Document OwVEBeb9gX6gkoor0o5x518p

PROCESS ENGINEERING DEPARTMENT TRAINING PROGRAM - 1975 GAS PROCESSING DIVISION OCTOBER 29, 1975 SAL 001C460 PROCESS ENGINEERING DEPARTMENT TRAINING PROGRAM - 1975 GAS PROCESSING DIVISION - OCTOBER 29, 1975 I. HISTORY The Gas Processing Division became a part of the Process Engineering Department in the Spring of 1969 under the direction of Mr. B. H. Williams. Mr. Williams had many years of experience in the gas processing industry and set out to train a nucleus of engineers to serve as a base for future growth. In August of 1973, Mr. Williams transferred to Production Engineering Services in Houston. This division's primary function began with feasibility studies and the engineering of new plants for the Natural Gas and Gas Products (NG&GP) Department in Houston. This work includes the engineering for Conoco's domestic gas plants which are listed in Table I on page 15. During 1970 we became involved in our first project for the Conoco Europe Production Department. This project developed as a result of the Conoco Viking Gas Field discovery in the North Sea. The project consisted of feasibility studies, plant design, working with the Ralph M. Parsons Company, detail design, and plant startup. The Theddlethorpe, England, terminal has been in operation since August of 1972 producing speci fication gas for sale to the British Gas Corporation and a stabilized condensate which is pipelined some 30 miles to the Humber Refinery. In 1972 we became involved through Production Engineering Services (PES) in Dubai Gas Lift Project. This facility is now in operation and Mike Morgan will relate some of his experience on this project later in the session. This describes our three main sources of work: (1) NG&GP in Houston, (2) PES in Houston, and (3) Conoco Europe Pro duction Department in London. Currently 75 percent of our work load is for the production departments which is a significant shift from our early work which was mainly for domestic gas plants. Our current manpower level is thirteen engineers. We will attempt to give you a feel for the types and processes we deal with by describing several which we have completed or now have in progress. of plants projects ,,,, SAL 0COC1C461 Training Program - 1975 Page 2 II. PROJECTS A. Hamlin Gas Plant We participated in this project from the beginning conceptual work to the plant startup which was in March of 1971. This consisted of process selection, preliminary design, economics, basis of bid prepa ration, contractor selection, review of contractor's design, following of construction, and plant startup. The Hamlin plant processes a very rich gas associated with oil production in West Texas. The plant is a straight refrigeration design. A simplified flow sheet of the process is shown in Figure 1, page 16. The gas is compressed from 5 psig to 564 psig treated with a monoethanol amine (MEA) unit to remove the hydrogen sulfide and a portion of the carbon dioxide. Typical specifications for sales gas are 0.25 grain H2S per 100 SCF, seven pounds of water per MMSCF, and a minimum of 1,000 BTU per SCF heating value. The gas is cooled against residue gas and propane refrigeration to 67F. Ethylene glycol is injected ahead of the ex changers to prevent hydrate formation. The gas is eventually cooled to -40F by two levels of propane refrigeration and exchanged with the demethanizer over head residue gas stream. Glycol is recovered in the glycol separator, regenerated with hot oil, and again reinjected. The residue gas from the demethanizer overhead is warmed against incoming feed gas and metered to sales. The demethanized product is cooled and provides side reboiler heat to the demethanizer before being pressured to surge and pumped to 2,000 psig into the Chapparal pipeline system. Approximately 130,000 GPD of demethanized product is produced from 20 MM SCFD of gas which is about 6.5 GPM (gallons per thousand standard cubic foot). This is a very rich gas which practically dictates a straight refrigeration process. The plant recovers 50 percent of the incoming ethane. To recover this much ethane from