-sieve dust is removed.

Gas from the mol-sieve dehydrators is cooled to -43 F. in three heat exchangers operating in parallel. One stream is cooled to -30 F. by exchange with cold residue gas.

The other two streams are cooled to -55 F. by exchange with condensed liquids in the de-ethanizer high-pressure and low-pressure feed preheaters. Streams then recombine and flow to the low-pressure expander inlet separator. This separator operates at -43 F. and 690 psig.

LOW-PRESSURE EXPANSION; DE-ETHANIZATION

Liquid from the separator is heated to -4 F. by exchange with the gas leaving the dehydrators and is then fed to the de-ethanizer on Tray No. 12. Gas from the separator flows through the low-pressure expander where its pressure is reduced to 170 psig and its temperature to -116 F.

Energy extracted from the expander is used to drive the low-pressure expander-compressor.

An expander bypass (Joule-Thomson valve) is also provided for the low-pressure expander.

Liquid condensed in the expander is separated from the gas in the low-pressure expander outlet separator. This separator operates at 116 F. and 170 psig. Vapor from the separator provides refrigeration to the de-ethanizer reflux condenser.

Liquid is pumped through the de-ethanizer low-pressure feed preheater where it is heated to -9 F. by exchange with gas leaving the dehydrators, then flows to the de-ethanizer of No. 5.

The de-ethanizer contains 28 valve trays and operates at 172 psig.

The column has one side reboiler and one bottom reboiler.

Reboil heat for the side reboiler is provided by inlet gas. Gas from residue gas compressors is used to provide heat to the bottom reboiler. A partial condenser is used to provide reflux to the column.

Cold gas from the de-ethanizer overhead combines with gas from the low-pressure expander, This combined residue gas stream provides refrigeration to the inlet gas, as discussed previously.

Liquid bottom product from the de-ethanizer is pumped to the pipeline that delivers the NGL product to the Jose fractionation facility.

COMPRESSION; CONDENSATE STABILIZATION

Residue-gas pressure is boosted to 190 psig in the low-pressure expander-compressor and to 220 psig in the high-pressure expander-compressor. The gas is then recompressed to 950 psig in two gas-turbine-driven centrifugal compressors operating in parallel.

Discharge of the residue-gas compressors combines with the residue gas from the second train and flows to Corpoven's gas-distribution pipeline.

The stabilization system serves to remove water and light hydrocarbons from liquids condensed in the high-pressure sections of the plant. These condensed liquids flow through a common line to the stabilizer-feed flash tank.

Operating at 69 F. and 410 psig, this tank is a three-phase separator in which any entrained water or glycol can be removed from the system. Gas from the flash tank flows to the stabilizer-overhead compressor, while the liquid is fed to the top of the stabilizer column.

The stabilizer contains 21 valve trays and operates at 389 psig. Reboil heat is provided by a hot-oil system. The stabilizer-overhead vapor combines with the vapor from the flash tank. This vapor is compressed to 995 psig in the stabilizer-compressor. Gas from the overhead compressor is cooled to 120 F. and returned to the inlet-gas stream upstream of the TEG contactor. Stabilizer bottom liquids are pumped to the NGL product pipeline along with the de-ethanizer-bottom liquids.

MODIFICATIONS-INLET SECTION

To analyze the results, let us look at the plant starting from the front end:

Excess capacity available at the inlet scrubbers is 170 MMscfd. The inlet scrubbers can therefore easily handle 150 MMscfd to bring the total to 1 bscfd (Table 1).

From calculations it was found that the El Toco-Santa Ana 20-in. line inlet scrubber and the Santa Rosa 20-in. line inlet scrubber, currently sized at 102 in., can move a maximum of 375 MMscfd each. Hence, if an increase in the capacity is to come from the 20-in. pipelines, serious considerations must be given to paralleling or unloading either of these inlet scrubbers (Table 2).

Therefore, an additional inlet scrubber should be installed in order to allow more than 1 bscfd gas flow avoiding a higher pressure drop at the front end.

MODIFICATIONS-GLYCOL SECTION

The inlet-filter separator is extremely crucial for the successful operation of the glycol contactor in terms of foaming-tendency prevention. It cannot be removed from operation or bypassed.