a lean gas, such as Grand Chenier with a GPM of 1.4 or Acadia with 2.4 ethane plus, would dictate that a cryogenic turboexpander process be used. The turbo expander plant will be discussed in Section F. SAL 0C0C1C462 Training Program - 1975 Page 3 B. Theddlethorpe or Mablethorpe, England, Gas Terminal We were involved in this project essentially to the same extent as for the Hamlin project from the con ceptual studies to plant startup. The major difference being this plant was on a much larger scale designed to process over 900 MM SCFD of gas. Another major difference being that the plant is located on the coast of Great Britain with the contractor's design work being carried out in London. This latter fact complicates the following of the contractor's work. A team was established in the contractor's office with back-up provided from Ponca City. The Conoco Theddlethorpe terminal resulted from our gas find some 87 miles off the British Coast in the North Sea. Conoco and the National Coal Board of England are 50/50 partners in leases held by Conoco in the British sector of the North Sea. Figure 2, page 17 shows the overall scheme. Conoco operates two platform complexes "A" and "B" some seven miles apart. The "A" complex consists of a drilling platform (AD), a production platform (AP), and a riser platform (AR). The "B" complex consists of two platforms BD and BP. The complexes are con nected by a seven-mile, 24-inch submarine pipeline. The combined gas stream from both complexes flows to shore in an 87-mile, 28-inch submarine pipeline. The pipeline pressure at the offshore "A" complex is 1,900 psig taking about a 700 psi pressure drop to deliver 920 MM SCFD to shore. Other gas terminals are at Bacton and Easington as shown in Figure 2. I won't go into the process at Theddlethorpe, it is very similar to the Hamlin process except that the gas is only chilled to -10F. Specification gas is delivered to the British Gas Corporation at 1,000 psig. By specification gas, I mean a gas meeting a calorific value of 970-1,070 BTU per SCF, a water dew point of 15F, and a hydrocarbon dew point of 30F in the winter months. Condensate at the rate of about five barrels per MM SCF is stablized and pipelined to the Humber Refinery. Methanol is in jected offshore into the 28-inch pipeline to prevent hydrate formation. The methanol is recovered at the Theddlethorpe plant and pumped back offshore through a 3-inch piggyback pipeline. Training Program - 1975 Page 4 Unusual design problems in the plant included (1) the catching of liquid slugs from the 87 miles of pipeline which results from rate changes or from "sphering" the line to increase pipeline capacity, (2) off-design conditions resulting from the demand changes by the BGC and by the varying inlet gas temperature caused by sea temperature changes from winter to summer, (3) the retrograde characteristics of this lean gas resulting in a maximum hydrocarbon dew point at about 400 psig, (4) the design of the safety relief system to handle over a billion SCFD of gas connected to 87 miles of high pressure pipe line . Lf.AL ^00104 Training Program - 1975 Page 5 C. MALJAMAR ETHANE RECOVERY The Maljamar plant located in Lea County, New Mexico, illustrates a refrigerated oil absorption process. In December 1972, a modification of the existing plant was completed which consisted of adding a demethanizer and associated facilities to deliver an ethane-propane product to the Chaparral pipe line. A simplified flow sheet is shown in Figure 3, page 18. The incoming compressed gas is treated for H2S and CO2 removal by an MEA unit, cooled to about 80 F in the gas-gas exchanger, and further cooled to 10F in the propane chiller. The gas then enters the bottom of the absorber column where about 40 per cent of the ethane, 85 percent of the propane, and 100 percent of the butanes plus are absorbed in the lean oil. Residue gas leaves the top of the absorber. A portion of the cool rich oil feeds the top of the demethanizer. The majority of the rich oil is heated in the oil-oil exchanger before feeding the demethanizer. The demethanizer overhead is used to subcool propane refrigerant and then is recycled to the front-end of the plant. The demethanizer reboil media is hot lean oil coming back through the system from the hot oil still. The deethanizer operates at 200 psig and the heavier components are absorbed by cool 10F lean oil. The overhead ethane-propane product is condensed and pumped to the Chaparral pipeline. The tower is re boiled against hot lean oil. The deethanizer bottoms feed the still where the lean oil is regenerated and a raw product containing propane plus is taken over head. The Maljamar plant also has a debutanizer which splits the raw product into butanes and a natural gasoline stream. SfiL C000IC465 Training Program - 1975 Page 6 D. DUBAI GAS LIFT AND GAS UTILIZATION During 1972 we became involved in projects for the Conoco Dubai facilities in the Persian Gulf. The gas lift project involves the injection of 100 MM SCFD of associated gas into the formation. The purpose of such a facility is to increase oil pro duction by lowering the density of the fluid coming up the well string. This does not increase ultimate field recovery. Our part in this project consisted of reviewing specific areas of the facility designed by Brown and Root in Houston. We currently are involved in two additional gas lift projects in Dubai. The first is 15 MM SCFD gas in jection from the Fateh "M" platform and the second is 6-8 MM SCFD injection for each of three platforms "F," "I," and "0." The engineering for the "M" platform is now essentially complete with startup scheduled for the summer of 1976. We are now involved in vendor selection of compressors for the MF," MI," and "O" platforms. Startup for "F," "I," and "0" is scheduled for January 1977. Also, we have been involved with EHPD in the evaluation of a project to recover LPG from the gas being flared. This project consists of offshore compression, de hydration, 64-mile pipeline to shore, slug catcher, refrigeration, gas treating, liquid treating, frac tionation storage, metering, and shipping. Figure 4, page 19, shows the overall scheme of the offshore operation at Dubai. Figure 5, page 20, is a simplified flow diagram of one of the two trains now operating for the gas lift operation. Each com pression train is powered by a GE frame three gas turbine site rated at about 11,600 BHP. A triethylene glycol contactor is shown for dehydration of the gas to minimize corrosion and hydrate problems. The TEG unit will not be installed with the initial gas lift installation. SAL 000010466 Training Program - 1975 Page 7 E. VIKING FIELD COMPRESSION As the Viking gas field is depleted, compression will be required to maintain the gas flow through the 28-inch pipeline to the Theddlethorpe gas terminal. In the summer of 1973, we became involved in studies to determine if the compression should be located onshore or at the offshore production platforms. The decision was reached to locate com pression offshore. We were part of a project team assembled in September of 1973 to work in London with an English Contractor to design the offshore compression facility. The design phase was com pleted in September 1974, and the project moved to Rotterdam, Holland, for the construction phase. Construction for the "A" complex compression is now complete and the new platform has been lifted into place adjacent to the "AP" platform. Mr. K. A. Poise is currently in England assisting with the startup. SAL 0C001C4671 Training ?rogam - 1975 Page 8 F. HENNESSEY ETHANE RECOVERY During the fall of 1973, we began feasibilitystudies to upgrade the Hennessey refrigerated absorption plant to recover about 60 percent of the ethane. Following the process evaluations, it was decided to proceed with a cryogenic turbo expander plant to replace the absorption process. A basis of bids was prepared and a Contractor selected in July of 1974. The construction is complete and the plant is in operation. Messrs. Joe Provine and Ron Heldenbrand were involved in this project from its inception to operation. Conoco have two turboexpander plants in operation, one processing 250 MM SCFD of gas at Grand