However, currently it is seen that it will develop pressure-drop problems because of high flows. The San Joaquin 1000 project has plans to install a parallel filter to alleviate this situation, and it will also provide an excellent backup to the existing filter.

For higher volumes, however, both filters must be in operation for one train.

The glycol contactor at 500 MMscfd will present a significant entrainment and glycol-loss problem as a result of carryover from high gas velocities in the contactor.

The tower is capable of handling up to a maximum flow of 440 MMscfd after which serious glycol losses will occur. This contactor has been evaluated with data from prominent tray manufacturers. The recommendation of this evaluation was that vapor flood was excessive and free flow vapor area is insufficient.

Reconfiguring the existing column by retraying with structured packing or other types of devices normally helps if liquid flooding problems exist. Liquid flooding is not seen to account for this situation. Therefore, the internals of the towers will not be changed.

In order to alleviate this bottleneck effectively, a parallel additional column sized for 100 MMscfd will be added to the existing column. This column will help slow gas velocities in both columns so that more efficient mass transfer is performed with good gas dehydrated and minimum glycol carryover and entrainment.

The dust filter separator located after the mol-sieve beds has excessive pressure drop at higher flows. An identical filter-separator will be installed in parallel with the original one, and both filters will be operational.

When filter elements need to be changed, however, there is no need for a plant shutdown. This will provide flexibility as well as ensure that excessive pressure drop does not jeopardize the low-pressure expander's performance.

TURBOEXPANDER SECTION

Since the exchangers in the original plant are not uniformly oversized, imbalances exist in the surface-area requirements for a heat-transfer point. For extraction of maximum cooling from the residue gas, a way must exist to force the inlet gas to flow in a certain direction.

This is normally achieved by a trim valve located in the inlet-gas lines after the flow splits. Currently, these valves are not installed for cooling after the mol-sieve bed. They provide better flow control on gas-flow splits, both in the critical cold gas/gas section.

Inlet gas being fed to the San Joaquin plant is becoming leaner as time passes; CO2 content will rise to 7.1 mol %. The effectiveness of the high-pressure expander had been evaluated which led to the conclusion that it does not contribute to any significant production of the liquids (Table 3).

The low-pressure expander as it is currently being operated represents a serious bottleneck to higher flow volume operations. The high-pressure expander must be converted to operate in parallel with the low-pressure expander. And, at the location of the high-pressure expander, a Joule-Thomson valve, a high-stress operation valve, must be installed.

In order not to bottleneck the plant, another Joule-Thomson valve must be placed in parallel to the current valve.

The expander will be modified as two 50% operation units (Fig. 2). The rest of the plant will be capable of handling the new flow volumes.

The residue-gas compressors are adequate to handle the lower suction pressure presented to them caused by the loss of horsepower from the high-pressure expander. And they should be able to process all the gas at higher volumes and lower suction pressures (Table 4).

MODEL PREDICTIONS

The San Joaquin plant was simulated with the original design conditions. The model predicted the actual NGL yields and CO2 freezing levels very well.

This model then was used to simulate increased flows flowing through the plant. The plant was simulated extensively with various compositions and varying levels of CO2 in the inlet gas.

Analysis indicated that the plant is capable of recovering more than 90% of the propane contained in the incoming feed. Table 5 shows the result of the recoveries.

The chief problem in recovering more than 90% is related to the CO2 content of the inlet gas. As the CO2 content rises, the freezing becomes problematic.

The main advantage for the low-pressure expander configuration is the possibility of handling 400 MMscfd/train even in case of a temporary failure of one low-pressure expander.

The total project, the major debottleneck modifications to be done on the inlet facilities, glycol, and turboexpander sections as described, will be completed by early next year. This project is being developed in three phases per train to reduce plant shutdowns and in accordance with the delivery time of the additional key equipment.

Total cost of the project has been estimated at close to $12 million and will represent an additional NGL production of 8,000 b/d, with a current market value of $25 million/year.

With the completion of this second project, Corpoven will have improved the design of the original plant, which was started in November 1985, by increasing propane recovery from 90% to 94% and gas-processing capacity from 800 to 1 bcfd.

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