Chenier, Lousiana, and another processing 90 MM SCFD of gas at Acadia, Louisiana. Figure 6, page 21 shows a simplified flow sheet of the Hennessey plant. Gas enters the plant at 600 psig and 140F, the stream is cooled, liquids separated, and dehydrated to a -150F dew point in a dry bed dehydrator. The gas is cooled to -45F in a series of gas-gas ex changers and propane chillers. A portion of the feed gas is used to reboil the demethanizer. The liquid at -45F is separated and feeds to the middle of the demethanizer. The vapor at -45F is sepa rated and feeds to the middle of the demethanizer. The vapor at -45F flows through the first stage of the turboexpander, provides side reboil heat to the demethanizer, and enters the second stage flash sepa rator at 455 psig and -107F. The liquid from the flash separator flows to the upper portion of the demethanizer. The vapor is expanded in the second stage of expansion to 210 psig and -150F and feeds the top of the demethanizer. The overhead from the demethanizer is warmed against incoming gas to 115F, compressed to 250 psig by en ergy recovery from the turboexpander and finally com pressed by reciprocating compressors to the sales gas pressure of 600 psig. The liquids from the demetha nizer is deethanized, treated, and pumped to product pipeline. SAL 000010463 Training Program - 1975 Page 9 G. NORTH SEA OIL We all hear and read a lot about oil in the North Sea. What does this mean to the World, to the European countries with holdings in the North Sea, to Conoco, and last, but not least, to PED and the Gas Processing Division. Proven oil reserves in the North Sea are about 15 billion barrels with an estimated potential for the area of about 68 billion barrels. To place these numbers in perspective, they represent about three percent of the comparable values for the World, and about 25 percent of the values for North America. The reserves are large enough, however, to make the area self-sufficient in crude oil for a reasonable period of time. Conoco is very active in the North Sea. Extensive acreage holdings of the Conoco group (Conoco, Gulf, NCB) place the company in a very favorable position. At the present time, the company owns a portion of several oil fields that have been discovered in the northern part of the North Sea. These fields are (1) Hutton, (2) Thistle, (3) Dunlin, (4) Statfjord, and (5) a new unnamed field close to the Statfjord field. Some pertinent environmental data for this part of the North Sea: Sewater Depth, Ft Waves, Maximum, Ft Wind Velocity, Maximum, Ambient Temperature, Maximum, F Minimum. F Seawater Temperature, Maximum, dF Minimum, F MPH 500 100 150 75 15 60 40 (70 ft. Crest, 30 ft Trough) How does PED and the Gas Processing Division get involved in oil production in an environment like this? Perhaps, a brief discussion of some of the pertinent parameters, sizes, and costs can best answer this question. Training Program - 1975 Page 10 At the present time, the preferred platform, or structure, to support the drilling and producing operations is probably the concrete ''gravity" type of structure. The structure is called a "gravity" structure because it is so big and heavy it sits on the sea bed without piling to hold it in place. A typical dead weight for a structure of this type is about 350,000 tons. The base area resting on the sea bed is about 75,000 square feet. The load carrying capacity on top of the structure is about 40.000 tons. The base area of the deck is about 50.000 square feet, or, in other words, the deck is about the size of a football field. The cost of one of these structures is about $150,000,000. The cost of the structure is not the only major cost item. Forty or more wells may be drilled from the structure at a cost of about $3,000,000 for each well. The drilling period may last from three to five years. With the investment required for structure and wells, it becomes obvious that the oil production rate will have to be large to justify these expenditures. This results in large and expensive oil, gas, water, and sand separation equipment and corresponding large ancillary and utility system equipment. To provide an example, we will quote a few numbers from the design of the production facilities for a platform in the Statfjord field. Oil Producing Rate Gas Producing Rate Gas Compression Horsepower Compressor Discharge Pressure Seawater Lifted (Cooling and Injection) Seawater Injected (Waterflood) Electric Power Consumption Waste Water Treating Capacity Fuel Consumption, as Gas Oil Off-Loading Capacity No. of Personnel on Platform 250,000 270,000 50,000 5,000 20,000 BPD MSCFD EHP psig GPM 400,000 40 262,500 20,000 76,000 150 BPD @ 2,000 psig MW (50,000 HP) BPD MSCFD BPH Since this is a completely self-contained facility, all the system that we normally encounter must be designed and installed. This includes potable water, instrument air, fire protection, chemicals, vent and SAL 000010470 Training Program - 1975 Page 11 flare, liquid fuel, etc. The estimated cost of all of these facilities for this platform is $250,000,000. In the design of these facilities, we prefer to be involved at the preliminary design (conceptual) stage and then to work closely with the detailed design contractor and other members of the Conoco team mechanical, civil, instrument, electrical, structural, etc. - in the detailed design and implementation of the conceptual design. In addition to the above costs for structure, wells, and facilities which total about $500 MM we have to add facilities to transport the crude to a refinery and/or the gas to a gas processing plant. The nearest landfall from Block 211 is the Shetland Islands about 120 miles away. A subsea pipeline to the Shetlands is projected to cost about $300 MM. Tankers to transport the crude run about $60 MM each. Therefore, a project of this sort requires about $1 billion of capital to get into operation, not including the storage/refining and other land facilities. sal acooio^ti j Training Proqram - 1975 Pag<? 12 H. LNG STUDIES Conoco have large acreage leased in various remote locations of the World. There have been various gas finds at some of these locations, and because the areas are often long distances from commercial users, it may prove economical to liquefy the gas for trans port by LNG carriers. We have made several studies for potential LNG projects since 1971. There are only a few large base load LNG plants in operation from which to obtain actual plant data. The first to go onstream was the Arzew, Algeria, plant designed to transport the equivalent of 150 MM SCFD of liquefied gas to Canvey Island, England, and Le Havre, France. The project continues to operate and it utilizes the more efficient Cascade refrigeration process. In 1968 the Esso, Libya, LNG plant was scheduled to go onstream, but due to the revolution, a fire in the slug catcher, and several other opera tional and political difficulties, the plant did not begin shipping LNG until the Spring of 1971. The plant is designed to liquefy 350 MM SCFD of gas for shipment to France and Spain. The plant utilizes the mixed refrigerant cycle as designed by Air Products, Inc. The Phillips-Marathon, Kenai, Alaska, plant came onstream in 1970. It uses the Cascade refrigeration cycle and is designed to deliver the equivalent of 140 MM SCFD of liquefied gas to Japan. The largest LNG facility is the Shell Brunei plant which started up in late 1972. This plant produces LNG equivalent to 750 MM SCFD of gas for shipment to Japan. The Brunei plant typifies the enormous size of these plants. Some data from the original design to deliver 580 MM SCFD of LNG are as follows: 340.000 420.000 Three 500.000 horsepower refrigeration compression pounds-per-hour steam generation 10,000-kw generator GPM cooling water circulation Our most recent study looked at an LNG plant for on shore Malaysia to process about 400 MM SCFD of gas that had been pipelined from offshore. We worked 0C0"I04?2 Training Program - 1975 Page 13 with Procon, Inc., the Contractor for the Shell Brunei LNG plant, in developing an investment estimate and project schedule for our plant. The amount of capital involved in LNG projects is considerable. For example, a grass roots LNG plant processing 500 MM SCFD requires about $400 MM invest ment. In addition each LNG tanker, operating at -260F and carrying 800,000 barrels (130,000 m3) , may cost $110 MM. Considering there may be 2-6 tankers involved, depending on the distance, and perhaps 100+ miles of undersea pipeline, an entire LNG project from an offshore operation can easily exceed $1 billion. A reasonable place to begin in deciding where LNG projects might develop is to look at the proven gas reserves in various areas that have comparatively low energy needs. Following is a list of natural gas reserves for possible LNG export projects. Trillion Cubic Feet* Algeria Iran Nigeria Libya Saudi Arabia Venezuela Pakistan U.S.S.R. 100 197 40 30 51 32 25 735 The Soviet Union's 7 35 trillion cubic feet of proven reserves are nearly three times that of the U.S. At least two one-billion SCF LNG projects from Russia to the U.S. are under negotiation. The last page of drawings shows a modern LNG plant flow sheet. *Source: Fourth International Conference on Liquefied Gas, June 24-27, 1974. - ~ ~-s given all of you some idea of the --r the Gas Processing Division. Also a types of processes used in gas plants -~;hh of the equipment. tzz oak appears encouraging since much -racing on in the production area and of gas increasing to its true mar--r I think more projects will develop c gas plant area. Also, the strong --rtne non-flaring of gas from oil pro- --Jducing some activity which should j^irrow in the future. SAL ocooio^74 Training Program - 1975 Page 15 TABLE I MAJOR CONOCO DOMESTIC GAS PROCESSING PLANTS Location Grand Chenier* Grand Chenier* Gillis* Gillis* Bayou Long Port* Rincon* Woodlawn Ramsey Maljamar* Short Junction* Hennessey* Medford* Fruita* Sussex* Hamlin* Design Gas Capacity, MM SCFD 500 250 200 100 60 175 25 200 12 26 60 30 30 25 is 20 Type Process Refrigerated Oil Absorption Cryogenic Turboexpander Refrigerated Oil Absorption Adsorption Adsorption Refrigerated Oil Absorption Adsorption Refrigerated Oil Absorption Straight Refrigeration Refrigerated Oil Absorption Refrigerated Oil Absorption Refrigerated Oil Absorption Cryogenic Turboexpander Refrigerated Oil Absorption Adsorption Straight Refrigeration Straight Refrigeration `Plants PED has worked on since 1969. Note: Not all of these plants continue to operate at design capacity. NCR-ler 9-24-75 SAL 000010475 Page 16 Page 17 u. UJ 5 vi <w CM UJ o *2 fc * 3 0> --i oL- VJ c -i < a: ki o 5 U 4-260 S CONTINENTAL OIL COMPANY Page 18 < 61*7010000 TVS Page 20 C00104B0 rage S^L 000010^1 SAL 000010^83  v ' Vl %/ .V\t:-i - AL^v-**, o;-i -. ^ViV^* i^'iSi-^'^f^N-'-'-^"7: ifev f"^.',;'::;;w^*y^/'--' /;i'f?^>'J' ": SAL 000010484 CONTINENTAL OIL COMPANY I o CO o rH o o o o < IS) 47-5 BPD > "O 3 W25FCC O PECAMT OIL GA5 6,300 MCPD 6A50UNJE 450 BPD N4 FCC DECAMT OIL 7,730 BPD PREM. COKE OPER. WEED ePRAT 3,400 BPD CARBOki BLACK OIL 1,300 BPD TO KJ5 FCC VRU TO Kje 2 REFORMERS HD5 TO NJ2.-5 PCC TO COKJTIKEKTTAL CARBOkl BLACK. PREMIUM COKE 420 T/p r W! JWU NAPHTHA HDS AND REFORMER UNIT HDS The naphtha hydrodesulfurization catalyst is a bimetallic cobalt molybdena or cobalt-nickel. The reactions are: RSH + H2-->H2S + R-N -C = N-C- + H2 --* 2R + NH3 RC1 + H2 -->HC1 + R ROH + H2 --*RH + H20 C = C + H2--> C-C All reactions are exothermic. Reaction conditions are: 300 - 600 psig 250 psi hydrogen partial pressure 650 - 750F The catalyst also absorbes metallic contaminants (arsenic, lead, copper, etc). Reformer The reformer unit upgrades straight run virgin naphtha to reformate. This represents an octane rating increase from about 50 to as high as 100. RONC. The catalyst is a platinum catalyst based on a hologenated alumina (usually chloride & fluoride). The reactions are: Platinum Catalyzed "" Halogen " "" 1) Aromatization (endo) 2) Cyclization (endo) 3) Cracking (exo) 4) Isomerization (slightly exo) The reaction requires a 6:1 hydrogen to naphtha mole ratio. Reaction conditions are: 925 - 1,025F 300 psig Operating conditions can be varied in severity to produce a reformate with the desired octane number. TJH-gml 9-22-75 SAL 00001C483  f SAL 000010490 i